CALCIUM LOOPING PROCESSES FOR CARBON CAPTURE
DISSERTATION
Presented in Partial Fulfillment of the Requirements for the Degree Doctor of
Philosophy in the Graduate School of The Ohio State University
By
Shwetha Ramkumar, B.Tech.
Graduate Program in Chemical and Biomolecular Engineering
The Ohio State University
2010
Dissertation Committee:
Dr. Liang-Shih Fan, Adviser
Dr. Bhavik R. Bakshi
Dr. Andre F. Palmer
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Copyright by
Shwetha Ramkumar
2010
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ABSTRACT
A growing need for the reduction in anthropogenic carbon dioxide (CO2)
emission has led to a global push toward the development of efficient, economical, and
reliable carbon capture and sequestration technologies (CCS) for application to fossil
fuel based power plants. Several options are being investigated for the implementation
of CCS on pre-combustion and post-combustion systems including using solvents,
sorbents, membranes and chemical looping processes. The calcium looping process
(CLP) which is a calcium sorbent based chemical looping process, has the potential to
reduce the cost and increase the efficiency of CCS implementation on post-combustion
and pre-combustion systems.
In the CLP, a regenerable calcium-based sorbent is used to chemically absorb
CO2, sulfur, and halide impurities from synthesis gas or hydrocarbon feedstock during
the production of hydrogen(H2) and electricity or only electricity. The removal of CO2
drives the water-gas shift reaction and hydrocarbon reforming reaction forward via Le
Chatelier’s principle enabling the production of high-purity H2. The process operates at
high temperature (e.g., 600-700 °C), eliminating the need for a water gas shift catalyst
and allowing the exothermic heat of the CO2 absorption reaction to be recovered for
use in generating steam. This significantly reduces the energy penalty associated with
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CO2 capture. The spent sorbent consisting mostly of calcium carbonate (CaCO3) is
heated in a calciner to regenerate calcium oxide (CaO) for reuse in the process and to
release a concentrated CO2 stream that can be dried and sequestered. Overall CO2
emissions from the process are essentially zero. The regenerated sorbent is reactivated
in a hydrator, to eliminate sintering and improve the recyclability of the sorbent, before
being reintroduced into the H2 production reactor.
Among various reaction and process factors that are of importance to the CLP,
the reactivity and recyclability of the calcium based sorbent are vital. The nature of
calcium sorbent sintering that has been observed during multicyclic operation could
pose a severe limitation to the commercialization of the process. In realistic calcination
conditions, the sorbent loses one third to half of its original reactivity in a single cycle
due to calcination at 950 ºC and 1000 ºC respectively. Several methods of improving
the recyclability of CaO sorbents have been investigated including sorbent
pretreatment, modification by addition of supports and reactivation. Hydration of the
sorbent as a reactivation method after every calcination cycle was found to be very
effective in improving sorbent performance. The Wt% capture of the sorbent was
found to be constant at 50% during multicyclic CO2 capture with sorbent hydration in
every cycle in both bench scale and subpilot scale tests.
The CLP for production of H2 from syngas was investigated and very high
purity H2 was produced with less than 1ppm of hydrogen sulfide (H2S) at high
temperatures and pressures. For near stoichiometric steam addition, high carbon
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monoxide (CO) conversion and H2 purity can be obtained at high pressures and an
optimal temperature of 600 ºC. At atmospheric pressure, the presence of a water gas
shift catalyst with CaO sorbent improves the purity of H2. At high pressures, typical of
commercial deployment, the absence of the catalyst and the reduction of excess steam
addition do not have any effect on CO conversion and high H2 purity is obtained.
For a hydrocarbon feed, the steam reforming of the hydrocarbon is integrated
with the water gas shift and carbonation reaction in a single reactor. In addition to
improving the conversion of the hydrocarbon to H2, the CLP also provides an efficient
mode of internal heat integration where the endothermic energy for the reforming
reaction is obtained from the exothermic energy released by the combined water gas
shift and carbonation reaction. Single cycle tests have shown that the conversion of
methane (CH4) is improved to a large extent by the addition of CaO sorbent at 650 ºC.
High purity H2 is obtained at low steam to carbon(S:C) ratios of 3:1 for various
pressures ranging from 1 to 11 atms. The effect of calcination conditions on the extent
of CH4 reforming was determined. The reactivity of the sorbent was found to decrease
over multiple cycles due to calcination in both pure nitrogen and in a mixture of steam
and CO2. Hydration was found to be effective in reducing the sintering of the sorbent.
System analysis using ASPEN Plus has shown that the CLP has a high
efficiency for conversion of both coal as well as natural gas to H2 and electricity. The
CLP is being scaled up to a 25 KW subpilot unit demonstration at the Ohio State
University and the unit is currently under construction. The subpilot scale unit design is
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based on the thermodynamic, kinetic and sorbent reactivity studies and cold flow tests.
This unit will be used to conduct continuous testing for the production of H2 from a
simulated syngas stream and a mixture of hydrocarbons.
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Dedicated to my parents and sister for their love and support.
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ACKNOWLEDGMENTS
I would like to sincerely express my gratitude toward my adviser, Professor
Liang-Shih Fan, for his invaluable guidance, support and encouragement throughout
my graduate education at The Ohio State University. His perseverance, enthusiasm and
quest for knowledge has been a constant source of motivation. I am indebted to him for
the trust he confided in me and the time he spent with me, enriching me with his
experiences that help me mature as a professional and an individual.
I am also grateful to Professors Bhavik R. Bakshi, Kurt W. Koelling, David L.
Tomasko and Andre F. Palmer for serving in my qualifier, candidacy, and dissertation
committees, and thereby providing valuable suggestions and comments in this research
study. I would also like to take this opportunity to thank all faculty in this department,
especially for their support and encouragement.
I would like to specifically thank Dr. Mahesh Iyer who served as my mentor
throughout this study. The insightful discussions with him formed an invaluable part of
this study. I would like to thank Dr. Robert Statnick, an invaluable member of our
extended research team, for his guidance, support and motivation. His vast industry
experience was immensely helpful in developing this study. I would like to thank Dan
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Connell from CONSOL Energy for his support with the techno-economic evaluations.
I would like to thank the members of my research team: Danny Wong, William Wang
for all their support with all the post combustion work. I would also like to thank Nihar
Phalak and Niranjani Deshpande for all their support during the later part of the study.
It was the tremendous team effort of this group that helped in the development of the
overall calcium looping process.
I would like to thank Dr. Alissa Park for her mentoring and support. She has
been a great friend and source of inspiration. I would also like to thank the members of
our research group Dr. Songgeng Li., Fuchen Yu, Zhenchao Sun, Siddharth Gumuluru,
Fanxing Li, Zhao Yu, Fei Wang, Deepak Sridhar, Ray Kim, Andrew Tong, Liang
Zeng, Dr. Puneet Gupta, and Dr. Luis Velazquez-Vargas for their support and
friendship. Finally, I greatly enjoyed working with the undergraduate research
assistants: Brittany Valentine, Jessica Huber, Theresa Vonder-Haar, Eric Sacia, and
Brian Stelzer and cherish their friendship.
My sincere thanks to Lynn Flanagan, Amy Dudley, Susan Tesfai, Angela Jones,
Kari Uhl, Bill Cory and Paul Green for helping me in administrative and other fronts. I
would like to thank my friends at OSU, for all their warmth and hospitality that made
my stay in Columbus a wonderful experience to cherish forever.
I am indebted to the Ohio Coal Development Office (OCDO) of the Ohio Air
Quality Development Authority (OAQDA) and the US Department of Energy for
viii
providing financial assistance throughout this study. My special gratitude goes to Mr.
Bob Brown and Mr. Dan Cicero for providing useful suggestions.
Most importantly, I would like to thank my parents for their unconditional love,
and support and for their faith in me. Without their constant encouragement and
motivation this study would not have been possible. I would like to thank my sister,
Shmita Ramkumar, who is my biggest source of strength and inspiration, for her love
and support.
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VITA
June 21, 1983....……………………………..Born – Chennai, India
July 2001 − June 2005...………….…………B.S. Chemical Engineering
Anna University,
Chennai, India
September 2005 − present………………….. Graduate Research Associate
Chemical and Biomolecular Engineering
The Ohio State University
Columbus, OH, USA
PUBLICATIONS
1. Wang W., Ramkumar S., Li S., Wong D., Iyer M.V., Gumuluru S., Sakadjian B,
Statnick R.M., and Fan L-S “Sub-Pilot Demonstration of the CarbonationCalcination Reaction (CCR) Process: High Temperature CO2 and sulfur capture
from Coal Fired Power Plants”, Ind. Eng. Chem. Res. (2010).
2. Fan L-S., Li F., and Ramkumar S. Utilization of chemical looping strategy in coal
gasification processes. Particuology, 6(3), 131-142 (2008)
3. Ramkumar S., Li S., Wang W., Gumuluru S., Sun Z., Phalak N., and Fan L.-S.
“Results from the Carbonation-Calcination Reaction (CCR) Process”, Proc. 26th
Intl. Pittsburgh Coal Conf., Pittsburgh, PA, September (2009)
4. Ramkumar S., and Fan L.-S. “Calcium Looping Process for Clean Fossil Fuel
Conversion”, Proc. 26th Intl. Pittsburgh Coal Conf., Pittsburgh, PA, September
(2009)
x
5. Ramkumar S., Phalak N., Sun Z., and Fan L.-S. “Calcium Looping Process
Enhanced Coal to Liquid Technology”, Proc. 26th Intl. Pittsburgh Coal Conf.,
Pittsburgh, PA, September (2009)
6. Ramkumar S., Connell D., and Fan L-S. “Calcium Looping process for clean fossil
fuel conversion” 1st Meeting of the High Temperature Solid Looping Cycles
Network, Oviedo, Spain, September (2009)
7. Ramkumar S., Wang W., Li S., Gumuluru S., Sun Z., Phalak N., Wong D., Iyer M.,
Statnick R.M., Fan L-S., Sakadjian B., and Sarv H. “Carbonation-Calcination
Reaction(CCR) Process for High Temperature CO2 and Sulfur Removal” 1st
Meeting of the High Temperature Solid Looping Cycles Network, Oviedo, Spain,
September (2009)
8. Ramkumar S., and Fan L-S., “Calcium looping process for clean fossil fuel
conversion”, 8th World Congress of Chemical Engineering, Montreal, Canada,
August (2009)
9. Ramkumar S., Wang W., Li S., Wong D., Iyer M.V., Gumuluru S., Statnick R.M.,
and Fan L-S., “Carbonation-Calcination Reaction Process for High Temperature
CO2 and sulfur Removal”, 8th World Congress of Chemical Engineering, Montreal,
Canada, August (2009)
10. Sakadjian B., Wang W., Li S., Ramkumar S., Gumuluru S., Fan L-S., and Statnick
R.M., “Sub-Pilot Demonstration of the CCR Process: High Temperature CO2
Capture Sorbents for Coal Fired Power Plants” Proc. Int. Tech. Conf. Coal
Utilization & Fuel Systems, Clearwater, FL (2009)
11. Fan L.-S., Li F., Velazquez-Vargas L.G., and Ramkumar S. “Chemical Looping
Gasification”. Proc. 9th International Conference on Circulating Fluidized Beds.
Hamburg, Germany. May (2008).
FIELDS OF STUDY
Major Field: Chemical Engineering
xi
TABLE OF CONTENTS
Page
Abstract………………………………………………………………………..……….ii
Dedication……………………………………………………………………...………vi
Acknowledgments………………………………………………………......…………vii
Vita……………………………………………………………………...……………...ix
List of Tables……………………………………………………………….………….xvi
List of Figures……………………………………………………………….……..….xviii
Chapters:
Chapter 1…………………………………………………………….……..…………... 1 Introduction…………………………………………………………….……..……... 1
Chapter 2…………………………………………………………….……..…………... 9 Literature Review: Processes for Enhanced H2 Production with CO2 capture……… 9 2.1 Introduction………………………………………………………….………... 9 2.2 CO2 Acceptor Process……………………………………………………….. 13 2.3. HyPr-RING Process……………………………………………………….. 20 2.4. Zero Emission Coal Alliance (ZECA) Process............................................... 26 2.5. ALSTOM Hybrid Combustion-Gasification Process………………………. 27 2.6. Fuel-Flexible Advanced Gasification-Combustion Process………………... 28
Chapter 3………………………………………………………….……..……………. 37 Reactivity and Recyclability of Calcium Based Sorbents for CO2 Capture……….. 37 3.1 Introduction………………………………………………………….………. 37 3.2 Sorbent Reactivity Over Multicyclic Reactions…………………………….. 38 3.3 Synthesis of High Reactivity Precipitated Calcium Carbonate (PCC)
Sorbent………………………………………………………….……..………… 42 3.4 Pretreatment of Calcium Based Sorbents and Addition of Supports………... 45 3.4.1 Reactivity Testing of Ca-based Sorbents for CO2 Capture……………... 45 xii
3.4.2 Recyclability of Natural, Pretreated and Supported Sorbents………….. 46 3.5 Effect of Realistic Calcination Conditions on Sorbent Reactivity………….. 48 3.5.1 Experimental Methods………………………………………………….. 49 3.5.2 Results and Discussion…………………………………………………. 50 3.6. Sorbent Reactivation by Hydration –Lab Scale Testing……………………. 51 3.6.1 Experimental Methods………………………………………………….. 51 3.6.2 Results and Discussion…………………………………………………. 52 3.7 Sub-Pilot Scale Demonstration of Reactivation of Calcium Sorbent by
Hydration…………………………………………………………………………54 3.7.1 Experimental Methods for the 120 KWth Subpilot Scale Testing……… 55 3.7.2 Results and Discussion…………………………………………………. 57 3.8 Conclusions………………………………………………………………….. 58
Chapter 4………………………………………………………….……..……………. 73 Enhanced Catalytic H2 Production from Syngas……………………………………73 4.1 Introduction………………………………………………………….………. 73 4.2 Calcium Looping Process(CLP) Configuration and Thermodynamics……... 74 4.2.1 The Carbonation Reactor……………………………………………….. 76 4.2.2 The Calciner………………………………………………………….…. 81 4.2.3 Sorbent Reactivation by Hydration…………………………………….. 82 4.3 Materials and methods………………………………………………………. 84 4.3.1 Chemicals, Sorbents, and Gases…………………………………………84 4.3.2 Fixed Bed Reactor Unit Setup………………………………………….. 84 4.3.3 Water Gas Shift Reaction Testing………………………………………. 86 4.3.4 Simultaneous Water Gas Shift and Carbonation………………………... 87 4.3.5 Catalyst Pretreatment…………………………………………………… 87 4.3.6 Combined H2 Production with H2S Removal…………………………... 88 4.4 Results and Discussion……………………………………………………… 89 4.4.1 Effect of Process Parameters on the Extent of Water Gas Shift Reaction
using HTS Catalyst………………………………………………………….…89 4.4.2 Enhancing the Water Gas Shift Reaction by In-situ CO2 Removal (HTS
Catalyst and CaO Sorbent)…………………………………………………… 91 4.4.3 Simultaneous Water Gas Shift, Carbonation and Sulfidation Reaction
Testing………………………………………………………………………… 94 4.4.4 Effect of Catalyst Type on the Water Gas Shift Reaction……………….95 4.5 Conclusions………………………………………………………………….. 99
Chapter 5…………………………………………………………………………….. 129 Enhanced Non-Catalytic H2 production from Syngas……………………………..129 5.1 Introduction………………………………………………………………… 129 5.2 Materials and Methods……………………………………………………... 130 5.2.1 Chemicals, Sorbents, and Gases………………………………………. 130 5.2.2 Experimental Setup: Fixed Bed Reactor………………………………. 130 xiii
5.2.3 Water Gas Shift Reaction in the Presence and Absence of HTS
Catalyst………………………………………………………….…………… 131 5.2.4 Simultaneous Water Gas Shift and CO2 Removal…………………….. 131 5.2.5 Combined H2 Production with H2S Removal…………………………. 132 5.3 Results and Discussion……………………………………………………... 132 5.3.1 Baseline Water Gas Shift Reaction Testing…………………………… 132 5.3.2 Water Gas Shift Reaction in the Presence of Only CaO Sorbent………133 5.3.3 H2 Production in the Presence of CaO Sorbent Only and a Mixture of CaO
Sorbent and Catalyst………………………………………………………….138 5.3.4 Multicyclic Investigation of H2 Production in the Presence of CaO
Sorbent Only. ………………………………………………………….……. 139 5.3.5 Enhanced H2 Production With CO2 and Sulfur Capture………………. 140 5.4. H2 Production From Coal Gasification Derived Syngas…………………... 143 5.4.1 Process Overview……………………………………………………… 143 5.4.2 System Thermodynamics Analysis……………………………………. 146 5.5 H2 Production From Syngas Derived From Natural Gas Feedstocks……... 149 5.5.1 Syngas from Steam Reforming of Natural Gas……………………….. 149 5.5.2 Syngas from Autothermal Reforming of Natural Gas………………… 151 5.5.3 Syngas from Partial Oxidation of Natural Gas…………………………151 5.6 Addressing The Issue of Sulfur in the Feedstock………………………….. 152 5.6.1 Experimental Analysis of the Regeneration of CaS…………………… 155 5.7. Conclusion………………………………………………………….………158
Chapter 6………………………………………………………….……..…………... 204 Process Simulation and Economics of the Calcium Looping Process (CLP) for
production of h2 from Coal……………………………………………………….. 204 6.1 Introduction………………………………………………………….……... 204 6.2 Production of Fuel Cell Grade H2 With a PSA…………………………….. 204 6.2.1 Cogeneration of H2 and Electricity……………………………………. 204 6.2.2 Production of Only H2 With Internal Heat Integration………………... 209 6.3 Production of H2 Having a Purity of 94–98% Without a PSA…………….. 211 6.3.1 Cogeneration of H2 and Electricity……………………………………. 211 6.3.2 Production of H2 With Internal Heat Integration……………………… 211 6.4 Comparison of the Process Efficiencies for Different Gasifiers…………… 212 6.5 Effect of Process Parameters on CLP Performance Using Syngas From a GE
Gasifier………………………………………………………….……..……….. 212 6.5.1 Approach………………………………………………………….…… 213 6.5.2 Sensitivity Analysis for the Yield and Purity of H2 Produced………… 213 6.5.3 Sensitivity Analysis for the Extent of Contaminant Removal from the
Product H2………………………………………………………….……..… 215 6.5.4 Sensitivity Analysis for the Cold Gas Efficiency and Overall Process
Efficiency………………………………………………………….……..….. 217 6.6 Effect of Addition of Sorbent Hydration to the CLP Process……………… 219 xiv
6.7 Techno-Economic Analysis of H2 Production From Coal…………………. 220 6.7 Conclusions………………………………………………………….……... 223
Chapter 7………………………………………………………….……..………….. 263 Enhanced Reforming of Hydrocarbons…………………………………………… 263 7.1 Introduction………………………………………………………….……... 263 7.2 Process Configuration and Thermodynamics……………………………… 264 7.2.1 The Carbonation Reactor System………………………………………264 7.2.2 Calciner or Sorbent Regeneration Reactor……………………………. 267 7.2.3. Hydrator or Sorbent Reactivation Reactor……………………………. 268 7.3. Experimental Methods…………………………………………………….. 268 7.3.1 Chemicals, Sorbents, and Gases………………………………………. 268 7.3.2 Bench Scale Experiment Setup………………………………………... 268 7.3.3 Steam Methane Reforming in the Presence of a Ni-based Catalyst……269 7.3.4 Simultaneous Steam Methane Reforming, Water Gas Shift and
Carbonation………………………………………………………….………. 270 7.3.5 Multicyclic Steam Methane Reforming and Spent Sorbent Calcination 270 7.4 Results and Discussion…………………………………………………….. 272 7.4.1 Base-line Steam Methane Reforming Testing………………………… 272 7.4.2 Simultaneous Reforming with In-situ CO2 Removal (Catalyst with CaO
Sorbent) ………………………………………………………….……..…… 273 7.4.3 Effect of Sorbent Calcination Conditions on the Extent of Steam
Reforming………………………………………………………….…………279 7.4.4 Calcination in N2 with Sorbent Hydration…………………………….. 280 7.4.5 Realistic Sorbent Calcination in a Steam/CO2 Atmosphere with Sorbent
Hydration………………………………………………………….……..….. 282 7.5 Applications of CLP in Hydrocarbon Reforming………………………….. 283 7.5.1 Steam Reforming of Natural Gas and Other Hydrocarbons for H2 and
Electricity Generation..…………………………………………….………... 283 7.5.2 Implementation of Carbon Capture in Liquid Fuels Production From
Coal………………………………………………………….……..………... 288 7.6 Conclusions………………………………………………………………… 293
Chapter 8………………………………………………………….……..…………... 332
Subpilot scale testing and recommendations for future work……………………. 332 8.1 Introduction………………………………………………………….…….. 332 8.2 Cold Flow Testing………………………………………………………….. 333 8.3 Design of the Subpilot Scale Unit………………………………………….. 336 8.4 Conclusions………………………………………………………………… 339 8.5 Recommendations for Future Work……………………………………….. 339
Appendix - A………………………………………………………….……..……… 350 xv
LCA Analysis - Comparison of the conventional coal to H2 process with the CLP
process………………………………………………………….……..………….. 350 References………………………………………………………….……..…………. 377 xvi
LIST OF TABLES
Table
Page
Table 2.1: A typical composition of the H2 rich synthesis gas from the gasifier ......... 30 Table 5.1: Typical fuel gas compositions obtained from different gasifiers (Stultz and
Kitto, 1992). ............................................................................................... 160 Table 5.2: Fuel gas composition entering the water gas shift reactor after steam
addition (S:C ratio =1:1) (adapted from Stultz and Kitto, 1992) ............... 161 Table 5.3: Fuel gas composition entering the water gas shift reactor after steam
addition (S:C ratio =3:1) (adapted from Stultz and Kitto, 1992) ............... 162 Table 5.4: Extent of equilibrium CO conversion and CO2 capture in the CLP from
Steam Methane Reforming (SMR) derived syngas ................................... 163 Table 5.5: Extent of equilibrium CO conversion and CO2 capture in the CLP from
Auto Thermal Reforming (ATR) derived syngas ...................................... 164 Table 5.6: Extent of equilibrium CO conversion and CO2 capture in the CLP from
partial oxidation (POX) derived syngas .................................................... 165 Table 6.1: Properties of Illinois # 6 coal .................................................................... 225 Table 6.2: Composition of the syngas exiting from the Shell gasifier ....................... 226 Table 6.3: Intermediated pressures for compression of the CO2 for sequestration .... 227 Table 6.4: Components list for the ASPEN Plus® simulation................................... 228 Table 6.5: ASPEN Plus® models used for the simulation of the CLP ....................... 229 Table 6.7: Power balance in the CLP process ............................................................ 237 Table 6.8: Process simulation results for the CLP process ........................................ 238 xvii
Table 6.9 Summary of the schemes investigated for the production of H2 alone and for
the coproduction of H2 and electricity with a PSA .................................... 239 Table 6.10 Summary of the schemes investigated for the production of H2 alone and
for the coproduction of H2 and electricity without a PSA ......................... 240 Table 6.11 Comparison of the efficiency of the H2 production process for different
gasifiers ...................................................................................................... 241 Table 6.12: Levelized annual costs and levelized cost of H2 for the conventional coal
to H2 plant (adapted from DOE, 2010) ...................................................... 242 Table 6.13: Levelized annual costs and levelized cost of H2 for the CLP plant ........ 243 Table 7.1: Thermodynamic extent of the various reactions occurring in the carbonator
................................................................................................................... 295 Table 7.2: Stream data for the integration of the CLP in a steam methane reforming
process ....................................................................................................... 296 Table 7.3: Energy balance for the production of H2 and electricity from natural gas
using the CLP. ........................................................................................... 298 Table 7.4: Heat required for the production of steam at 650 ºC from water at 15 atms
................................................................................................................... 299 Table 7.5: Heat required for preheating the natural gas at 15 atms to 650 ºC ........... 300 Table 7.6: Heat released by the solids from the calciner in the H2 production reactor.
................................................................................................................... 301 Table 7.7: Heat generated from the H2 production reactor ........................................ 302 Table 7.8: Heat released on cooled the H2 from 650 ºC to ambient temperature ...... 303 Table 7.9: Heat required for preheating the PSA tail gas from 650 ºC to 900 ºC ...... 304 Table 7.10: Heat required for preheating the oxygen from ambient temperature to 900
ºC ............................................................................................................... 305 Table 7.11: Heat absorbed by the solids from the H2 production reactor in the calciner
................................................................................................................... 306 Table 7.12: Heat released from the calciner ............................................................... 307 Table A.1: Energy balance for the conventional process ........................................... 362 xviii
Table A.2: Water balance for the conventional process............................................. 363 Table A.3: Quantification of the inputs for the conventional process ....................... 364 Table A.4: Quantification of outputs for the conventional process ........................... 365 Table A.5: Energy balance for the CLP ..................................................................... 366 Table A.6: Water balance for the CLP ....................................................................... 367 Table A.7: Quantification of the inputs for the CLP .................................................. 368 Table A.8: Quantification of the outputs for the CLP ................................................ 369 Table A.9: Quantification of the Global Warming Potential(GWP) for the inputs and
outputs for the conventional process. ........................................................ 370 Table A.10: Quantification of the Global Warming Potential(GWP) for the inputs and
outputs for the CLP .................................................................................... 371 xix
LIST OF FIGURES
Table
Page
Figure 1.1: Historical data and projections of the world energy consumptions till 2030
(EIA, 2009). ...................................................................................................... 6 Figure 1.2: Projections of the world energy supply by different fuel types including
fossils fuels and renewable (EIA, 2009). .......................................................... 7 Figure 1.3: Implementation of Carbon Capture and Sequestration (CCS) in fossil fuel
based power plants............................................................................................ 8 Figure 2.1: Schematic diagram of the reactor system in the gas synthesis block of the
CO2 Acceptor process (Dobbyn et al., 1978) ................................................. 32 Figure 2.2: Schematic diagram of the HyPr-RING process .......................................... 33 Figure 2.3: Schematic of the ZECA process ................................................................. 34 Figure 2.4: Schematic of the ALSTOM process ........................................................... 35 Figure 2.5: Schematic of the GE process....................................................................... 36 Figure 3.1: Comparison in the CO2 capture capacity of CaO sorbents obtained from
different precursors. (Calcination conditions: T = 700 ºC, P = 1 atm, pure N2
carrier gas; Carbonation conditions: T = 650 ºC, P = 1 atm, 10% CO2/90%
N2 feed gas) .................................................................................................... 60 Figure 3.2: Comparison in the multicyclic conversion of PCC powder sorbent PCC
pelletized and broken sorbent (Calcination conditions: T = 700 ºC, P = 1 atm,
pure N2 carrier gas; Carbonation conditions: T = 650 ºC, P = 1 atm, 10%
CO2/90% N2 feed gas) .................................................................................... 61 xx
Figure 3.3: CO2 capture capacity of pretreated and supported Ca-based sorbents over
multiple carbonation –calcination cycles (Calcination conditions: T = 700 ºC,
P = 1 atm, pure N2 carrier gas; Carbonation conditions: T = 650 ºC, P = 1
atm, 10% CO2/90% N2 feed gas) .................................................................... 62 Figure 3.4: Effect of steam concentration in the calcination carrier gas on the CO2
capture capacity of CaO sorbent (Calcination conditions: T = 900 ºC, P =
1atm) ............................................................................................................... 63 Figure 3.6: Effect of steam calcination on multicyclic carbonation and calcination of
CaO sorbent (Calcination conditions: T = 900 ºC, P = 1 atm, carrier gas =
50%H2O/50% CO2; Carbonation conditions: T = 650 ºC, P = 1 atm, 10%
CO2/90% N2 feed gas) .................................................................................... 65 Figure 3.7: Effect of hydration conditions on sorbent reactivity ................................... 66 Figure 3.8: Effect of hydration pressure on sorbent reactivity (Hydration temperature =
600 ºC) ............................................................................................................ 67 Figure 3.9: Effect of steam hydration on sorbent reactivity over multiple calcinationhydration-carbonation cycles (Calcination conditions: T = 900 ºC, P = 1 atm,
carrier gas = pure CO2; Carbonation conditions: T = 650 ºC, P = 1 atm, 10%
CO2/90% N2 feed gas, Hydration conditions: T = 500 ºC, P ~ 1 atm, 90%
H2O/10% N2 feed gas) .................................................................................... 68 Figure 3.10: Process flow diagram of the CLP for CO2 and SO2 removal from
combustion flue gas ........................................................................................ 69 Figure 3.11: Snapshot of the sub-pilot scale facility of the CLP integrated with a coal
fired combustor. .............................................................................................. 70 Figure 3.12: Effect of hydration on the % CO2 removed from the flue gas over multiple
cycles .............................................................................................................. 71 Figure 3.13: Wt.% CO2 capture achieved by the hydrated sorbent over multiple cycles
........................................................................................................................ 72 Figure 4.1: Schematic of the CLP................................................................................ 100 xxi
Figure 4.2: Thermodynamic data illustrating the equilibrium constants of the water gas
shift reaction and the combined water gas shift and carbonation reaction ... 101 Figure 4.3: Thermodynamic data for the hydration and carbonation of CaO sorbent . 102 Figure 4.4: Equilibrium H2 purity in the carbonator at varying temperatures, pressures
and S: C ratios. ( Feed gas: 10% CO and balance nitrogen) ........................ 103 Figure 4.5: Thermodynamic data for the sulfidation (H2S) of CaO with varying steam
partial pressures. (PTotal = 30 atm) ................................................................ 104 Figure 4.6: Thermodynamic data for predicting the equilibrium COS concentration for
CaO sulfidation with varying CO2 concentration (PTotal = 30 atm) .............. 105 Figure 4.7: Thermodynamic data for predicting the equilibrium HCl concentration for
CaO reaction with HCl with varying steam concentration (PTotal = 30 atm) 106 Figure 4.8: Thermodynamic data for the carbonation of CaO..................................... 107 Figure 4.9: Thermodynamic data for the hydration of CaO ........................................ 108 Figure 4.10: Simplified flow sheet of the bench scale experimental setup ................. 109 Figure 4.11: X-ray diffraction patters of the HTS catalyst before pretreatment
(hematite) ...................................................................................................... 110 Figure 4.12: X-ray diffraction patters of the HTS catalyst after pretreatment (magnetite)
...................................................................................................................... 111 Figure 4.13: Effect of reaction temperature and S:C ratio on the conversion of CO by
the water gas shift reaction in the presence of HTS catalyst at (a) 1 atm (b) 21
...................................................................................................................... 112 Figure 4.14: Effect of reaction temperature and pressure on the observed partial
pressure ratio for the water gas shift reaction in the presence of HTS catalyst
at a S:C ratio of (a)1:1 (b)3:1........................................................................ 113 Figure 4.15: Typical curves for the combined water gas shift and carbonation reaction
in the presence of CaO sorbent and HTS catalyst depicting (a) Gas
composition (mol%) and (b) CO conversion (650 ºC, 1 atm, S:C ratio of 3:1)
...................................................................................................................... 114 xxii
Figure 4.16: Effect of pressure on purity of H2 produced during the combined water gas
shift and carbonation reaction in the presence of CaO sorbent and HTS
catalyst at a S:C ratio of (a) 3:1 (b) 1:1 (650 ºC) .......................................... 115 Figure 4.17: Effect of S:C ratio on the combined water gas shift and carbonation
reaction in the presence of CaO sorbent and HTS catalyst at 650 ºC (a) CO
conversion at 1 atm (b) H2 gas composition at 1 atm (c) CO conversion at 21
atm (d)H2 gas composition at 21 atm ........................................................... 116 Figure 4.18: Effect of temperature on CO conversion by the combined water gas shift
and carbonation reaction in the presence of CaO sorbent and HTS catalyst at
1 atm and S:C ratio of 3:1 ............................................................................. 118 Figure 4.19: Effect of temperature on CO conversion by the combined water gas shift
and carbonation reaction in the presence of CaO sorbent and HTS catalyst at
21 atm and S:C ratio of (a) 3:1 (b) 1:1 ......................................................... 119 Figure 4.20: Effect of S:C ratio on (a) the composition of H2S in the H2 stream and (b)
CO conversion in the presence of the catalyst and sorbent during the
simultaneous water gas shift, carbonation and sulfidation reaction(600 ºC, 1
atm) ............................................................................................................... 120 Figure 4.21: Effect of S:C ratio on the composition of H2S in the H2 stream during the
combined water gas shift, carbonation and sulfidation reaction in the presence
of CaO sorbent and HTS catalyst (600 ºC, 1 atm) ........................................ 121 Figure 4.22 : Effect of S:C ratio and temperature on CO conversion during the water
gas shift reaction in the presence of STC and HTS catalyst ......................... 122 Figure 4.23: Effect of reaction temperature on CO conversions for various pressures at
an S:C ratio of 1:1 for the STC (0.25g STC, Total flow = 0.725 slpm) ....... 123 Figure 4.24: Effect of reaction temperature on CO conversions for the HTS and STC at
11 atms and S:C ratio of 1:1(Total flow = 0.725 slpm) ................................ 124 Figure 4.25: Effect of reaction temperature on CO conversions for the HTS and STC at
21 atms and S:C ratio of 1:1(Total flow = 0.725 slpm) ................................ 125 xxiii
Figure 4.26: Effect of S:C ratio, type of catalyst and effect of H2S on CO conversion
during the water gas shift reaction(650 ºC, 1atm) ........................................ 126 Figure 4.27: Effect of temperature on CO conversion (Temperature=650°C, Pressure
= 1 atm, S:C ratio= 1:1) ................................................................................ 127 Figure 4.28: Comparison in the CO conversion obtained at different S:C ratios for
different sorbent and catalyst mixtures (650 ºC, 1atm) ................................ 128 Figure 5.1: Simplified flow sheet of the bench scale experimental setup ................... 166 Figure 5.2: Effect of reaction temperature and S:C ratio on the conversion of CO by the
water gas shift reaction at 1 atm ................................................................... 167 Figure 5.3: Effect of reaction temperature and S:C ratio on the conversion of CO by the
water gas shift reaction at 21 atm ................................................................. 168 Figure 5.4: Typical breakthrough curves for the production of H2 in the presence of
CaO sorbent without catalyst (a) Gas composition (mole%) and (b) CO
conversion (600 °C, 21 atm, S:C ratio of 3:1) .............................................. 169 Figure 5.5: Effect of pressure on CO conversion obtained in the presence of CaO
sorbent without catalyst (650°C, S:C ratio of 3:1) ....................................... 170 Figure 5.6: Effect of S:C ratio on CO conversion obtained in the presence of CaO
sorbent without catalyst at (a) 1 atm, (b) 11 atm, (c) 21 atm (650°C) .......... 171 Figure 5.7: Effect of temperature on CO conversion obtained in the presence of CaO
sorbent without catalyst at a S:C ratio of (a) 1:1 and (b) 3:1 (1 atm) ........... 173 Figure 5.8: Effect of CO concentration in the feed on the (a) CO conversion and (b)
purity of H2 produced in the presence on CaO sorbent without catalyst (11
atm, 600°C, S:C ratio of 3:1) ........................................................................ 174 Figure 5.9: SEM image of the (a) initial CaCO3 sorbent (b) CaO sorbent obtained from
the calcination of CaCO3 .............................................................................. 175 Figure 5.10: SEM image of sorbent at the end of the water gas shift and carbonation
reaction in the absence of a catalyst at (a) 1 atm (b) 21 atm (S:C ratio of 3:1,
600°C)........................................................................................................... 176 xxiv
Figure 5.11: Comparison in the product H2 purity in the presence of the sorbent and in
the presence of the sorbent and catalyst mixture at 1 atm (650°C, S:C ratio of
1:1) ................................................................................................................ 177 Figure 5.12: Comparison in the product H2 purity in the presence of the sorbent and in
the presence of the sorbent and catalyst mixture (650°C, 21 atm) ............... 178 Figure 5.13: Product H2 purity obtained over multiple reaction and regeneration cycles
in the presence of CaO sorbent without catalyst at 4.5 atms. (600°C, S:C ratio
of 3:1) ........................................................................................................... 179 Figure 5.14: Product H2 purity obtained over multiple reaction - regeneration cycles in
the presence of CaO sorbent without catalyst at 21 atms (600°C, S:C ratio of
3:1) ................................................................................................................ 180 Figure 5.15: Effect of S:C ratio on the (a) extent of H2S removal and (b) the purity of
H2 produced during the combined water gas shift, carbonation and sulfidation
reaction in the presence of CaO sorbent (1atm, 600oC) ............................... 181 Figure 5.16: Effect of temperature on the (a)extent of H2S removal and (b) purity of H2
produced during the combined water gas shift, carbonation and sulfidation
reaction in the presence of CaO sorbent (1 atm, S:C ratio of 1:1) ............... 182 Figure 5.17: Effect of pressure on the (a) extent of H2S removal (b) purity of H2
produced during the combined water gas shift, carbonation and sulfidation
reaction in the presence of CaO sorbent (S:C ratio of 1:1, 600oC) .............. 183 Figure 5.18: SEM image of the (a) initial CaCO3 sorbent (b) CaO sorbent obtained
from the calcination of CaCO3 (c) sorbent at the end of the water gas shift,
carbonation and sulfidation reaction at 1 atm (c) CaO sorbent obtained from
the calcination of CaCO3 (600 oC, S:C ratio of 1:1) (c) sorbent at the end of
the water gas shift, carbonation and sulfidation reaction at 21 atm (600oC, S:C
ratio of 1:1) ................................................................................................... 184 Figure 5.19: (a) Conventional process for H2 production from coal (b) Integration of
the CLP in a conventional process for H2 production from coal .................. 185 xxv
Figure 5.20: Integration of the CLP in a coal gasification system for the production of
electricity, H2 and liquid fuels ...................................................................... 187 Figure 5.21: Comparison of the PCO2 in the carbonator with the equilibrium PCO2 for
the carbonation of CaO for a S:C ratio of (a)1:1 (b)3:1 ............................... 188 Figure 5.22: Comparison of the PH2O in the carbonator with the equilibrium PH2O for
the hydration of CaO for a S:C ratio of (a)1:1 (b)3:1 .................................. 190 Figure 5.23: Effect of temperature on equilibrium CO conversion in the water gas shift
reactor at a S:C ratio of (a) 1:1 (b) 3:1 ......................................................... 192 Figure 5.24: Effect of temperature on equilibrium CO conversion in the presence of
CaO in the carbonation reactor of the CLP at a S:C ratio of (a) 1:1 (b) 3:1 . 193 Figure 5.25: Effect of temperature on equilibrium H2 purity in the presence of CaO at a
S:C ratio of (a) 1:1 (b) 3:1 ............................................................................ 194 Figure 5.26: Effect of temperature and S:C ratio on the % of carbon captured in the
CLP using syngas from different gasifiers as the feed ................................. 195 Figure 5.27: Conventional steam reforming of natural gas for H2 production with a
methanator .................................................................................................... 196 Figure 5.28: Conventional steam reforming of natural gas for H2 production with a
PSA ............................................................................................................... 197 Figure 5.29: CLP integrated in the conventional steam reforming of natural gas process
...................................................................................................................... 198 Figure 5.30: Conventional partial oxidation process for conversion of natural gas to H2
...................................................................................................................... 199 Figure 5.31: CLP integrated in the partial oxidation of natural gas for H2 production200 Figure 5.32: Effect of the change in temperature and steam composition on the
regeneration of CaS with H2O ...................................................................... 201 Figure 5.33: Effect of the change in steam and CO2 composition on the regeneration of
CaS in the presence of H2O and CO2 ........................................................... 202 Figure 5.34: H2S evolved in the presence of H2O and CO2 from spent sorbent produced
during combined CO2 and H2S removal at 1 and 21 atms............................ 203 xxvi
Figure 6.1: The CLP for coproduction of fuel cell grade H2 and electricity from coal
...................................................................................................................... 244 Figure 6.2: ASPEN simulation flow diagram for the CLP process with a PSA .......... 245 Figure 6.3 Aspen simulation for the production of H2 using the CLP without a PSA.246 Figure 6.4: Aspen model used for sensitivity analysis of the combined reactions
occurring in the H2 production reactor of the CLP. ...................................... 247 Figure 6.5: Effect of temperature on the H2 purity produced at the outlet of the
carbonation reactor (S:C ratio = 3, Pressure = 10 atms)............................... 248 Figure 6.6: Effect of pressure on the H2 purity produced at the outlet of the carbonation
reactor( S:C ratio = 3, Temperature = 600 ºC) ............................................. 249 Figure 6.7: Effect of S:C ratio on the H2 purity produced at the outlet of the
carbonation reactor ( Pressure = 10 atms, Temperature = 600 ºC)............... 250 Figure 6.8: Effect of temperature and S:C ratio on the extent of H2S removal. .......... 251 Figure 6.9: Effect of temperature and S:C ratio on the extent of COS removal.......... 252 Figure 6.10: Effect of temperature and S:C ratio on the amount of CO impurity present
in the H2 stream. ........................................................................................... 253 Figure 6.11: Effect of temperature and S:C ratio on the extent of CO2 removal. ....... 254 Figure 6.12: Effect of temperature and S:C ratio on the amount of CH4 impurity present
in the H2 product stream. .............................................................................. 255 Figure 6.13: Effect of pressure on the cold gas efficiency, process efficiency and H2
purity obtained from the H2 production reactor at various S:C ratios. ......... 256 Figure 6.14: Effect of S:C ratio on H2 purity, cold gas efficiency and process efficiency
...................................................................................................................... 257 Figure 6.15: Effect of temperature on H2 purity, cold gas efficiency and process
efficiency (1:1, 10 atms) ............................................................................... 258 Figure 6.16: Effect of Ca:C ratio on H2 purity, cold gas efficiency and process
efficiency (600 ºC, 1:1, 10 atms) .................................................................. 259 Figure 6.17: Effect of the addition of sorbent hydration to the CLP ........................... 260 xxvii
Figure 6.18: Process flow diagram of the conventional coal to H2 plant used for the
economical analysis ( DOE, 2010) ............................................................... 261 Figure 6.19: Process flow diagram of the CLP plant used for the economical analysis
...................................................................................................................... 262 Figure 7.1: Schematic of the CLP for the conversion of hydrocarbons to H2 ............. 309 Figure 7.2: Thermodynamic data illustrating the equilibrium constants of the steam
reforming of CH4, water gas shift and carbonation reaction ........................ 310 Figure 7.3: Simplified schematic of the bench scale experimental setup .................... 311 Figure 7.4: Effect of temperature and S:C ratio on (a)H2 purity and (b) the amount of
CO, CO2 and CH4 remaining in the product gas for the steam methane
reforming reaction in the presence of Ni-based catalyst ( P = 1 atm) .......... 312 Figure 7.5: Breakthrough curve in the composition of the product gases obtained
during the simultaneous reforming, water gas shift and carbonation reaction.
(T = 650 ºC, P = 1 atm) ................................................................................ 313 Figure 7.6: CH4 conversion obtained during the simultaneous reforming, water gas
shift and carbonation reaction. (T = 650 ºC, P = 1 atm) ............................... 314 Figure 7.7: Effect of temperature and S:C ratio on (a) H2 purity (b) conversion of CH4
(P = 1atm) ..................................................................................................... 315 Figure 7.8: Effect of temperature and S:C ratio on the amount of (a) CO and (b) CO2
remaining in the product gas for H2 production from methane with/without
sorbent. ( P = 1 atm) ..................................................................................... 316 Figure 7.9: Effect of pressure on (a) H2 purity and (b) CH4 concentration in the product
stream. (T = 650 ºC, S:C ratio = 3) ............................................................... 317 Figure 7.10: Effect of pressure on (a) CO2 and (b) CO concentration in the product
stream. (T = 650 ºC, S:C ratio = 3) ............................................................... 318 Figure 7.11: Effect of pressure on the prebreakthrough and postbreakthrough
concentration of CH4, CO and CO2 in the product stream. (T = 650 ºC, S:C
ratio = 3) ....................................................................................................... 319 xxviii
Figure 7.12: Effect of calcination conditions on (a) H2 purity and (b) CH4 composition
in the product gas for cycles 1,2,3 and 4. [(Reforming reaction conditions :T=
650 ºC, P = 1 atm, S:C ratio = 3) (Calcination conditions: sorbent for cycles
1, 2 and 3 are calcined in pure N2 at 950 and the sorbent for cycle 4 is
calcined in a 50:50 CO2/H2O atmosphere at 950 ºC.)] ................................. 320 Figure 7.13: Effect of hydration on (a) H2 purity and (b) CH4 composition in the
product gas for cycles 1, 2, 3 and 4. [(Reforming reaction conditions :T= 650
ºC, P = 1 atm, S:C ratio = 3) (Calcination conditions: sorbent for cycles 1, 2,
3 and 4 are calcined in pure N2, T = 950, P = 1 atm)(Hydration conditions:
hydration of calcined sorbent from the 3rd cycle in a 80:20 H2O/N2
atmosphere, T = 600, P = 11 atm)] ............................................................... 321 Figure 7.14: Effect of hydration on H2 purity for cycles 1,2,3 and 4. [(Reforming
reaction Conditions: T= 650 ºC, P = 1 atm, S:C ratio = 3) (Calcination
conditions: sorbent for cycles 1, 2, 3 and 4 are calcined in pure N2, T = 950, P
= 1 atm)(Hydration conditions: hydration for cycles 1, 2, 3 and 4 in a 80:20
H2O/N2 atmosphere, T = 600, P = 11 atm)] ................................................. 323 Figure 7.15: Effect of hydration on (a) H2 purity and (b) CH4 content in the product gas
for cycles 1,2,3 and 4. [(Reforming reaction Conditions: T= 650 ºC, P = 1
atm, S:C ratio = 3) (Calcination conditions: sorbent for cycles 1, 2, 3 and 4
are calcined in a 50:50 CO2/H2O atmosphere, T = 950, P = 1 atm) (Hydration
conditions: hydration for cycles 1, 2, 3 and 4 in a 80:20 H2O/N2 atmosphere,
T = 600, P = 11 atm)] ................................................................................... 324 Figure 7.16: Integration of the CLP in a natural gas reforming system ...................... 326 Figure 7.17: Detailed schematic for H2 production from natural gas .......................... 327 Figure 7.18: Conventional CTL plant .......................................................................... 328 Figure 7.19: Integration of the CLP in a CTL plant in two configurations ................. 329 Figure 7.20: Integration of the CLP in a CTL plant – configuration 1 ........................ 330 Figure 7.21: Integration of the CLP in a CTL plant – configuration 2 ........................ 331 Figure 8.1: Standard deviation of pressure in the fluidized bed .................................. 342 xxix
Figure 8.2 (a): Schematic diagram of the cold flow model for the CLP ..................... 343 Figure 8.2 (b): Snapshot of the cold flow model for the CLP ..................................... 344 Figure 8.3: Cold flow model for the hydrator.............................................................. 345 Figure 8.4: Schematic of the subpilot scale unit being constructed at OSU for testing
the Calcium Looping Concept for H2 production ......................................... 346 Figure 8.5: Sorbent hopper and screw feeder .............................................................. 347 Figure 8.6: Water cooled heat exchanger .................................................................... 348 Figure A.1: Schematic of a conventional gasification plant for the cogeneration of H2
and electricity ............................................................................................... 372 Figure A.2: Input-output diagram for the conventional coal to H2 process................. 373 Figure A.3: Schematic of a CLP plant for the cogeneration of H2 and electricity ...... 374 Figure A.4: Flow sheet developed for the CLP using ASPEN plus simulator ............ 375 Figure A.5: Input-output diagram for the coal to H2 process using the CLP .............. 376 xxx
CHAPTER 1
INTRODUCTION
The world energy demand, as shown in Figure 1.1, is projected to increase by
40% at a rate of 1.5% per year from 2007 to 2030 (EIA, 2009). Although the energy
generation from renewable resources is projected to grow, as illustrated in Figure 1.2,
fossil fuels are still projected to contribute a major portion of the energy needs in the
near future (EIA, 2009). A growing need for the reduction in anthropogenic carbon
dioxide (CO2) emission has led to a global push towards the development of efficient,
economical, and reliable carbon capture and sequestration technologies (CCS) for
application to fossil fuel based power plants. The implementation of CO2 capture in
fossil fuel based systems could be through post-combustion capture, oxy-combustion
and pre-combustion capture as illustrated in Figure 1.3.
Post-combustion capture technology involves the combustion of coal or natural
gas to produce hot flue gas which is used to generate steam. The CO2 from the flue gas
is then captured. The capture of CO2 from flue gas results in a large increase in
parasitic energy and cost of electricity (COE) due to the large volumes of flue gas and
the low concentration of CO2 (13-14%) for coal combustion and 3-4% for natural gas
1
combustion). In oxy-combustion, the fuel is burnt in oxygen and recycled flue gas, to
produce a concentrated stream containing CO2 and steam which is then dried,
compressed and transported for sequestration. Although oxy-combustion obviates the
need for a separate CO2 capture stage, it requires an Air Separation Unit (ASU) which
is energy intensive and expensive. Pre-combustion capture involves the gasification of
coal or the reforming of natural gas to produce syngas. The syngas is then cleaned and
sent to shift reactors to convert the carbon monoxide (CO) to hydrogen (H2) and CO2 in
the presence of steam. Downstream of the shift reactors, the CO2 is removed using
solvents like amines, rectisol, selexol, etc., and the H2 stream is further purified in a
Pressure Swing Adsorber (PSA) for high H2 purity applications.
The application of CCS to gasification systems has been found to be more
efficient and economical when compared to CCS for post-combustion systems. It has
been estimated that with the implementation of CCS using solvent based systems, the
increase in the COE for an Integrated Gasification Combined Cycle (IGCC) will be 25
to 40 % while that for Pulverized Coal (PC) boilers will be 60 to 85%.(MIT, 2007) In a
carbon constrained scenario, it has been estimated that the cost of a super critical PC
boiler will be $2140/KWe while that of an IGCC will be $1890/KWe (MIT, 2007). In
addition to being more economical and efficient, gasification is also very versatile and
capable of producing H2 and liquid fuels in addition to electricity.
Several options are being investigated for the implementation of CCS on precombustion systems including using solvents, sorbents, membrane and chemical
2
looping processes. The calcium looping process (CLP) which is a calcium sorbent
based chemical looping process, has the potential to reduce the cost and increase the
efficiency of CCS implementation. (Abanades et al, 2007, Ramkumar et al, 2009) In
this study, the CLP concept and application to various feed streams (syngas, natural gas
and other hydrocarbons) has been studied using thermodynamic analysis, lab, bench
and subpilot scale experimental studies and system analysis and preliminary process
economics.
Chapter 2 gives an overview of the calcium sorbent based processes that have
been developed for the enhanced conversion of fossil fuels to hydrogen and electricity
with simultaneous CO2 capture. In Chapter 3, the study conducted at the Ohio State
University on calcium sorbent reactivity and recyclability is detailed. Sorbent
modification, pretreatment and reactivation methods to improve recyclability have been
described. Sorbent reactivation by hydration has been found to be very effective in
maintaining the sorbent reactivity over multiple cycles in the lab, bench and subpilot
scale investigations.
Chapter 4 describes the thermodynamics and experimental analysis of the
application of the CLP to catalytic H2 production in the presence of a water gas shift
catalyst with insitu CO2 and sulfur capture from syngas. The effect of different water
gas shift catalysts and process conditions on the purity of H2 is discussed.
3
Chapter 5 describes the non catalytic production of H2 by the CLP. The process
conditions for which the water gas shift catalyst can be eliminated without a decrease
in H2 purity have been identified. The application of the CLP to syngas produced from
a gasification system, steam methane reforming, partial oxidation and autothermal
reforming process has been discussed.
Chapter 6 discusses the system analysis and techno-economical analysis
conducted for the production of H2 from gasifier derived syngas. Sensitivity analysis
for the effect of various process parameters on system efficiency has been conducted.
Techno-economic analysis predict that the CLP has a potential to reduce the cost of H2
and electricity production from gasifier syngas.
Chapter 7 describes the application of the CLP to H2 and electricity production
from hydrocarbons. The CLP aids in combining several unit operations including
steam reforming of the hydrocarbons, water gas shift reaction and CO2 capture in a
single reactor. The integration of the CLP enhanced steam reforming process to H2 and
electricity production from natural gas and to a coal to liquid fuel (CTL) process has
been described.
Chapter 8 describes the scaleup of the CLP for H2 production from bench to
subpilot scale at the Ohio State University. Cold flow tests have been conducted and a
subpilot scale unit has been designed based on the information discussed in the
4
previous chapters. The subpilot scale unit is currently under construction. Chapter 8
also provides recommendations for future work.
5
Figure 1.1: Historical data and projections of the world energy consumptions till 2030
(EIA, 2009).
6
Figure 1.2: Projections of the world energy supply by different fuel types including
fossils fuels and renewable (EIA, 2009).
7
Post Combustion
Coal
Or
Natural Gas
Steam
Generator
Boiler
Air
CO2
Capture
Electricity
Oxy Combustion
Coal
Or
Natural Gas
CO2 free Flue Gas
Electricity
CO2 Compression,
Transportation
and Sequestration
Steam
Generator
Boiler
ASU
Coal
Gasifier
Gas
Cleanup
WGSR
CO2
capture
or
Natural
Gas
Reformer
Gas
Turbine
PSA
FT Reactor
Electricity
Steam
Turbine
Pre Combustion
Hydrogen
•Ammonia Synthesis Liquid
Fuels
•Hydrogenation
•Other chemicals
N2 and Steam synthesis
Figure 1.3: Implementation of Carbon Capture and Sequestration (CCS) in fossil fuel
based power plants.
8
CHAPTER 2
LITERATURE REVIEW: PROCESSES FOR ENHANCED H2 PRODUCTION
WITH CO2 CAPTURE
2.1 INTRODUCTION
Hydrogen can be produced conventionally from coal by the gasification
process, natural gas by the steam methane reforming process and higher hydrocarbons
by the partial oxidation process. In a typical coal gasification system, the coal is fed
along with steam and/or oxygen to the gasifier to produce syngas. The syngas is then
cooled using a gas cooler or a water quench. The quench system also provides the
excess steam required for the water gas shift reaction.(Holt, 2005, MIT, 2007)) While
higher temperatures enhance the kinetics of the water gas shift reaction, the equilibrium
limitation of the water gas shift reaction adversely affects H2 production and the H2
yield falls with rising temperature. Hence, a high steam: CO (S:C) ratio is required to
enhance CO conversion and the consequent H2 yield. The S:C ratio required at 550 oC
can be as high as 50 in a single-stage operation or 7.5 for a more expensive dual-stage
process to obtain 99.5 % pure H2. (Haussinger et al, 2000) Numerous research studies
have focused on the development of low temperature catalysts to improve H2
9
production. (Haussinger et al, 2000) Commercially, the dual stage sweet water gas shift
reaction is carried out in series, with a HTS (300-450 oC) stage containing iron oxide
catalyst and a LTS (180-270 oC) stage containing copper catalyst. (Loyd et al, 1996)
The commercial iron oxide catalyst has a sulfur tolerance of about several hundred
ppms while the copper catalyst has a lower tolerance to sulfur and chloride impurities.
(Haussinger et al) Hence syngas clean up is required upstream of the shift reactors
which is achieved in conventional scrubbing towers using physical solvents like selexol
or, rectisol or chemical solvents like amine based solvents. This low temperature
syngas cleanup process is energy intensive due to the gas cooling and reheating
requirements. In a sour gas shift system, where the sulfur content of synthesis gas is
greater than 1000 ppm, a sulfided catalyst is used in a series of reactors at a
temperature of 250–500 °C and the desulfurization unit is located downstream of the
water gas shift reactors.(Loyd et al, 1996, Hiller et al, 2007) After the shift reaction, the
syngas is subjected to scrubbing using solvents to remove the CO2 and is sent to the
PSA unit to produce a pure stream of H2. The tail gas from the PSA unit is then used as
fuel for power generation.
Several methods to enhance the purity of H2 with the simultaneous separation
of CO2 have been cited in literature. A slight advancement in the commercial method
of H2 production has been to remove the CO2 from the reaction mixture between the
two stages of the shift reaction. However solvents operate at ambient temperatures and
this method involves severe energy penalties due to cooling and reheating of the
10
reaction gas mixture.
An effective technique to shift the water gas shift reaction to the right for
enhanced H2 generation has been to remove H2 from the reaction mixture. This concept
has led to the development of H2 separation membranes. Kreutz et al, 2002 have
described the integration of these membranes in a commercial coal gasification unit.
(Kreutz et al, 2002) The syngas produced from the gasifier is shifted at a high
temperature over a STC followed by a water gas shift H2 membrane reactor which aids
in producing more H2 and separating it from the gas mixture. (Kreutz et al, 2002)
However, ceramic membranes have a very low H2 permeability and the intermediate
temperature composites inspite of having a high H2 flux are difficult to fabricate and
are very susceptible to poisoning. The cermet membranes are superior to the other two
classes of membranes but again they are susceptible to poisoning and are expensive.
(Roark et al, 2002) Donghao Ma and Carl R. F. Lund (2003) have reported the
investigation of a Pd membrane reactor system packed with HTS catalyst. (Ma and
Lund, 2003) For optimum performance these reactors require 2 stages and a S:C ratio
of 3. These reactors also suffer from inhibition effects of CO2, which reduces the yield
of H2 from 90% to 50%. (Ma and Lund, 2003)
In addition, membranes cannot
completely remove H2 from the mixture and suffer from a considerable pressure drop
across them. (Roark et al, 2002) Any remaining H2 in the main stream would dilute the
CO2 and would lead to poor process economics.
High temperature CO2 membranes have been developed which operate in the
11
same temperature range as that of the water gas shift reaction. Although polymeric
membranes for the removal of CO2 from H2 have been found to have several
advantages like simplicity of operation, high energy efficiency and lower cost, most
polymers have a poor H2/CO selectivity. Hence they are not very effective in shifting
the equilibrium of the water gas shift reaction and producing high purity H2. (Chung et
al, 2006)
An alternative concept to drive the water gas shift reaction forward has been to
remove the CO2 from the reaction mixture using solid sorbents which either physisorb
or react with the CO2 in the water gas shift reactor. The separation of CO2 from the
reaction mixture at high temperatures removes the equilibrium constraint of the water
gas shift reaction and enhances H2 production. Sorbents that operate at higher
temperatures are beneficial to the process as the water gas shift reaction has superior
kinetics due to the high temperature and enhanced thermodynamic extent due to the
removal of CO2 from the product gas stream. Faster rates of reaction and larger CO2
capture capacities allow the use of smaller reactors and require a smaller amount of
solids circulation through the system. CaO has a high CO2 capture capacity and
removes CO2 to ppm levels at a high temperature of 600 ºC making it one of the most
suitable sorbents for this application.(Gupta and Fan, 2002)
The concept of utilizing CaO for CO2 capture has existed for well over a
century. It was first introduced by DuMotay and Marechal in 1869 for enhancing the
gasification of coal (Squires, 1967) and followed by CONSOL’s CO2 Acceptor process
12
(Curran et al, 1967) a century later when this concept was tested in a 40 t/day plant. A
variation of this process, the HyPr-RING process, (Lin et al, 2005, Lin et al, 2002) was
developed in Japan for the production of H2 at high pressures. Several other processes
have also been developed to enhance H2 production using calcium based sorbents such
as the ZECA (Ziock et al, 2001), Alstom (Andrus, 2006) and GE process (Rizeq et al,
2001). A detailed description of these processes is provided in the following sections.
2.2 CO2 ACCEPTOR PROCESS
The CO2 Acceptor process was developed by the Consolidation Coal Company
and later Conoco Coal Development Company (Dobbyn et al., 1978). The American
Gas Association, U. S. Department of Interior, and Energy Research and Development
Administration were among the major sponsors for the development of this process.
The CO2 Acceptor process was designed to produce synthetic pipeline gas from
pulverized lignite or sub-bituminous coal. There are five main operational blocks
comprising this process, i.e. the feedstock preparation block, the gas synthesis block,
the gas cleanup block, the methanation block, and the utility block. Figure 2.1 shows
the schematic of the gas synthesis block, where Ca-based sorbent circulating between
two fluidized bed reactors – a gasifier and a regenerator operating at high pressures of
~11 atms(150 psig) aids in the conversion of coal to H2. (Curran et al, 1967)
In the reactor system of the gas synthesis block shown in Figure 2.1, preheated
coal is ground to ~150 μm before it is fed to the gasifier. The gasifier is a fluidized bed
13
with steam as the fluidizing gas. It is operated at 800 ~ 850 ºC and 10 atm. The
relatively low temperatures in the gasifier enables the use of high sodium coal which
tends to fuse forming silica and alumina solid aggregates at high temperatures (>870
ºC). The “acceptor” is the calcined limestone or dolomite sorbent, which is fed at the
top of the gasifier. The sorbent reacts with CO2 produced from the combined steamcarbon reaction and water-gas shift reaction. The exothermic carbonation reaction of
the sorbent provides the heat for the endothermic steam-carbon reaction and also drives
the water-gas shift reaction towards the forward direction, thereby increasing the H2
content in the product gas. The major reactions that occur in the gasifier include:
Steam Carbon Reaction:
C + H2O Æ CO+H2
(2.1)
Water Gas Shift Reaction:
CO + H2O Æ CO2 + H2
(2.2)
Carbonation:
CaO + CO2 Æ CaCO3
(2.3)
CaO.MgO + CO2 Æ CaCO3.MgO
(2.4)
The H2 rich gas obtained from the gasifier, containing ~20% CO and CO2, is
subsequently quenched, purified, methanated, and transported by pipelines. The spent
sorbent discharged from the bottom of the gasifier is then fed to the regenerator where
the spent sorbent is regenerated at 1010 ºC and 10 atm through the calcination reaction:
Calcination:
CaCO3 Æ CaO + CO2
(2.5)
The heat for calcination is provided by combusting the residual char, which is
discharged from the gasifier to the regenerator. The regenerator is also a fluidized bed.
14
Air is used as both fluidizing and oxidizing gas for the regenerator. The regenerated
sorbent is sent back to the gasifier to complete the loop. The CO2 containing flue gas is
generated in the regenerator.
The CO2 Acceptor process was proved to be technically feasible after being
successfully tested in a pilot plant facility located at Rapid City, South Dakota. The
designed capacity of the pilot plant was 40 ton/day. Over a period of six years that
ended in 1977, an accumulative operation of 13,000 hours was achieved with the
longest continuous run of ~2300 hours. A total of 6500 tons of dry coal of various
ranks including three North Dakota lignites, one Texas lignite and three subbituminous coals were tested in the facility. As noted, the gas synthesis reactor system,
which produces H2 rich gas from coal with the aid of the calcium based CO2 sorbent or
“acceptor”, characterizes the key innovation of the CO2 Acceptor process. The detailed
description of a pilot scale gas synthesis system, which includes the regenerator and the
gasifier, is given below.
Regenerator
The regenerator decomposes the spent sorbent from the gasifier via coal char
combustion. It was estimated that 2.04 – 2.34 MWth was released by char combustion
in the pilot regenerator. The fluidized bed regenerator provides some, though not
complete, mixing of the char and the spent sorbent. The heat released by char reactions
coupled with endothermic calcination leads to ~17ºC temperature gradients throughout
15
the fluidized bed regenerator. The feed to the regenerator contains calcium sulfide
(CaS), part of which will be oxidized to Calcium Sulfate (CaSO4) in the presence of
CO2 and O2 as shown by:
1/4 CaS +CO2 Æ1/4 CaSO4 +CO
(2.6)
1/4 CaS +1/2O2 Æ1/4 CaSO4
(2.7)
The concurrent presence of CaS and CaSO4 in the regenerator generates CaO
and SO2 as shown by:
1/4 CaS +3/4 CaSO4 ÆCaO + SO2
(2.8)
This reaction occurs through a series of intermediate steps. At temperatures above
955oC, a transient liquid of a eutectic mixture of CaSO4 and CaS is formed as an
intermediate which solidifies and is deposited on the regenerator walls. Thus, a
reducing environment with a CO concentration of 1 – 5% is maintained in the
regenerator in order to prevent the formation of the transient liquid. Consequently, the
gas exiting from the regenerator will contain a small amount of CO in the CO2 stream.
Further treatment of the CO containing exit stream will be needed. The temperature in
the regenerator is controlled by adjusting the air flow rate. During the regenerator
operation, nearly all the char fed to the regenerator is consumed. The coal ash is
entrained with the spent air and collected in the external cyclone-lock hopper located at
the gaseous product outlet of the regenerator. It was determined that 99% of the carbon
in coal was converted by the gasifier and the regenerator.
16
A minimum temperature of 471 ºC in the regenerator is required for immediate
combustion of char. At the start up, part of the flue gas from the regenerator, re-heated
at natural gas furnaces, was used to increase the temperature of the regenerator to ~538
°C in order to initiate char combustion. After the initiation of combustion, the heat
released from combustion increased the temperature of the regenerator. At steady state,
the regenerator was operated at 1010 °C. Under such a high temperature, spent sorbent
is decomposed, releasing CO2 in the calcination reaction. The regenerated sorbent is
discharged from the regenerator through two outlets: one outlet purges a predetermined
amount of sorbent while the other outlet discharges the remainder into the gasifier
where the regenerated hot sorbent is used for coal gasification. The heat from the hot
sorbent is partially used to balance the heat requirement of the endothermic steam
gasification reaction in the gasifier.
Gasifier
The gasifier is operated using a fluidized bed with continuous solids feeding.
The hot sorbent particles from the regenerator are fed to the upper part of the gasifier
while coal and steam are injected to the middle part and the lower part of the gasifier,
respectively. In the gasifier, the CaO sorbent is converted to CaCO3. The exothermic
heat of the carbonation reaction is used to compensate for the endothermic gasification
reaction. Examining the converted sorbent or spent sorbent settled at the bottom of the
gasifier before its transport by gravity to the engager pot, it is found that the extent of
conversion of the sorbent in the gasifier from CaO to CaCO3 is high. The spent sorbent
17
in the engager pot is then pneumatically transported back to the regenerator by air. The
fresh makeup sorbent is also provided to the engager pot.
The presence of the reaction products such as H2, CO, CO2, and methane (CH4)
was found to limit the rate of gasification (Dobbyn et al., 1978) as noted by the rate of
gasification initially at the lower part of the gasifier to be much larger than that at the
higher part due to the difference in the gas composition. For example, 62% of the char
was gasified at the bottom section of the bed while only 12% additional char was
gasified in the middle section. In order to create a more uniform reaction rate
throughout the gasifier, a portion of the product gas from the gasifier was recycled to
the bottom of the gasifier. Such a product gas recycling step assists in moderating the
rate of the reaction in the gasifier where solids are not well mixed. Further, the
presence of the recycled gases decreases the partial pressure of the steam and hence,
decreases the formation of calcium hydroxide (Ca(OH)2) in the gasifier. Thus, the
formation of the eutectic mixture of CaO- Ca(OH)2-CaCO3 is minimized. It was found
that lignite char was distinctively more reactive than sub-bituminous char.
Synthesis gas and synthetic natural gas are continuously produced in this
process. A typical synthesis gas composition obtained from the gasifier is given in
Table 2.1. The synthesis gas produced from the gasifier is used in the methanation
system, which is not shown in Figure 2.1, for the production of synthetic natural gas. A
typical composition of the synthetic natural gas obtained from the CO2 Acceptor
process is given in Table 2.2. As can be seen in Table 2.2 the heating value of the
18
synthetic natural gas obtained from the pilot plant exceeds 900 Btu/SCF
(33.5MJ/NM3). The pilot plant studies also examined the factors that affected the
activity of the sorbent and its environmental impact.
Some Findings from the Process Testing
The activity of the CO2 sorbent, which is expressed by the ratio of the weight of
CO2 absorbed to the weight of the fresh (unreacted) sorbent, was found to be the key
parameter to determining the technical feasibility of this process. The average acceptor
activity must exceed a certain level in order for the process to be in heat balance since
the system heat requirements are met by the sensible and reaction heat released by the
acceptor at a given CO2 removal rate. The minimum activity for the CO2 acceptor
process was found to be 0.26 for dolomite sorbent and 0.14 for limestone sorbent.
Through the pilot plant testing, an activity of 0.35 was achieved using the dolomite
sorbent, which exceeds the minimum activity requirement.
In order for the process to be economically attractive, the sorbent needs to
maintain a high activity with a minimal purge rate. It was observed that the activity of
the acceptor decreases as the number of carbonation-calcination cycles increase. Other
important factors that affect the activity of the acceptor include the acceptor residence
time in the gasifier and the reactor operating temperatures. The decrease in the acceptor
reactivity was attributed to the CaO crystalline growth. It was hypothesized that the
calcium atom is relatively mobile at the gasifier/regenerator operating conditions,
19
especially when the operating temperature is high (Dobbyn et al., 1978). The high
mobility of the calcium atom leads to fast CaO crystalline growth which forms a
“bridge” between closely placed CaO crystals during the carbonation reaction. The
crystals formed in this bridging effect are highly stable and have slow carbonation
reaction kinetics. Such CaO crystals remain intact during the calcination reaction and
tend to continuously grow in size and hence, reduce the gas diffusion rate. Thus, the
reactivity of the acceptor decreases over time, especially at high temperatures. The
crystalline growth effect was found to be more significant in the limestone acceptor
than in the dolomite acceptor. To increase the recyclability of the acceptor, two
approaches, i.e. acceptor reactivation and acceptor structure modifications, were tested
in the pilot scale facility. Satisfactory results for both approaches were reported.
Besides these attempts in improving the acceptor reactivity and recyclability, several
strategies were adopted to enhance the energy conversion efficiency of the CO2
Acceptor process. These strategies include the utilization of high pressure exhaust gas
for air compression and the recovery of heat from exhaust gas for steam generation.
Through the pilot testing, the metallurgical aspect of the reactor materials and its
feasibility in usage were determined.
2.3. HYPR-RING PROCESS
The H2 Production by Reaction Integrated Novel Gasification Process (HyPrRING) currently under development in Japan, is similar to the CO2 Acceptor process.
Both processes promote fuel conversion using CaO and/or Ca(OH)2 sorbents. While
20
the CO2 Acceptor process was aimed at synthetic natural gas production, the goal of
the HyPr-RING process is the production of high purity H2. (Lin, et al. 2004) The
HyPr-RING process comprises principally two units, i.e., gasifier and regenerator, as
shown in Figure 2.2.
Coal is introduced, along with CaO and steam, in the gasifier where the
following reactions take place:
CaO + H2O Æ Ca(OH)2
(2.9)
C + H2O Æ CO + H2
(2.1)
CO + H2O Æ CO2 + H2
(2.2)
CaO + CO2 Æ CaCO3
(2.3)
Ca(OH)2 + CO2 Æ CaCO3 + H2O
(2.10)
The solids mixture from the gasifier, which contains unreacted coal char, CaCO3,
Ca(OH)2 and ash, are fed into the regenerator where the following reactions take place:
C + O2 Æ CO2
(2.11)
CaCO3 Æ CaO + CO2
(2.5)
CaO sorbent and ash from the regenerator are then recycled back to the gasifier.
Before re-entering the gasifier, a portion of the solids mixture is discharged and the
fresh makeup is added. This purge step helps prevent the ash accumulation and
maintain the sorbent reactivity.
21
The concentration of H2S, NH3 and HCN in the syngas stream were reported to
be 2.2 ppm, ~0 ppm and 3.2 ppm respectively in the pilot unit located at Japan’s Coal
Energy Center with a coal feeding rate of 3.5 kg/hr.
The HyPr-RING process has been extensively studied for H2 production. A
comparison of the quantity of the synthesis gas produced from the pyrolysis of coal
mixed with CaO and Ca(OH)2 revealed that the extent of pyrolysis of coal is improved
in the order, from high to low, of coal/Ca(OH)2 offering the best extent of reaction,
followed by a coal/CaO mixture, and lastly with pure coal at pressures of 30-60 atm.
As this demonstrates, there are advantages of using Ca(OH)2 as a sorbent. Water is
supplied at a high temperature so that calcium is in the form of Ca(OH)2 and aids in the
reforming reaction. Further, CaO from the decomposition of Ca(OH)2 enhances the
pyrolysis reaction by the removal of CO2 from the product gases. CaO also has a
catalytic effect on the decomposition of tar, which further increases the gaseous
product yield. The composition of H2 in the gaseous products was found to be the
highest in the temperature range of 650-700 °C. This temperature range conforms to
the optimal temperature for the combined water gas shift and carbonation reaction. An
increase in pressure results in an increase in the H2 purity since the water gas shift and
carbonation reactions are kinetically favored at higher pressures (Lin et al, 2003).
Studies of H2 generation from a mixture of pulverized coal and CaO with high
pressure steam in a fixed bed reactor revealed that at a temperature of 700 °C, the
hydration of CaO occurs at a steam partial pressure higher than 30 atm. The yield of H2
22
was found to be doubled with an increase in temperatures from 650 to 700 °C and the
yield of H2 increases by 1.5 times with an increase in the total pressures from 10-60
atm and the steam partial pressures from 7-42 atm.(Lin et al, 2002). Gasification using
pellets containing a mixture of coal and CaO in a fixed bed reactor revealed that
although there is a decrease in the volume of the product gas, the composition of the
product gas from pellet gasification is similar to that from gasification using the
pulverized coal and CaO mixture. Further, in the gasifier, pellets retain their size and
morphology at a gasification temperature of 650 °C. At 700 °C, however, the pellets
are separated into two distinct parts: a dark part containing carbon and a white part
containing a mixture of CaO, Ca(OH)2 and CaCO3 which form a eutectic melting
mixture of solids. Recycling the CaO pellets between the reaction and the regeneration
results in a constant H2 yield over 4 cycles beyond which CaO is significantly
deactivated due to the deposition of ash and inerts on the surface of the pellet. (Lin et
al, 2004).
Studies of the HyPr-RING process in a fluidized bed reactor at 650 °C and 50
atm revealed that the hydration of CaO and the carbonation of the Ca(OH)2 occurred in
series, resulting in a gaseous product containing 76% H2, 17% CH4, 2%C2H4, 3% C2H6
and 2% CO2. As the time scale for the combined hydration, water gas shift and
carbonation reactions is 1-2 sec, which is much shorter than that for the gasification
reaction, CO in the product gases is completely converted to CO2 and almost all CO2 is
removed by CaO or Ca(OH)2. In the continuous flow gasifier, the increase in the rate
23
of the combined water gas shift, reforming, and carbonation reactions is higher than the
increase in the rate of the methanation reaction when the total and steam partial
pressures are increased, as noted earlier. This behavior leads to the enhancement of H2
production and the inhibition of CH4 formation.(Lin et al, 2004).
The carbon
conversion was found to be 60% near the entrance area of the fluidized bed reactor and
80% at the outlet of the reactor. The eutectic melting of the Ca(OH)2, CaCO3 and CaO
mixture, which occurs at 700 °C in the fixed bed experiments with the pelletized coal
and CaO, was not present in the fluidized bed at 650 °C. However, in the fluidized bed
reactor, even at a low temperature of 650 °C, particle growth occurs due to
crystallization and cohesion of calcium compounds.(Lin et al, 2006). Studies of the
effect of various sorbents including CaCO3, CaOSiO2, MgO, SnO and Fe2O3 on H2
production indicate that high purity H2 is obtained only with CaCO3 and CaOSiO2
sorbents and CO2 cannot completely be removed from the product gas using the other
sorbents (Lin et al, 2005). For different Ca-based sorbents, the rate of hydration was
found to decrease with an increase in the CaO content. Further, the initial rate of
hydration increases with an increase in the surface area of the sorbent while the final
rate increases with an increase in the porosity. (Lin et al, 2008).
Studies of the regeneration of the spent calcium sorbent in a 100% CO2
environment and the reactivity of the calcined sorbent for the hydration and
carbonation reactions reveal that for a residence time of 70 min for the calcination
sorbent in a fluidized bed reactor, 73% of CaCO3 calcined at 920 °C , 95% calcined at
24
1020 °C and almost 100% above 1020 °C. The rates of the hydration and carbonation
reactions decrease with an increase in the calcination temperature. Further, the extent
of carbonation of CaO decreases from 60% at 950°C to 52% at 1000°C, and 40% at
1020°C. Thus, to improve the extent of carbonation, the hydration of CaO is desired in
order to improve the porosity of the sorbent (Yin et al, 2007). Calcination in the
presence of steam yields a sorbent that requires only half the time for hydration,
compared to a sorbent obtained from calcination in the presence of 100% CO2. The
extent of carbonation of completely calcined CaO is also increased from 40% for 100%
CO2 calcination to 70% for steam calcination (0.4 atm CO2 partial pressure and 30 atm
total pressure) (Yin et al, 2008). Thus, by the combination of steam calcination and
hydration, the sorbent loading in the process can be significantly reduced.
For generating H2 with high purity for fuel cell applications, extensive cleaning
of H2S, CH4, and other pollutants or byproducts from the H2 stream will be necessary.
The energy efficiency, defined as the high heating value (HHV) of the H2 produced
divided by the HHV of the coal converted, for this process was reported to be 77% (Lin
et al, 2005). It can be noted that the difference between the CO2 Acceptor process and
the HyPr-RING process lies in the gasifier operating conditions. Comparing the
operating conditions of the gasifier used in the CO2 Acceptor process, i.e., 800 – 850
ºC and 10 atm, the operating conditions of the gasifier used in the HyPr-RING process
have a lower operating temperature (650 ºC) and a higher operating pressure (30 atm).
The lower temperature and higher pressure in the HyPr-RING process gasifier
25
thermodynamically favors the carbonation reaction, thereby further enhancing the H2
production. Moreover, an excess of steam is used in the HyPr-RING process to
enhance the reactivity of the CaO sorbent by refreshing the pore structure of the
particles.
2.4. ZERO EMISSION COAL ALLIANCE (ZECA) PROCESS
The Zero Emission Coal Alliance Process, or ZECA process, was proposed by
Klaus Lackner and H. Ziock (ZECA Corporation, 2002). Figure 2.3 shows the
schematic diagram of the process. In this process, coal is first converted to methane by
reacting with H2 in a gasifier:
C + 2H2 Æ CH4
(2.12)
This hydrogasification step also produces light hydrocarbons. The CH4 and
light hydrocarbons are then sent to the reformer. Steam and CaO sorbent are introduced
to the reformer to convert the hydrocarbons into H2 via sorbent enhanced reforming
reactions:
CH4 + H2O Æ CO + 3H2
(2.13)
CO + H2O Æ CO2 + H2
(2.2)
CO2 + CaO Æ CaCO3
(2.3)
The H2 gas generated in the reformer is split into two streams: one stream is
recycled to the hydrogasifier and the other is sent to a solid oxide fuel cell (SOFC) for
26
power generation. The spent sorbent, consisting mainly of CaCO3, is regenerated in a
calciner. CO2 is readily separated in this step:
CaCO3 Æ CaO + CO2
(2.5)
The heat required for the calcination reaction is provided by the waste heat
from the solid oxide fuel cell system, which is operated using H2 from the reformer.
The fuel cell also generates steam, which is recycled back to the gasifier and the
reformer for CH4 and H2 generation. Stoichiometrically, to convert one mole of carbon,
2 moles of H2 gas are consumed in the gasifier and 4 moles of H2 gas are generated in
the reformer. Therefore, there is a net gain of 2 moles of H2 gas per mole of carbon
converted. The excess H2 stream is used to meet the process heat requirement and to
generate electricity.
2.5. ALSTOM HYBRID COMBUSTION-GASIFICATION PROCESS
In a typical configuration of the ALSTOM chemical looping process for H2
production, calcium based sorbents and bauxite ore are used to carry oxygen, CO2, and
heat in three loops. The first loop is the CaSO4-CaS loop in which coal is gasified using
CaSO4, an oxygen carrying agent, to produce CO. CO is then converted to CO2 and H2
by the water gas shift reaction. The CaS produced in this process is regenerated in air
to produce CaSO4 through an exothermic oxidation reaction. The second loop consists
of the CaO-CaCO3 loop in which the CaO sorbent is used to remove CO2 during the
water gas shift reaction, forming CaCO3 while producing a pure stream of H2. The
27
third loop is a heat transfer loop in which hot CaSO4 or bauxite is used to transfer the
heat from the exothermic CaS oxidation reaction to the calciner to support the
endothermic calcination of CaCO3. (Andrus et al, 2006)
2.6. FUEL-FLEXIBLE ADVANCED GASIFICATION-COMBUSTION PROCESS
The GE process comprises two loops, an oxygen transfer loop and a carbon
transfer loop, and involves three reactors. In the first reactor, coal is gasified to produce
CO and H2 along with CO2, which is constantly removed by the CaO sorbent. The
reacted CaCO3 product along with the unconverted char is then routed to the second
reactor where hot oxygen transfer material from the third reactor is reduced while
converting the char to CO2. The hot solids also provide heat for the calcination of
CaCO3. In the third reactor, the reduced oxygen transfer material is reoxidized,
releasing a considerable amount of heat that heats up the solids and generates steam for
power production. The GE process obtains a H2 concentration of only 80%. (Rizeq et
al, 2001)
In the processes discussed above, CO2 is removed in the gasifier by the CaO
sorbent. Brun-Tsekhovoi et al., 1988, Fan et al., 2007, Ortiz and Harrison, 2001, Han
and Harrison, 1994, Johnsen et al., 2006, Balasubramanian et al., 1999, Hufton et al.,
1999, and Akiti et al., 2004, have also applied CO2 removal by CaO to the removal of
CO2 and the production of H2 from syngas through the water-gas shift reaction and
from CH4 through the sorption-enhanced steam methane reforming reaction. Chapter 4,
28
5 and 6 describe the conversion of syngas to H2 in the presence of CaO sorbent.
Chapter 4 discusses the production of H2 in the presence of CaO sorbent and a water
gas shift catalyst while Chapter 5 discusses the non-catalytic H2 production in the
presence of CaO. Chapter 6 is a system analysis of the process. The conversion of CH4
to H2 in the presence of CaO sorbent is described in Chapter 7.
29
Gas Type
CH4
H2
CO
CO2
N2
H2O
HHV Btu/SCF ,
(MJ/NM3)
Percentage (%)
11.4
65.6
15.7
4.7
0.7
1.9
379.1 (14.1)
Table 2.1: A typical composition of the H2 rich synthesis gas from the gasifier
30
Gas Type
CH4
H2
CO
CO2
N2
HHV
Btu/SCF,
(MJ/NM3)
Percentage (%)
92.6
4.7
0.01
0.5
2.2
> 900 (33.5)
Table 2.2: A typical composition of the synthetic natural gas from the methanation
system
31
Flue Gas
Product Gas
Ash
Spent
Sorbent
CaO
1010ºC
Regenerator
823ºC
Gasifier
Coal Fuel
char
Lift gas
CaCO3
Sorbent
makeup
Steam
Engager Pot
Air
Figure 2.1: Schematic diagram of the reactor system in the gas synthesis block of the
CO2 Acceptor process (Dobbyn et al., 1978)
32
Water
H2
Water
CO2
Regenerator
Gasifier
O2
C/CaCO3/Ca(OH)2
Steam
CaO
Coal
Steam
Ash/
CaO
Steam
Turbine
Figure 2.2: Schematic diagram of the HyPr-RING process
33
N2
A.S.U.
Figure 2.3: Schematic of the ZECA process
34
CaO
H2O
CO + H2O + CaOÎ CaCO3 + H2
Loop 2
CaCO3
CaCO3 Î CaO + CO2
Hot Loop 3 Cold Bauxite
Bauxite
CO
CaCO3
CaSO4
4C + CaSO4 Î 4CO + CaS
Loop 1
CaS
Coal
Figure 2.4: Schematic of the ALSTOM process
35
CaS + 2O2 Î CaSO4
H2
Coal
CaCO3
Gasification
C + 2H2O ÎCO2 + 2H2 Loop 1
CaO + CO2 Î CaCO3
CaO
CO2
FeO
Regeneration
C + Fe2O3 Î CO2 + FeO Loop 2
CaCO3 Î CaO + CO2
Fe2O3
Steam
Steam
Figure 2.5: Schematic of the GE process
36
N2
Oxidation
4FeO + O2 Î 2Fe2O3
Air
CHAPTER 3
REACTIVITY AND RECYCLABILITY OF CALCIUM BASED SORBENTS
FOR CO2 CAPTURE
3.1 INTRODUCTION
The successful operation of the CLP is highly dependent on the performance of
the CaO particles for CO2 and sulfur capture. In the CLP, the sorbent participates in
several reactions in at least two reactors: the carbonation reactor and the calciner. The
carbonation reaction occurs in the temperature range of 500 to 750 ºC and the
calcination reaction occurs at higher temperatures. The reactivity of the sorbent over
multiple cycles is very important for the economics of the process since it affects the
size of the reactors and the amount of solid circulation and sorbent makeup. Some of
the major factors that affect the solid circulation and makeup are the reactivity of the
sorbent, the recyclability, which depends on the temperature and gas atmosphere of
calcination, the amount of sulfur in the feed gas and the extent of attrition of the
sorbent. In this chapter, the reactivity and recyclability of natural and synthetic sorbents
is investigated. In addition, the effect of realistic calcination conditions on sorbent
37
reactivity is explored and the effectiveness of sorbent reactivation by hydration is
determined on the bench and subpilot scale.
3.2 SORBENT REACTIVITY OVER MULTICYCLIC REACTIONS
The reaction between CaO and CO2 occurs in two distinct stages. The first
stage occurs rapidly and is kinetically controlled, while the second stage is slower and
diffusion controlled. For any commercial application, only the first stage of the
reaction should be considered in order to use a compact reactor for the removal of CO2.
Abanades et al studied the rate and the extent of the carbonation reaction and the
variation
of
these
parameters
with
multiple
carbonation
and
calcination
cycles.(Abanades and Alvarez, 2003) The CaO conversion at the end of the rapid
kinetically controlled regime is found to decay sharply for naturally occurring
limestone with an increase in the number of cycles. Although the initial decay is
smoother for dolomite and other modified sorbents, it is intrinsic to most sorbents used
in the CLP. In addition to the decay in CO2 capture capacity, dolomite and other
supported sorbents also have the disadvantage of carrying more inert material in the
loop thereby increasing the parasitic energy requirement of the regeneration process.
Since the cost of the supported and modified sorbents is also higher, their performance
over multiple cycles also needs to be significantly higher in order to compete with
natural limestone. The decay in lime conversion over multiple cycles has been reported
by numerous researchers including Curran et al, Shimizu et al, Silaban and Harrison,
Barker , and Aihara et al (Curran et al, 1967, Shimizu et al, 1999, Saliban and
38
Harrison, 1995, Barker, 1973, Aihara et al, 2001). Using these data, Abanades and
Alvarez concluded that the decay in conversion is dependent only on the number of
cycles and independent of the reaction times and conditions (Abanades and Alvarez,
2003). Using a simple relationship given in Equation (3.1), Abanades related the
conversion of lime for any given cycle number (xc,N) to fitted constants (f, b) and the
cycle number (N) as given by:
xc,N = fN+1 + b
(3.1)
where the fitted parameters f and b have a numerical value of .782 and .184,
respectively (Abanades, 2002). Taking into consideration the sorbent conversion decay
over multiple cycles, the kinetics of the reaction, and mass and energy flows, Abanades
developed Equation (3.2) to determine the maximum capture efficiency of CO2 in a
system containing a continuous purge of solids and a make up of fresh sorbent.
(Abanades, 2002)
E CO 2
⎡
⎤⎡
⎤
F
⎛F
⎞
1 + ⎛⎜ 0 ⎞⎟
⎢
⎥⎢ f ⎜ 0 F ⎟
F
R
R
⎝
⎠
⎝
⎠ + b⎥
⎥⎢
=⎢
⎥
⎢ ⎛ F0 ⎞ ⎛ FCO 2 ⎞ ⎥ ⎢ ⎛ F0 ⎞ + 1 − f
⎥
⎟ ⎥⎢⎜ F ⎟
⎢ ⎜⎝ FR ⎟⎠ + ⎜
F
R
⎝
⎠
⎦⎥
⎣
R
⎝
⎠⎦
⎣
(3.2)
where E CO 2 is the maximum obtainable efficiency, F0 is the fresh feed added to the
system (mol CaO/s); FR is the total amount of sorbent required to react with the CO2 in
the system (mol CaO/s); FCO 2 is the flow of CO2 (mol/s); and f and b are constants as
defined in Equation (3.1). b is the residual carbonation conversion due to the formation
39
of a product layer of carbonate inside the macropores in highly sintered sorbents. This
residual carbonation of the lime sorbent is beneficial as it aids in reducing the amount
of fresh sorbent to be added. From an economic standpoint, it is desirable to minimize
the ratios FR/ FCO 2 and F0/FR in order to minimize the energy required for calcination
and the amount of fresh sorbent required (Abanades, 2002). For FO and FR to be low,
the sorbent should have a high resistance to sintering.
The CaCO3 product layer formation and pore pluggage during carbonation and
the sintering of CaO during calcination are both attributed to the decay and
irreversibility of limestone. Abanades et al concluded that micropores contribute to the
fast stage of the carbonation reaction (Abanades and Alvarez, 2003). The fast reaction
stage ceases when the micropores connecting the crystal grains are plugged due to the
increase in the molar volume during the formation of CaCO3 from CaO, where CaCO3
has greater than twice the molar volume as CaO. In the larger pores (mesopores and
micropores), CaCO3 forms a layer on the CaO wall (Alvarez and Abanades, 2005).
Although the pore is sufficiently large to handle the increase in pore volume, the
resistance of CO2 diffusion through the CaCO3 layer dramatically increases. The
increased resistance forms the boundary between the two stages of carbonation.
Sintering of CaO during calcination over multiple cycles results in grain growth which
drastically reduces the CaO microporosity while increasing the mesoporosity. This
leads to a reduced fast carbonation reaction zone, and therefore, a decrease in CO2
capture capacity over multiple cycles (Abanades and Alvarez, 2002).
40
Sun et al also investigated the sintering mechanism of limestone with increasing
number of cycles and attributed sintering to be due to CO2 released during the
calcination process (Sun et al, 2007). They showed that the increase in the carbonation
time did not have any effect on the structure of the calcine as the calcination process
eliminates the changes caused by carbonation. However, an increase in the calcination
time resulted in a decrease in the pore volume for pores <220 nm (Sun et al, 2007).
Similar to the observation made by Abanades et al with the increase in the number of
cycles the pore volume decreased for pores < 220nm and consequently increased for
pores >220 nm (Abanadez and Alvarez, 2003). A sintering model has been developed
by Sun et al based on the packed bed model, shrinking core model and a modified
sintering kinetic model and the average CO2 conversion is given below ( Sun et al,
2007):
Xcarb = 1.07 (n+1)-0.49
(3.3)
To be commercially viable, the CaO sorbent must maintain its reactivity
towards CO2 over multiple cycles. Additives and processed sorbents have been
investigated, but these techniques undermine the main advantage of using natural
limestone, which is its low cost. Using natural limestone has its challenges, which must
be overcome.
The effect of doping CaO with NaCl and Na2CO3 has also been investigated in
a Thermo Gravimetric Analyzer (TGA) (Salvador et al, 2003). The addition of NaCl
41
increased the CO2 removal capacity of the sorbent to 40% over 13 cycles due to
favorable changes in the pore structure and surface area of the sorbent while the
addition of Na2CO3 did not have any effect on the extent of carbonation. When the
doped sorbents were tested in the fluidized bed, both NaCl and Na2CO3 caused a
decrease in the CO2 removal capacity of the CaO sorbent which might be attributed to
the coating of the surface of the sorbent, leading to pore blockage during the
calcination stage ( Salvador et al, 2003). This chapter focuses on modification and
reactivation methods that could be used to improve the reactivity of Ca-based sorbents
over multiple cycles.
3.3 SYNTHESIS OF HIGH REACTIVITY PRECIPITATED CALCIUM CARBONATE (PCC)
SORBENT
One method of improving the recyclability of Ca-based sorbents is to modify the
pore structure of the sorbent to increase the pore volume and surface area. Fan et al.
have developed a wet precipitation process to synthesize a high surface area
Precipitated Calcium Carbonate (PCC) (Fan et al, 1998, Fan and Gupta, 2006) The
PCC - CaO sorbent can achieve almost complete conversions (> 95%) due to presence
of mesopores (5-30 nm). PCC is synthesized by bubbling CO2 through a slurry of
Ca(OH)2. The surface properties of the sorbent are tailored by the addition of anionic
surfactants (Agnihotri et al, 1999, Ghosh-Dastidar et al, 1996, Wei et al., 1997). The
system reaches an optimum only when the zeta potential equals zero. The sorbent
42
optimization process results in production of a sorbent with a surface area of 60 m2/g
and a pore volume of 0.18 cc/g.
CaCO3 primarily occurs in three different polymorphs, each of which may have
multiple morphologies depending on the arrangement of the atoms and ions in the
crystal structure. These polymorphs are all present in nature as well as in synthesized
PCC and can be classified as calcite, aragonite and vaterite.
Calcite is the most stable polymorph and typically occurs in the triagonalrhombohedral (acute to obtuse), scalenohedral, tabular and prismatic morphologies.
Calcite crystals also display intergrowth or twinning to form fibrous, granular, lamellar
and compact structures. The rhombohedral and prismatic forms find applications in
paper coating and in polymer strength enhancing agents while the scalenohedral form
is used in paper filling due to its light scattering ability. Calcite exhibits a unique
property by which its solubility in water decreases with increasing temperature.
The aragonite polymorph has an orthorhombic morphology with needle shaped
or acicular crystals. Twinning of these crystals results in the formation of pseudohexagonal structures which could be in a columnar or fibrous matrix. Aragonite is
unstable at standard temperatures and pressures and eventually gets converted to calcite
over geological timescales. Aragonite also exhibits a higher density and solubility
than calcite. The needle shaped morphology of aragonite is beneficial for high gloss
43
paper coating applications as well as for strength enhancing additives in polymeric
materials.
Vaterite is the most unstable form of CaCO3 at ambient conditions and readily
gets converted to calcite (at lower temperatures) and aragonite (at higher temperatures
of 60 ºC) on exposure to water. Vaterite is usually spherical in shape and has a higher
solubility in water than the other polymorphs. The transformation of aragonite and
vaterite to calcite is accelerated with temperature (Yamaguchi and Murakawa, 1981)
Although PCC predominantly contains calcite, various factors in the synthesis
procedure like the extent of saturation of the Ca(OH)2 solution, pH of the solution,
concentration of CO2, etc dictate the type and size of its morphology. For example,
PCC synthesized from highly saturated aqueous Ca(OH)2 solutions contains aragonite
at 70 ºC and vaterite at 30 ºC (Wary and Daniels, 1957). Cizer et al (2008) have shown
that rhombohedral calcite crystals formed by the exposure of Ca(OH)2 to 100% CO2
are micrometer sized while that precipitated with 20% CO2 are submicrometer sized. In
addition, it was also found that during the initial stages of carbonation, when the
concentration of Ca2+ ions in the solution is greater than the concentration of CO3- ions,
a scalenohedral calcite is precipitated. The scalenohedral morphology gets transformed
into the rhombohedral form during the later stages of precipitation, when the CO3concentration in the solution is high.
44
3.4 PRETREATMENT OF CALCIUM BASED SORBENTS AND ADDITION OF SUPPORTS
The CO2 capture capacity of CaO obtained from several precursors was
determined for a single cycle and for multiple cycles in a TGA. The experimental
procedure and the details of the TGA setup can be obtained elsewhere (Iyer et. al,
2004).
3.4.1 Reactivity Testing of Ca-based Sorbents for CO2 Capture
Figure 3.1 illustrates the comparison in the CO2 capture capacity of the CaO
sorbent obtained from different precursors. The CO2 capture capacity has been defined
by the weight % capture which is the grams of CO2 removed/ gram of the CaO sorbent.
The Wt% capture of CaO obtained from limestone is 58%. It can be seen that the
weight % capture attained by the sorbent obtained from PCC powder is 74% when
compared to that of 60% attained by the Ca(OH)2 hydroxide sorbent and 20% attained
by the ground lime sorbent. In order to improve the strength of the PCC particles, the
PCC powder was pelletized into 2mm pellets and then ground to a size of 150 microns.
The CO2 capture capacity of the PCC pellets as well as the pelletized and broken
sorbent was also determined. The CO2 capture capacity of the pelletized and broken
PCC is almost the same (71%) as the PCC powder as shown in Figure 3.1. The PCC
pellet requires a very large residence time due to mass transfer resistance but reaches
45
the same final CO2 capture capacity of 71% as that of the PCC pelletized and broken
sorbent.
3.4.2 Recyclability of Natural, Pretreated and Supported Sorbents
Since the PCC powder as well as the PCC pelletized and broken sorbents have
very high CO2 capture capacities and require almost the same residence time for
carbonation, a multicyclic calcination and carbonation experiment was conducted on
the two sorbents. Figure 3.2 illustrates the comparison in the conversion attained by the
PCC powder and PCC pelletized and broken sorbents over 5 calcination and
carbonation cycles. It can be seen that during the first cycle the PCC powder and PCC
pelletized and broken sorbent both achieve the same CaO conversion. As the number
of cycles increases, the conversion falls for both sorbents due to sintering but the
conversion for the PCC powder sorbent falls more than the pelletized and broken PCC
sorbent. This shows that the sintering of the PCC sorbent could be reduced by
pelletizing the PCC sorbent and grinding it to the size range of 150 microns. This not
only improves the multicyclic conversion but also improves the strength of the sorbent.
To improve the recyclability of CaO for CO2 capture, several pretreatment
methods as well as the addition of metal oxide supports were evaluated in a TGA as
shown in Figure 3.3. The calcination of the sorbent precursor during the first cycle as
well as calcination of the spent sorbent every cycle was conducted at 700 ºC in pure
nitrogen. The wt% capture which is the weight of CO2 captured/ unit weight of the
46
sorbent was determined by reacting the calcined sorbent with CO2 in a feed gas
containing 10% CO2 and 90% N2 at 650 ºC. Linwood Carbonate (LC) is naturally
occurring limestone and its reactivity decreases from 59 wt% capture to 30% in 18
cycles. As the simplest method of pretreatment, a freshly calcined sample of LC was
hydrated with water to produce Linwood Hydrate (LH) at ambient temperature which
was then tested in the TGA for 18 carbonation and calcination cycles. As shown in
Figure 3.3, the first cycle reactivity of LH is lower than LC but LH performs better in
the cyclic tests. The reactivity of LH only decreases from 53% to 43% and LH has a
13% higher reactivity at the end of 18 cycles than LC. This shown that pretreatment of
the sorbent by hydration has the potential to improve the recyclability of calcium
sorbents. As described in the section above, PCC was synthesized by the addition of a
surfactant and the first cycle reactivity as well recyclability of PCC was found to be
higher than most of the other sorbents. PCC powder has a first cycle reactivity of 67%
and a wt% capture of 49% at the end of 18cycles. The effect of pretreatment of the LC
sorbent with formic and acetic acid was also investigated. The pretreatment of the LC
sorbent with formic and acetic acid prior to the first calcination step aids in increasing
the pore volume of the sorbents due to the formation of calcium acetate and calcium
formate which have a higher molar volume than CaCO3. The calcium formate
precursor has the highest first cycle reactivity of 70% but its reactivity decreases
steeply to 27% at the end of 18 cycles. Calcium acetate precursor has a lower first
cycle wt% capture than the calcium formate precursor of 63% but it has good
recyclability over multiple cycles similar to the LH precursor. In addition to sorbent
47
pretreatment, the effect of addition of metal oxides like MgO, SiO2 and Al2O3 to CaO
sorbent was also investigated. The synthetic sorbents were prepared in the laboratory
from LC. Calcined LC was mixed with water to make a slurry of Ca(OH)2. The metal
oxide was added to the slurry and CO2 was bubbled through the mixture to precipitate
out a mixture of CaCO3 and the metal oxide. The slurry was filtered and the solid
sorbent was then dried in an oven. The modified sorbent which is a mixture of CaCO3
and the metal oxide was then subjected to 18 carbonation and calcination cycles in the
TGA. All the sorbents with the metal oxide supports have a low first cycle wt% capture
due to the presence of the metal oxide which behaves as an inert during the carbonation
reaction. The wt% capture of the sorbent with MgO decreases from 51% to 41% while
that with SiO2 decreases from 47% to 41% over 18 cycles. The sorbent with Al2O3 has
a lower wt% capture over 18 cycles than the other two supported sorbents which
decreases from 44% to 34% over 18 cycles.
3.5 EFFECT OF REALISTIC CALCINATION CONDITIONS ON SORBENT REACTIVITY
In the previous sections, the multicyclic reactivity of sorbents was investigated
in a TGA with calcination conducted in ideal conditions (at a low temperature of 700
ºC in a pure stream of N2 carrier gas). In a CO2 constrained scenario, carrier gases like
N2 and air cannot be used in the calciner as they will mix with the CO2 produced by the
calcination of the spent sorbent. Hence in the absence of these carrier gases the
temperature of calcination is increased significantly. From thermodynamic analysis, a
minimum temperature of 890 ºC is required to calcine CaCO3 in a pure CO2
48
atmosphere. The temperature of calcination can be reduced by using a condensable gas
like steam as a carrier gas in the calciner. This will result in the production of a wet
CO2 stream which can then be dried and compressed for sequestration. The following
section describes the effect of realistic calcination conditions on the reactivity of the
sorbent in a bench scale calciner.
3.5.1 Experimental Methods
A detailed description of the bench scale rotary bed calciner is provided
elsewhere and consists of a stainless steel reactor tube rotating within a horizontal
furnace ( Sakadjian et al, 2007). The carrier gas consisting of pure CO2 or a mixture of
steam and CO2 was fed to the reactor and the outlet of the reactor was connected to a
CO2 analyzer. The sorbent was loaded in the reactor tube and the temperature was
increased to the calcination temperature. At the end of calcination, the CO2 capture
capacity of the sorbent was determined in a TGA apparatus procured from PerkinElmer
Corp. In the TGA a small sample of the sorbent (15-20 mg) was placed in a quartz boat
suspended from a platinum wire. The sorbent was brought to a reaction temperature of
650 °C in flowing nitrogen. Subsequently, the flow was switched to the reaction gas
stream which contained 10% CO2 and balance N2. The TGA records the increase in the
sample weight with respect to time, which signifies the CO2 capture by the sorbent.
The Wt% CO2 capture capacity of the sorbent was then determined as the grams of
CO2 captured *100 /gram of CaO sorbent.
49
3.5.2 Results and Discussion
Figure 3.4 illustrates the effect of realistic calcination conditions on the
reactivity of limestone sorbent. The Wt% CO2 capture of the original limestone sorbent
calcined in ideal conditions in a 100% nitrogen stream at 700 ºC in the TGA is 50%.
The realistic calcination of the limestone sorbent in the bench scale rotary bed calciner
at 900 ºC in a pure CO2 atmosphere produced CaO sorbent with a Wt% CO2 capture of
28%. Hence, the sorbent only retains half of its original reactivity to CO2 after a single
cycle of realistic calcination at 900 ºC. The effect of calcination in the presence of a
mixture of steam and CO2 at 900 oC was also determined on the reactivity of the
sorbent. Almost complete calcination of the sorbent was obtained in every case. It is
found that on calcination of the limestone sorbent in an atmosphere of 33% steam and
67% CO2, a CaO sorbent with a Wt% CO2 capture of 35% is obtained. A further
increase in the steam concentration to 50% resulted in the production of a more
reactive sorbent with a Wt% CO2 capture of 45%. Hence, the addition of steam aids in
reducing the extent of sintering of the sorbent and results in the production of a CaO
sorbent that is more reactive to CO2.
To determine the recyclability of the sorbent with steam calcination, a
multicyclic carbonation- calcination test was conducted and the results are shown in
Figure 3.6. The Wt% capture of the original sorbent calcined in ideal conditions at a
temperature of 700 ºC in pure nitrogen is 50%. During the cyclic testing, the sorbent
calcination was conducted in the bench scale rotary bed calciner while the carbonation
50
was conducted in a bench scale fixed bed reactor shown in Figure 3.5. The sorbent was
packed in the fixed bed reactor and carbonation was conducted in a 10% CO2/90% N2
stream at 650 ºC.
On calcining the limestone sorbent at 900 ºC in a 50%/50%
H2O/CO2 atmosphere, the Wt% capture of the sorbent reduces to 45%. During the
second and third cycles the reactivity further decreases to 30% and 25%. Hence
although steam calcination reduces the extent of sintering, the sorbent reactivity is not
maintained a constant and it continues to fall over multiple cycles.
3.6. SORBENT REACTIVATION BY HYDRATION –LAB SCALE TESTING
The use of sorbent hydration as a pretreatment method was investigated in the
TGA and found to improve the recyclability of the sorbent as shown in Figure 3.3. In
order to maintain the reactivity of the sorbent a constant over multiple cycles, sorbent
hydration was also investigated as a reactivation process. Sorbent reactivation by
hydration was included as a step in every carbonation- calcination cycle.
3.6.1 Experimental Methods
The effect of sorbent hydration with water at 25 oC, and steam at high
temperatures was investigated on sorbent reactivity. Limestone sorbent was calcined in
realistic calcination conditions at 1000 oC in a 100% CO2 atmosphere in the bench
scale rotary bed calciner described earlier (Sakadjian et al, 2007). At the end of
calcination, the sorbent was hydrated. Water hydration was conducted by spraying
water at 25 oC on the sorbent with vigorous stirring. Steam hydration was conducted in
51
the bench scale reactor shown in Figure 3.5 in a 20% nitrogen and 80% steam
atmosphere. At the end of hydration, the CO2 capture capacity of the sorbent was
determined in the TGA. A small sample of the sorbent (15-20 mg) was placed in the
quartz boat suspended from the platinum wire. The hydrated sorbent was brought to a
reaction temperature of 650 °C in flowing nitrogen. Complete dehydration of the
sorbent occurred by the time the sample was heated to 650 oC. Subsequently, the flow
was switched to the reaction gas stream containing 10% CO2 and balance N2. The TGA
records the increase in the sample weight with respect to time, which signifies the CO2
capture by the sorbent. The Wt% CO2 capture capacity of the sorbent was then
determined as the grams of CO2 captured *100 /gram of CaO sorbent.
3.6.2 Results and Discussion
Figure 3.7 illustrates the effect of water hydration (at 25 oC) and steam
hydration (at 150 oC and 500 ºC) on the CO2 capture capacity of the sorbent. The
original limestone sorbent calcined in the TGA in the presence of 100% nitrogen at 700
o
C has a Wt% CO2 capture capacity of 52%. Calcination of the limestone at 1000 oC in
pure CO2 in the bench scale calciner reduces its Wt% CO2 capture capacity to 20%. On
hydration of the sorbent with water at ambient temperature, the Wt% CO2 capture
capacity increases to >55%. Another method of hydration at atmospheric pressure, in
the presence of steam at 150 oC yielded in the production of a sorbent with 52 Wt%
CO2 capture. Steam hydration at atmospheric pressure and 500 ºC yielded in a sorbent
with a Wt% capture of 45%. While the extent of hydration obtained with water and
52
with steam at 150 ºC is greater than 95%, it is only 80% when the sorbent is hydrated
at 500 ºC. The reduction in extent of hydration might have resulted in the lower CO2
capture capacity observed for the sorbent hydrated at 500 ºC.
Figure 3.8 shows the effect of hydration at a high temperature of 600 oC for
total pressures ranging from 8 atms to 21 atms. The Wt% CO2 capture of the sorbent
calcined in 100% CO2 at 1000 oC increases from 20% to 45% by pressure hydration at
600 oC and 8 atms. The reactivity of the sorbent is found to decrease to a small extent
on increasing the pressure of hydration at 600 oC. Further investigation is required to
determine if this decrease in reactivity is due to an increase in the sintering of the
sorbent at high pressures.
Multicyclic calcination – hydration – carbonation tests were also conducted as
shown in Figure 3.9. The calcination was conducted at a temperature of 950 ºC. The
calcined sorbent was hydrated at 500 ºC and carbonated in a 10%CO2 / 90%N2 stream
at 650 ºC. The Wt% capture was calculated on the basis of the Ca(OH)2 in the sorbent
sample and not on the basis of entire solid sample weight as before. As illustrated in
Figure 3.9, the Wt% capture of the sorbent is maintained a constant over multiple
cycles. Hence sorbent hydration is a promising method of completely reactivating the
sorbent and improving its recyclability.
53
3.7 SUB-PILOT SCALE DEMONSTRATION OF REACTIVATION OF CALCIUM SORBENT
BY HYDRATION
The effectiveness of hydration on improving the recyclability of the sorbent
was tested in a subpilot scale demonstration for CO2 and SO2 capture from combustion
flue gas using the CLP process. Figure 3.10 illustrates the process flow diagram of the
calcium based CO2 and SO2 capture process from combustion flue gas. The CaO
sorbent or Ca(OH)2 sorbent is injected into the carbonator, which is an entrained bed
reactor, where it reacts with the CO2 and SO2 to form CaCO3 and CaSO4 at a high
temperature between 450 °C and 650 °C. Thermodynamic limitations prevent greater
than 90% CO2 removal from a coal combustion flue gas stream at temperatures greater
than 650 °C. The CaO sorbent could be obtained from such precursors as natural
limestone, hydrated lime, and reengineered and supported sorbents. The spent sorbent
mixture is then regenerated by calcining it at a high temperature between 850 °C-1300
°C where the CaCO3 decomposes to yield CaO and a pure, dry stream of CO2 when
calcined. The calciner could be a flash or entrained bed calciner, a fluidized bed or a
rotary kiln. While energy has to be provided for the calcination reaction, the
carbonation reaction is exothermic and releases high quality heat. Hence, a good
indirect heat integration strategy aids in reducing the parasitic energy consumption of
the process. With Ca(OH)2 as the sorbent, the CaO is further reactivated by hydration
and re-circulated to the carbonator, while the CO2 is compressed and transported for
sequestration. Since CaSO4 begins to decompose only at temperatures greater than
54
1450 °C, under the conditions experienced in the calciner, CaSO4 is stable and a small
amount of solids must be continuously purged out of the system to prevent complete
conversion of sorbent to CaSO4 .The amount of solid purge from the CLP will depend
on the amount of sulfur and flyash that are fed to the carbonator to prevent the
accumulation of inert solids in the process. Based on a preliminary economic analysis,
the purge percentage will be in the range of 2% to 10%.Thus the CLP process captures
CO2 in the flue gas stream and converts it into a concentrated sequestration ready CO2
stream. The CLP process is capable of capturing CO2 from flue gas streams produced
from various fuels including coal, oil, natural gas, biomass, etc,.
3.7.1 Experimental Methods for the 120 KWth Subpilot Scale Testing
The effectiveness of sorbent reactivation by hydration was tested in a 120KWth
subpilot scale demonstration of the calcium based CO2 capture process at the Ohio
State University (Wang et al, 2009). Coal was stored in a coal hopper, which is
connected to an underfeed stoker, provided by Babcock & Wilcox Co., Barberton,
OH,. The underfeed stoker has two Forced Draft (FD) fans that provide combustion air
to the stoker. Natural gas is connected to the inlet of the stoker for start-up and to
maintain gas temperature. The flue gas stream is transported through the ductwork via
an Induced Draft (ID) fan. Connected to the ductwork are a hopper and screw-feeder,
two sets of gas analyzers, multiple temperature monitoring ports, multiple pressure
measurement ports, a cyclone and a baghouse. Figure 3.11 illustrates a snapshot of the
sub-pilot scale facility.
55
A Schenck-Accurate mid-range volumetric hopper, is the main sorbent feeder
and is connected to the calciner feed inlet. An electrically-heated rotary calciner
manufactured by FEECO that has a maximum operating temperature of 980 °C is used
to calcine the spent calcium sorbent. The calcined sorbent was hydrated offline for the
data reported in this study and injected into the flue gas duct. Once injected into the
ductwork, the sorbent is entrained by the flue gas, and it simultaneously reacts with the
CO2 and SO2 present in the flue gas. At the end of the process, a Donaldson Torit
downflow baghouse is used to separate the solid sorbent from the CO2/SO2 free flue
gas which is emitted to the outside atmosphere.
To monitor the gas composition, two sets of gas analyzers are employed. One
set of gas analyzers is located upstream of the sorbent injection port and is used as the
baseline. The other set of gas analyzers is located downstream of the sorbent injection.
The difference, after correcting for air in leakage and other factors, between the two
measurements determines the percent removal. The gas analyzers are CAI 600
analyzers and continuously monitor the concentrations of CO2, SO2, and CO. In
addition, a CAI NOxygen analyzer monitors the upstream oxygen and nitrogen oxides
concentrations, while a Teledyne Analytical 3000P analyzer monitors the downstream
oxygen concentration. All data are continuously recorded via a data acquisition system.
Multiple Type K thermocouples continuously monitor the temperature throughout the
entire system to determine the proper operating temperature for the carbonation
56
reaction, which occurs at a reasonable rate between 450 °C and 650 °C (Gupta and Fan,
2002, Koji et al, 2003, Abanadez et al, 2003, Lee, 2004, Wong, 2007)
Prior to each experimental run, all analyzers were calibrated. The stoker was
heated and operated according to the start-up procedures. Once the flue gas
temperature at the sorbent injection location reached approximately 650 °C, which is
sufficiently high to allow both the carbonation and sulfation reaction to proceed at a
high rate and achieve greater than 90% removal of the CO2. The flowrate of the
sorbent was set via the control panel. After the sorbent reacted with the CO2 and SO2 in
the carbonator, the gas temperature was lowered and the spent sorbent was collected in
the bag house. To calcine the spent sorbent, the calciner temperature was set to 950 °C
and the calcined solids were then reactivated using offline hydration. The carbonationcalcination-hydration cycle was repeated.
At the completion of each experiment, solids from the baghouse were collected
and analyzed via a TGA.
3.7.2 Results and Discussion
The effect of sorbent reactivation by hydration was investigated on the %CO2
removal from the flue gas and on the Wt% capture of the sorbent over multiple cycles.
Figure 3.12 illustrates the carbonation -calcination cycles for pulverised lime over 3
cycles and the calcination-carbonation-hydration cycles for Ca(OH)2 over 4 cycles. A
maximum of only 50% CO2 removal is achieved in all the tests shown in Figure 3.12
57
since a substoichiometric calcium to carbon (Ca:C) mole ratio of 0.75 was used for
testing. Greater than 90% CO2 removal is achieved for a Ca:C ratio of 1.3 (Wang et al,
2009). For the pulverized lime sorbent the CO2 and SO2 capture in all the cycles was
conducted using commercially available pulverized lime sorbent obtained from
Greymont. As illustrated in Figure 3.12 the % CO2 removal decreases from 50% in the
first cycle to 20% in the second cycle. No CO2 was captured by the sorbent in the third
cycle. This shows that the sorbent sinters to a large extent in the system and loses all its
reactivity in three cycles. For the cycles with Ca(OH)2, the CO2 and SO2 capture in the
first cycle was conducted using commercially available Ca(OH)2 from Graymont. The
CO2 and SO2 capture in the remaining cycles was achieved using sorbent that was
reactivated by hydration. As can be seen from Figure 3.12, uniform CO2 removal was
achieved at a temperature of 625 ºC, a Ca:C ratio of 0.75 and a constant residence time.
Figure 3.13 illustrates the Wt% CO2 capture achieved by the Ca(OH)2 sorbent during
the 4 cycles shown in Figure 3.12. The Wt% capture of the sorbent is maintained a
constant at 52% over the 4 cycles. Hence hydration is very effective in maintaining
sorbent reactivity over multiple cycles even in the subpilot scale facility.
3.8 CONCLUSIONS
Among various reaction and process factors that are of importance to the CLP,
the reactivity and recyclability of the calcium based sorbent are vital. The nature of
CaO/CaCO3 sintering that has been observed during multicyclic operation could pose a
severe limitation to the commercialization of the process. In this Chapter, several
58
methods of improving the recyclability of CaO sorbents have been investigated
including sorbent pretreatment, modification by addition of supports and reactivation.
Reengineering the sorbent morphology by increasing the pore volume and surface area
of the precursor has been found to be effective in improving the reactivity and
recyclability of the sorbent. PCC sorbent that is synthesized from natural limestone has
an improved performance due to presence of a greater surface area and pore volume. A
similar improvement was observed by pretreating natural limestone with acetic acid
which also increases the pore volume and surface area. Hydration of the sorbent also
showed an improvement in sorbent performance. Although the addition of metal oxide
supports to natural CaO sorbent improves the recyclability of the sorbent it reduces the
amount of CO2 that can be captured by a certain amount of sorbent due to the presence
of the inert metal oxide. The effect calcination conditions of sorbent reactivity was also
investigated and the sorbent loses one third to half of it original reactivity in a single
cycle due to calcination at 950 ºC and 1000 ºC respectively. Modification of
calcination conditions by the addition of steam in the calciner was found to improve
the reactivity of the sorbent although a loss in reactivity was still observed over
multiple cycles. Hydration of the sorbent as a reactivation method after every
calcination cycle was found to be very effective in improving sorbent performance.
The Wt% capture of the sorbent was found to be constant at 50% during multicyclic
CO2 capture with sorbent hydration in every cycle in both bench scale and subpilot
scale tests.
59
PCC
PCC-Pelletised and Broken
PCC-Whole Pellet
Calcium Hydroxide
Ground Lime
Limestone
100
Weight% Capture
80
60
40
20
0
0
20
40
60
80
100
120
Time(sec)
Figure 3.1: Comparison in the CO2 capture capacity of CaO sorbents obtained from
different precursors. (Calcination conditions: T = 700 ºC, P = 1 atm, pure
N2 carrier gas; Carbonation conditions: T = 650 ºC, P = 1 atm, 10%
CO2/90% N2 feed gas)
60
Comparison in the multicyclic conversion of the PCC powder sorbent
and the PCC pelletised and crushed sorbent
1.0
PCC
PCC pelletised and crushed
Conversion
0.8
0.6
0.4
0.2
0.0
0
100
200
300
Time (sec)
Figure 3.2: Comparison in the multicyclic conversion of PCC powder sorbent PCC
pelletized and broken sorbent (Calcination conditions: T = 700 ºC, P = 1
atm, pure N2 carrier gas; Carbonation conditions: T = 650 ºC, P = 1 atm,
10% CO2/90% N2 feed gas)
61
CO2 Capture Capacity (%)
80
60
40
20
0
0
2
4
6
8
10
12
14
16
18
Number of Cycles
Formic Acid
Acetic Acid
MgO
SiO2
Al2O3
PCC
LC
LH
Figure 3.3: CO2 capture capacity of pretreated and supported Ca-based sorbents over
multiple carbonation –calcination cycles (Calcination conditions: T = 700
ºC, P = 1 atm, pure N2 carrier gas; Carbonation conditions: T = 650 ºC, P
= 1 atm, 10% CO2/90% N2 feed gas)
62
60
Wt% Capture
50
40
30
20
10
0
Original
Sorbent
0%
33%
50%
Steam Concentration in Carrier Gas
Figure 3.4: Effect of steam concentration in the calcination carrier gas on the CO2
capture capacity of CaO sorbent (Calcination conditions: T = 900 ºC, P =
1atm)
63
Thermocouple
And
Pressure Guage
Steam Generator
Steam &
Gas Mixture
Gas
Gas
Mixture
Mixture
Water In
Sorbent
MFC
64
H2
Heated Steel
Tube Reactor
Hydrocarbon
Analyzer
Back Pressure
Regulator
Analyzers (CO,
CO2, H2, H2S)
Water Syringe
Pump
Heat
Exchanger
Water Trap
Figure 3.5: Simplified flow sheet of the bench scale fixed bed experimental setup
64
MFC
MFC
CO
MFC
CO2 Hydro
carbons
60
Wt% Capture
50
40
30
20
10
0
Original
Sorbent
1
2
3
Number of Cycles
Figure 3.6: Effect of steam calcination on multicyclic carbonation and calcination of
CaO sorbent (Calcination conditions: T = 900 ºC, P = 1 atm, carrier gas =
50%H2O/50% CO2; Carbonation conditions: T = 650 ºC, P = 1 atm, 10%
CO2/90% N2 feed gas)
65
60
150C
Wt % Capture
50
500C
40
30
20
10
0
Original
Sorbent
Calcined
Sorbent
Water
Hydration
Steam
Hydration
Figure 3.7: Effect of hydration conditions on sorbent reactivity
66
Steam
Hydration
60
Wt% Capture
50
600C
600C
600C
40
30
20
10
0
Calcined
Sorbent
8 atms
100
psig
11
150atms
psig
21
300atms
psig
Hydration Pressure
Figure 3.8: Effect of hydration pressure on sorbent reactivity (Hydration temperature =
600 ºC)
67
60
Wt% Capture
50
40
30
20
10
0
1
2
3
4
5
Cycles
Figure 3.9: Effect of steam hydration on sorbent reactivity over multiple calcinationhydration-carbonation cycles (Calcination conditions: T = 900 ºC, P = 1
atm, carrier gas = pure CO2; Carbonation conditions: T = 650 ºC, P = 1
atm, 10% CO2/90% N2 feed gas, Hydration conditions: T = 500 ºC, P ~ 1
atm, 90% H2O/10% N2 feed gas)
68
Carbonator
Pure CO2
Co 2 Free
Flue Gas
Fresh
Sorbent
Spent
Sorbent Purge
69
Calciner
Flue
Gas
Calcined
sorbent
Bag
House
Boiler
Steam
Hydrator
Coal
Air
Figure 3.10: Process flow diagram of the CLP for CO2 and SO2 removal from combustion flue gas
69
Carbonation
Reactor
Figure 3.11: Snapshot of the sub-pilot scale facility of the CLP integrated with a coal
fired combustor.
70
Cyclic CO2 Removals
60
Calcium Hydroxide
Pulverized Ground Lime
% CO2 Removal
50
Hydration
Hydration
40
Hydration
30
20
10
0
0
1
2
3
4
Cycle Number
Figure 3.12: Effect of hydration on the % CO2 removed from the flue gas over multiple
cycles
71
Wt % Capture
50
40
30
20
10
0
10
21
32
43
Cycle Number
Figure 3.13: Wt.% CO2 capture achieved by the hydrated sorbent over multiple cycles
72
CHAPTER 4
ENHANCED CATALYTIC H2 PRODUCTION FROM SYNGAS
4.1 INTRODUCTION
This chapter describes the CLP for high purity H2 production in the presence of
a water gas shift catalyst from syngas. The CLP combines H2 production with CO2,
sulfur and chloride capture from the syngas stream in a single stage reactor.
Most H2 production processes reported in literature require a separate sulfur
clean up unit to prevent poisoning of the sorbent used for CO2 capture. Sulfur is present
in syngas in the form of H2S and carbonyl sulfide (COS). According to equilibrium
calculations, at temperatures below 1027C (1300K) which exists in the gasifier, all
sulfur radicals combine to form predominantly H2S which is close to 95% of the total
sulfur content and COS forms the other 5%. (Jazbec et al, 2004) There have been
studies conducted on the simultaneous calcination and sulfidation of calcium based
sorbents at temperatures higher than 600 ºC. ( De Diego et al, 2004) There have also
been studies on the sulfidation of CaCO3 in the presence of CO2 but the CO2 was used
only to maintain a high enough partial pressure to prevent the calcination of CaCO3
(Fenouil et al, 1994, Fenouil, 1995, Zevenhoven et al, 1998, De Diego et al, 1999).
However there is no mention of studies conducted on simultaneous CO2 and sulfur
73
capture integrated with H2 production in the literature. In the CLP described in the
sections below, simultaneous CO2 and H2S capture is achieved during the production
of H2.
4.2 CALCIUM LOOPING PROCESS (CLP) CONFIGURATION AND THERMODYNAMICS
Several options are being investigated for the implementation of CCS on coal
gasification systems including using solvents, sorbents, membrane and chemical
looping processes. The CLP which is a calcium sorbent based chemical looping
process, has the potential to reduce the cost and increase the efficiency of H2 and/or
electricity production from coal derived syngas by implementing the principles of
process intensification (Fan et al, 2007, Fan et al, 2008, Ramkumar et al, 2009,
Ramkumar et al, 2010). The CLP integrates the water-gas shift reaction with in-situ
CO2, sulfur, and halide removal at high temperatures in a single stage reactor. It utilizes
a high temperature regenerable CaO sorbent which in addition to capturing the CO2,
enhances the yield of H2 and simultaneously captures sulfur and halide impurities.
The advantages of the CLP include:
1) The simplification of the coal to H2 process by integration of the reaction and
separation steps. This results in a decrease in the number of process units and
combines the two staged water gas shift reactors (HTS and LTS), the CO2,
sulfur and halide capture units into a single stage reactor.
2) The enhancement in H2 yield at high temperatures due to elimination of the
74
equilibrium limitation of the water gas shift reaction.
3) The potential to reduce excess steam requirement for the water gas shift
reaction due to the enhanced thermodynamics of H2 production by the
combined water gas shift and carbonation reactions.
4) The potential to eliminate the requirement for water gas shift reaction catalyst
due to H2 production at high temperatures.
5) Although energy needs to be supplied for the endothermic calcination reaction,
the carbonation reaction is exothermic at high temperatures of 500 – 750oC
resulting in the production of high quality heat. By using a good strategy of heat
integration it is possible to achieve high process efficiencies.
6) The calcination reaction results in the production of a pure sequestration ready
CO2 stream.
A schematic of the CLP is shown in Figure 4.1. The CLP comprises the
carbonation reactor, the calciner and the hydrator. In the carbonation reactor highpurity H2 is produced while contaminant removal is achieved, in the calciner the
calcium sorbent is regenerated and a sequestration-ready CO2 stream is produced and
in the hydrator the sorbent is reactivated. Thermodynamic analyses are conducted for
the reactions occurring in each reactor using HSC Chemistry v 5.0 (Outokumpu
Research Oy, Finland) software. All the reactions shown in Figure 4.1 are found to be
thermodynamically spontaneous but reversible and the extent of each of these reactions
depends on the partial pressure of the respective gas species and the reaction
75
temperature. The following sections give a description of the three reactors using
thermodynamic analyses.
4.2.1 The Carbonation Reactor
The carbonation reactor comprises either a fluidized bed or an entrained flow
reactor that operates at pressures ranging from 1 to 30 atm and temperatures of 500-750
o
C. The exothermic heat released from the carbonation reactor can be used to generate
steam or electricity. In the carbonation reactor, the thermodynamic constraint of the
water gas shift reaction is overcome by the incessant removal of the CO2 product from
the reaction mixture, which enhances H2 production and obviates the need for excess
steam addition. This is achieved by the concurrent water gas shift reaction and
carbonation reaction of CaO to form CaCO3 thereby removing the CO2 product from
the reaction mixture. In addition, the CaO sorbent is also capable of reducing the
concentration of sulfur and halides in the outlet stream to ppm levels. The in-situ
removal of CO2 removes the equilibrium limitation of the water gas shift reaction
thereby obviating the need for excess steam addition. Thermodynamic analysis,
presented subsequently, predicts that the removal of H2S using CaO is inhibited by the
presence of steam. Since almost all the steam is consumed in the enhanced water gas
shift reaction, the removal of H2S is favored in the system. The reactions occurring in
the carbonation reactor are as follows:
Water gas shift reaction: CO + H2O Ù H2 +CO2
76
(ΔH = -41 kJ/mol)
(4.1)
Carbonation: CaO + CO2Ù CaCO3
(ΔH = -178 kJ/mol)
(4.2)
Sulfur capture (H2S) : CaO + H2S Ù CaS + H2O
(4.3)
Sulfur capture (COS) : CaO + COS Ù CaS + CO2
(4.4)
Halide capture(HCl) : CaO + 2HCl Ù CaCl2 +H2O
(4.5)
Thermodynamic analysis of reactions occurring in the carbonation reactor
The equilibrium constants for the water gas shift reaction and the combined
water gas shift and carbonation reaction for various temperatures are shown in Figure
4.2. The equilibrium constants are obtained using HSC Chemistry v 5.0 (Outokumpu
Research Oy, Finland). The equilibrium constant for the water gas shift reaction can be
defined as shown below:
Keq1
eq =
PH2PCO
C 2
PCO
C PH2O
(4.6)
where PCO2, PH2, PCO, PH2O are the partial pressures of CO2, H2, CO and H2O at
equilibrium. The combined water gas shift and carbonation reaction is as follows:
Combined water gas shift and carbonation: CO + H2O + CaO Ù H2 + CaCO3
(6)
The equilibrium constant for the combined water gas shift and carbonation
reaction is defined as shown below:
77
Keq2 =
PH2
PCOPH2O
(4.7)
where Keq2 = Keq1 * Kcarb and Kcarb is the equilibrium constant of the carbonation
reaction.
The equilibrium of the water gas shift reaction decreases with an increase in the
temperature resulting in low H2 yields at higher temperatures. Hence, in the
conventional water gas shift system, a LTS is used after the HTS to convert the CO slip
and to increase the yield of H2 in the presence of a LTS catalyst. The equilibrium
constant of the combined water gas shift and carbonation reaction is significantly
higher than the equilibrium constant of the water gas shift reaction alone, in the desired
temperature of operation ranging from 500 to 750 oC. Hence, the CLP is capable of
producing a much higher H2 yield, and hence, purity due to almost complete CO
conversion, when compared to the conventional H2 production process.
Equilibrium curves for the partial pressures of H2O (PH2O), CO2 (PCO2) and H2S
(PH2S) as a function of temperature, for the hydration, carbonation and sulfidation
reactions with CaO were also obtained using HSC Chemistry v 5.0 (Outokumpu
Research Oy, Finland). The relationship between reaction temperature and equilibrium
partial pressure of CO2 and H2O for the carbonation and hydration reaction with CaO
sorbent is shown in Figure 4.3.
Hydration : CaO + H2O Ù Ca(OH)2
(4.8)
78
Carbonation and hydration of CaO are reversible reactions which occur
depending on the conditions of temperature and partial pressures of CO2 and H2O
respectively. Carbonation of CaO occurs at conditions above the equilibrium PCO2
curve while calcination of CaCO3 occurs at conditions below the curve. Similarly
hydration of CaO occurs above the PH2O curve while dehydration occurs at conditions
below the curve.
Figure 4.4 illustrates the equilibrium H2 purity that can be obtained in the
carbonation reactor for a feed gas containing 10% CO and balance nitrogen. Pure H2
can be obtained even for stoichiometric S:C ratios and atmospheric pressure at
temperatures below 500 ºC. The purity of H2 begins to decrease with an increase in the
temperature. Increase in the S:C ratio and pressure favor the production of pure H2 at
all temperatures. Bench scale experiments at various temperatures, pressures and S:C
ratios have been conducted for a feed gas stream containing 10% CO and balance
nitrogen. The results of these tests have been discussed in sections 4.3 and 4.4 of this
chapter.
For the reversible sulfidation of CaO, the extent of H2S removal will depend on
the temperature and PH2O in the carbonation reactor. Figure 4.5 depicts the equilibrium
H2S concentrations in the product H2 stream, in ppm, for varying moisture
concentrations (PH2O) at 30 atm total system pressure. It can be seen that the
equilibrium H2S concentration in the product H2 stream increases with the increase in
PH2O. At a temperature of 600 oC, the H2S concentration is 0.1 ppm for a PH2O of 0.02
79
atm and 1ppm for a PH2O of 0.2 atm. By operating the carbonation reactor at nearstoichiometric steam requirement, it is possible to obtain low concentrations of steam
in the reactor system leading to low H2S concentrations of less than 1 ppm in the
product stream. It can also be seen that the reactor system will favor H2S removal using
CaO at around 500-650 oC, which is a suitable temperature for the carbonation reaction
as well.
Similarly, the removal of COS is dependent of temperature and partial pressure
of CO2 in the carbonation reactor. Figure 4.6 illustrates the equilibrium COS
concentrations in the product H2 stream for varying CO2 concentrations at 30 atm total
system pressure. The equilibrium COS concentration in the product H2 stream
increases with the increase in temperature and partial pressure of CO2. Since CO2 in the
carbonation reactor can be reduced to very low concentrations in the product stream by
the CaO sorbent, COS capture by the CaO sorbent will occur to large extents. The
equilibrium partial pressure of CO2 for the carbonation reaction with CaO sorbent at
600 ºC is less than 0.01 atms as shown in Figure 4.3. Hence the equilibrium COS
concentration in the product H2 stream can be predicted from Figure 4.6 to be lower
than 0.001 atms.
CaO sorbent is also capable of capturing HCl in the carbonation reactor. Similar
to the removal of H2S by CaO sorbent, the extent of HCl capture also dependent on the
temperature and the partial pressure of steam in the carbonator. The equilibrium HCl
concentration in the product H2 stream from the carbonation reactor for varying
80
temperatures and steam partial pressures is shown in Figure 4.7. Hence the reduction of
S:C ratio in the carbonation reactor will improve the extent of HCl removal.
4.2.2 The Calciner
The spent sorbent at the exit of the carbonation reactor is a mixture consisting
of CaCO3, CaO, CaS and calcium chloride (CaCl2). The CaCO3 in the spent sorbent
mixture is regenerated back to CaO in the calciner. The calciner is operated at
atmospheric pressure in a rotary or a fluidized bed system. The heat can be supplied
directly or indirectly using a mixture of fuel and oxidant. From the thermodynamic
curve for CaO and CO2 shown in Figure 4.8, calcination is found to occur at
temperatures above 890 oC in the presence of 1 atm of CO2. Dilution of CO2 in an
indirectly fired calciner with steam or oxy-combustion of a fuel (syngas, natural gas,
coal, etc) in a direct fired calciner will permit the calcination reaction to be conducted
at temperatures lower than 890 oC. The reaction occurring in the calciner is:
Calcination: CaCO3→ CaO + CO2
(4.9)
The regenerability of CaO sorbents over multiple cycles has been the major
drawback of high temperature calcium based CO2 capture processes. CaO sorbents are
prone to sintering during the high-temperature calcination step. There is a decrease in
sorbent reactivity even when steam is present in the calcination atmosphere. Over
multiple cycles, the percentage of sintered CaO progressively increases and reduces the
CO2 capture capacity of the sorbent (Curran et al, 1967, Iyer et al, 2004, Sun et al,
81
2008, Abanades and Alvarez, 2003, Barker, 1973, Bhatia and Perlmutter, 1983, Saliban
et al, 1996, Koji et al, 2003, Wang et al, 2005, Sun et al, 2007). Due to sintering, higher
solid circulation or make-up rates need to be used to maintain a high level of CO2
removal ( Romeo et al, 2009). Pretreatment methods have been developed to reduce the
decay in reactivity, which involve hydration of the sorbent (Koji et al, 2003, Iyer, 2003,
Manovic et al, 2007,Fennell et al, 2007, Sun et al, 2008), preheating and grinding of
the sorbent (Manovic and Anthony, 2008) and synthesis of novel sorbents by physical
or chemical modification of the precursor (Sun et al, 2008, Gupta and Fan, 2002,
Sakadjian, 2004, Salvador, 2003, Reddy and Smirniotis, 2004, Lu et al, 2009. In the
CLP process, the addition of a sorbent reactivation step by hydration, as part of the
carbonation-calcination cycle is used to reverse the effect of sintering during each cycle
and thus maintain the sorbent reactivity. (Fan et al, 2008) Sorbent hydration has been
found to be effective in maintaining sorbent reactivity as shown in Chapter 3.
4.2.3 Sorbent Reactivation by Hydration
The calcination process causes sintering of the sorbent which results in a
reduction in its reactivity and hence, the overall CO2 capture capacity. The hydration
process reverses this effect by increasing the pore volume and surface area available
for reaction with the gas mixture. Figure 4.9 shows the partial pressure of steam
required for hydration of the sorbent at various temperatures. Hydration occurs at
atmospheric pressure at temperatures below 500 oC. At temperatures of 600 oC and
above hydration occurs at steam partial pressures of above 4 atms. Operation of the
82
hydrator at high temperatures reduces the extent of cooling and reheating of the solids
required between the calciner and the carbonation reactor. This aids in reducing the
parasitic energy consumption of the process. Hydration at higher temperatures also
produces high quality heat which can be used to produce steam or electricity.
Depending of the reactivity of the calcined sorbent, a fraction of the calcined
sorbent or the entire stream of sorbent could be hydrated. The reactivity of the calcined
sorbent will depend on a variety of reasons including the type of calciner (direct or
indirect), mode of calcination (rotary kiln, fluidized bed or entrained bed), the
temperature of calcination and the gas atmosphere within the calciner. The reaction
occurring in the hydrator is shown below:
Hydration: CaO + H2O = Ca(OH)2
(4.8)
The Ca(OH)2 from the hydrator is conveyed to the carbonation reactor where it
dehydrates to produce high reactivity CaO and steam. The steam obtained from the
dehydration reaction is consumed in the water gas shift reaction. The advantage of this
reactivation process is that no excess steam is required for hydration. Part or all of the
steam required for the water gas shift reaction is supplied to the hydrator depending on
the fraction of the calcined sorbent that is sent to the hydrator for reactivation.
83
4.3 MATERIALS AND METHODS
4.3.1 Chemicals, Sorbents, and Gases
The HTS and STC catalyst were procured from Süd-Chemie Inc., Louisville,
KY. The HTS catalyst consists of iron (III) oxide supported on chromium oxide while
the STC catalyst consists of cobalt-molybdenum on alumina support. The CaO
precursor for the tests conducted in this chapter was PCC. PCC was synthesized from
Ca(OH)2 obtained from Fisher Scientific (Pittsburgh, PA). The high surface area PCC
(BET analysis; SA 49.2 m2/g; PV 0.17 cm3/g) was synthesized using a dispersant
modified wet precipitation technique. The anionic dispersant used in this process was
N40V, supplied by Ciba Specialty Chemicals (Basel, Switzerland). PCC was
synthesized by bubbling CO2 through a slurry of hydrated lime. The neutralization of
the positive surface charges on the CaCO3 nuclei by negatively charged N40V
molecules forms CaCO3 particles characterized by a higher surface area/pore volume
and a predominantly mesoporous structure. Details of this synthesis procedure have
been reported elsewhere (Fan and Gupta, 2006, Fan et al, 1998). The feed gas for all
the H2 production tests was a mixture of 10%CO and 90% Nitrogen (N2).
4.3.2 Fixed Bed Reactor Unit Setup
Figure 4.10 shows the bench scale, fixed bed reactor system, used for studying
H2 production at various process conditions. The bench scale reactor is coupled with a
set of continuous gas analyzers which detect concentrations of CO, CO2, H2S, CH4 and
84
H2 in the product stream. The reactor setup is capable of handling high pressures and
temperatures of up to 21 atms and 900 ºC respectively, which are representative of the
conditions in a commercial syngas to H2 system.
The mixture of gases from the cylinders is regulated and sent into the fixed bed
reactor by means of mass flow controllers that can handle pressures of about 21 atms.
From the mass flow controllers the reactant gases flow to the steam generating unit.
The steam generating unit is maintained at a temperature of 200 oC and contains a
packing of quartz chips which provide a large surface area of contact and mixing
between the reactant gases and steam. The steam generating unit not only facilitates the
complete evaporation on the water being pumped into the steam generating unit but it
also serves to preheat the reactant gases entering the reactor. The reactor, which is
heated by a tube furnace, is provided with a pressure gauge and a thermocouple to
monitor the pressure and temperature within. The rector consists of two concentric
sections; the inner section is filled with the catalyst or sorbent-catalyst mixture and the
outer section provides a preheating zone for the gases before they come in contact with
the bed of solids. The sorbent and catalyst loading section of the reactor is detachable
which enables easy removal and loading of the sorbent. The reactant gases leaving the
reactor enter a back pressure regulator which builds pressure by regulating the flow
rate of the gases and is capable of building pressures of up to 68.9 atm. The back
pressure regulator is very sensitive and the pressure within the reactor can be changed
quickly without any fluctuations. In addition, the back pressure regulator is also
85
capable of maintaining a constant pressure for a long period of time. The valve seat
material of the regulator is made of PEEK which is corrosion resistant to acidic H2S
vapors, which makes it suitable for conducting sulfur removal experiments. As shown
in Figure 4.10, the inlet of the back pressure regulator is connected to the reactor rod
and the outlet is connected to a heat exchanger. Since the entire section of the
equipment setup upstream of the backpressure regulator will be exposed to high
pressures, flexible stainless steel lines are used to withstand the pressure and the
reactor is constructed from inconel which is resistant to corrosion due the high pressure
high temperature steam and H2S gas.
The product gas mixture exiting the back pressure regulator is then cooled in a
heat exchanger using chilled ethylene glycol-water mixture to condense the
unconverted steam. The product gas at the exit of the heat exchanger is dried in a
desiccant bed and is sent to a set of continuous analyzers capable of determining the
concentrations of CO, CO2, H2S, CH4 and H2 in the gas stream.
4.3.3 Water Gas Shift Reaction Testing
The water gas shift reaction was conducted using the catalysts obtained from
Süd-Chemie. These experiments were conducted as base line experiments to determine
the conditions for maximum water gas shift catalytic activity at different ranges of
temperatures (450-800 oC), S:C ratios and pressures, which are beyond the commercial
mode of operation, but are of interest for the CLP. Catalyst particles were used in a
86
fixed bed reactor setup for all the experiments. The total flow rate of the gases through
the reactor was maintained a constant at 725 sccm for all the experiments and the
concentration of CO in the reaction mixture was maintained at 10.3 %. 0.25 g of the
catalyst was loaded into the reactor and the pressure, temperature and gas flow rates
were adjusted for each run. The dry gas compositions at the outlet of the reactor were
monitored continuously using the CO, CO2, H2S, CH4 and H2 gas analyzers.
4.3.4 Simultaneous Water Gas Shift and Carbonation
The combined water gas shift and carbonation reaction was conducted using a
sorbent (CaO) to catalyst ratio of 10:1 by weight. The CaCO3 sorbent was calcined by
heating the sorbent-catalyst mixture to 700 oC in a stream of N2 until the CO2 analyzer
confirmed the absence of CO2 in the outlet stream. At the end of calcination, the feed
gas was switched from nitrogen to a mixture of 10% CO and 90% nitrogen for the
combined water gas shift and carbonation reaction. The combined water gas shift and
carbonation reaction experiments were conducted at 600, 650, and 700°C with a S:C
ratio of 3:1, 2:1, 1:1 at various pressures ranging from 1-21 atm.
4.3.5 Catalyst Pretreatment
It is imperative to understand the HTS catalyst composition during calcination
of the sorbent which occurs in the presence of a CO2 atmosphere at high temperature.
Iron oxide occurs in three different phases: Hematite (Fe2O3), magnetite (Fe3O4) and
wustite (FeO). The active phase of the HTS catalyst is magnetite. However, in the
87
presence of an oxidizing atmosphere, like CO2 or steam, the magnetite phase gets
oxidized to hematite which is likely during the calcination step. This is evident from
the iron oxide phase diagram for a CO-CO2 system (Ross, 1980). Thus, a pretreatment
procedure was developed which consists of treating the oxidized catalyst in a 20%/80%
of H2/H2O atmosphere at 600oC which reduces the hematite to magnetite. The
effectiveness of the pretreatment procedure was confirmed by X-ray diffraction
analyses of the HTS catalyst before and after the pretreatment procedure. The HTS
catalyst as obtained contains hematite phase as shown in Figure 4.11. The catalyst is
subsequently subjected to the pretreatment procedure which changes its phase to the
active magnetite form as shown in Figure 4.12. In the commercial deployment of the
CLP, pretreatment of the catalyst can be avoided by using a fixed fluidized bed reactor
for the carbonation reactor in which the catalyst remains in the carbonation reactor
while the CaO sorbent is looped between the carbonation reactor and the calciner. In
this configuration the, catalyst is never exposed to oxidizing gases in the calciner. No
deactivation of the STC catalyst was observed during calcination.
4.3.6 Combined H2 Production with H2S Removal
To study the effect of sulfur on the CLP, 5000 ppm of H2S was mixed with CO,
N2 and steam before being sent to the reactor. The H2 production tests were conducted
in the presence of the catalyst and CaO sorbent.
88
4.4 RESULTS AND DISCUSSION
4.4.1 Effect of Process Parameters on the Extent of Water Gas Shift Reaction using
HTS Catalyst
An investigation of the water gas shift reaction in the presence of a HTS
catalyst was conducted in the bench scale fixed bed reactor to determine the effect of
temperature, pressure and S:C ratio on the extent of reaction. Figure 4.13a shows the
CO conversion profiles for increasing reaction temperatures and S:C ratios at ambient
pressures. The CO conversion increases with increasing temperature as it approaches
the equilibrium value at an optimal temperature (600 - 650 oC) beyond which it begins
decreasing monotonically. At a pressure of 1 atm and a S:C ratio of 3:1, the conversion
increases from 45.8 % at 450 ºC to 83.2 % at 600oC. Beyond 600 oC, the conversion
decreases and at 800 ºC, it is 69.4%. This decrease in conversion with the increase in
temperature is observed due to the thermodynamic limitation of the water gas shift
reaction. Thus at lower temperatures although the equilibrium constant is high, the
reaction rate is low. At high temperatures, although the reaction is very fast, the
equilibrium constant is low. Consequently maximum conversion is reached at an
optimum temperature at which both the kinetics and the reaction equilibrium are
favorable. From Figure 4.13a, it can also be seen, as expected, that the conversion
increases with the increase in the S:C ratio for all temperatures. At a temperature of
650oC, the conversion is 63.5% for a S:C ratio of 1:1, 71.6% for 2:1 and 80.28% for
3:1. As can be seen in Figure 4.13b, the effects of reaction temperatures and S:C ratios
89
on CO conversion at 21 atm follow the same trend as that at 1 atm. In addition, below
600-650 ºC, the CO conversion at 21 atm is greater than at 1 atm due to an increase in
the rate of the reaction with increase in pressure.
The observed partial pressure ratios were computed for different S:C ratios,
temperatures and pressures and were compared with the equilibrium values obtained
from HSC Chemistry v 5.0 (Outokumpu Research Oy, Finland). The observed partial
pressure ratio (Kobs) was computed from the experimental data and is defined as the
ratio of the product of partial pressures of the products to that of the reactants as given
below:
Kobs =
PH PCO
2
2
PCOPH O
(4.6)
2
As shown in Figures 4.14a and 4.14b it was found that each value of the
observed partial pressure ratio (Kobs) was within the equilibrium value. From the
figures it can be seen that the partial pressure ratio increases with an increase in the
temperature till it approaches equilibrium and then decreases along the equilibrium
curve. Besides, as the pressure increases, the system is closer to equilibrium for both
S:C ratios of 1:1 and 3:1. This can be explained by the increase in the rate of the
reaction with increase in pressure.
90
4.4.2 Enhancing the Water Gas Shift Reaction by In-situ CO2 Removal (HTS
Catalyst and CaO Sorbent)
From Figures 4.13a and b, it can be observed that the CO conversion achieved
in the presence of the HTS catalyst is only 80-90% even at a high pressure of 21 atms
and a high S:C ratio of 3:1. At atmospheric pressure and a stoichiometric S:C ratio, a
low CO conversion of 20-60% is obtained. In order to enhance the H2 yield, CaO
sorbent could be introduced into the H2 production reactor for in-situ CO2 removal
from the reaction zone. This increase in H2 yield can be explained by the LeChatlier’s
principle where the simultaneous CO2 removal drives the equilibrium limited water gas
shift reaction forward. This concept was demonstrated by conducting the combined
water gas shift and carbonation reaction in the presence of the calcined PCC sorbent
and HTS catalyst in the fixed bed reactor. Figures 4.15a illustrates the typical
breakthrough curves obtained during the combined water gas shift reaction and
carbonation reaction for the N2 free dry product gas compositions. High purity H2 is
produced in the pre-breakthrough region due to in-situ CO2 removal by the sorbent. As
the sorbent gets exhausted the breakthrough region occurs followed by the postbreakthrough region in which all the sorbent has been converted to CaCO3 and H2
production occurs in the presence of the HTS catalyst. The concentrations of CO and
CO2 in the product gas mixture are very low in the pre-breakthrough region and
increase in the breakthrough region due to the depletion of the sorbent. Figure 4.15b
illustrates the typical breakthrough curve obtained for CO conversion.
91
4.4.2.1 Effect of Pressure
The effect of pressure on the combined water gas shift and carbonation reaction
for S:C ratios of 3:1 and 1:1 are shown in Figures 4.16a and 4.16b. The purity of H2
produced in the pre-breakthrough region of the curves increases with the increase in
pressure. From Figure 4.16a it can be observed that during the initial pre-breakthrough
period for a S:C ratio of 3:1, 95.6 % H2 is produced at 1 atm, 99.7% pure H2 is
obtained at 11 atm, and 99.8% pure H2 is produced at 21 atm. The extent of the prebreakthrough region, which signifies the extent of conversion of the CaO sorbent, also
increases with the increase in pressure. A similar observation is made for a lower S:C
ratio of 1:1 as shown in Figure 4.16b. It can be inferred that higher pressure results in
increased partial pressure of CO2 which enhances the extent and rate of carbonation
due to higher driving force. This consequently results in enhanced CO conversion,
sorbent conversion and H2 yield.
4.4.2.2 Effect of S:C Ratio
The effect of S:C ratio on the CO conversion and H2 purity for the combined
water gas shift and carbonation reaction are shown in Figures 4.17a, 4.17b, 4.17c and
4.17d. The effect of S:C ratio at atmospheric pressure is shown in Figures 4.17a/b
while that at 21 atm is shown in Figures 4.17c/d. It can be seen that at atmospheric
pressure a reduction in the S:C ratio results in a decrease in the CO conversion and
associated H2 purity. However, at a higher pressure of 21 atm, almost 100 % CO
92
conversion and H2 purity is achieved for all the three S:C ratios in the pre-breakthrough
region. This can again be attributed to the higher partial pressure of CO2 contributing to
enhanced carbonation kinetics which plays a key role in driving the water gas shift
reaction to completion. Besides, from a process design and cost perspective, operation
at high pressures clearly illustrates the benefit of using a smaller amount of steam for a
high CO conversion, resulting in cost savings.
4.4.2.3 Effect of Temperature
The effect of temperature on the CO conversion achieved in the presence of the
CaO sorbent and HTS catalyst at atmospheric pressure and a S:C ratio of 3:1 is
depicted in Figure 4.18. In the pre-breakthrough region, the CO conversion decreases
with the increase in temperature due to equilibrium limitations of the combined water
gas shift and carbonation reaction. In the initial pre-breakthrough region, a CO
conversion of 95% is obtained at 600 and 650 ºC and it decreases to 90% at 700 ºC.
Figures 4.19(a) and (b) illustrate the effect of temperature on the combined
reactions at 21 atm and S:C ratios of 3:1 and 1:1 respectively. At a high S:C ratio of
3:1, there is almost no change in the CO conversion with the change in temperature as
can be seen in Figure 4.19a. On decreasing the S:C ratio to the stoichiometric amount,
it is observed in Figure 4.19b, that temperature plays a significant role in the extent of
CO conversion and a temperature of 600 ºC is optimum for achieving high CO
conversions of 99.7%. Thus, from a process design perspective this defines the
93
operating temperature for achieving high CO conversions and H2 yield while
maintaining low steam requirements.
4.4.3 Simultaneous Water Gas Shift, Carbonation and Sulfidation Reaction Testing
Since syngas obtained from the gasifier contains 0.5 to 4% sulfur mostly in the
form of H2S, the effect of sulfur and the extent of its removal by the CaO sorbent were
determined on the combined water gas shift and carbonation reaction. Integrated H2
production, CO2 and H2S removal using calcium sorbent and HTS catalyst was
investigated by the addition of 5000 ppm of H2S to the fixed bed reactor feed. The
calcium sorbent was used to simultaneously capture H2S and CO2 while enhancing H2
production in the presence of the HTS catalyst.
As illustrated in Figure 4.20(a), it was found that H2S concentration in the
outlet H2 stream is reduced to a few ppm in the pre-breakthrough region by the reaction
of H2S with the CaO sorbent. In the thermodynamics section of the sulfidation of CaO,
illustrated in Figure 4.5, it was observed that the extent of H2S removal is inhibited by
the presence of a high partial pressure of steam in the system. This concept is
demonstrated in the experimental results depicted in Figures 4.20a and 4.21. Figure
4.20a illustrates the entire breakthrough curve of H2S concentration in the product H2
stream with the pre-breakthrough and breakthrough regions. Figure 4.21 is a magnified
image of the pre-breakthrough region in Figure 4.20a and it shows that with the
increase in S:C ratio, the H2S concentration in the H2 product increases. At a lower S:C
94
ratio of 1:1, the H2S in the outlet stream is lower than 1 ppm while at an S:C ratio to
3:1, the H2S concentration increases from 2 ppm to 30 ppm in 750 secs during the prebreakthrough region. At a S:C ratio of 1:1, in the pre-breakthrough region, the
carbonation reaction enhances the water gas shift reaction which results in the
consumption of most of the steam. Hence H2S removal by the calcium sorbent is
enhanced and the H2S composition in the outlet stream is low.
As the reaction proceeds, the CaO sorbent gets consumed to form CaCO3 and
CaS resulting in the breakthrough curve seen in Figure 4.20a.
Since the steam
composition in the system is higher for an S:C ratio of 3:1 the H2S concentration in the
product stream is higher. During the breakthrough region, H2S reacts with both CaO
and CaCO3. The post-breakthrough region is not visible in Figure 4.20a as the H2S
concentration in the product will keep increasing with time until all the CaCO3 is also
converted to CaS. In the post-breakthrough region the H2S concentration in the product
will be equal to the H2S concentration in the feed stream.
Figure 4.20b illustrates the change in CO conversion with respect to time for
S:C ratios of 3:1 and 1:1. In the pre-breakthrough region, the CO conversion for an S:C
ratio of 3:1 is slightly higher than that for 1:1.
4.4.4 Effect of Catalyst Type on the Water Gas Shift Reaction
A STC procured from Sud Chemie was also tested for its suitability in the CLP.
The water gas shift reaction was conducted in the presence of the STC at a range of
95
temperatures (400-800 ºC), pressures (1-21 atms) and
S:C ratios (1:1-3:1) . The
performance of the HTS catalyst was compared with that of the STC catalyst. As
illustrated in Figure 4.22, it was found that there is an increase in the CO conversion
with an increase in the S:C ratio for both the STC as well as the HTS catalyst. It was
also found that at temperatures below 650 ºC the CO conversion in the presence of the
HTS catalyst is higher that the CO conversion obtained in the presence of the STC. 550
ºC-650 ºC is found to be the optimum temperature of operation in the presence of the
HTS catalyst and 700-800 ºC is found to the optimum temperature of operation for the
STC.
The water gas shift reaction was conducted in the presence of the STC at a
range of temperatures (400 ºC to 800 ºC) and a range of pressures (1-21 atms) as
shown in Figure 4.23. It was found that the conversion increases with increase in
temperature due to improved kinetics and beyond a particular temperature decreases
since the reaction is exothermic. The conversion increases with an increase in pressure
and at each pressure the conversion reaches a maximum value at a particular
temperature. As shown in Figure 4.23, with increase in pressure the temperature for
maximum conversion decreased from 750 ºC at 1 atm, 700 ºC at 11 atms to 600 ºC at
21 atms.
As shown earlier, at atmospheric pressures, the HTS catalyst gives higher
conversion than the STC at all temperatures. As shown in Figure 4.24, at a pressure of
11 atms the HTS catalyst gives maximum CO conversion at a temperature of 550 ºC
96
while the STC gives maximum conversion at a temperature of 700 and 750 ºC. It was
found that with the increase in temperature above 675 ºC the conversion in the
presence of the STC increases above the conversion obtained in the presence of the
HTS catalyst. At a temperature of 600 ºC the conversion in the presence of STC is 50%
while in the presence of the HTS the conversion is 64%. At a temperature of 700 ºC,
the conversion in the presence of the sulfur tolerant is 60% while in the presence of
HTS is 58%. Hence in this case it is very important to determine the rate of carbonation
at different temperatures as equilibrium conversion for carbonation decreases at
temperatures above 650 ºC. The temperature of operation of the sorbent and the
catalyst should be similar for production of the highest purity H2. Hence combined
water gas shift and carbonation reactions need to be conducted in the presence of the
catalyst and CaO sorbent to determine the catalyst best suited for the reaction.
A similar trend was observed at 21 atms and as the temperature was increased
the conversion obtained in the presence of the STC increased and was equal to the
conversion in the presence of the HTS catalyst at temperatures above 650 ºC. As
shown in Figure 4.25, at a temperature of 600 ºC, conversion in the presence of the
HTS catalyst is 70% while that in the presence of the STC is 65%. But at a temperature
of 700 ºC both catalysts give the same conversion of 65%.
The effect of H2S on the activity of the HTS and STC catalyst was investigated.
Figure 4.26 depicts the comparison in CO conversion achieved at atmospheric pressure
in the presence and absence of H2S in the inlet gas steam. It was found that at 650 ºC,
97
the CO conversion decreases in the presence of H2S for both the HTS catalyst and the
STC catalyst. It has been shown in literature, that the HTS catalyst still retains half its
original activity in its sulfided form and the same inference is obtained from Figure
4.26 at both S:C ratios of 3:1 and 1:1 (Hla et al, 2009). Although the decrease in the
conversion obtained in the presence of the STC catalyst is very low when compared to
that in the HTS catalyst, it was found that even in the presence of H2S, the HTS
catalyst shows higher CO conversion at a temperature of 650 ºC.
The combined water gas shift and carbonation reaction was conducted at
different temperatures at atmospheric pressure in the presence of the STC and CaO
sorbent. As shown in Figure 4.27 it can be seen that the conversion decreases as the
temperature is increased and it is highest at 650 ºC. In the 650 to 750 ºC temperature
range, although conversion of CO in the presence of STC increases with temperature
till 750 ºC (in Figure 4.22), the combined reaction conversion decreases with increase
in temperature (Figure 4.27). This is because maximum CO2 removal occurs at 650 ºC
and at temperatures higher than 650 ºC, the equilibrium conversion for the carbonation
reaction decreases.
The enhancement in CO conversion on the addition of CaO sorbent to the STC
catalyst is illustrated in Figure 4.28. At both ratios of 3:1 and 1:1, the CO conversion
was found to be the highest in the presence of the HTS catalyst and CaO sorbent.
Although the CO conversion is increased by the addition of CaO to the STC, it is still
98
lower than the conversion obtained in the presence of the mixture of HTS catalyst and
CaO sorbent.
4.5 CONCLUSIONS
Enhancement in the production of high purity H2 from syngas can be achieved
using CaO sorbent that can drive the equilibrium limited water gas shift forward by insitu removal of CO2. Thermodynamic analyses for the reactions occurring in the
carbonation reactor, calciner and hydrator were conducted to determine the operating
window for various process parameters. Operating at near stoichiometric steam
conditions is advantageous for simultaneous sulfur removal to low levels in the product
H2 stream. Bench scale experimental data demonstrate that greater than 99% pure H2
can be produced at high temperatures and pressures. For near stoichiometric
conditions, high CO conversion and H2 purity can be obtained at high pressures and an
optimal temperature of 600 ºC. This operating temperature was also found to be
favorable for simultaneous H2S removal to <1ppm in the product H2 stream. At
atmospheric pressure, a water gas shift catalyst which has high activity in the optimal
temperature range of the carbonation reaction (500-750 ºC) in the presence of sulfur is
beneficial. The HTS catalyst with CaO sorbent results in the production of a high
purity H2 stream at atmospheric pressure. This is important in situations where the
conversion of fuel gas to H2 at atmospheric pressure is beneficial. Further investigation
conducted to determine whether a catalyst is required for the production of H2 at high
pressures
within
a
short
residence
99
time
is
described
in
Chapter
5.
Purge Stream
Fresh Sorbent
Reaction
Regeneration
Hydrogen
Pure CO 2 gas
Integrated
Hydrogen
reactor
Net Heat
Output
Heat
Input
Calciner
Syngas
Dehydration :
WGSR
:
CO2 removal :
Sulfur
:
Halide
:
Ca(OH) 2 Æ CaO + H 2O
CO + H2O Æ CO2 + H2
CaO + CO2 Æ CaCO3
CaO + H 2 S Æ CaS + H 2O
CaO + 2HX ÆCaX 2 + H2O
Calcination: CaCO3 Æ CaO + CO2
100
Reactivation
Heat
Output
Hydrator
H2O
Hydration : CaO + H2O Æ Ca(OH) 2
Figure 4.1: Schematic of the CLP
100
100000
K(Equilibrium Constant)
10000
WGSR + Carbonation
1000
100
10
WGSR
1
0.1
300
400
500
600
700
800
900
1000
Temperature (ºC)
Figure 4.2: Thermodynamic data illustrating the equilibrium constants of the water gas
shift reaction and the combined water gas shift and carbonation reaction
101
CaO + CO2
CaO + H2O
Equlibrium Partial Pressure (atm)
102
CaCO3
Ca(OH)2
101
100
Hydration
Carbonation
Dehydration
10-1
Calcination
10-2
10-3
10-4
10-5
10-6
P H2O
2
P CO22
10-7
10-8
400
600
800
1000
Temperature (C)
Figure 4.3: Thermodynamic data for the hydration and carbonation of CaO sorbent
102
100
Hydrogen purity (%)
80
60
S/C = 1:1
40
20
0
400
21 atm
300
psig
11
atm
150 psig
atm
01psig
S/C = 3:1
21 atm
300
psig
11 atm
150
psig
0 1psig
atm
500
600
700
800
900
1000
Temperature (C)
Figure 4.4: Equilibrium H2 purity in the carbonator at varying temperatures, pressures
and S: C ratios. (Feed gas: 10% CO and balance nitrogen)
103
Equilibrium H2S Conc (ppm)
with 30 atm total pressure
10000
1000
CaO+ H2 S
CaS + H2O
20 atm (PH2O)
2 atm (PH2O)
0.2 atm (PH2O)
0.02 atm (PH2O)
Typical
Gasifier
100
10
1
0.1
0.01
400
CLP
500
600
700
800
900
1000
Temperature (oC)
Figure 4.5: Thermodynamic data for the sulfidation (H2S) of CaO with varying steam
partial pressures. (PTotal = 30 atm)
104
Equilibrium COS Conc (ppm)
with 30 atm total pressure
10
1 atm (PCO2)
0.1 atm (PCO2)
1
0.01 atm (PCO2)
0.001 atm (PCO2)
0.1
0.01
0.001
0.0001
400
500
600
700
800
900
1000
Temperature (oC)
Figure 4.6: Thermodynamic data for predicting the equilibrium COS concentration for
CaO sulfidation with varying CO2 concentration (PTotal = 30 atm)
105
Equilibrium HCl Conc (ppm)
with 30 atm total pressure
10000
1000
20 atm
2 atm
0.2 atm
0.02 atm
100
10
1
0.1
0.01
400
500
600
700
800
900
1000
Temperature (oC)
Figure 4.7: Thermodynamic data for predicting the equilibrium HCl concentration for
CaO reaction with HCl with varying steam concentration (PTotal = 30 atm)
106
Equlibrium Partial Pressure (atm)
102
P CO22
101
100
Carbonation
10-1
10-2
Calcination
10-3
10-4
10-5
10-6
10-7
10-8
400
600
800
o
Temperature ( C)
Figure 4.8: Thermodynamic data for the carbonation of CaO
107
1000
Equlibrium Partial Pressure (atm)
102
101
Hydration
0
10
10-1
Dehydration
-2
10
10-3
10-4
10-5
10-6
10-7
10-8
400
600
800
o
Temperature ( C)
Figure 4.9: Thermodynamic data for the hydration of CaO
108
1000
Thermocouple
And
Pressure Guage
Steam Generator
Steam &
Gas Mixture
Sorbent
&
Catalyst
Powder
Vent
Water In
MFC
109
Back Pressure
Regulator
Analyzers (CO,
CO2, H2, H2S)
Water Syringe
Pump
Heat
Exchanger
Water Trap
Figure 4.10: Simplified flow sheet of the bench scale experimental setup
109
MFC
H2
Heated Steel
Tube Reactor
Hydrocarbon
Analyzer
Gas
Gas
Mixture
Mixture
MFC
CO
MFC
CO2 Hydro
H2S
carbons
FE2O3HTS
160
150
140
130
120
110
Lin (Counts)
100
90
80
70
60
50
40
30
20
10
0
10
20
30
40
50
60
70
80
2-Theta - Scale
FE2O3HTS - File: HT S.RAW - Type: 2T h/Th unlocked - Start: 10.000 ° - End: 85.000 ° - Step: 0.030 ° - Step time: 1.8 s - Temp.: 25 °C (Room) - T ime Started: 0 s - 2-Theta: 10.000 ° - Theta: 0.000 ° - Chi: 0.00 ° Operations: Smooth 0.150 | Back ground 1.000,1. 000 | Bac kground 1. 000,1.000 | Import
84-0308 (C) - Iron Oxide - Fe2O3 - Y: 47.91 % - d x by: 1. - WL: 1.54056 - 0 - I/Ic PDF 3.2 -
Figure 4.11: X-ray diffraction patters of the HTS catalyst before pretreatment
(hematite)
110
HTS6001HR
Lin (Counts)
380
370
360
350
340
330
320
310
300
290
280
270
260
250
240
230
220
210
200
190
180
170
160
150
140
130
120
110
100
90
80
70
60
50
40
30
20
10
15
20
30
40
50
60
70
2-Theta - Scale
HTS6001HR - File: HTS600~1.RAW - Type: 2Th/Th unlocked - Start: 15.000 ° - End: 79.980 ° - Step: 0.030 ° - Step time: 1.5 s - Temp.: 25 °C (Room) - Time Started: 0 s - 2-Theta: 15.000 ° - Theta: 0.000 ° - Chi:
Operations: Smooth 0.150 | Smooth 0.150 | Import
80-0389 (C) - Magnetite - Fe.99Fe1.97Cr.03Ni.01O4 - Y: 89.58 % - d x by: 1. - WL: 1.54056 - Cubic - a 8.39500 - b 8.39500 - c 8.39500 - alpha 90.000 - beta 90.000 - gamma 90.000 - Face-centred - Fd-3m (227)
Figure 4.12: X-ray diffraction patters of the HTS catalyst after pretreatment (magnetite)
111
1.0
CO Conversion
0.8
0.6
0.4
1:1
2:1
3:1
0.2
0.0
400
500
600
700
800
Temperature (oC)
(a)
1.0
CO Conversion
0.8
0.6
0.4
0.2
0.0
400
1:1
2:1
3:1
500
600
700
800
Temperature (oC)
(b)
Figure 4.13: Effect of reaction temperature and S:C ratio on the conversion of CO by
the water gas shift reaction in the presence of HTS catalyst at (a) 1 atm (b)
21
112
Partial Pressure Ratio
8
1 atm
11 atm
21 atm
Theoretical
6
4
2
0
400
500
600
700
800
Temperature (oC)
(a)
Partial Pressure Ratio
8
1 atm
11 atm
21 atm
Theoretical
6
4
2
0
400
500
600
700
800
Temperature (oC)
(b)
Figure 4.14: Effect of reaction temperature and pressure on the observed partial
pressure ratio for the water gas shift reaction in the presence of HTS
catalyst at a S:C ratio of (a)1:1 (b)3:1
113
100
CO
CO2
H2
Gas Compositions
80
60
40
20
0
1000
2000
3000
4000
3000
4000
Time (sec)
(a)
1.0
CO Conversion
0.9
0.8
0.7
0.6
0.5
0
1000
2000
Time (sec)
(b)
Figure 4.15: Typical curves for the combined water gas shift and carbonation reaction
in the presence of CaO sorbent and HTS catalyst depicting (a) Gas
composition (mol%) and (b) CO conversion (650 ºC, 1 atm, S:C ratio of
3:1)
114
H2 Gas Composition (%)
100
80
60
40
20
1 atm
11 atm
21 atm
0
0
500
1000
1500
2000
2500
Time(sec)
(a)
H2 Gas Composition (%)
100
80
60
40
1 atm
11 atm
21 atm
20
0
0
500
1000
1500
2000
2500
Time (sec)
(b)
Figure 4.16: Effect of pressure on purity of H2 produced during the combined water gas
shift and carbonation reaction in the presence of CaO sorbent and HTS
catalyst at a S:C ratio of (a) 3:1 (b) 1:1 (650 ºC)
115
1.0
CO Conversion
0.8
0.6
0.4
0.2
1:1
2:1
3:1
0.0
0
500
1000
1500
2000
2500
Time (sec)
(a)
H2 Gas Composition(%)
100
80
60
40
20
1:1
3:1
0
0
500
1000
1500
2000
2500
Time (sec)
(b)
Continued
Figure 4.17: Effect of S:C ratio on the combined water gas shift and carbonation
reaction in the presence of CaO sorbent and HTS catalyst at 650 ºC (a) CO
conversion at 1 atm (b) H2 gas composition at 1 atm (c) CO conversion at
21 atm (d)H2 gas composition at 21 atm
116
Figure 4.17 continued
1.0
0.6
0.4
1:1
2:1
3:1
0.2
0.0
0
500
1000
1500
2000
2500
Time (sec)
(c)
100
H2 Gas Composition (%)
CO Conversion
0.8
80
60
40
20
1:1
2:1
3:1
0
0
500
1000
1500
TIme (sec)
(d)
117
2000
2500
1.0
CO Conversion
0.9
0.8
0.7
0.6
600C
650C
700C
0.5
0.4
0
500
1000
1500
2000
2500
3000
Time(sec)
Figure 4.18: Effect of temperature on CO conversion by the combined water gas shift
and carbonation reaction in the presence of CaO sorbent and HTS catalyst
at 1 atm and S:C ratio of 3:1
118
1.0
CO Conversion
0.9
0.8
0.7
0.6
600 C
650 C
700 C
0.5
0.4
0
500
1000
1500
2000
2500
Time (sec)
(a)
1.0
CO Conversion
0.9
0.8
0.7
0.6
600 C
650 C
700 C
0.5
0.4
0
500
1000
1500
2000
2500
Time(sec)
(b)
Figure 4.19: Effect of temperature on CO conversion by the combined water gas shift
and carbonation reaction in the presence of CaO sorbent and HTS catalyst
at 21 atm and S:C ratio of (a) 3:1 (b) 1:1
119
1400
1:1
3:1
H2S Concentration (ppm)
1200
1000
800
600
400
200
0
0
1000
2000
3000
4000
Time(sec)
(a)
1.0
1:1
3:1
CO Conversion
0.8
0.6
0.4
0.2
0.0
0
1000
2000
3000
Time(sec)
(b)
Figure 4.20: Effect of S:C ratio on (a) the composition of H2S in the H2 stream and (b)
CO conversion in the presence of the catalyst and sorbent during the
simultaneous water gas shift, carbonation and sulfidation reaction (600 ºC,
1 atm)
120
H2S concentration (ppm)
30
20
10
0
-10
1-1
3-1
-20
200
400
600
800
1000
1200
1400
Time(sec)
Figure 4.21: Effect of S:C ratio on the composition of H2S in the H2 stream during the
combined water gas shift, carbonation and sulfidation reaction in the
presence of CaO sorbent and HTS catalyst (600 ºC, 1 atm)
121
1.0
Conversion
0.8
0.6
0.4
STC- 3:1(S:C)
STC - 1:1(S:C)
HTS - 3:1(S:C)
HTS - 1:1(S:C)
0.2
0.0
400
450
500
550
600
650
700
750
800
Temperature (C)
Figure 4.22: Effect of S:C ratio and temperature on CO conversion during the water
gas shift reaction in the presence of STC and HTS catalyst
122
70
CO Conversion
60
50
40
30
20
1 atm
0 psig
11 atm
150
psig
21 atm
300
psig
10
0
300
400
500
600
700
800
900
Temperature (C)
Figure 4.23: Effect of reaction temperature on CO conversions for various pressures at
an S:C ratio of 1:1 for the STC (0.25g STC, Total flow = 0.725 slpm)
123
0.8
CO Conversion
0.6
0.4
0.2
Sulfur Tolerant Catalyst
HTS
0.0
400
500
600
700
800
900
Temperature (C)
Figure 4.24: Effect of reaction temperature on CO conversions for the HTS and STC at
11 atms and S:C ratio of 1:1(Total flow = 0.725 slpm)
124
0.8
CO Conversion
0.7
0.6
0.5
0.4
0.3
0.2
0.1
300
Sulfur Tolerant Catalyst
HTS
400
500
600
700
800
Temperature (C)
Figure 4.25: Effect of reaction temperature on CO conversions for the HTS and STC at
21 atms and S:C ratio of 1:1(Total flow = 0.725 slpm)
125
1.0
CO Conversion
0.8
0.6
0.4
0.2
0.0
0
500
1000
1500
2000
Time (secs)
STC - 1:1(S:C) - In the presence of H2S
STC - 1:1(S:C) - In the absence of H2S
STC - 3:1(S:C) - In the presence of H2S
STC - 3:1(S:C) - In the absence of H2S
HTS - 1:1(S:C) - In the presence of H2S
HTS - 1:1(S:C) - In the absence of H2S
HTS - 3:1(S:C) - In the presence of H2S
HTS - 3:1(S:C) - In the absence of H2S
Figure 4.26: Effect of S:C ratio, type of catalyst and effect of H2S on CO conversion
during the water gas shift reaction (650 ºC, 1atm)
126
1.0
Conversion
0.8
0.6
0.4
0.2
650C
700C
750C
0.0
0
500
1000
1500
2000
2500
Time (sec)
Figure 4.27: Effect of temperature on CO conversion (Temperature=650°C, Pressure
= 1 atm, S:C ratio= 1:1)
127
1.0
CO Conversion
0.8
0.6
0.4
3:1(S:C)-STC with CaO
1:1(S:C)-STC with CaO
3:1(S:C)-STC
1:1(S:C)-STC
3:1(S:C)-HTS with CaO
1:1(S:C)-HTS with CaO
0.2
0.0
0
500
1000
1500
2000
2500
Time (sec)
Figure 4.28: Comparison in the CO conversion obtained at different S:C ratios for
different sorbent and catalyst mixtures (650 ºC, 1atm)
128
CHAPTER 5
ENHANCED NON-CATALYTIC H2 PRODUCTION FROM SYNGAS
5.1 INTRODUCTION
In Chapter 4, H2 production with contaminant removal in the presence of CaO
sorbent and a water gas shift catalyst was investigated. The presence of the sorbent and
catalyst in the carbonation reactor results in the production of high purity H2 with low
levels of CO, CO2 and sulfur but introduces issues and costs associated with the
separation of the sorbent and catalyst prior to calcination or pretreatment of the catalyst
to the active form after its deactivation in the presence of CO2 in the calciner at high
temperatures, replacement of the spent catalyst, deactivation of the catalyst in the
presence of sulfur impurities (H2S) and the use of expensive STC. In an attempt to
further simplify the process, the non catalytic CLP was investigated (Iyer et al, 2006,
Iyer et al, 2006, Ramkumar et al, 2008). The feasibility of enhancing the purity of H2
and the optimum process conditions for H2 production in the absence of a water gas
shift catalyst were determined.
129
5.2 MATERIALS AND METHODS
5.2.1 Chemicals, Sorbents, and Gases
The HTS catalyst was procured from Süd-Chemie Inc., Louisville, KY and
consists of iron (III) oxide supported on chromium oxide. The CaO sorbent was
obtained from a PCC precursor which was synthesized from Ca(OH)2 obtained from
Fisher Scientific (Pittsburgh, PA). The high surface area PCC (BET analysis; SA 49.2
m2/g; PV 0.17 cm3/g) was synthesized using a dispersant modified wet precipitation
technique. The anionic dispersant used in this process was N40V, supplied by Ciba
Specialty Chemicals (Basel, Switzerland). PCC was synthesized by bubbling CO2
through a slurry of hydrated lime. The neutralization of the positive surface charges on
the CaCO3 nuclei by negatively charged N40V molecules forms CaCO3 particles
characterized by a higher surface area/pore volume and a predominantly mesoporous
structure. Details of this synthesis procedure have been reported elsewhere (Fan et al,
1998, Gupta and Fan, 2002). The feed gas for all the H2 production tests was a mixture
of 10%CO and 90% Nitrogen (N2).
5.2.2 Experimental Setup: Fixed Bed Reactor
Figure 5.1 shows the integrated experimental setup, used for the bench scale
studies of the non-catalytic CLP for H2 production from syngas. The setup is similar to
the one used earlier to study the catalytic CLP system, The bench scale reactor is
coupled with a set of continuous gas analyzers which detect concentrations of CO,
130
CO2, H2S, CH4 and H2 in the product stream. The reactor setup is capable of handling
high pressures and temperatures of up to 21 atms and 900oC respectively, which are
representative of the conditions in a commercial syngas to H2 system.
5.2.3 Water Gas Shift Reaction in the Presence and Absence of HTS Catalyst
The extent of the water gas shift reaction was determined at different
temperatures in an empty stainless steel reactor. The reactant gases were made to flow
through the empty heated reactor and the product gases were analyzed by means of
continuous analyzers. The extent of the water gas shift reaction was also determined in
the presence of the HTS catalyst obtained from Süd-Chemie. 0.25 g of the catalyst was
loaded into the reactor and the pressure, temperature and gas flow rates were adjusted
for each run. The dry gas compositions at the outlet of the reactor were monitored
continuously using the CO, CO2, H2S, CH4 and H2 gas analyzers. The total flow rate
of the gases through the reactor was maintained a constant at 725 sccm for all the
experiments and the concentration of CO in the reaction mixture was maintained at
10.3 %.
5.2.4 Simultaneous Water Gas Shift and CO2 Removal
The combined water gas shift and carbonation reaction was conducted either
using a catalyst-sorbent mixture or non-catalytically, using only the sorbent without the
water gas shift catalyst. The combined experiments were conducted using a sorbent
(CaO) to catalyst ratio of 10:1 by weight or only CaO sorbent. The effect of various
131
temperatures (600, 650, and 700°C), S:C ratios (3:1, 2:1, 1:1) and pressures (1-21 atm)
was investigated. The CaCO3 sorbent was calcined by heating the sorbent-catalyst
mixture or only the sorbent to 700 oC in a stream of N2 until the CO2 analyzer
confirmed the absence of CO2 in the outlet stream. Multicyclic experiments were
conduct in the fixed bed reactor with only CaO sorbent by alternating the carbonation
and calcination steps and switching between the above mentioned temperatures and
feed gas streams. The total flow rate of the gases through the reactor was maintained a
constant at 725 sccm for all the experiments and the concentration of CO in the
reaction mixture was maintained at 10.3 %.
5.2.5 Combined H2 Production with H2S Removal:
To study the effect of sulfur on the CLP, 5000 ppm of H2S is mixed with the
CO, N2 and steam before being sent to the reactor. The H2 production tests are
conducted in the presence CaO sorbent as described in the section above.
5.3 RESULTS AND DISCUSSION
5.3.1 Baseline Water Gas Shift Reaction Testing
Base line experiments in an empty stainless steel reactor and in the presence of
a HTS catalyst were conducted to study the kinetics of the water gas shift reaction. A
comparison of the extent of the water gas shift reaction in the presence and absence of
a catalyst gives a perspective of the feasibility of eliminating the need for the water gas
132
shift catalyst in the carbonation reactor of the CLP. Figure 5.2 shows the CO
conversion obtained when a 10%CO and 90% N2 feed stream is reacted with steam at
different temperatures in an empty stainless steel reactor and in a stainless steel reactor
with HTS catalyst at atmospheric pressure. The CO conversion in the presence of HTS
catalyst was higher than in the empty reactor at temperatures lower than 800°C. In both
the presence and absence of the catalyst, the CO conversion increases with the increase
in temperature due to higher kinetics of the water gas shift reaction. Beyond a
particular optimum temperature, the CO conversion decreases with the increase in
temperature due to the thermodynamic limitation of the water gas shift reaction. It can
be seen that, as expected, the CO conversion increases with the increase in S:C ratio.
The effects of reaction temperatures and S:C ratios on CO conversion at 21 atm, shown
in Figure 5.3, follow the same trend as that at 1 atm. These baseline experiments show
that CO conversion occurs even in the absence of a catalyst due to rapid kinetics of the
water gas shift reaction in the temperature range of 500 to 750oC which is the
temperature range at which CO2 removal occurs with CaO sorbent. Hence this CO
conversion achieved in the empty reactor can be further improved by the addition of
CaO sorbent to the reaction system and removing the thermodynamic constraint of the
water gas shift reaction.
5.3.2 Water Gas Shift Reaction in the Presence of Only CaO Sorbent
The results obtained above lead to the conclusion that the water gas shift
reaction takes place to a considerable extent even in the absence of a catalyst at
133
relatively higher temperatures than the conventional water gas shift reaction. Hence it
is possible to increase the yield of H2 from the water gas shift reactor by shifting the
equilibrium of the reaction in the forward direction by removing the CO2 product
formed. The CO2 formed by the water gas shift reaction is removed using CaO sorbent.
Figure 5.4 (a) shows the N2 and steam free gas concentration at the outlet of the reactor
due to the combined water gas shift and carbonation reaction at 600oC and 21 atm.
High purity H2 is produced with very low levels of CO and CO2 during the prebreakthrough region of the curve when the CaO sorbent is active. As the CaO sorbent
gets consumed, the purity of H2 reduces and the concentration of CO and CO2 increase
in the breakthrough region of the curves. In the post-breakthrough region of the curve,
the CaO sorbent is completely consumed and the composition of the outlet gas is
similar to the composition at the outlet of the non catalytic water gas shift reaction.
Figure 5.4 (b) illustrates the CO conversion obtained with time for the gas
compositions obtained in Figure 5.4 (a). Almost complete conversion of CO is
obtained in the pre-breakthrough region of the curve where the combined water gas
shift and carbonation reaction takes place.
5.3.2.1 Effect of Pressure and S:C Ratio
Pressure has been found to have an important role in increasing the purity of H2
by the combined water gas shift and carbonation reaction in the presence of CaO
sorbent. Figure 5.5 shows the effect of the change in pressure on CO conversion at a
temperature of 650 oC and S:C ratio of 3:1. The CO conversion is found to increase
134
with the increase in pressure. At 1 atm, a clear pre-breakthrough region is not obtained
and a 90 to 95% CO conversion is obtained in the initial part of the breakthrough
curve. As the pressure is increased to 4.5 atms, a pre-breakthrough CO conversion of
greater than 98% is observed and at a pressure of 21 atms, almost 100% CO conversion
is observed in the pre-breakthrough region.
Since pressure has been found to be an important variable, the combined effect
of pressure and S:C ratio was investigated to determine conditions where the S:C ratio
can be decreased without causing a large decrease in CO conversion or H2 purity.
Combined water gas shift and carbonation experiments were conducted in the absence
of a catalyst for various S:C ratios and pressures ranging from 1 to 21 atm. When the
S:C ratio is decreased from 3:1 to 1:1 at ambient pressure, the CO conversion decreases
in the breakthrough curve as shown in Figure 5.6 (a). At higher pressures of 11 and 21
atms, there is almost no decrease in the initial pre-breakthrough CO conversion with
the decrease in S:C ratio. As illustrated in Figure 5.6 (b) at a pressure of 11 atms, the
CO conversion remains at 98 to near 100% for both S:C ratios of 3:1 and 1:1. At 21
atms, a near 100% CO conversion is obtained in the pre-breakthrough curve for all S:C
ratios of 3:1, 2:1 and 1:1 as shown in Figure 5.6(c). Hence by operating the carbonation
reactor at high pressures it is possible to reduce the excess steam addition without
causing a decrease in the CO conversion and the corresponding H2 purity.
135
5.3.2.2 Effect of Temperature
The effect of temperature was investigated at various S:C ratios for the
combined water gas shift and carbonation reaction. Figures 5.7 (a) and (b) illustrate the
change in CO conversion obtained when the temperature is varied from 600 to 700oC
at two S:C ratios of 3:1 and 1:1 at atmospheric pressure. At both S:C ratios, it can be
seen that the highest CO conversion in the pre-breakthrough region is obtained at
600oC and the CO conversion decreases with the increase in temperature due to the
highly exothermic nature of the combined water gas shift and carbonation reaction. A
reverse trend is obtained in the post-breakthrough region where the water gas shift
reaction occurs in the absence of both a sorbent and a catalyst and its rate increases
with the increase in temperature.
5.3.2.3 Effect of CO Concentration in the Feed Gas
The effect of CO concentration in the reactant gas was investigated at a
pressure of 11 atms on the CO conversion and purity of H2 produced for the same
amount of sorbent loaded. As shown in Figures 5.8 (a) and (b), near 100% CO
conversion and high purity H2 was produced for both 10% and 15% CO in the feed
stream. With an increase in the CO concentration, the pre-breakthrough region of the
curve becomes shorter. This is due to the higher flow rate of CO2 produced from the
136
CO in the feed by the water gas shift reaction which results in the faster conversion of
the CaO bed to CaCO3.
5.3.2.4 Sorbent Characterization and Morphology Analysis
Scanning Electron Microscopy (SEM) analysis was conducted on the sorbent
samples, to visualize the changes in the physical structure of the sorbent. PCC sorbent
was examined using SEM as shown in Figure 5.9 (a). It can be seen that the surface of
PCC is rough and the structure is porous and not dense like the structure of the
limestone sample observed by Abanades and Alvarez (Abanadez and Alvarez, 2003). It
can be clearly observed that the structure of PCC has been modified to improve the
porosity by introducing mesopores in the structure using surface modifying agents. The
PCC was then calcined to form PCC- CaO which was also examined under the SEM.
Figure 5.9 (b) is the image of a freshly calcined sample of PCC and it shows smaller
sized clusters than the PCC precursor. The calcined sorbent is then used in the water
gas shift reactor at atmospheric pressure to remove the CO2 produced and to shift the
equilibrium of the water gas shift reaction in the forward direction, thereby increasing
the yield and purity of H2. Figure 5.10 (a) shows the surface characteristics and pore
structure of the PCC sorbent which has undergone carbonation during the water gas
shift reaction at atmospheric pressure. During H2 production, ~ 70% conversion of CaO
to CaCO3 was obtained. It can be seen that the structure and surface of the first
carbonated sample is different from the fresh PCC sample shown in Figure 5.9 (a).
Elongated structures can be observed on the surface of the first carbonated sample.
137
Figure 5.10 (b) shows the surface structure for calcium sorbent which has undergone
carbonation during H2 production at 21 atms. At 21 atms it is found that ~85%
conversion of CaO to CaCO3 is obtained. It can be seen that the surface structure
formed during carbonation at 21 atms is similar to that formed during carbonation at 1
atm, but is denser due to the compaction at higher pressures.
5.3.3 H2 Production in the Presence of CaO Sorbent Only and a Mixture of CaO
Sorbent and Catalyst
The effect of the presence of water gas shift catalyst in the carbonation reactor
was investigated at atmospheric pressure and a high pressure of 21 atms. Figure 5.11
depicts the H2 purity obtained in the presence and absence of the catalyst at
atmospheric pressure. It can be seen that the H2 purity obtained in the presence of the
catalyst is 90% while it is 70% in the absence of the catalyst. In addition, a clear prebreakthrough region is observed in the presence of the catalyst for H2 purity while there
is almost no pre-breakthrough region in the absence of the catalyst. In contrast, at a
high pressure of 21 atm, there is no difference in the purity of H2 produced in the
absence and presence of the catalyst. Almost 100% pure H2 is produced in both cases.
The same effect is observed at both S:C ratios of 3:1 and 2:1 as shown in Figure 5.12.
Hence although the catalyst can be eliminated without causing a decrease in H2 purity
at high pressures the same is not true at atmospheric pressure. However, in commercial
facilities most of the H2 production applications are typically deployed at high
pressures.
138
5.3.4 Multicyclic Investigation of H2 Production in the Presence of CaO Sorbent
Only
Multicyclic reaction and regeneration of the calcium sorbent was conducted to
determine the effect of the number of cycles on the purity of H2 produced in the fixed
bed reactor. During the reaction step, H2 was produced from a 10% CO/90% N2 feed
stream in the presence of CaO sorbent. The gas compositions for CO, CO2, H2 and
hydrocarbons were recorded using continuous analyzers connected to a computer
program. At the end of the reaction step, the sorbent was calcined at 750oC in N2.
Figure 5.13 illustrates the purity of H2 obtained when the reaction step is conducted at
a pressure of 4.5 atms for 10 cycles. The purity of H2 in the product stream is found to
decrease with sorbent cycling from near 100% to 97% at the end of 10 cycles. In
addition, it can be observed that for each additional cycle, the pre-breakthrough region
is shorter than the previous one. This trend might be due to the reduction in useful
porosity available for the carbonation of CaO due to sintering of the sorbent.
Figure 5.14 illustrates the N2 and steam free H2 purity obtained from a 10%
CO/90% N2 feed stream in the presence of CaO sorbent at an operating pressure of 21
atm. The purity of H2 in the pre-breakthrough region remains almost constant for 10
cycles. However, the time for which the pre-breakthrough region lasted decreased with
the increase in the cycle number but to a lower extent than at 4.5 atms. The shortening
of the pre-breakthrough region with each progressive cycle again might be attributed to
sorbent sintering.
139
5.3.5 Enhanced H2 Production With CO2 and Sulfur Capture
In the CLP, the CaO sorbent assumes the role of a multipollutant capture
sorbent, in addition to enhancing the water gas shift reaction. Hence, the influence of
various process variables like temperature, S:C ratio and pressure on the purity of H2
produced and the extent of H2S removed during the combined water gas shift,
carbonation and sulfidation reaction of CaO was determined.
5.3.5.1 Effect of S:C Ratio
Figures 5.15 (a) and (b) illustrate the effect of varying S:C ratio on the extent of
H2S removal and the purity of H2 produced in the combined water gas shift,
carbonation and sulfidation reaction. The extent of H2S removal by the CaO sorbent is
found to increase with the decrease in S:C ratio in the carbonation reactor. As shown in
Figure 5.15 (a), the concentration of H2S in the product H2 stream decreases from 100
ppm to <1ppm with the decrease in S:C ratio from 3:1 to 0.75:1. This decrease in H2S
concentration is due to a reduction in the inhibiting effect of steam on the reaction
between H2S and CaO. The same inference was drawn from the thermodynamic
analysis shown in Chapter 4. The effect of the change in S:C ratio on the purity of H2 is
illustrated in Figure 5.15 (b). Similar to observations made earlier in this chapter, at
atmospheric pressure the purity of H2 is found to decrease with the decrease in S:C
ratio during the breakthrough region of the curves.
140
5.3.5.2 Effect of Temperature
Figures 5.16 (a) and (b) illustrate the effect of temperature on the extent of H2S
removal and the purity of H2 produced respectively via breakthrough curves. A low
concentration of H2S in the order of ~10 ppm is detected in the outlet H2 stream at
temperatures ranging from 560 to 600 oC at atmospheric pressure. With the increase in
the temperature above 600 oC, the H2S concentration in the outlet is found to increase
to 50 ppm at 650 oC and 90 ppm at 700 oC. The effect of temperature on the purity of
H2 has been illustrated in Figure 5.16 (b). The H2 purity is found to be the highest
(70%) within the temperature range of 600-650 oC. The purity of H2 is found to
decrease to 60% with the decrease in temperature to 560 oC. A similar effect is
observed with the increase in temperature to 700 oC. Hence, from Figures 5.16 (a) and
(b), it can be inferred that the optimum temperature of operation for simultaneous H2
production and H2S removal is ~600 oC.
5.3.5.3 Effect of Pressure
Pressure has been found to be a very important variable for the non catalytic
production of high purity H2 at low S:C ratios. The effect of the increase in pressure on
the extent of H2S removal and the purity of H2 produced is illustrated in Figures 5.17
(a) and (b). The concentration of H2S in the product H2 stream is found to decrease
from 10 ppm to <1ppm when the pressure is increased from 1 atm to 21 atm. Hence,
the combined effect of operating at a low S:C ratio and high pressure results in the
141
production of a H2 stream with <1 ppm of sulfur impurities. Figure 5.17 (b) illustrates
the effect of the increase in pressure on the purity of H2 produced during the combined
water gas shift, carbonation and sulfidation reaction. At a temperature of 600 oC and a
stoichiometric S:C ratio, the purity of H2 is found to increase from 70% to >99% with
the increase in pressure of 1 atm to 21 atm. Hence, the CLP is capable of producing
high purity H2 (>99%) with <1 ppm of sulfur impurities in it at stoichiometric S:C
ratios.
5.3.5.4 Sorbent Characterization and Morphology Analysis
The scanning electron microscopy (SEM) analysis is conducted on the sorbent
samples, to visualize the changes in the physical structure of the sorbent. PCC sorbent
is examined using SEM as shown in Figure 5.18 (a). It can be seen that the surface of
PCC is rough and the structure is porous. The PCC is then calcined to form PCC- CaO
which is also examined under the SEM. Figure 5.18 (b) is the image of a freshly
calcined sample of PCC and it shows smaller sized clusters than the PCC precursor.
The calcined sorbent is then used in the carbonation reactor at atmospheric pressure to
remove the CO2 and H2S and shift the equilibrium of the water gas shift reaction in the
forward direction, thereby increasing the yield and purity of H2. Figure 5.18 (c) shows
the surface characteristics and pore structure of the PCC sorbent which has undergone
carbonation and sulfidation during the water gas shift reaction at atmospheric pressure.
It can be seen that the structure and surface of the first carbonated sample is different
from the fresh PCC sample shown in Figure 5.18 (a). Elongated structures can be
142
observed on the surface of the first carbonated sample. Figure 5.18 (d) shows the
surface structure for calcium sorbent which has undergone carbonation and sulfidation
during H2 production at 21 atms. It can be seen that the surface structure formed during
carbonation at 21 atms is similar to that formed during carbonation at 1 atm, but is
denser due to the compaction at higher pressures.
5.4. H2 PRODUCTION FROM COAL GASIFICATION DERIVED SYNGAS
5.4.1 Process Overview
In the conventional coal gasification process, H2 can be produced from coal
through the sweet shift or the sour shift route (Stiegel and Ramezan, 2006). Figure
5.19(a) illustrates the conventional coal to H2 process in which coal is fed along with
steam and/or oxygen to the gasifier to produce syngas. In the sweet shift route, the
syngas is cooled in a radiant cooler. The ash is then separated from the cool syngas
which is fed to a syngas scrubber for ammonia and HCl removal. Following this, sulfur
is removed using a solvent based system as the commercial HTS catalyst has a sulfur
tolerance of about several hundred ppms while the LTS catalyst has a lower tolerance
to sulfur and chloride impurities. Ash, ammonia, HCl and sulfur removal is conducted
at low temperatures of 40 to 200oC which is energy intensive due to the gas cooling
and reheating requirements (Haussinger et al, 2000). The syngas temperature is then
raised for the water gas shift reaction. Higher temperatures enhance the kinetics of the
water gas shift reaction. However, as shown in Chapter 4, the equilibrium limitation of
143
the water gas shift reaction adversely affects H2 production, with the H2 yield falling
with rising temperature. Hence, a high S:C ratio is required to enhance CO conversion
and the consequent H2 yield. The S:C ratio required at 550 oC can be as high as 50 in a
single-stage operation or 7.5 for a more expensive dual-stage process to obtain 99.5 %
pure H2 (David, 1980). Numerous research studies have focused on the development of
low temperature catalysts to improve H2 production (David, 1980). Commercially, the
dual stage sweet water gas shift reaction is carried out in series, with a HTS (300-450
o
C) stage containing iron oxide catalyst to convert bulk of the CO and a LTS (180-270
o
C) stage containing copper catalyst ( Haussinger et al, 2000). Following the shift
reactors, the syngas is fed to a mercury removal unit and a CO2 capture unit based on
physical solvents like selexol or, rectisol or chemical solvents like amine based
solvents. For high purity H2 production, a PSA is used as the final step and the tail gas
from the PSA is combusted to produce electricity. In a sour gas shift system, syngas is
cooled using a water quench which provides the excess steam required for the water
gas shift reaction and removes impurities like ash, HCl and ammonia (Holt, 2005, MIT,
2007). Since the sulfur content of synthesis gas is greater than 1000 ppm, a sulfided
catalyst is used in a series of reactors at a temperature of 250–500 °C (Lloyd et al,
1996, Hiller et al, 2007). CO2 and sulfur removal is achieved in a dual stage acid gas
removal system and the H2 is finally purified in a PSA.
Figures 5.19(b) and 5.20 show the integration of the CLP in a typical coal or
biomass gasification system with the cogeneration of electricity and H2. The syngas
144
from the gasifier is cooled in a radiant heater and fed along with steam and CaO to the
carbonation reactor in the CLP. The water gas shift reaction almost goes to completion
in the presence of the CaO sorbent. The CaO sorbent reacts with the CO2, sulfur and
halide impurities and removes them from the product stream. The product gas stream
from the reactor contains predominantly H2 which is purified further in a PSA for ultra
pure applications (eg. Fuel cells). The H2 stream upstream of the PSA could also be
converted to electricity in a combined cycle system for the generation of electricity or
used for the production of fuels and chemicals. The spent sorbent from the carbonation
reactor is then regenerated in the calciner where a sequestration ready CO2 stream is
produced. When calcination is conducted in the presence of steam, a CO2 stream
containing a small concentration of H2S is produced from the calciner, which can then
be sequestered as is (Smith et al, 2007). Since CaS and CaCl2 cannot be regenerated
completely, a portion of the sorbent mixture is purged at the exit of the carbonation
reactor. Fresh sorbent make up is added upstream of the calciner. The amount of purge
and makeup will depend on the sulfur and chloride content of the coal syngas and the
extent of sintering of the sorbent. The sorbent makeup and purge will result in the
production of a H2 stream with constant purity and will prevent the accumulation of
inert material (CaCl2 and CaS) in the circulating sorbent mixture. On comparison of
Figures 5.19(a) and (b) it can be seen that by using the CLP, the unit operations in the
coal to H2 process can be significantly reduced.
145
5.4.2 System Thermodynamics Analysis
Thermodynamic analysis was conducted for the reactions occurring in the
carbonation reactor with syngas feed compositions from different gasifiers. The
equilibrium constants for the reactions occurring in the carbonator were obtained using
HSC Chemistry v 5.0 (Outokumpu Research Oy, Finland) software and are shown in
Chapter 4. Table 5.1 shows the syngas compositions from different gasifiers. It is noted
that air is used as the oxidant in the moving bed, dry gasifier while oxygen is used in
all the other gasifiers. Using the composition of the syngas, the feasibility of the
simultaneous water gas shift, carbonation and sulfidation reaction in the temperature
range of 500-750 ºC was determined. The temperature range of 500-750 ºC was chosen
as the preferred operating range primarily because the kinetics of CO2 capture by CaO
is high and the equilibrium partial pressure of CO2 is low as shown in Chapter 4. Steam
is added to the syngas before it is fed to the carbonation reactor to adjust the S:C ratio
for the water gas shift reaction. Tables 5.2 and 5.3 list the compositions of the syngassteam mixture for S:C ratios of 1:1 and 3:1 respectively. Depending on the
thermodynamic extent of the water gas shift reaction occurring in the carbonator, the
PCO2 in the carbonator can be determined. Figure 5.21(a) illustrates the PCO2 in the
carbonator after the water gas shift reaction has occurred in the syngas-steam mixture
with a S:C ratio of 1:1 from various gasifiers. It also shows the equilibrium PCO2
required for the carbonation reaction with CaO. It can be seen that the PCO2 in the
carbonator for various gasifiers is higher than the equilibrium PCO2 for the carbonation
146
of CaO and hence it can be inferred that in the temperature range of 500-750 ºC, CO2
removal is achieved by CaO to greater than 90%. Figure 5.21(b) illustrates the PCO2 in
the shifted syngas-steam mixture for a S:C ratio of 3:1. It can be seen that Figure
5.21(b) is very similar to Figure 5.21(a) and CO2 removal is achieved by CaO to high
extents for a S:C ratio of 3:1.
In order to determine whether the CaO sorbent will undergo hydration, the
PH2O in the carbonation reactor was determined after the combined water gas shift and
carbonation reaction has occurred to equilibrium for S:C ratios of 1:1 and 3:1. Figure
5.22(a) shows the PH2O in the carbonator for a S:C ratio of 1:1 and since the PH2O in
the carbonator is lower than equilibrium PH2O for the hydration of CaO, hydration will
not occur for any syngas composition in the temperature range under consideration.
Figure 5.22(b) shows the PH2O in the carbonator for a S:C ratio of 3:1 and it can be
seen that hydration will occur at lower temperatures. Hydration of CaO will occur at
temperatures below 600 ºC for fluidized bed and moving bed ( dry), below 670 ºC for
entrained flow (dry), below 680 ºC for moving bed (slagging) and below 700 ºC for
entrained flow ( slurry) gasifier syngas.
Figures 5.23 (a) and (b) depict the equilibrium CO conversion obtained in a
conventional water gas shift reactor. The CO conversion for all syngas compositions
and S:C ratios decreases with the increase in temperature due to the equilibrium
limitation of the exothermic water gas shift reaction. The CO conversion increases with
the increase in S:C ratio from 1:1 to 3:1 for all compositions of syngas. Figures 5.24
147
(a) and (b) illustrate the CO conversion obtained in the carbonation reactor of the CLP
for different syngas feed compositions. It can be seen that the CO conversion is
enhanced in the presence of the CaO sorbent in the CLP in comparison to the
conventional water gas shift reactor. It can be seen that greater than 95% CO
conversion can be obtained for syngas from all the gasifiers by operating in the
temperature range of 550 to 650°C. Although greater CO conversions can be obtained
at temperatures lower than 550°C, the kinetic of the water gas shift reaction and CO2
removal by CaO will decrease, resulting in the need for the use of larger reactors.
Figures 5.25 (a) and (b) illustrate the purity of H2 produced by the carbonation
reactor in the CLP at S:C ratios of 1:1 and 3:1. High H2 purities can be obtained at both
S:C ratios of 3:1 and 1:1 in all the gasifiers where oxygen is used as the oxidant. In the
moving bed, dry gasifier, lower H2 purities are obtained due to dilution by nitrogen
since air is used as the oxidant in the gasifier.
Figure 5.26 depicts the total equilibrium carbon capture obtained in the CLP at
S:C ratios of 1:1 and 3:1. The % carbon captured is defined as the total moles of carbon
in the form of CO, CO2 and CH4 that is removed in the process. The % carbon captured
for all syngas compositions and S:C ratios decreases with the increase in temperature.
This decrease is due to the decrease in not only CO conversion but also CO2 capture by
CaO at high temperatures. It can be seen that greater than 95% carbon capture can be
obtained from all the gasifiers by operating in the temperature range of 550 to 650 oC .
Although greater CO conversions can be obtained at temperatures lower than 550 oC ,
148
the kinetic of the water gas shift reaction and CO2 removal by CaO will decrease,
resulting in the need for the use of larger reactors. The % carbon captured increases
with the increase in S:C ratio from 1:1 to 3:1 for all compositions of syngas.
5.5 H2 PRODUCTION FROM SYNGAS DERIVED FROM NATURAL GAS FEEDSTOCKS
5.5.1 Syngas from Steam Reforming of Natural Gas
Steam reforming has been used commercially since 1931 for the conversion of
hydrocarbons to H2 as shown in Figure 5.27 (Gunardson, 1998). Catalytic steam
reforming is conducted at high temperatures of 800 to 1000 ºC and pressures of 7.9 to
24.7 atms ( 8 – 25 bars)in the presence of a catalyst. Nickel based catalyst is most
commonly used and is resistant to sulfur poisoning in hydrocarbon feeds containing
less than 0.1 ppm of sulfur. Since the steam reforming process is endothermic, heat is
supplied by burning natural gas in air which produces a flue gas. The hydrocarbon feed
to the reformer is preheated by the product syngas and the reformer flue gas. As a first
step in the reforming process the feed is hydrogenated at 290- 370 ºC to convert all the
organic sulfur to H2S. For hydrocarbon feeds containing high concentrations of organic
sulfur, the feed is scrubbed with amine solution to reduce the concentration of H2S to
25 ppm. The feed is then passed over a ZnO catalyst at 340-370 ºC to further reduce
the H2S concentration to 0.01 ppm and enters the reformer. At the exit of the reformer
the hot gases are cooled to 340 – 455 ºC and sent to the HTS reactor. For maximizing
the H2 yield from this process, the gases from the exit of the HTS reactor are further
149
shifted in a LTS stage. The gases are then cooled and fed to a CO2 removal unit where
the CO2 is scrubbed using amine based solvents or selexol, rectisol, etc. Since CO and
CO2 are poisons for applications like NH3 synthesis, refinery processes and
petrochemical processes, CO and CO2 need to be removed to less than 5 ppm in the H2
stream. This is achieved in the methanation reactor at 310 ºC. (Gunardson, 1998)
High purity H2 is also produced by using a Pressure Swing Absorption unit
which replaces the CO2 removal units and the methanator. Figure 5.28 illustrates the
process flow for the production of H2 using a steam reformer and PSA assembly. The
syngas from the reformer is shifted in a HTS stage since the performance and
efficiency of the PSA is dependent on the purity of H2 fed to it. Addition of a LTS
stage will improve the recovery of H2 and the efficiency of the PSA while increasing
the cost of H2 production. (Gunardson, 1998)
Figure 5.29 illustrates the integration of the CLP in an existing steam methane
reformer for the conversion of syngas to H2 with simultaneous CO2 capture. The
syngas produced in the reformer consists of a mixture of CO, CO2, H2, CH4 and steam.
The dry gas composition of the syngas is shown in Table 5.4. (Gunardson, 1998) The
syngas is fed to the carbonation reactor in the CLP along with CaO sorbent and steam.
The CO in the syngas is converted to H2 and the CO2 product is simultaneous removed
by the CaO sorbent. The equilibrium product gas composition at the exit of the
carbonator at a temperature of 650 ºC and a pressure of 25 atms is also shown in Table
150
5.4. From thermodynamic analysis, about 99% of CO conversion to H2 and 99% CO2
removal by CaO sorbent is achieved in the carbonation reactor.
In addition to converting syngas feed to H2, the CLP can also be used to convert
hydrocarbons like natural gas directly to H2 in a single stage reactor. This configuration
of the CLP is elaborated in Chapter 8.
5.5.2 Syngas from Autothermal Reforming of Natural Gas
In autothermal reforming, a portion of the natural gas is combusted in the
reformer with pure oxygen from an ASU. The hot gases (1200-1250 ºC) are then
passed through a catalyst bed where the steam methane reforming and water gas shift
reactions take place to equilibrium. Table 5.5 shows the composition of syngas from
the exit of the autothermal reformer. (Gunardson, 1998) If the CLP is added to the
autothermal reformer for the production of H2 in a carbon constrained scenario, then
99% of CO conversion and CO2 removal can be obtained in the process. The
equilibrium product gas from the carbonation reactor of the CLP is shown in Figure
5.5.
5.5.3 Syngas from Partial Oxidation of Natural Gas:
The partial oxidation of natural gas is an uncatalyzed reaction with steam and
oxygen to produce syngas. A schematic of the H2 production process from syngas
derived from the partial oxidation of natural gas is shown in Figure 5.30. The syngas is
151
then cooled in a radiant heater or using a water quench system. The water quench in
addition to cooling the gas also saturates the gas with water which is used for shifting
the CO in the syngas. The syngas is then scrubbed with water to remove carbon and is
fed to a high and low temperature shift system. The CO2 in the shifted syngas is then
removed using amine solvents, rectisol, selexol, or other solvents. (Gunardson, 1998)
The HTS, LTS and CO2 capture unit operations can be replaced by the CLP.
Table 5.6 illustrates the composition of syngas obtained from the partial oxidation of
natural gas. (Gunardson, 1998) When this syngas is fed to the carbonation reactor,
99% CO conversion and 97% CO2 removal is achieved in the presence of steam of the
CaO sorbent. Figure 5.31 illustrates the integration of the CLP with a partial oxidation
system for the production of high purity H2.
5.6 ADDRESSING THE ISSUE OF SULFUR IN THE FEEDSTOCK
The effect of sulfur on the CLP is significant when the CLP is applied to precombustion systems. During the production of H2 and electricity from syngas
containing H2S and COS, as in coal gasification derived syngas, the CaO reacts with
H2S and COS to form CaS.
The CaS is stable in the carbonation reactor under reducing conditions. In a
direct oxyfired calciner, the CaS is stable if the calciner is maintained under reducing
conditions. This method was used in the CO2 acceptor process where the air fired coal
calciner was maintained under reducing conditions by the addition of 5% CO. If the
152
calciner atmosphere is oxidizing, then the CaS is converted to CaSO4 in the presence of
O2, CO2 and H2O, as shown below, depending on the residence time of the solids in the
calciner.
CaS oxidation: CaS + 2O2 Î CaSO4
(5.1)
CaS + 4CO2 Î CaSO4 + 4CO
(5.2)
CaS + 4H2O Î CaSO4 + 4 H2
(5.3)
In a direct oxyfired coal calciner, CaSO4 may also be produced from the direct
reaction of CaO with SO2 in the presence of O2 as shown below:
CaO + SO2 + 0.5 O2 Î CaSO4
(5.4)
In an oxidizing atmosphere in the calciner, the CaS and CaSO4 may form a
eutectic mixture producing CaO and SO2 as shown below:
CaS + 3CaSO4 Î 4CaO + 4SO2
(5.5)
The SO2 produced in the calciner exits with the CO2 and the entrained sorbent
mixture. The sorbent and gas mixture is cooled at the exit of the calciner before it is fed
to a particle capture device. The SO2 in the gas is captured by the sorbent when the
sorbent and gas mixture is cooled before the particle capture device. Hence almost no
SO2 is present in the CO2 stream that is sent for sequestration.
153
Thermodynamic analyses predict the complete conversion of CaS to CaO in the
calciner and the mixture of solids entering the hydrator contains only CaO, CaCO3,
CaSO4 and inerts (flyash from the calciner and inerts in the limestone sorbent). During
actual operation, the extent of CaS conversion to CaO will depend on the residence
time in the calciner. The residence time in the calciner can be as short as 2 secs in a
commercial flash calciner. Hence if CaS is present in the solid mixture entering the
hydrator then the CaS may be converted to CaO and H2S in the presence of the steam
as shown below.
CaS + H2O Î CaO + H2S
(5.6)
The solids at the exit of the hydrator will be a mixture of Ca(OH)2, CaO,
CaSO4, CaCO3 and CaS. When these solids are fed into the carbonator, the CaSO4 will
get reduced in the presence of CO and H2 as shown below.
CaSO4 + 4CO Î CaS + 4CO2
(5.7)
CaSO4 + 4H2 Î CaS + 4H2O
(5.8)
The reduction of CaSO4 results in a decrease of about 10% of H2 yield from the
carbonation reactor due to the consumption of CO and H2. Although H2 yield is
reduced in the carbonation reactor, the subsequent oxidation of the CaS in the calciner
is exothermic and the energy released in the calciner aids in the calcination of CaCO3.
This will help in reducing the amount of coal that needs to be added to the oxyfired
154
calciner. Hence the efficiency of the overall process will not be changed significantly
due to the reactions involving sulfur in the CLP.
Depending on whether the calciner is operated in an oxidizing or reducing mode, the
composition of the sorbent mixture circulating through the CLP and the composition of
the purged solids will vary. If CaS is present in the purged solids it will have to be
converted to CaSO4 or CaO before disposal.
It has been shown in the CO2 acceptor process and the HyPr-RING process that
the complete oxidation of CaS does not occur in one stage in the calciner and hence
CaS is always present in the solids mixture. The oxidation of CaS does not go to
completion since the CaSO4 formed has a higher molar volume than CaS. The
oxidation reaction is slowed down by diffusional resistance and CaSO4 forms an outer
layer leaving CaS in the core. Squires et al and Keairns et al suggested using a mixture
of H2O and CO2 to convert CaS to CaCO3 and H2S (Squires et al,1971, Keairns et al,
1976). They found that the rate of the regeneration reaction increased with the increase
in temperature and complete regeneration could be achieved at 650 ºC. An
investigation of CaS regeneration with H2O and CO2 is described in the following
sections.
5.6.1 Experimental Analysis of the Regeneration of CaS:
The regeneration of a mixture of CaS and CaCO3 was investigated in a fixed
bed reactor setup. The samples for the regeneration experiments were obtained from
155
the combined water gas shift, carbonation and sulfidation experiments conducted for
H2 production at 600 ºC and at different pressures illustrated in Chapter 5. The
regeneration was studied in the presence of steam alone and in the presence of steam
and CO2. 1.5g of the spent sorbent containing a mixture of CaCO3 and CaS was packed
in the fixed bed reactor. High pressure steam was produced by pumping water into a
heated tube and the steam produced was carried into the preheating section of the
reactor by nitrogen in the case of regeneration in the presence of steam alone or by a
mixture of nitrogen and CO2 in the case of regeneration in the presence of steam and
CO2. The preheated gas mixture was then fed into the heated fixed bed reactor
containing the spent sorbent. At the exit of the reactor, the mixture of gases was cooled
and conditioned and fed into a continuous analyzer system for recording the
concentration of H2S, CO2 and CO. These concentrations were then recorded
continuously for every second to yield the breakthrough curve.
5.6.1.1 Regeneration in the Presence of Steam
The regeneration of CaS in the presence of steam is achieved by the following
reaction:
CaS + H2O Î CaO + H2S
(5.6)
A small amount of CaS is also oxidized by steam to give CaSO4 as given below:
CaS + 4H2O Î CaSO4 + 4H2
(5.3)
156
Figure 5.32 illustrates the H2S evolved from CaS at different temperatures of
650 and 700 ºC and steam compositions of 31% and 15%. The spent sorbent samples
for the three sets of data given below were obtained by conducting simultaneous water
gas shift, carbonation and sulfidation reaction at 600 ºC and 21 atms in the presence of
5000ppm of H2S. As shown in Figure 5.32, it was found that when the regeneration is
conducted in the presence of 31% steam and 69% nitrogen, a larger concentration of
H2S is obtained at the outlet of the reactor at 700 ºC than at 650 ºC. In addition it was
also found that when the regeneration reaction is conducted at 700 ºC the H2S
concentration at the outlet of the reactor is higher when a 31 % steam concentration is
used than when the 15% steam concentration is used. This suggests that the increase in
steam concentration increases the conversion of CaS to CaO.
5.6.1.2 Regeneration in the Presence of Steam and CO2
During the regeneration of CaS with H2O and CO2, the following reactions
occur:
CaS + CO2 + H2O Î CaCO3 + H2S
(5.9)
CaS + H2O Î CaSO4 + H2
(5.3)
CaS + H2O Î CaO +H2S
(5.6)
The regeneration of CaS was conducted in the presence of a mixture of steam
and CO2 and the effect of the change in concentration of steam and CO2 was studied on
the conversion of CaS to CaO. As shown in Figure 5.33, it was found that when a 15%
157
steam and 15% CO2 mixture is used the extent of conversion of CaS is higher than
when a mixture of 31% steam and 31% CO2 is used. In contrast it can be seen that the
rate of removal of H2S is higher when a 31% steam and CO2 mixture is used than when
the 15% steam and CO2 mixture is used. This suggests that with the 31 % steam and
CO2 mixture the initial rate of decomposition of H2S is very high but with time this
reaction is limited. This might be due to the sintering of the sorbent due to the presence
of steam or CO2.
Figure 5.34 illustrates the comparison in the regenerability of the sorbent which
has undergone carbonation and sulfidation at a high pressure of 21 atms and that which
has undergone carbonation and sulfidation at atmospheric pressure. Since the
sulfidation reaction is favored by the increase in pressure the sample obtained from the
carbonation and sulfidation experiment at 21 atms releases more H2S than the sample
obtained from the carbonation and sulfidation experiment conducted at atmospheric
pressure. The rate of release of H2S from the sample is also higher for the sample
treated at 21 atms.
5.7. CONCLUSION
The feasibility and optimum process conditions for the production of H2 in the
absence of a water gas shift catalyst were determined. Experimental analysis revealed
that CaO sorbent was found to enhance the thermodynamics of the water gas shift
reaction and H2 purity at a high reaction rate in the absence of the catalyst. Pressure
158
was found to have a large effect on H2 purity. At high pressures, typical of commercial
deployment, the absence of the catalyst and the reduction of excess steam addition did
not have any effect on CO conversion and high H2 purity (>99%) was obtained. A
greater enhancement in H2 purity was found to occur at lower temperatures of 600 and
650oC and the effect of CaO sorbent was found to diminish with the increase in
temperature. The effects of sintering of the CaO sorbent were observed on H2 purity
during multiple reaction and regeneration cycles without hydration. The effect of S:C
ratio, temperature, and pressure was also studied on H2 purity and the extent of H2S
removal by CaO sorbent. Lowering the S:C ratio in the carbonator was found to
improve the extent of H2S removal by the CaO sorbent. Greater than 99% H2 purity
with less than 1 ppm of H2S was obtained at a stoichiometric S:C ratio at high
pressures. The integration of the CLP in coal gasification and natural gas reforming
systems is also discussed and the advantages of the CLP process have been
highlighted.
159
Oxidant
Fuel
Pressure (atm)
CO (mole %)
H2 (mole %)
CO2 (mole %)
H2O (mole %)
N2 (mole %)
CH4+ HCs (mole %)
H2S + COS (mole %)
Moving Bed, Moving Bed Fluidized Entrained Entrained
dry
slagging
Bed
Flow, slurry Flow, dry
Air
Oxygen
Oxygen
Oxygen
Oxygen
Sub
Bituminous
Bituminous
Lignite
Bituminous Bituminous
20.1
31.6
9.9
41.8
24.8
17.4
46
48.2
41
60.3
23.3
26.4
30.6
29.8
30
14.8
2.9
8.2
10.2
1.6
…
16.3
9.1
17.1
2
38.5
2.8
0.7
0.8
4.7
5.8
4.2
2.8
0.3
…
0.2
1.1
0.4
1.1
1.3
Table 5.1: Typical fuel gas compositions obtained from different gasifiers (Stultz and
Kitto, 1992).
160
Oxidant
Fuel
Total Pressure (atm)
CO (atm)
H2 (atm)
CO2 (atm)
H2O (atm)
N2 (atm)
CH4+ HCs (atm)
H2S + COS (atm)
Moving Bed, Moving Bed Fluidized Entrained
dry
slagging
Bed
Flow, slurry
Oxygen
Oxygen
Oxygen
air
Sub
Bituminous
Bituminous
Lignite
Bituminous
20.1
31.6
9.9
41.8
2.97
11.24
3.42
13.81
3.98
6.45
2.17
10.04
2.53
0.71
0.58
3.44
2.97
11.24
3.42
13.81
6.58
0.68
0.05
0.27
0.99
1.03
0.20
0.10
0.03
0.27
0.03
0.37
Entrained
Flow, dry
Oxygen
Bituminous
24.8
9.46
4.71
0.25
9.46
0.74
0.00
0.20
Table 5.2: Fuel gas composition entering the water gas shift reactor after steam
addition (S:C ratio =1:1) (adapted from Stultz and Kitto, 1992)
161
Oxidant
Fuel
Pressure (atm)
CO (atm)
H2 (atm)
CO2 (atm)
H2O (atm)
N2 (atm)
CH4+ HCs (atm)
H2S + COS (atm)
Moving Bed,
dry
air
Sub
Bituminous
20.1
2.29
3.07
1.95
6.88
5.08
0.76
0.03
Moving Bed,
slagging
Oxygen
Bituminous
31.6
6.57
3.77
0.41
19.72
0.40
0.60
0.16
Fluidized Entrained Entrained
Bed
Flow, slurry Flow, dry
Oxygen
Oxygen
Oxygen
Lignite
9.9
2.02
1.28
0.34
6.06
0.03
0.12
0.02
Bituminous
41.8
8.32
6.05
2.07
24.96
0.16
0.06
0.22
Bituminous
24.8
5.37
2.67
0.14
16.11
0.42
0.00
0.12
Table 5.3: Fuel gas composition entering the water gas shift reactor after steam
addition (S:C ratio =3:1) (adapted from Stultz and Kitto, 1992)
162
CH4
CO
CO2
H2
Syngas from SMR(%) CLP Product Gas(%) % removal
8.373
9.276
11.058
0.082
99.000
9.479
0.096
99.000
71.090
90.545
Table 5.4: Extent of equilibrium CO conversion and CO2 capture in the CLP from
Steam Methane Reforming (SMR) derived syngas
163
CH4
CO
CO2
H2
N2
Syngas from ATR (%) CLP Product Gas (%) % removal
0.983
1.057
23.112
0.243
99.000
7.048
0.073
99.000
68.536
98.281
36.202
14.316
Table 5.5: Extent of equilibrium CO conversion and CO2 capture in the CLP from Auto
Thermal Reforming (ATR) derived syngas
164
CH4
CO
CO2
H2
N2
Syngas from POX (%) CLP Product Gas (%) % removal
0.300
0.003
35.000
0.202
99.432
2.600
0.075
97.166
61.100
98.703
1.000
1.017
Table 5.6: Extent of equilibrium CO conversion and CO2 capture in the CLP from
partial oxidation (POX) derived syngas
165
Thermocouple
And
Pressure Guage
Steam Generator
Steam &
Gas Mixture
Gas
Gas
Mixture
Mixture
Water In
Sorbent
MFC
H2
Heated Steel
Tube Reactor
Hydrocarbon
Analyzer
Back Pressure
Regulator
Analyzers (CO,
CO2, H2, H2S)
MFC
Water Syringe
Pump
Heat
Exchanger
Water Trap
Figure 5.1: Simplified flow sheet of the bench scale experimental setup
166
MFC
CO
MFC
CO2 Hydro
H2S
carbons
1.0
CO Conversion
0.8
0.6
0.4
0.2
0.0
400
500
600
700
800
900
Temperature (oC)
Empty Reactor, S:C - 3:1
With HTS Catalyst, S:C - 3:1
Empty Reactor, S:C - 2:1
With HTS Catalyst, S:C - 2:1
Empty Reactor, S:C - 1:1
With HTS Catalyst, S:C - 1:1
Figure 5.2: Effect of reaction temperature and S:C ratio on the conversion of CO by the
water gas shift reaction at 1 atm
167
1.0
CO Conversion
0.8
0.6
Empty reactor, S:C - 3:1
With HTS catalyst, S:C - 3:1
Empty reactor, S:C - 2:1
With HTS catalyst, S:C - 2:1
Empty reactor, S:C - 1:1
With HTS catalyst, S:C - 1:1
0.4
0.2
0.0
500
550
600
650
700
750
800
Temperature (C)
Figure 5.3: Effect of reaction temperature and S:C ratio on the conversion of CO by the
water gas shift reaction at 21 atm
168
Outlet Gas Compositions (%)
100
H2
CO2
CO
80
60
40
20
0
0
500
1000
1500
2000
2500
2000
2500
Time (sec)
(a)
1.0
CO Conversion
0.8
0.6
0.4
0.2
0.0
0
500
1000
1500
Time (sec)
(b)
Figure 5.4: Typical breakthrough curves for the production of H2 in the presence of
CaO sorbent without catalyst (a) Gas composition (mole%) and (b) CO
conversion (600 °C, 21 atm, S:C ratio of 3:1)
169
1.0
CO Conversion
0.8
0.6
0.4
1 atm
4.5 atm
21 atm
0.2
0.0
0
500
1000
1500
2000
2500
Time(sec)
Figure 5.5: Effect of pressure on CO conversion obtained in the presence of CaO
sorbent without catalyst (650°C, S:C ratio of 3:1)
170
1.0
CO Conversion
0.8
0.6
0.4
0.2
3:1
2:1
1:1
0.0
0
500
1000
1500
2000
2500
Time (sec)
(a)
1.0
CO Conversion
0.8
0.6
0.4
0.2
3:1
1:1
0.0
0
500
1000
1500
2000
2500
Time (sec)
(b)
Continued
Figure 5.6: Effect of S:C ratio on CO conversion obtained in the presence of CaO
sorbent without catalyst at (a) 1 atm, (b) 11 atm, (c) 21 atm (650°C)
171
Figure 5.6 continued
1.0
CO Conversion
0.8
0.6
0.4
0.2
3:1
2:1
1:1
0.0
0
500
1000
1500
Time (sec)
(c)
172
2000
2500
1.0
CO Conversion
0.8
0.6
0.4
600C
650C
700C
0.2
0.0
0
1000
2000
3000
Time(sec)
(a)
1.0
CO Conversion
0.8
0.6
0.4
600C
650C
700C
0.2
0.0
0
1000
2000
3000
Time(sec)
(b)
Figure 5.7: Effect of temperature on CO conversion obtained in the presence of CaO
sorbent without catalyst at a S:C ratio of (a) 1:1 and (b) 3:1 (1 atm)
173
1.0
CO Conversion
0.8
0.6
0.4
0.2
10% CO
15 % CO
0.0
0
500
1000
1500
2000
2500
TIme (sec)
(a)
100
10% CO
15% CO
H2 Purity (%)
80
60
40
20
0
0
500
1000
1500
2000
2500
TIme(sec)
(b)
Figure 5.8: Effect of CO concentration in the feed on the (a) CO conversion and (b)
purity of H2 produced in the presence on CaO sorbent without catalyst (11
atm, 600°C, S:C ratio of 3:1)
174
(a)
(b)
Figure 5.9: SEM image of the (a) initial CaCO3 sorbent (b) CaO sorbent obtained from
the calcination of CaCO3
175
(a)
(b)
Figure 5.10: SEM image of sorbent at the end of the water gas shift and carbonation
reaction in the absence of a catalyst at (a) 1 atm (b) 21 atm (S:C ratio of
3:1, 600°C)
176
100
CaO and HTS catalyst
CaO
H2 Purity (%)
80
60
40
20
0
0
500
1000
1500
2000
2500
Time (sec)
Figure 5.11: Comparison in the product H2 purity in the presence of the sorbent and in
the presence of the sorbent and catalyst mixture at 1 atm (650°C, S:C ratio
of 1:1)
177
100
H2 Gas
Composition
H2 Purity
(%)
80
60
40
Without catalyst - 3-1
Without catalyst - 2-1
With catalyst - 3:1
With catalyst - 2:1
20
0
0
500
1000
1500
2000
Time(sec)
Figure 5.12: Comparison in the product H2 purity in the presence of the sorbent and in
the presence of the sorbent and catalyst mixture (650°C, 21 atm)
178
100
Cycle 1
Cycle 3
Cycle 5
Cycle 7
Cycle 9
H2 Purity (%)
80
60
40
20
0
0
500
1000
1500
2000
Time (sec)
Figure 5.13: Product H2 purity obtained over multiple reaction and regeneration cycles
in the presence of CaO sorbent without catalyst at 4.5 atms. (600°C, S:C
ratio of 3:1)
179
1.0
H2 Purity (%)
0.8
0.6
Cycle 1
Cycle 3
Cycle 5
Cycle 7
Cycle 9
0.4
0.2
0.0
0
500
1000
1500
2000
Time (sec)
Figure 5.14: Product H2 purity obtained over multiple reaction - regeneration cycles in
the presence of CaO sorbent without catalyst at 21 atms (600°C, S:C ratio
of 3:1)
180
H2S concentration (ppm)
800
0.75:1
1:1
3:1
600
400
200
0
1000
2000
3000
4000
Time(sec)
(a)
100
0.75:1
1:1
3:1
H2 Purity (%)
80
60
40
20
0
0
1000
2000
3000
Time (sec)
(b)
Figure 5.15: Effect of S:C ratio on the (a) extent of H2S removal and (b) the purity of
H2 produced during the combined water gas shift, carbonation and
sulfidation reaction in the presence of CaO sorbent (1atm, 600oC)
181
800
H2S concentration (ppm)
o
560 C
o
600 C
o
650 C
o
700 C
600
400
200
0
500
1000
1500
2000
2500
3000
Time(sec)
(a)
100
o
560 C
o
600 C
650oC
700oC
H2 Purity (%)
80
60
40
20
0
0
500
1000
1500
2000
2500
3000
Time(sec)
(b)
Figure 5.16: Effect of temperature on the (a)extent of H2S removal and (b) purity of H2
produced during the combined water gas shift, carbonation and sulfidation
reaction in the presence of CaO sorbent (1 atm, S:C ratio of 1:1)
182
800
H2S concentration (ppm)
300 psig
0 psig
1 atm
600
400
200
21 atm
< 1 ppm
0
0
1000
2000
3000
4000
5000
Time (sec)
(a)
100
0 psig
300 psig
2
H2 GasHComposition
Purity (%) (%)
High purity H2
80
60
21 atm
40
1 atm
20
0
0
1000
2000
3000
4000
Time(sec)
(b)
Figure 5.17: Effect of pressure on the (a) extent of H2S removal (b) purity of H2
produced during the combined water gas shift, carbonation and sulfidation
reaction in the presence of CaO sorbent (S:C ratio of 1:1, 600oC)
183
PCC
(a)
PCC- Calcined
(b)
PCC – Carbonated andsulfided at 300 psig
PCC – Carbonated andsulfided at 0 psig
(c)
(d)
Figure 5.18: SEM image of the (a) initial CaCO3 sorbent (b) CaO sorbent obtained
from the calcination of CaCO3 (c) sorbent at the end of the water gas shift,
carbonation and sulfidation reaction at 1 atm (c) CaO sorbent obtained
from the calcination of CaCO3 (600 oC, S:C ratio of 1:1) (c) sorbent at the
end of the water gas shift, carbonation and sulfidation reaction at 21 atm
(600oC, S:C ratio of 1:1)
184
CO2
Sequestration
CO2
Compression
Water
Quench
185
Sulfuric
Acid
Plant
H2S
Removal
Mercury
Removal
Dual
Stage
Acid Gas
Removal
Sour Shift
Scrubber
Raw
Syngas
Shift
Reactors
Air
Radiant
Cooler
ASU
Gasifier
Coal
Feed
Coal Water
Ash/HCl/
Ammonia/
Sulfur
Removal
HTS/LTS
Reactors
Steam
Slag
Mercury
Removal
Sweet Shift
PSA
H2
Boiler
Flue
Gas
CO2
Removal
Air
Steam
Turbine
(a)
Continued
Figure 5.19: (a) Conventional process for H2 production from coal (b) Integration of the CLP in a conventional process for H2
production from coal
185
Figure 5.19 continued
CO2
Sequestration
CO2
Compression
Solids Waste
Calcium
Calciner
Radiant
Cooler
Looping
Carbonator
Mercury
Removal
Process
Air
PSA
ASU
Water Limestone
Gasifier
Coal
Feed
Coal Water
Steam
Turbine
Slag
(b)
186
H2
Steam
I NTEGRATED
WGS +H 2 S
+COS + HC L
C APTURE
Hydrogen
Air
Fuel
Cell
CaO
To Steam
Turbine
CaCO3
Gas Turbine
Steam
H2+O2
Coal/Biomass
Rotary
Calciner
BFW
Air
Compressor
Generator
CO2
HRSG
Air
Oxygen
Gasifier
Stack
Slag
Air
Separation
Fuels &
Chemicals
Steam
Turbine
Figure 5.20: Integration of the CLP in a coal gasification system for the production of
electricity, H2 and liquid fuels
187
Partial Pressure of CO2, atm
100
10
1
0.1
0.01
0.001
0.0001
500
550
600
650
700
750
800
Temperature(C)
Equilibrium PCO2 for Carbonation of CaO
Moving Bed, dry
Moving Bed, slagging
Fluidized Bed
Entrained Flow, slurry
Entrained Flow, dry
(a)
Continued
Figure 5.21: Comparison of the PCO2 in the carbonator with the equilibrium PCO2 for
the carbonation of CaO for a S:C ratio of (a)1:1 (b)3:1
188
Figure 5.21 continued
Partial Pressure of CO2, atm
100
10
1
0.1
0.01
0.001
0.0001
500
550
600
650
700
750
Temperature(C)
Equilibrium PCO2 for Carbonation of CaO
Moving Bed, dry
Moving Bed, slagging
Fluidized Bed
Entrained Flow, slurry
Entrained Flow, dry
(b)
189
800
Partial Pressure of H2O, atm
100
10
1
0.1
0.01
500
550
600
650
700
750
800
Temperature(C)
Equilibrium PH2O for Hydration of CaO
Moving Bed, dry
Moving Bed, slagging
Fluidized Bed
Entrained Flow, slurry
Entrained Flow, dry
(a)
Continued
Figure 5.22: Comparison of the PH2O in the carbonator with the equilibrium PH2O for
the hydration of CaO for a S:C ratio of (a)1:1 (b)3:1
190
Figure 5.22 continued
Partial Pressure of H2O, atm
100
10
1
500
550
600
650
700
750
Temperature(C)
Equilibrium PH2O for Hydration of CaO
Moving Bed, dry
Moving Bed, slagging
Fluidized Bed
Entrained Flow, slurry
Entrained Flow, dry
(b)
191
800
1.0
Moving Bed, dry
Moving Bed, slagging
Fluidized Bed
Entrained Flow, slurry
Entrained Flow, dry
CO Conversion
0.8
0.6
0.4
0.2
0.0
550
600
650
700
750
700
750
Temperature (C)
(a)
1.0
CO Conversion
0.8
0.6
0.4
0.2
0.0
550
Moving Bed, dry
Moving Bed slagging
Fluidized Bed
Entrained Flow, slurry
Entrained Flow, dry
600
650
Temperature (C)
(b)
Figure 5.23: Effect of temperature on equilibrium CO conversion in the water gas shift
reactor at a S:C ratio of (a) 1:1 (b) 3:1
192
1.00
CO Conversion
0.95
0.90
0.85
0.80
0.75
550
Moving Bed, dry
Moving Bed, slagging
Fluidized Bed
Entrained Flow, slurry
Entrained Flow, dry
600
650
700
750
Temperature (C)
(a)
1.00
CO Conversion
0.99
0.98
0.97
0.96
0.95
550
Moving Bed, dry
Moving Bed slagging
Fluidized Bed
Entrained Flow, slurry
Entrained Flow, dry
600
650
700
750
Temperature (C)
(b)
Figure 5.24: Effect of temperature on equilibrium CO conversion in the presence of
CaO in the carbonation reactor of the CLP at a S:C ratio of (a) 1:1 (b) 3:1
193
100
90
H2 purity (%)
80
Moving Bed, dry
Moving Bed, slagging
Fluidized Bed
Entrained Flow, slurry
Entrained Flow, dry
70
60
50
40
500
550
600
650
700
750
800
Temperature (C)
(a)
100
H2 Purity (%)
90
80
Moving Bed, dry
Moving Bed, slagging
Fluidized Bed
Entrained Flow, slurry
Entrained Flow, dry
70
60
50
40
500
550
600
650
700
750
800
Temperature(C)
(b)
Figure 5.25: Effect of temperature on equilibrium H2 purity in the presence of CaO at a
S:C ratio of (a) 1:1 (b) 3:1
194
% Carbon Captured (mole %)
100
95
90
85
80
550
1:1, Moving Bed, dry
1:1, Moving Bed, slagging
1:1, Fluidized Bed
1:1, Entrained Flow, slurry
1:1, Entrained Flow, dry
3:1, Moving Bed, dry
3:1, Moving Bed slagging
3:1, Fluidized Bed
3:1, Entrained Flow, slurry
3:1, Entrained Flow, dry
600
650
700
750
Temperature ( C)
Figure 5.26: Effect of temperature and S:C ratio on the % of carbon captured in the
CLP using syngas from different gasifiers as the feed
195
Steam
Reformer
Sulfur
Removal
Low
Temperature
Shift Reactor
High
Temperature
Shift Reactor
CO2
Absorber
196
Reformer
Flue Gas
CO2
Natural Gas
Figure 5.27: Conventional steam reforming of natural gas for H2 production with a methanator
196
Methanator
H2
Steam
Reformer
Sulfur
Removal
High
Temperature
Shift Reactor
Reformer
Flue Gas
Natural Gas
PSA
H2
Purge Gas
used as
reformer fuel
Figure 5.28: Conventional steam reforming of natural gas for H2 production with a
PSA
197
Steam
Reformer
Sulfur
Removal
Calcium
Looping
Process
Polishing
PSA
H2
Reformer
Flue Gas
Natural Gas
Figure 5.29: CLP integrated in the conventional steam reforming of natural gas process
198
Oxygen
Reformer
Steam
Sulfur
Removal
Water
Quench
Scrubber
High
Temperature
Shift Reactor
Low
Temperature
Shift Reactor
CO2
Absorber
199
CO2
Natural Gas
Figure 5.30: Conventional partial oxidation process for conversion of natural gas to H2
199
H2
Oxygen
Reformer
Steam
Sulfur
Removal
Water
Quench
Scrubber
Calcium
Looping
Process
CO2
Natural Gas
Figure 5.31: CLP integrated in the partial oxidation of natural gas for H2 production
200
H2
350
CaS_700C_15%H2O
CaS_650C_31%H2O
CaS_700C_31%H2O
300
H2S (ppm)
250
200
150
100
50
0
0
2000
4000
6000
8000
10000
Time(sec)
Figure 5.32: Effect of the change in temperature and steam composition on the
regeneration of CaS with H2O
201
50
CaS_700C_15%H2O_15%CO2
45
CaS_700C_31%H2O_31%CO2
40
H2S (ppm)
35
30
25
20
15
10
5
0
0
500
1000
1500
2000
2500
3000
3500
4000
Time (sec)
Figure 5.33: Effect of the change in steam and CO2 composition on the regeneration of
CaS in the presence of H2O and CO2
202
350
CaS_700C_31%H2O_0psig
300
CaS_700C_31%H2O_300psig
Spent sorbent from carbonation and sulfidation at 21 atms
H2S (ppm)
250
200
150
100
Spent sorbent from carbonation and sulfidation at 1 atms
50
0
0
1000
2000
3000
4000
5000
6000
7000
8000
9000
Time (sec)
Figure 5.34: H2S evolved in the presence of H2O and CO2 from spent sorbent produced
during combined CO2 and H2S removal at 1 and 21 atms
203
CHAPTER 6
PROCESS SIMULATION AND ECONOMICS OF THE CALCIUM LOOPING
PROCESS (CLP) FOR PRODUCTION OF H2 FROM COAL
6.1 INTRODUCTION
The CLP for high purity H2 production from coal, described in Chapters 4 and
5, has been analyzed in this Chapter based on the Aspen Plus® process simulation
through two schemes. The first scheme is based on the cogeneration of H2 and
electricity in the same facility. In the second scheme, the only product is H2 and all the
energy produced in the process is used internally for the parasitic energy requirement.
Both simulations are conducted for the production of 280 t/day of H2 from Illinois #6
coal.
6.2 PRODUCTION OF FUEL CELL GRADE H2 WITH A PSA
6.2.1 Cogeneration of H2 and Electricity
The CLP produces high purity H2 and electricity with high efficiency by
integrating various unit operations like the water gas shift reaction, and CO2, sulfur and
hydrogen halide removal from synthesis gas into a single stage reactor at high
204
temperatures. The CLP-assisted coal to H2 process comprises of six major unit
operations:
•
Shell gasifier
•
ASU
•
Calcium looping reactor
•
Calciner
•
Pressure Swing Absorber (PSA)
•
Heat recovery and steam generation block
6.2.1.1 Process Configuration
The process flow diagram for the CLP for co-production of fuel cell grade H2
and electricity is given in Figure 6.1. A Shell gasifier is used to gasify 2,405 tonnes/day
of Illinois # 6 coal in the presence of oxygen supplied by the ASU. The properties of
the coal are given in Table 6.1. The Shell gasifier produces 951,624 m3/day of syngas
at a temperature of 1538 ºC and a pressure of 36 atm. The composition of syngas
produced at the outlet of the gasifier is given in Table 6.2. 79% of the syngas produced
at the outlet of the gasifier is fed to the H2 production reactor or carbonator for the
production of high purity H2 while 21% of the syngas is combusted in the calciner to
provide the energy required for the endothermic calcination reaction.
205
79% of the hot syngas from the gasifier is cooled in a radiant heater and is fed
to the H2 production reactor or carbonator along with CaO sorbent and steam. In the
carbonation reactor, H2 production and purification are achieved by the integrated
water gas shift reaction, carbonation and sulfidation of the CaO sorbent at a
temperature of 600°C and pressure of 21 atm. A Ca:C mole ratio of 1.32 and S:C ratio
of 3 are used to achieve high CO conversions and almost 100% carbon and sulfur
capture. The heat produced in the carbonation reactor through the exothermic water gas
shift and carbonation reactions is used to produce high temperature and high pressure
steam which is used to generate electricity. The H2 rich product stream is further
purified in a PSA to produce 99.999% H2 which can be used either in H2 fuel cells or
for the production of fuels and chemicals. Since the purity of the dry H2 feed stream to
the PSA is very high (~94%) the energy consumption in the PSA is considerably
reduced. The spent sorbent, which is separated from the H2 product in a particulate
capture device (PCD), is regenerated in the calciner at 850 °C to produce a
sequestration ready CO2 stream. All the energy required for the calcination of the
sorbent is supplied by the combustion of the syngas and PSA tail gas with oxygen in
the calciner. At this stage, 6.2 wt% of the spent sorbent is purged and an equivalent
moles of calcium in the form of CaCO3 is added as make up to maintain the high
reactivity of the sorbent mixture towards CO2 and sulfur capture. In this process, the
pure H2 stream is produced at 21 atm which is compressed to 60 atm for transportation
and the CO2 is compressed to 150 atm for sequestration. The intermediate pressures
used for the multistage CO2 compression are shown in Table 6.3.
206
In this scenario, 280 tonnes/day of H2 is produced with an efficiency of 56.5%
(HHV) and 81MW electricity is produced with an efficiency of 10%.
6.2.1.2 Components and Physical Properties
Table 6.4 lists the components used in the simulation. Syngas obtained from a
Shell gasifier using Illinois #6 coal is selected as the specific feedstock. RKS-BM is
selected as the property method for conventional components.
6.2.1.3 Operating Conditions and Premises
The basis for the conceptual process evaluation of the CLP is given below:
•
Analysis has been conducted for a production rate of 280 tonnes/day of H2 and
81 MWe of electricity
•
Illinois #6 coal has been used at a feed rate of 2,405 tonnes/day.
•
A Shell gasifier is used to generate the syngas at a temperature of 1538C and
pressure of 36 atm
•
A Ca:C ratio of 1.32 is used to achieve almost 100% CO2 and sulfur capture.
•
A S:C ratio of 3:1 is used to achieve high CO conversion in the CLP.
•
A solids purge of 6.2 wt% is used for the calcium sorbent and an equivalent
moles of calcium in the form of CaCO3 is added as make up.
•
100% calcination occurs in the calciner and the heat is supplied by the complete
combustion of syngas and PSA tail gas with oxygen.
207
•
Mercury removal is conducted using an activated carbon bed.
•
The temperature is maintained constant during the individual unit operations.
•
H2 purity of greater than 99.999% is obtained by using a PSA.
•
H2 is produced from the calcium looping reactor at 21 atm and is compressed to
60 atm for transportation while the sequestration ready CO2 is compressed to
150 atm.
•
The mechanical efficiency of pressure changers such as compressors, turbines,
and expanders is 1 while their isentropic efficiency is 0.72.
•
A part of the solid waste which includes fly ash, bottom ash, gasifier slag, etc
is reused in the plant while the rest of it is disposed in the coal mine.
•
Waste water is treated before discharge to meet effluent guidelines
Table 6.5 shows the Aspen Plus® models used for the various important unit
operations in the CLP.
6.2.1.4 Results and Discussions
Based on the description of the process and the assumptions given above, the
ASPEN Plus® simulation for the production of H2 and electricity using the CLP is
shown in Figure 6.2.
Table 6.6 shows the material and energy balance for the entire process. Table
6.7 illustrates the power balance in the CLP process. A total of 141.8 MWe of power is
produced in the CLP process out of which 61MWe is consumed in the process for coal
208
preparation, pumps and compressors for the ASU, CO2 and H2. .Net electricity of 81
MWe is produced which can be exported.
Table 6.8 illustrates the results of the ASPEN simulation conducted for the
conversion of coal to H2 and electricity. The CLP process produces 280 tonnes/day of
H2 and 81 MWe of net electricity with 100% CO2 and sulfur capture. The overall
process efficiency from coal to H2 and electricity is 66.5%, which is much higher than
the conventional coal to H2 process. The CLP has a higher efficiency than the
conventional H2 production process from coal using solvents due to the integration of
various unit operations in a single stage reactor and the high temperature gas clean up
achieved in the system.
6.2.2 Production of Only H2 With Internal Heat Integration
For only H2 production, a Shell gasifier is used to gasify 2,180 t/day of Illinois
#6 coal in the presence of oxygen supplied by the ASU. The properties of the coal are
given in Table 6.1.
The Shell gasifier produces 847,200 m3/day of syngas at the temperature of
1538 °C and the pressure of 36 atm. Due to the high content of sulfur in the coal, the
syngas contains 1.15% of H2S and 848 ppm COS. Since the CLP is capable of in situ
sulfur capture during the production of H2, it can handle high sulfur coal effectively.
The composition of syngas produced at the outlet of the gasifier used in the simulation
is given in Table 6.2. 88.7% of the syngas produced at the outlet of the gasifier is fed to
209
the calcium looping reactor for the production of high purity H2 while 11.3% of the
syngas is combusted in the calciner to provide the energy required for the endothermic
calcination reaction.
The hot syngas is cooled in a radiant heater and is fed to the calcium looping
reactor along with high temperature and high pressure steam and PCC-CaO sorbent. In
the carbonation reactor, H2 production, purification and sulfur removal are achieved by
the integrated water gas shift reaction, carbonation, and sulfidation of the CaO sorbent
at a temperature of 600 °C and pressure of 21 atm. The H2 rich product stream is then
further purified in a PSA to produce up to 99.999% pure H2 which can be used either in
H2 fuel cells or for the production of fuels and chemicals. Since the purity of the H2
feed stream to the PSA is very high (94–98%) the energy consumption in the PSA is
considerably reduced. The spent sorbent, which is separated from the H2 product in a
cyclone, is regenerated in the calciner at 850 °C to produce a sequestration ready CO2
stream. At this stage, 8% of the spent sorbent is purged and a makeup of PCC sorbent
is added to maintain the high reactivity of the sorbent mixture toward CO2 and sulfur
capture. In this process, the pure H2 stream is produced at a high pressure of 20 atm
and the CO2 is compressed to a pressure of 150 atm for transportation to the
sequestration site. A Ca:C ratio of 1.3 is used to achieve almost 100% carbon and
sulfur capture and sequestration from coal. This process leads to the production of 280
t/day of H2 from coal with an efficiency of 62.3% (HHV) as shown in Table 6.9.
210
6.3 PRODUCTION OF H2 HAVING A PURITY OF 94–98% WITHOUT A PSA
6.3.1 Cogeneration of H2 and Electricity
The Aspen Plus® flow sheet for the CLP for the production of 94–98% pure H2
without a PSA is given in Figure 6.3. The syngas obtained from the Shell gasifier is
split into two streams, one fed to the calcium looping reactor and the other to the
calciner. In this scenario, for the cogeneration of H2 and electricity in the absence of a
PSA, 2,350 t/day of coal is used for the production of 280 t/day of H2 at an efficiency
of 57.8% (HHV). In addition to the H2, 67.56 MW of electricity is produced with an
efficiency of 8.5% from coal.
6.3.2 Production of H2 With Internal Heat Integration
In the case of H2 production without a PSA, 19% of the syngas is required to
supply the energy for the endothermic calcination reaction and the remaining 81% is
fed to the calcium looping reactor for the production of high purity H2. This process in
the absence of a PSA also leads to the production of 280 t/day of H2 with an efficiency
of 63% from coal (HHV).
The two scenarios for the production of H2 or cogeneration of H2 and electricity
from coal by the CLP in the absence of a PSA are summarized in Table 6.10.
Comparing the processes with and without the PSA unit, it can be seen that the H2
generation efficiencies for these processes with internal heat integration are almost the
211
same. For the cogeneration of H2 and electricity, the overall efficiencies of the
processes with and without the PSA unit are also similar. Although, with the PSA unit,
the H2 generation efficiency (56.5%) is lower than without the PSA unit (57.8%), more
electricity is produced with the PSA unit (81 MW) than without the PSA unit (67.5
MW).
6.4 COMPARISON OF THE PROCESS EFFICIENCIES FOR DIFFERENT GASIFIERS
The CLP is optimized for high purity H2 production using syngas obtained from
three different gasifiers, the Shell, Lurgi, and GE gasifiers (Zheng and Furinsky, 2005).
A comparison of the efficiencies obtained for the different gasifiers is given in Table
6.11. The type of gasifier used has a large effect on the process efficiency due to the
composition of the syngas obtained and the inherent efficiency of the gasifier for the
conversion of coal to syngas. It is seen in Table 6.11 that the CLP in combination with
the Shell gasifier has the highest efficiency due to the high efficiency of the dry feed
Shell gasifier. The co-generation of H2 and electricity yields a higher efficiency than
the case optimized for the production of H2 alone.
6.5 EFFECT OF PROCESS PARAMETERS ON CLP PERFORMANCE USING SYNGAS
FROM A GE GASIFIER
As shown in the previous section the CLP integrated with the Shell gasifier has the
highest efficiency closely followed by the GE gasifier. However from an economic
standpoint, the cost of a GE gasifier is lower than a Shell gasifier and hence the
212
sensitivity analysis conducted in this section is based on the simulation for the CLP
with the GE gasifier,
6.5.1 Approach
The thermodynamics of the combined reactions occurring in the H2 production
reactor or carbonation reactor has been investigated using ASPEN Plus® software. The
effect of temperature, pressure and S:C ratio has been investigated on the combined
water gas shift, carbonation and sulfidation reaction. The analysis has been conducted
using syngas obtained from a GE or Texaco gasifier. In the ASPEN model shown in
Figure 6.4, the syngas from the GE gasifier is fed to the H2 production reactor along
with steam and CaO (from the calciner). At the outlet of the H2 production reactor, the
H2 product is separated from the solids and the mixture of solids containing CaO,
CaCO3, Ca(OH)2 and CaS is regenerated in the calciner. 8% of the solids obtained at
the exit of the H2 production reactor is purged and a makeup stream containing an
equivalent quantity of CaCO3 is added to maintain the fraction of CaS in the circulating
solids stream at equilibrium.
6.5.2 Sensitivity Analysis for the Yield and Purity of H2 Produced
Sensitivity analyses have been conducted for the H2 production reactor, to
investigate the effect of temperature, pressure and amount of steam addition on the
purity and yield of H2 produced.
213
6.5.2.1 Effect of Temperature
As it can be seen in Figure 6.5, the purity of H2 is very high in the temperature
range of 550 to 650 ºC. In this temperature interval, the thermodynamic limitation of
the water gas shift reaction is removed due to the incessant removal of the CO2 product
and a very high yield of H2 is produced. At temperatures of 650 ºC and above, the
purity of H2 decreases as the equilibrium conversion of the carbonation reaction
decreases. At temperatures of 880 ºC and above, the carbonation reaction does not
occur because at these temperatures equilibrium favors the reverse or calcination
reaction. Hence at temperatures above 880 ºC, H2 is produced only due to the water gas
shift reaction.
6.5.2.2 Effect of Pressure
The effect of pressure in the range of 1 to 40 atms was investigated and it was
found that the change in pressure results in a comparatively smaller change in H2 purity
when compared to the change in temperature. Figure 6.6 illustrates the effect of
pressure on H2 purity. Pressure influences the combined reaction for H2 production in
two ways. Although high pressure has a positive effect on the thermodynamics of the
carbonation reaction which increases the equilibrium conversion of the water gas shift
reaction, it also favors the methanation reaction which results in a decrease in the
overall H2 yield. The formation of one mole of CH4 results in the loss of 3 moles of H2
and hence the yield of H2 decreases with the increase in pressure.
214
From Figure 6.6 it can be seen that the purity of H2 increases with the increase
in pressure up to 10 atms. This is due to the decrease in CO and CO2 in the final
product due to the improved thermodynamics of the combined carbonation and water
gas shift reaction at high pressures. Above a pressure of 10 atms, the thermodynamics
of the methanation reaction is very favorable and hence there is a loss in H2 and an
increase in the CH4 in the final product. It can be seen that according to thermodynamic
evaluation, pressure has a small effect on the purity of H2 produced and with the
increase in pressure from 1 to 40 atms the H2 purity changes only from 96.1 to 97.3%.
6.5.2.3 Effect of S:C Ratio
The effect of S:C ratio was first investigated at 600 ºC. It can be seen from
Figure 6.7 that there is an increase in the purity of H2 produced with the increase in S:C
ratio. This is because excess steam favors the equilibrium of the water gas shift
reaction in the forward direction. It was also found that with an increase in steam
addition, the CH4 composition in the product stream decreases. From Figure 6.7 it can
be seen that there is a substantial increase in the H2 purity when the S:C ratio is
increased from the 1 to 2. Beyond this the increase in H2 yield is very small.
6.5.3 Sensitivity Analysis for the Extent of Contaminant Removal from the Product
H2
As shown in Figure 6.8, it can be seen that the H2S in the outlet stream
increases with the increase in the S:C ratio due to the inhibiting effect of steam on the
215
sulfidation reaction of CaO. It was also found that with the increase in temperature, the
H2S in the outlet stream increases and maximum removal is achieved in the
temperature range of around 600 ºC.
The effect of steam concentration and temperature on the removal of COS by
CaO was also studied and it was found that the COS in the outlet increases with the
increase in temperature. This is due to the increase in the CO2 concentration which
inhibits the removal of COS by CaO. As shown in Figure 6.9, almost all the COS in the
syngas stream is removed by the CaO sorbent at temperatures lower than 800 ºC. The
concentration of COS in the outlet stream was also found to increase with the increase
in steam addition at temperatures above 800 ºC. This is also due to the increase in the
CO2 flow rate with the increase in steam addition.
The concentration of CO was found to increase with the increase in temperature
at high temperatures of above 700 ºC due to the equilibrium limitation of the water gas
shift reaction and due to the decrease in the CO2 removal by the CaO sorbent at high
temperatures as illustrated in Figure 6.10. At temperatures below 700 ºC it was found
that the S:C ratio does not have an effect on the CO concentration at the outlet of the
reactor and very low concentration of CO is obtained even at low S:C ratios. This is
due to the removal of CO2 by the CaO which enhances the equilibrium of the water gas
shift reaction. With the increase in temperature of above 700 ºC it was found that the
CO concentration increases with the decrease in steam addition. This is due to the low
CO2 removals at temperatures of above 700 ºC.
216
Figure 6.11 illustrates the change in the outlet flow rate of CO2 with the
increase in temperature and S:C ratio. Almost all the CO2 in the outlet H2 product
stream is removed by the CaO sorbent at a temperature of 600 ºC and the CO2
concentration increases with the increase temperature.
Figure 6.12 depicts the change in the outlet flow rate of CH4 with the change in
temperature and S:C ratio. The flow rate of CH4 in the outlet H2 product is found to
decrease with the increase in steam addition. It was also found that equilibrium favors
the formation of CH4 at low temperatures of 600 ºC.
6.5.4 Sensitivity Analysis for the Cold Gas Efficiency and Overall Process Efficiency
6.5.4.1 Effect of Pressure
Figure 6.13 depicts the effect of pressure and S:C ratio on H2 purity, cold gas
efficiency and the process efficiency. The cold gas efficiency is the efficiency with
which the energy in coal is converted to H2 and is defined as the ratio of the HHV of
the product H2 stream to the HHV of coal. The process efficiency is the total efficiency
of H2 and electricity production which is obtained from a detailed heat integration
within the process and includes the parasitic energy required for the sorbent
regeneration, the gasifier, ASU, PSA etc and the energy obtained from the exothermic
H2 production reaction, cooling of hot streams, etc. At a S:C ratio of 2, the purity of H2
increases with the increase in pressure to 5 atms. With a further increase in pressure to
20 atms, the H2 purity falls by <1%. In contrast, the H2 purity at a S:C ratio of 1
217
decreases by 3.5% with the increase in pressure from 1 to 20 atms. This decrease in H2
purity with the increase in pressure is due to the increase in the formation of CH4 in the
H2 production reaction from the CO and H2 in the syngas. The cold gas efficiency as
well as the process efficiency increase with the increase in pressure from 1 to 10 atms
and then decrease by a small amount with the further increase in pressure to 20 atms.
6.5.4.2 Effect of S:C Ratio
With the decrease in S:C ratio, although a small decrease in the H2 purity is
observed in Figure 6.14 due to the increased production of CH4, there is almost no
change in the cold gas efficiency or the process efficiency. This is due to the heat
integration within the process which utilizes all the CH4 in the tail gas of the PSA to
provide a part of the parasitic energy requirement of the process.
6.5.4.3 Effect of Temperature
The effect of the increase in temperature of the H2 production reactor on the
purity of H2 produced, and efficiency of the process was also investigated and as
illustrated in Figure 6.15 it was found that the purity of H2 decreased with the increase
in the temperature of the combined water gas shift, carbonation and sulfidation
reaction. With the increase in temperature the removal efficiency of CO2 by the CaO
sorbent is reduced and hence the purity of H2 is also decreased. However, the
efficiency of the process remained constant with the increase in temperature of the H2
production reactor.
218
6.5.4.4 Effect of Ca:C Ratio
The performance of the CLP depends to a large extent on the reactivity and
recyclability of the calcium sorbent. For highly reactive sorbents the Ca:C ratio
required is lower than that required for naturally occurring limestone and hence the
amount of solid circulation will also be low. As illustrated in Figure 6.16, the increase
in Ca:C ratio results in a decrease in the cold gas efficiency as well as process
efficiency. Hence by using a highly reactive sorbent, the amount of solids circulation
can be reduced and the efficiency can be improved.
6.6 EFFECT OF ADDITION OF SORBENT HYDRATION TO THE CLP PROCESS
Figure 6.17 illustrates the ASPEN Plus model for the CLP with sorbent
hydration as a part of the carbonation - calcination cycle. The sorbent at the exit of the
calciner is hydrated at a high temperature of 600 ºC and a pressure of 21 atms. The
hydrated sorbent is then fed to the carbonation or H2 production reactor. In the
carbonation reactor the Ca(OH)2 sorbent is converted to CaCO3 and the steam produced
from the Ca(OH)2 is consumed in the water gas shift reaction. The hydration of CaO is
exothermic and hence heat is extracted in the hydration reactor. A part of the
exothermic energy released in the carbonation reactor due to the exothermic of
carbonation and the water gas shift reaction is consumed by the endothermic
decomposition of the Ca(OH)2. The addition of sorbent hydration aids in reducing the
219
Ca:C ratio and the solids circulation in the CLP process. The reduction in Ca:C ratio
aids in improving the overall efficiency of the process as shown in Figure 6.16.
6.7 TECHNO-ECONOMIC ANALYSIS OF H2 PRODUCTION FROM COAL
A comparison of the economics of the conventional coal to H2 process with the CLP
process has been discussed in this section. Figure 6.18 is a flow diagram of the
conventional coal to H2 process that was used as the base case in this study. The
design, assumptions and economic analysis for the base case is based on a draft US
Department of Energy (DOE) study (DOE, 2009). In the conventional plant which is
the base case in this study, H2 is produced from a GE gasifier followed by a water
quench, syngas scrubber, water gas shift reactors, syngas coolers, mercury removal
system, dual stage selexol system for the removal of CO2 and H2S and a PSA. The PSA
produces 99.9% pure hydrogen and the tail gas is combusted in a boiler to generate
steam for electricity production. The H2S is sent to a Claus plant for the production of
elemental sulfur and the CO2 is dried and compressed for transportation and
sequestration. The assumptions used for the conventional process are listed below.
Key parameters for the conventional process:
1) 249 tons/hr of Illinois #6 coal is fed to the gasifier along with 243 tons/hr of O2
from an ASU for hydrogen production.
2) 26 tons/hr of solid waste is generated from the process.
3) 26 tons/hr of 99.9% pure H2 at 21 atms and 31 MWe is produced.
220
4) The process results in a net CO2 emission of 60 tons/hr and 517 tons/hr of CO2
is sequestered.
Figure 6 illustrates a flow diagram of the CLP integrated with a GE gasifier. The
syngas from the gasifier is cooled in a radiant cooler and its pressure is reduced in an
expander. Since the syngas sent to the expander should be free from all particulate
matter, a metallic filter is used to remove the flyash from the syngas. The syngas is
then sent to the carbonation reactor along with Ca(OH)2 sorbent from the hydrator. The
Ca(OH)2 dehydrates in the carbonation reactor producing steam which is consumed in
the water gas shift reaction. The H2 rich stream with the spent sorbent is then sent to a
cyclone and metallic filter assemble to separate the sorbent from the H2 stream. The H2
rich stream is then further purified to 99.9% in a PSA. The spent sorbent is sent to the
calciner. The energy for the calciner is supplied by the direct oxy combustion of coal
and the PSA tail gas. The calcined sorbent is separated from the CO2 stream in a
cyclone and 95% of the sorbent is recovered. The remaining 5% of the sorbent is
cooled with the CO2 stream and is finally separated from the CO2 stream in a fabric
filter at low temperature. The CO2 is then dried and compressed for transportation and
sequestration. The calcined sorbent is hydrated with steam at 500C and sent to the
carbonation reactor. The assumptions for the CLP are listed below.
Key parameters for the CLP process:
1) The CLP process is a co-generation facility which results in the production of
23 tons/hr of 99.9% pure H2 at 21 atms and 324 MWe of electricity.
221
2) A total of 367 tons/hr of Illinois #6 coal is fed to the over all process. 249
tons/hr of coal is fed to the gasifier while the rest is fed to the calciner to
provide the energy for calcination.
3) A total of 592 tons/hr of O2 is fed to the process out of which 243 tons/hr is fed
to the gasifier and the rest is fed to the calciner to combust the coal and PSA
tailgas.
4) A Ca:C ratio of 1.3 is used in the carbonation reactor and 5 wt% of the solids at
the exit of the calciner is purged and an equivalent moles of limestone is added
to the calciner as makeup.
5) 139 tons/hr of solid waste is generated from the process.
6) The process results in almost zero CO2 emissions and 865 tons/hr of CO2 is
sequestered.
Based on the process simulation and the economic assumptions used for the base case,
the economic analysis was conducted for the CLP. All costs are “overnight” costs in
2008 dollars and the cost estimates were prepared for an Nth-of-a-kind plant. The
plants are assumed to be located in Midwest US. The economic comparison of the
conventional plant and the CLP is expressed on the basis of the levelized cost of H2 in
dollars per Kg of H2. The capital charge factor used to levelize the capital costs and the
levelization factors for coal, electricity and O& M costs are provided in Tables 6.12
and 6.13. The annual levelized costs for the conventional and the CLP plants are
provided in Tables 6.12 and 6.13 respectively. As shown in the tables, the CLP plant
has higher capital costs, fixed O&M and variable O&M costs but it also produces more
222
than 10 times the amount of electricity produced by the conventional plant. In addition,
the CLP plant has almost no CO2 emission while the conventional plant emits 10% of
the CO2. The increased costs of capital and O&M are offset by the large amount of
electricity produced and the credit obtained for the amount of CO2 avoided. Hence the
CLP plant has a levelized cost of H2 of $1.81/Kg of H2 while the conventional plant
has a levelized cost of $2.03/Kg of H2. A more detailed description of the two
processes, assumptions for the technical and economic analysis, and the results is
provided elsewhere. (Connell et al, 2010)
6.7 CONCLUSIONS
The CLP integrated in a gasification system was simulated for the production of
H2 from coal using ASPEN Plus software. Two cases were explored for the production
of only H2 from the entire process and for the cogeneration of H2 and electricity in the
process. The effect of the addition of a PSA at the end of the process for the production
of high purity fuel cell grade H2 was also evaluated. Different types of gasification
systems in conjunction with the CLP were evaluated and it was found that the
efficiency of the coal to H2 process depends on the composition of syngas obtained
from the gasifier as well as the efficiency of the gasifier. The syngas obtained from the
Shell gasifier was found to form the least amount of CH4 in the calcium looping H2
reactor when compared to the syngas obtained from the other gasifiers. In addition to
the reduction in CH4 production, the Shell gasifier also has a higher efficiency for the
conversion of coal to syngas and hence the highest efficiency for the conversion of coal
223
to H2 was obtained for the integration of the CLP with the Shell gasifier. The effect of
S:C ratio, temperature and pressure were also investigated. The purity of H2 produced
from the H2 production reactor was found to decrease by a small amount with the
decrease in S:C ratio and the increase in temperature. The decrease in H2 purity with
the decrease in S:C ratio especially at high pressures was due to the increase in the
formation of CH4 in the H2 production reactor. However, the decrease in S:C ratio and
temperature did not result in a significant change in the process efficiency and the cold
gas efficiency. The increase in pressure from 1 to 10 atms resulted in an increase in
process efficiency. A further increase in pressure did not result in a change in the
process efficiency. A decrease in the Ca:C ratio resulted in an increase in the efficiency
of the process and hence a sorbent with a higher reactivity and recyclability is
beneficial for the process. A techno-economic comparison of the CLP with the
conventional coal to H2 process shows that the CLP has the potential to reduce the cost
of H2 production from coal.
224
Proximate
Analysis
Wt% (AsReceived)
Moisture
11.12
Fixed Carbon
44.19
Volatiles
Ash
Total
HHV (MJ/kg)
Wt%, dry Ultimate
Wt% (AsReceived)
Wt%, dry
Moisture
11.12
49.72
Ash
9.7
10.91
34.99
9.7
100
39.37
10.91
100
Carbon
Hydrogen
Nitrogen
Chlorine
63.75
4.5
1.25
0.29
71.72
5.06
1.41
0.33
27.13
29.2
Sulfur
2.51
2.82
Oxygen
6.88
7.75
Table 6.1: Properties of Illinois # 6 coal
225
Syngas Composition
Mole %
H2O
2.5
N2
4.1
O2
0
H2
27.6
CO
61.4
CO2
2.2
Ar
0.8
COS (ppm)
884
H2S
1.2
CH4
0.1
Temperature (°C)
1538
Pressure (atm)
36
Mass Flow Rate (Kg/hr)
198934
Table 6.2: Composition of the syngas exiting from the Shell gasifier
226
Stage
Discharge
Pressure (Mpa)
1
2
3
4
5
0.4
0.9
2.3
5.9
15.3
Table 6.3: Intermediated pressures for compression of the CO2 for sequestration
227
Component
ID
CH4
CO2
CO
H2
C2H6
C2H4
H20
CACO3
CAO
C
H2S
CAS
COS
N2
O2
AR
NH3
COAL
S
CL2
HCL
ASH
O2S
Type
CONV
CONV
CONV
CONV
CONV
CONV
CONV
SOLID
SOLID
SOLID
CONV
CONV
CONV
CONV
CONV
CONV
CONV
NC
CONV
CONV
CONV
NC
CONV
Component name
METHANE
CARBON-DIOXIDE
CARBON-MONOXIDE
HYDROGEN
ETHANE
ETHYLENE
WATER
CALCIUM-CARBONATE-CALCITE
CALCIUM-OXIDE
CARBON-GRAPHITE
HYDROGEN-SULFIDE
CALCIUM-SULFIDE
CARBONYL-SULFIDE
NITROGEN
OXYGEN
ARGON
AMMONIA
Formula
CH4
CO2
CO
H2
C2H6
C2H4
H2O
CACO3
CAO
C
H2S
CAS
COS
N2
O2
AR
H3N
SULFUR
CHLORINE
HYDROGEN-CHLORIDE
S
CL2
HCL
SULFUR-DIOXIDE
O2S
Table 6.4: Components list for the ASPEN Plus® simulation
228
Unit Operation
Aspen Plus®
Model
CLP Hydrogen Reactor
RGibbs
Purge for solids
FSplit
Solids Make-up
Mixer
Calciner
RGibbs
Gas-Solid Separation
SSplit
HRSG
MHeatX
PSA
Sep
CO2 Compression
Comments / Specifications
1.32:1 Calcium:Carbon molar ratio based on
active calcium sorbent and total carbon
content in syngas, 3:1 S:C ratio,
thermodynamic modeling of the water gas
shift, carbonation and sulfidation reaction of
CaO, isothermal operation with heat
extraction
Splits and purges 6.2% of the solids based on
molar fraction
Combines recycle stream and fresh feed
stream in terms of material and heat
Thermodynamic modeling of limestone
calcination with syngas and PSA tail gas
combustion operates isothermally at 850 °C
with 100% conversion of CaCO3 to CaO and
complete combustion of syngas and PSA
tailgas in oxygen
Operates with 100% separation efficiency
Modeling of heat exchange among multiple
streams
90% yield of H2 obtained from the PSA,
remaining 10% H2 and other gas components
removed in the PSA tailgas stream
105 kWh electricity/tonne CO2 to compress
to 150 atm
Table 6.5: ASPEN Plus® models used for the simulation of the CLP
229
230
Temperature C
Pressure bar
Mass Flow kg/hr
Volume Flow cum/hr
Mole Flow kmol/hr
CH4
CO2
CO
H2
H2O
H2S
COS
N2
O2
AR
O2S
CaCO3
CAO
C
CAS
S
1
2
25
838.6
21.28
1
255817 691561.3
338.556 258.863
0
0
0
0
0
0
0
0
14200
0
0
0
0
0
0
0
0
0
0
0
0
0
0 4891.995
0 1387.27
0
0
0 1720.735
0
0
3
4
57.6
600
21.28
21.28
255817 42429.55
346.881
15.823
0
0
0
0
14200
0
0
0
0
0
0
0
0
0
0
0
5
6
7
600
70
70
21.28
21.28
21.28
641918 13270.21 162137.5
239.382
18.179 222.112
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
290.567 4395.995
91.696 1387.27
0
0
113.737 1720.735
0
0
0
0.003
0.001
0.007
0
0
0.006
0.073
736.598 8999.861
0.004
0.048
0
0
0
0.002
0
0
0
0.005
0
0
0
0
0
0
0
0
0
0
0
0
Continued
Table 6.6: Material and energy balance for the CLP
230
Table 6.6 Continued
231
Temperature C
Pressure bar
Mass Flow kg/hr
Volume Flow cum/hr
Mole Flow kmol/hr
CH4
CO2
CO
H2
H2O
H2S
COS
N2
O2
AR
O2S
CaCO3
CAO
C
CAS
S
8
10
12
13
14
18
19
840
70
900
550
250
70
70
36
21.28
10
1
1
21.28
1
157038.1 26687.45 31014.94 313741.2 313741.2 202095.1 313741.2
19308.84 9428.859 13068.33 558040.1 354370.3 9669.15 231700.2
7.452
54.997
163.934
3.363
4567.793
0.834
2056.625 6433.311
186.288
91.176
85.693
2.431
6.587
0
305.513 305.511
0
0
59.612
59.606
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
55
0
59.193 6200.809 6200.809
3.371 6200.809
0
16.051
16.051
0.834
16.051
0
3.482
3.482 6433.39
3.482
846.933 1466.272 1466.272 9827.635 1466.272
0
2.438
2.438
2.484
2.438
0
0.522
0.522
0
0.522
305.511 387.024 387.024 305.513 387.023
64.277
0
0
0
0
59.606
75.511
75.511
59.612
75.511
2.431
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
Continued
231
Table 6.6 continued
232
Temperature C
Pressure bar
Mass Flow kg/hr
Volume Flow cum/hr
Mole Flow kmol/hr
CH4
CO2
CO
H2
H2O
H2S
COS
N2
O2
AR
O2S
CaCO3
CAO
C
CAS
S
20
22
23
24
25
26
27
742.1
214.8
515.9
476.7
550
250
620.4
21.28
21.28
21.28
1
21.28
21.28
21.28
157038.1
255817
255817 49643.25 202095.1 202095.1 3062.598
29696.1 18942.53 43117.13
18.473 53641.68 33460.21 588.675
7.452
163.934
4567.793
2056.625
186.288
85.693
6.587
305.513
0
59.612
0
0
0
0
0
0
0
0
0
0
14200
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
14200
0
0
0
0
0
0
0
0
0
0
0
0
55
55
0
3.371
3.371
0
0.834
0.834
0 6433.39 6433.39
0 9827.635 9827.635
0
2.484
2.484
0
0
0
0 305.513 305.513
0
0
0
0
59.612
59.612
0
0
0
496
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
170
0
0
0
0
0
0
0
0
0
0
0
Continued
232
Table 6.6 continued
233
Temperature C
Pressure bar
Mass Flow kg/hr
Volume Flow cum/hr
Mole Flow kmol/hr
CH4
CO2
CO
H2
H2O
H2S
COS
N2
O2
AR
O2S
CaCO3
CAO
C
CAS
S
29
30
31
32
34
100
25
620.4
625
50
1
10
21.28
1
21.28
25919.03 15999.4 252754.4 476650.2 13270.21
25132.3 1234.115 48583.02 158.714
17.892
0
0
0
0
0
0
0
0
810
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
500
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
14030
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0 6255.324
0
0
0 1744.676
0
0
0
0.001
0
0.006
736.598
0.004
0
0
0
0
0
0
0
0
0
0
35
36
57.6
50
21.28
1
255817 313741.2
346.881 199008.8
0
0
0 6200.733
0
16.051
0
3.482
14200 1466.272
0
2.438
0
0.522
0 387.024
0
0
0
75.511
0
0
0
0
0
0
0
0
0
0
0
0
Continued
233
Table 6.6 continued
234
Temperature C
Pressure bar
Mass Flow kg/hr
Volume Flow cum/hr
Mole Flow kmol/hr
CH4
CO2
CO
H2
H2O
H2S
COS
N2
O2
AR
O2S
CaCO3
CAO
C
CAS
S
37
38
40
41
42
43
44
20
20
70
600
589.9
1537.8
850
1.01
1.01
21.28
21.28
1
36
1
289402.3 24338.87 175407.7 684347.6 691561.3 41898.56 476650.2
162798.7
32.108
240.29 255.205 257.8654
8351 158.7155
0
0
0.003
0
0
1.988
0
6200.733
0.076
0.008
0
0
43.74
0
16.051
0
0
0
0 1218.711
0
3.482
0
0.079
0
0 548.718
0
115.445 1350.827 9736.459
0
0
49.702
0
2.438
0
0.052
0
0
22.863
0
0.522
0
0
0
0
1.757
0
387.024
0
0.002
0
0
81.512
0
0
0
0
0
0
0
0
75.511
0
0.006
0
0
15.905
0
0
0
0
0
0
0
0
0
0
0 4686.561 4891.995
0
0
0
0
0 1478.97 1387.27
0 6255.324
0
0
0
0
0
0
0
0
0
0 1834.473 1720.735
0 1744.676
0
0
0
0
0
0
0
Continued
234
Table 6.6 continued
235
Temperature C
Pressure bar
Mass Flow kg/hr
Volume Flow cum/hr
Mole Flow kmol/hr
CH4
CO2
CO
H2
H2O
H2S
COS
N2
O2
AR
O2S
CaCO3
CAO
C
CAS
S
45
46
47
48
49
50
51
850
850
600
850
625
70
600
1
1
21.28
10
1
21.28
21.28
790393.8 313741.2 202095.1 31014.94 313741.2 11671.91 202095.1
761619.5 761460.9 56954.17 12510.02 608906.6 7857.965 56954.17
0
0
55
6200.731 6200.809
3.371
16.104
16.051
0.834
3.494
3.482 6433.39
1466.284 1466.284 9827.635
2.438
2.438
2.484
0.522
0.522
0
387.026 387.024 305.513
0
0
0
75.512
75.511
59.612
0
0
0
0
0 4686.561
6255.173
0
1477.2
0
0
0
1744.827
0 1836.239
0
0
0
0
0
59.193 6200.809
0
16.051
0
3.482
846.933 1466.272
0
2.438
0
0.522
305.511 387.024
64.277
0
59.606
75.511
2.431
0
0
0
0
0
0
0
0
0
0
0
0
55
0
3.371
0
0.834
5789.98 6433.39
0 9827.635
0
2.484
0
0
0 305.513
0
0
0
59.612
0
0
0
0
0
0
0
0
0
0
0
0
Continued
235
Table 6.6 continued
236
Temperature C
Pressure bar
Mass Flow kg/hr
Volume Flow cum/hr
Mole Flow kmol/hr
CH4
CO2
CO
H2
H2O
H2S
COS
N2
O2
AR
O2S
CaCO3
CAO
C
CAS
S
52
53
54
55
56
25
50
620.4
1537.8
70
1
21.28
21.28
36
21.28
49643.25 42431.31
255817 157038.1 15015.54
18.293
15.695 49171.69
31300 1470.796
0
0
0
0
0
0
0
0
0
0
0
496
0
0
0
0
0
0
0
0
0
0
0
0
0
0
0
290.567
91.586
0
113.847
0
0
7.452
0 163.934
0 4567.793
0 2056.625
14200 186.288
0
85.693
0
6.587
0 305.513
0
0
0
59.612
0
0
0
0
0
0
0
0
0
0
0
0
236
54.997
3.363
0.834
643.331
91.176
2.431
0
305.511
0
59.606
0
0
0
0
0
0
MWe
Electricity from Steam Turbine
Electricity from Syngas Expander
Electricity Output
134.964
6.842
141.806
Coal handling, milling and coal slurry pumps
ASU air compressor and oxygen compressor
CO2 compressor
Feed water pumps
H2 Compression
Electricity Used in the Plant
Net Electricity Produced
1.01088
30
17.8218
0.71934
11.6
61.1521
80.6539
Table 6.7: Power balance in the CLP process
237
Hydrogen and electricity
2405
100
280
457
81
66.5
Coal feed (tonnes/day)
Carbon Capture(%)
Hydrogen(tonnes/day)
Hydrogen (MW)
Net Power(MW)
Overall Efficiency(%HHV)
Table 6.8: Process simulation results for the CLP process
238
Hydrogen
Coal Feed (t/day)
Carbon Capture (%)
Hydrogen (t/day)
Hydrogen (MW, HHV)
Net Power (MW)
Overall Efficiency (%HHV)
2,180
100
280
457
0
62.3
Hydrogen and
Electricity
2,405
100
280
457
81
66.5
Table 6.9 Summary of the schemes investigated for the production of H2 alone and for
the coproduction of H2 and electricity with a PSA
239
Coal Feed (t/day)
Carbon Capture(%)
Hydrogen (t/day)
Hydrogen (MW)
Net Power (MW)
Overall Efficiency (%HHV)
Hydrogen
2,155
100
280
457
0
63
Hydrogen and Electricity
2,350
100
280
457
67.56
66.3
Table 6.10 Summary of the schemes investigated for the production of H2 alone and for
the coproduction of H2 and electricity without a PSA
240
Hydrogen
and
Hydrogen
Electricity
Shell
Lurgi
(BGL)
GE
62.3%
66%(81 MW)
55%
56(32 MW)
60%
63.6(104.2 MW)
Table 6.11 Comparison of the efficiency of the H2 production process for different
gasifiers
241
Capital (TPC)
Fixed O&M
Coal
Electricity
CO2 Emission Allowances
Other Variable O&M
TOTAL
Cost ($ or $/y)
1444612000
34447025.6
81439239.83
‐23448072
0
6188890.725
Capital Charge Factor and Levelization Levelized factors
Annual Cost 0.175
252807100
1.1757
40499368
1.2485 101676890.9
1.1907 ‐27919619.33
1
0
1.1757 7276278.825
374340018.4
Levelized cost of H2=374340018.14/ 183970841 = $2.03/kg H2
Table 6.12: Levelized annual costs and levelized cost of H2 for the conventional coal to
H2 plant (adapted from DOE, 2010)
242
Capital (TPC)
Fixed O&M
Coal
Electricity
CO2 Emission Allowances
Other Variable O&M
TOTAL
Cost ($ or $/y)
2085092000
49221169.6
119926980.7
‐246429756
‐24616772.44
34004142.92
Capital Charge Factor and Levelization factors
0.175
1.1757
1.2485
1.1907
1
1.1757
Levelized Annual Cost 364891100
57869329.1
149728835.4
‐293423910.5
‐24616772.44
39978670.84
294427252.4
Levelized cost of H2= 294427252.4/ 162678456 = $1.81/kg H2
Table 6.13: Levelized annual costs and levelized cost of H2 for the CLP plant
243
Calciner
Radiant Cooler
Coal
GE Gasifier
Gasifier
Slag
H2 Production
Reactor
PCD
Steam
Calcium Looping
Process
CO2
Compression
Hg Removal
CO2 to
Sequestration
PSA
Pure H2
Slurry
Water
ASU
N2 Rich Stream
Air
Figure 6.1: The CLP for coproduction of fuel cell grade H2 and electricity from coal
244
Shell Gasifier
Coal
Q-DECOMP
DRY - CO AL
Steam Generation
D ECOMP
GAS IFIER
INBU RNER
RYI E LD
STEA M3
RGI BBS
32
A IR
1
B12
B7
H OTSEQCO
ASU
CA O
9
A SU
O2C
2
CO2 H2S -O
25
13
Integrated
Reactor
Calciner
CAC O3SO L
23
N23
P1
CA LC
B10
COO L
2
C AO,CO2
Q
W
CYC CO2
16
2
4
B20
20
B2
B 13
SOL MA KUP
245
22
Q
MI XE R
29
COMP A IRO
15
CARB
11
B1
8
S YNGA S
CA LCS Y NG
24
B11
CYCH2
5
B5
SOL PUR GE
4
H OTOFFGA
B4
FS
P LI T
33
C ACO3
PUREH2
19
3
COM BUS TO
Q
TAILGA S
B14
B8
6
34
B6
30
53
H2 at 20 bar
COMPC O2
CO2
B9
PSA
10
B16
18
1
B15
PSA
B3
31
35
Figure 6.2: ASPEN simulation flow diagram for the CLP process with a PSA
245
Shell Gasifier
Coal
Steam Generation
40
STEAM3
Q-DECOMP
32
B12
DECOMP
B7
HOTSEQCO
CAO,CO2
9
GASIFIER
39
INBURNER
RYIELD
RGIBBS
CAO
25
13
CO2H2S-O
CACO3SOL
23
CALCSYNG
B22
17
20
CALC
B10
CARB
2
AIR1
W
38
Q
B19
CYCCO2
14
26
B2
B13
Calcine
r
5
B5
35
B21
B11
SOLPURGE
ASU
11
B1
SYNGAS
24
28
B20
MIXER
4
B4
FSPLIT
246
CACO3
18
Q
1
O2C
CYCH2
ASU
Integrated
Reactor
2
Q
P1
COOL
COMP
PUREH2
3
B8
19
H2 at 21
20atm
bar
B6
FLASH2
B3
7
B18
COMPCO2
B16
136 atm
psi
CO2 at 2000
53
HYDROGEN
33
Figure 6.3 Aspen simulation for the production of H2 using the CLP without a PSA.
246
27
B8
B22
STEAM3
32
39
B12
B7
HOTSEQCO
W=-13591
31
CAO
9 38
Hydrogen production reactor
CO2H2S-O
25
29
23
Q
Duty (Gcal/hr)
W
Power(kW)
CACO3SOL
CALC
13
B10
15
CARB
2
CAO,CO2
Q
W
CYCCO2
B20
26
14
B2
B13
SOLMAKUP
22
Q=7
B11
MIXER
Q=132
Q
24
Q=-315
CYCH2
SOLPURGE
B5
B1
W=-8891
11
B15
SYNGAS
Q=-90
HOTOFFGA
5
4
Q=-7
B4
FSPLIT
33
CACO3
19
10
18
W=311 B17
B16
W=52636
3
COMBUSTO
Q=-2
Q
B6
B18
Q=-1
TAILGAS
Q=-0
B14
B21
B9
PSA
PUREH2
Q=-37
30
53
6
Q=-11
B3
35
Figure 6.4: Aspen model used for sensitivity analysis of the combined reactions
occurring in the H2 production reactor of the CLP.
247
100
Effect of Temperature on Hydrogen Purity
80
75
70
65
60
55
50
248
Purity of Hydrogen
85
90
95
PURE
550
575
600
625
650
675
700
725
750
775
800
825
850
875
900
925
950
975
1000
Temperature (C)
Figure 6.5: Effect of temperature on the H2 purity produced at the outlet of the carbonation reactor (S:C ratio = 3, Pressure = 10 atms)
248
96.8
96.7
96.6
96.5
96.4
96.3
96.2
249
Purity of Hydrogen
96.9
97
97.1
97.2
97.3
Effect of Pressure on Hydrogen Purity
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
18
19
20
21
22
23
24
25
26
27
28
29
30
31
32
33
Pressure (bar)
(atm)
Figure 6.6: Effect of pressure on the H2 purity produced at the outlet of the carbonation reactor( S:C ratio = 3,
Temperature = 600 ºC)
249
34
35
36
37
38
39
40
98
Effect of Steam to Carbon Ratio on Hydrogen Purity
94
93
92
91
90
89
88
87
86
85
250
Hydrogen Purity (%)
95
96
97
PURE
0.8
1
1.2
1.4
1.6
1.8
2
2.2
2.4
Steam to Carbon Ratio
2.6
2.8
3
3.2
3.4
3.6
Figure 6.7: Effect of S:C ratio on the H2 purity produced at the outlet of the carbonation reactor ( Pressure = 10 atms, Temperature =
600 ºC)
250
Steam Ratio vs. H2S Output
90
Outlet H2S Flowrate (Kmoles/hr)
80
600
70
650
60
700
750
50
800
40
850
30
900
20
950
1000
10
0
0
0.5
1
1.5
2
2.5
3
3.5
4
4.5
Steam Ratio
Figure 6.8: Effect of temperature and S:C ratio on the extent of H2S removal.
251
Steam Ratio vs. COS Output
Outlet COS Flowrate (Kmoles/hr)
1.2
1
600
650
700
750
800
850
900
950
1000
0.8
0.6
0.4
0.2
0
0
0.5
1
1.5
2
2.5
3
3.5
4
4.5
Steam Ratio
Figure 6.9: Effect of temperature and S:C ratio on the extent of COS removal.
252
Steam Ratio vs. CO Output
4000
Outlet CO Flowrate (Kmoles/hr)
3500
3000
600
650
700
750
800
850
900
950
1000
2500
2000
1500
1000
500
0
0
0.5
1
1.5
2
2.5
3
3.5
4
4.5
Steam Ratio
Figure 6.10: Effect of temperature and S:C ratio on the amount of CO impurity present
in the H2 stream.
253
Steam Ratio vs. CO2 Output
5000
Outlet CO2 Flowrate (Kmoles/hr)
4500
600
4000
650
3500
700
3000
750
2500
800
2000
850
900
1500
950
1000
1000
500
0
0
0.5
1
1.5
2
2.5
3
3.5
4
4.5
Steam Ratio
Figure 6.11: Effect of temperature and S:C ratio on the extent of CO2 removal.
254
Steam Ratio vs. CH4 Output
800
Outlet CH4 Flowrate (Kmoles/hr)
700
600
600
500
650
700
400
750
300
800
850
200
900
950
100
1000
0
0
0.5
1
1.5
2
2.5
3
3.5
4
4.5
Steam Ratio
Figure 6.12: Effect of temperature and S:C ratio on the amount of CH4 impurity present
in the H2 product stream.
255
65
98
97
60
55
95
Efficiency (%)
H2 Purity (%)
96
94
50
93
92
45
0
5
10
15
20
Pressure (bar)
atm
S/C = 1 - H2 Purity
S/C = 2 - H2 Purity
S/C = 1 - H2 Cold gas Efficiency
S/C = 2 - H2 Cold gas Efficiency
S/C = 1 -Process Efficiency
S/C = 2 -Process Efficiency
Figure 6.13: Effect of pressure on the cold gas efficiency, process efficiency and H2
purity obtained from the H2 production reactor at various S:C ratios.
256
97.5
65
97
H2 Purity
96
55
95.5
95
Efficiency (%)
60
96.5
50
94.5
94
45
0
0.5
1
1.5
2
2.5
3
3.5
S/C Ratio
H2 Purity
H2 Cold gas Efficiency
Process Efficiency
Figure 6.14: Effect of S:C ratio on H2 purity, cold gas efficiency and process efficiency
257
94.4
65
94.2
H2 Purity (%)
93.8
55
93.6
93.4
Efficiency (%)
60
94
50
93.2
93
580
600
620
640
660
680
700
45
720
Temperature (C)
H2 Purity
H2 Cold gas Efficiency
Process Efficiency
Figure 6.15: Effect of temperature on H2 purity, cold gas efficiency and process
efficiency (1:1, 10 atms)
258
94.9
65
94.8
60
94.6
94.5
55
94.4
94.3
Efficiency(%)
H2 Purity(%)
94.7
50
94.2
94.1
45
0
0.5
1
1.5
2
2.5
Ca/C Ratio
3
3.5
4
4.5
H2 Purity
H2 Cold gas Efficiency
Process Efficiency
Figure 6.16: Effect of Ca:C ratio on H2 purity, cold gas efficiency and process
efficiency (600 ºC, 1:1, 10 atms)
259
Hydration
Q
Duty (Gcal/hr)
W
Power(kW)
Steam generation
B29
W=57291
9
B25
31
B7
Q=-8
Q=14
Q=38
Q=22
B27
12
43
3
B6
Q=6
B10
Q=54
38
2
Calcination
33
Integrated Reactor
CO2H2S-O
CALC
19
22
15
39
W=105 B3
49
CARB
CAO,CO2
18
CYCCO2
B1
8
21
Q=124
W=-8891
Q
Q=-90 B15
Q=-173
B2
24
B20
20
Q=22
Q=6
B13
W
B9
Q
17
SOLMA KUP
M I XER
TAILGA S
1
260
SOLPURGE
B5
H2
4
B4
F S P L I T
Q=-6
CYCH2
PSA
CACO3
5
B28
44
34
7
Q=-35
W=-10796
B8
40
Q=-161
Q=8
B22
B18
36
47
13
B19
B11
B17
14
32
B21
Q=5
Q=-10
B12
CAO
B14
B26
42
27
30
54
45
PUREH2
10
Q=-1
Q=-3
Q=14
53
Figure 6.17: Effect of the addition of sorbent hydration to the CLP
260
HYDROGEN
B16
23
Q
SYNGA S
CO 2
Sequestration
CO2
Compression
Sulfuric
Acid
Plant
H2S
Removal
Mercury
Removal
Dual
Stage
Acid Gas
Removal
Steam
Scrubber
261
Air
Radiant
Cooler
Raw
Syngas
Shift
Reactors
Syngas
Cooling
Final
Syngas
Scrubber
ASU
PSA
H2
Boiler
Flue
Gas
Gasifier
Air
Coal
Feed
Coal Water
Steam
Turbine
Slag
Figure 6.18: Process flow diagram of the conventional coal to H2 plant used for the economical analysis ( DOE, 2010)
261
Metallic
Filter
Air
Expander
Carbonator
H2
Cooling
Metallic
Filter
PSA
H2
Radiant
Cooler
ASU
ASU
Gasifier
Coal
Coal
Feed
262
Coal Water
Filter
Sorbent
Makeup
Calci ner
Solids
Purge
Slag
Filter
Hydrator
Figure 6.19: Process flow diagram of the CLP plant used for the economical analysis
262
CO2
Condenser
CHAPTER 7
ENHANCED REFORMING OF HYDROCARBONS
7.1 INTRODUCTION
The CLP for the reforming of hydrocarbons is similar to the CLP for the
conversion of syngas to H2. For a hydrocarbon feed, the steam reforming of the
hydrocarbon is integrated with the water gas shift and carbonation reaction in a single
reactor. In addition to improving the conversion of the hydrocarbon to H2, the CLP also
provides an efficient mode of internal heat integration where the endothermic energy
for the reforming reaction is obtained from the exothermic energy released by the
combined water gas shift and carbonation reaction. A schematic of the CLP for the
reforming of hydrocarbons is shown in Figure 7.1. The CLP comprises of three
reactors; the carbonation reactor where the thermodynamic constraint of the reforming
and water gas shift reaction is overcome by the in-situ removal of the CO2 product by a
calcium based sorbent, the calciner where the spent calcium sorbent is regenerated and
a sequestration-ready CO2 stream is produced and the hydrator where the calcined
sorbent is reactivated to improve its recyclability.
263
7.2 PROCESS CONFIGURATION AND THERMODYNAMICS
7.2.1 The Carbonation Reactor System
The carbonation reactor comprises either a fluidized bed, fixed fluidized bed or
an entrained flow reactor that operates at pressures ranging from 1 to 30 atm and
temperatures of 500-750 oC. In the carbonation reactor, the reforming reaction, water
gas shift reaction and CO2 removal occur in a single reactor in the presence of a
reforming catalyst and CaO sorbent. The steam reforming of the hydrocarbon occurs in
the presence of the reforming catalyst, and the CO2 produced by the combined
reforming and water gas shift reaction is removed by the CaO sorbent. The
concomitant carbonation of the CaO leading to the formation of CaCO3 incessantly
drives the equilibrium-limited water gas shift and reforming reaction forward by
removing the CO2 product from the gas mixture.
Various reactions occurring in the carbonator are as follows:
Hydrocarbon reforming:
CxHy + xH2O Æ
xCO + (y/2+x) H2
(7.1)
CH4 + H2O
Æ
CO + 3H2
(7.2)
Water Gas Shift Reaction:
CO + H2O
Æ
H2 + CO2
(7.3)
Carbonation Reaction:
CaO + CO2
Æ
CaCO3
(7.4)
The CLP offers several advantages. By improving the equilibrium conversion
of the reforming and water gas shift reaction, steam addition can be greatly reduced. In
264
addition, since the combined reforming, water gas shift and carbonation reaction occurs
at a high temperature of 500 to 750 ºC, the water gas shift catalyst can also be
eliminated. A major advantage of the CLP is the internal heat integration that it
provides to the reforming of hydrocarbons. The exothermic carbonation and water gas
shift reactions convert the highly endothermic reforming of hydrocarbons into a heat
neutral process thus simplifying the reforming process and reducing the temperature of
reforming from >900 ºC to 650 ºC. The heat of reaction of the combined steam
methane reforming, water gas shift and carbonation reaction occurring in the
carbonator is shown below:
Steam Methane Reforming
and Water Gas Shift:
CH4 + 2H2O = CO2 + 4H2
H = +165 KJ/mole
(7.5)
Carbonation Reaction:
CaO + CO2 = CaCO3
H = -178 KJ/mole
(7.4)
H = -13 KJ/mole
(7.6)
Combined Reaction: CH4 + 2H2O+CaO = CaCO3 + 4H2
Thermodynamic analysis of reactions occurring in the carbonation reactor
The equilibrium constants for the steam methane reforming (equation 7.2),
steam methane reforming and water gas shift reaction (equation 7.5), and the combined
reforming and carbonation reaction (equation 7.6) for various temperatures are shown
in Figure 7.2. The equilibrium constants are obtained using HSC Chemistry v 5.0
(Outokumpu Research Oy, Finland). The equilibrium constant for the steam methane
reforming reaction can be defined as shown below:
265
Keq
eq1=
PH3PCPCO
O
PPCH
C 4 PP
HHOO
2
2
2 2
(7.7)
where PCH4, PH2, PCO, PH2O are the partial pressures of CH4, H2, CO and H2O at
equilibrium.
The equilibrium constant of the combined reforming and water gas shift
reaction is as follows:
Keq 2
PHH224PCPCO
2 2
O
=
2
PPCH
C 4 PP
HH22OO
(7.8)
The equilibrium constant for the combined reforming, water gas shift and
carbonation reaction is defined as shown below:
P H2 4
Keq 3 =
PCH PH O2
4
2
(7.9)
Where Keq3 = Keq2 * Kcarb and Kcarb is the equilibrium constant of the CaO carbonation
reaction.
From Figure 7.2 it can be seen that the steam methane reforming reaction does
not occur at temperatures below 600 ºC. The equilibrium constants of the steam
methane reforming and the combined reforming and water gas shift reactions increase
266
with the increase in temperature beyond 600 ºC. The equilibrium constant of the
combined CaO carbonation, steam methane reforming and water gas shift reaction is
higher than that of the steam methane reforming reaction alone at temperatures of 0 to
800 ºC. Beyond 800 ºC, the carbonation of CaO does not occur. Hence, the CLP is
capable of reducing the temperature of the steam methane reforming reaction with the
aid of the conventional steam methane reforming catalyst.
7.2.2 Calciner or Sorbent Regeneration Reactor
The spent sorbent at the exit of the carbonation reactor is a mixture consisting
of CaCO3 and CaO. The CaCO3 in the spent sorbent mixture is regenerated back to
CaO in the calciner. The calciner is operated at atmospheric pressure in a rotary or a
fluidized bed system. The heat can be supplied directly or indirectly using a mixture of
fuel and oxidant. For a directly fired calciner, the heat of calcination can be provided
by the combustion of natural gas in oxygen. From the thermodynamic curve for CaO
and CO2, calcination is found to occur at temperatures above 890 oC in the presence of
1 atm of CO2. Dilution of CO2 in an indirectly fired calciner with steam or oxycombustion of natural gas in a direct fired calciner will permit the calcination reaction
to be conducted at temperatures lower than 890 oC . The reaction occurring in the
calciner is:
Calcination : CaCO3→ CaO + CO2
267
(7.10)
7.2.3. Hydrator or Sorbent Reactivation Reactor
The hydrator is similar to that described in Chapter 4 for H2 production from
syngas and the reaction occurring in the hydrator is shown below:
Hydration : CaO + H2O = Ca(OH)2
(7.11)
The Ca(OH)2 from the hydrator is conveyed to the carbonation reactor where it
dehydrates to produce high reactivity CaO and steam. The steam obtained from the
dehydration reaction is consumed in the combined reforming and water gas shift
reaction.
7.3. EXPERIMENTAL METHODS
7.3.1 Chemicals, Sorbents, and Gases
The reforming catalyst was procured from Süd-Chemie Inc., Louisville, KY
and consists of a nickel oxide catalyst supported on calcium aluminate. The CaO
sorbent is obtained from a PCC precursor which is synthesized from Ca(OH)2 obtained
from Fisher Scientific as described in Chapter 4.
7.3.2 Bench Scale Experiment Setup
Figure 7.3 shows the integrated experimental setup, used for the bench scale
studies of the CLP. The bench scale reactor is coupled with a set of continuous gas
analyzers which detect concentrations of CO, CO2, H2S, CH4 and H2 in the product
268
stream. The reactor setup is capable of handling high pressures and temperatures of up
to 21 atms and 900 oC respectively, which are representative of the conditions in a
commercial syngas to H2 system. A description of the bench scale setup is provided in
Chapter 4.
7.3.3 Steam Methane Reforming in the Presence of a Ni-based Catalyst
The extent of the steam methane reforming reaction was determined in the
presence of the reforming catalyst obtained from Süd-Chemie. Catalyst particles were
used in a fixed bed reactor setup for all the experiments. 3 g of the catalyst was loaded
into the reactor and the pressure, temperature and gas flow rates were adjusted for each
run.
Pure CH4 was used as the feed gas for all the tests and it was metered into the
reactor by a mass flow controller. From the mass flow controller, the CH4 flows to the
steam generating unit which also serves to preheat the CH4 entering the reactor. The
product gas mixture exiting the back pressure regulator is then cooled in a heat
exchanger using chilled ethylene glycol-water mixture to condense the unconverted
steam. The product gas at the exit of the heat exchanger is dried in a desiccant bed. The
dry gas compositions are monitored continuously using the CO, CO2, H2S, CH4 and H2
gas analyzers.
269
7.3.4 Simultaneous Steam Methane Reforming, Water Gas Shift and Carbonation
The simultaneous reforming, water gas shift and carbonation reaction was
conducted in the presence of the CaO sorbent and nickel based reforming catalyst. 2.5
grams of PCC was mixed with 2.5 grams of reforming catalyst and loaded into the
reactor. The mixture was calcined by heating the reactor to 700 oC in a stream of N2
until the CO2 analyzer confirmed the absence of CO2 in the outlet stream. At the end of
calcination, the temperature of the reactor was reduced to 650 ºC and H2 was made to
flow through the catalyst and sorbent bed for an hour to reduce the catalyst to the active
form for the steam methane reforming reaction. At the end of catalyst pretreatment, the
temperature of the reactor was set for the combined reforming water gas shift and
carbonation reaction. Pure CH4 was used as the feed gas for all the tests and it was
metered into the steam generator by a mass flow controller. The CH4 and steam
mixture was introduced into the reactor and dry gas composition of the product gas was
monitored continuously using the CO, CO2, H2S, CH4 and H2 gas analyzers.
7.3.5 Multicyclic Steam Methane Reforming and Spent Sorbent Calcination
7.3.5.1 Effect of Sorbent Calcination Conditions on the Extent of Steam Reforming
In these tests the effect of two sorbent calcination conditions was investigated
on the extent of CH4 reforming obtained in the presence of CaO sorbent and Ni based
catalyst. In the first set of tests, the PCC sorbent and Ni-based catalyst mixture was
loaded into the reactor and the calcination was conducted at 950 ºC in pure N2. The
270
mixture was then exposed to H2 to activate the catalyst at 650 ºC. The combined
reforming, water gas shift and carbonation reaction was conducted at 650 ºC and 1 atm
with a 100% CH4 feed stream and a S:C ratio of 3:1. At the end of the H2 production
stage, the sorbent was calcined again at 950 ºC in pure N2 and 3 cycles were repeated.
For the 4th cycle the spent sorbent at the end of the 3rd cycle was calcined in a 50:50
CO2/H2O mixture at 950 ºC. The sorbent catalyst mixture was then reduced in H2 at
650 ºC and the combined reforming, water gas shift and carbonation reaction was
conducted at 650 ºC.
7.3.5.2 Calcination in N2 with Sorbent Hydration
2.5 grams of PCC sorbent was loaded into the reactor and the calcination was
conducted at 950 ºC in pure N2. At the end of calcination, the sorbent was hydrated in a
80:20 H2O/N2 gas mixture at 600 ºC and 11 atms. The hydrated sorbent was then
mixed with the 2.5 g of reforming catalyst and loaded into the reactor. The mixture was
then exposed to H2 to activate the catalyst at 650 ºC. The combined reforming, water
gas shift and carbonation reaction was conducted at 650 ºC and 1 atm with a 100% CH4
feed stream and a S:C ratio of 3:1. At the end of the H2 production stage, the sorbent
was separated from the catalyst. The sorbent was calcined again at 950 ºC in pure N2
and then hydrated and 4 cycles were repeated in a similar manner.
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7.3.5.3 Realistic Sorbent Calcination in a Steam/CO2 Atmosphere With Hydration
2.5 grams of PCC sorbent was loaded into the reactor and the calcination was
conducted at 950 ºC in a 50:50 CO2/H2O mixture. At the end of calcination, the sorbent
was hydrated in a 80:20 H2O/N2 gas mixture at 600 ºC and 11 atms. The hydrated
sorbent was then mixed with 2.5 grams of the reforming catalyst and loaded into the
reactor. The mixture was then exposed to H2 to activate the catalyst at 650 ºC. The
combined reforming, water gas shift and carbonation reaction was conducted at 650 ºC
and 1 atm with a 100% CH4 feed stream and a S:C ratio of 3:1. At the end of the H2
production stage, the sorbent was separated from the catalyst. The sorbent was calcined
again at 950 ºC in a 50:50 CO2/H2O mixture and 4 cycles were repeated in a similar
manner.
7.4 RESULTS AND DISCUSSION
7.4.1 Base-line Steam Methane Reforming Testing
Base line tests were conducted in the bench scale reactor for the steam
reforming of CH4 in the presence of the nickel based catalyst procured from Sud
Chemie. The reactor was filled with a mixture 5 gms of catalyst and quartz chips. Pure
CH4 from the mass flow controller was mixed with steam in the steam generation
section and sent to the reactor. Figure 7.4(a) illustrates the composition of H2 in the
product gas at the outlet of the reactor on a steam and nitrogen free basis. It can be seen
in Figure 7.4(a) that for both S:C ratios of 3:1 and 5:1, the composition of H2 increases
272
with the increase in temperature till the temperature reaches 700 ºC. At temperatures
above 700 ºC, the purity of H2 remains constant. With the increase in S:C ratio from
3:1 to 5:1, it was found that the purity of H2 increases at all the temperatures
investigated. Figure 7.4(b) depicts the composition of CO, CO2 and CH4 in the product
gas at the outlet of the reactor on a steam and nitrogen free basis. The amount of CH4
in the stream decreases with temperature and is in the ppm range at temperatures above
850 ºC. It can also be seen that with the increase in the S:C ratio the conversion of CH4
to CO, CO2 and H2 increases. The CO content of the product stream increases with the
increase in temperature as the conversion of CH4 to CO increases. In addition, the
conversion of CO to CO2 decreases with an increase in temperature due to the
equilibrium limitation of the water gas shift reaction. The amount of CO2 decreases
with an increase in temperature due to the thermodynamic constraint of the water gas
shift reaction. The amount of CO2 increases with the increase in S:C ratio due to the
higher conversion of CO to CO2 by the water gas shift reaction at higher steam addition
rates.
7.4.2 Simultaneous Reforming with In-situ CO2 Removal (Catalyst with CaO
Sorbent)
In order to study the improvement in CH4 conversion and H2 purity, the steam
methane reforming reaction was conducted in the presence of the reforming catalyst
and CaO sorbent. The mixture of sorbent and catalyst was fed into the reactor and the
sorbent was calcined at 700 ºC in pure nitrogen. Following this, the catalyst and
273
sorbent mixture was exposed to pure H2 at 650 ºC in order to reduce the catalyst to its
active form. Pure CH4 and steam were then fed into the reactor at 650 ºC and
atmospheric pressure. Figure 7.5 depicts the concentration of H2, CO, CO2 and CH4 in
the product gas for the steam methane reforming reaction conducted in the presence of
CaO sorbent and Ni based catalyst. It was found that >99% pure H2 can be obtained in
the pre-breakthrough region of the curve when the CaO sorbent is active. It was also
found that the CH4 is almost completely converted and the concentration of CH4, CO
and CO2 in the product stream is only a few ppm. The removal of CO2 by the CaO
sorbent enhances the water gas shift reaction and the reforming reaction resulting in the
production of a pure H2 product stream. As the sorbent gets consumed in the fixed bed,
the concentrations of CO2, CH4 and CO begin to increase in the product H2 stream.
This region of the curve is the breakthrough period. At the end of the breakthrough
period, the CaO sorbent is completely converted to CaCO3 and no further CO2 capture
is obtained. The conversion of CH4 to H2 occurs in the presence of the reforming
catalyst in the post-breakthrough period.
As illustrated in Figure 7.6, it can be seen that >99% conversion of CH4 can be
obtained during the pre-breakthrough period of the combined reforming, water gas
shift and reforming reaction. As the sorbent gets consumed, the conversion of CH4
decreases forming the breakthrough region of the curve.
During the post-breakthrough period, the sorbent is in the form of CaCO3 and
the reforming reaction takes place in the presence of the catalyst alone.
274
7.4.2.1 Effect of Temperature and S:C Ratio
The effect of S:C ratio and temperature was investigated on the reforming of
CH4 in the presence of the reforming catalyst and the CaO sorbent. The purity of H2
produced is greatly enhanced by the presence of the sorbent as shown in Figure 7.7(a).
Purity of H2 increases from <80% in the presence of the catalyst alone to >90% in the
presence of the catalyst and sorbent. Higher H2 purity is obtained at 650 ºC than at 700
ºC due to the favorable thermodynamics of CO2 removal by the CaO sorbent and the
water gas shift reaction at lower temperatures. In the presence of the catalyst and CO2
sorbent, the H2 purity increases with the increase in S:C ratio from 2:1 to 3:1. A further
increase in the S:C ratio does not produce an appreciable increase in H2 purity.
Figure 7.7(b) illustrates the effect of temperature and S:C ratio on the
conversion of CH4. The presence of the sorbent enhances the conversion of CH4 to a
large extent especially at 650 ºC when the presence of the sorbent increases the
conversion of CH4 from 83% to 94% at a S:C ratio of 3:1. At a S:C ratio of 5:1, the
enhancement in CH4 conversion due to addition of sorbent is not significant. The
conversion of CH4 with/without the sorbent almost reaches 100%. The conversion of
CH4 in the presence of the sorbent and catalyst mixture at 650 ºC is similar to that at
700 ºC.
From Figure 7.8(a), it can be observed that at a particular temperature, the CO
composition in the product stream for the sorbent-enhanced reaction is lower than the
275
case with the catalyst alone, which can be attributed to the fact that the presence of CO2
sorbent enhances the water gas shift reaction that converts CO to CO2. The CO
composition is lower at 650 ºC than at 700 ºC as the extent of CO2 removal and the
water gas shift reaction are both thermodynamically more favorable 650 ºC than at 700
ºC. At 650 ºC, the CO composition does not change with S:C ratio as the removal of
CO2 as well as the water gas shift reaction almost reach completion even at a low S:C
ratio. But at a higher temperature of 700oC, the composition of CO in the product
stream is more sensitive to the S:C ratio. On increasing the S:C ratio from 2:1 to 3:1 at
700 ºC, the CO concentration decreases from 12% to 3% and remains constant with the
further increase in steam addition.
Figure 7.8(b) depicts the composition of CO2 in the product gas mixture for the
reforming and the sorbent-enhanced reforming reactions. The CO2 concentration
decreases from >12% to <3% in the presence of the CO2 sorbent. In the presence of the
sorbent, the CO2 composition increases with the increase in S:C ratio at 700 ºC. This
results from a decrease in the partial pressure of CO2 in the reactor due to the presence
of excess steam which reduces the extent of CO2 removal by the sorbent. However, at
650 ºC the CO2 is completely removed from the gas mixture for S:C ratios ranging
from 2:1 to 5:1 as CO2 is removed to very low partial pressures at low temperatures. In
the absence of the sorbent, the CO2 composition in the product stream increases both at
650 ºC and 700 ºC with the increase in S:C ratio due to the production of a larger
amount of CO2 by the water gas shift reaction.
276
7.4.2.2 Effect of Pressure
The effect of pressure was studied on the combined reforming, water gas shift
and carbonation reaction in the presence of the catalyst and sorbent. Thermodynamics
predicts a decrease in the purity of H2 in the presence of a catalyst alone with the
increase in pressure according to the Le Chatlier principle. Figure 7.9(a) shows that the
H2 purity remains almost a constant in the presence of the sorbent due to the
simultaneous removal of CO2. In the presence of the sorbent and catalyst, a high H2
purity of >95% is obtained at pressures ranging from 1 to 11 atms.
Figure 7.9(b) shows the effect of pressure on the concentration of CH4 in the
product stream. With an increase in pressure there is a small increase in CH4
concentration from 3% at 1 atm to 5% at 11 atms in the pre-breakthrough region of the
curves. This is due to the thermodynamics of the combined reforming, water gas shift
and carbonation reactions governed by the Le Chatlier’s principle. In the postbreakthrough region of the curve when the sorbent no longer captures CO2, the increase
in pressure of 1 to 11 atms results in a large increase from 5 to 17% of CH4 in the
product stream. Hence the presence of the sorbent results in a large increase in CH4
conversion even at high pressures.
CO2 in the product steam was reduced to undetectable levels in the presence of
the CaO sorbent in the pre-breakthrough curve at all pressures as shown in Figure
7.10(a). Figure 7.10(b) depicts the effect of pressure on the concentration of CO in the
277
product stream. Due to insitu CO2 removal by the sorbent, the increase in pressure
thermodynamically and kinetically improves the conversion of CO by the water gas
shift reaction. Hence with the increase in pressure from atmospheric to 11 atms, the CO
in the product gas decreases from 1.5% to ppm levels in the pre-breakthrough regions
of the curves.
Figure 7.11 illustrates the change in product gas composition with the increase
in pressure in the pre-breakthrough and post-breakthrough regions. The prebreakthrough compositions are characteristic of the combined reforming, water gas
shift and carbonation reaction as they are produced in the presence of the reforming
catalyst and active CaO sorbent. The post-breakthough compositions are characteristic
of only the reforming and water gas shift reaction as they are produced in the presence
of the reforming catalyst and the spent CaCO3 sorbent. There is almost no change in
the pre-breakthrough CH4, CO2 and CO gas compositions with the change in pressure.
An increase in the post-breakthrough CH4 concentration is observed with the increase
in pressure and this change is predictable from the Le Chatelier’s principle while there
is a decrease in CO concentration. The pre-breakthrough concentrations are
substantially lower than the post-breakthrough concentrations due to the removal of
CO2 by the CaO sorbent. The advantage of the CLP is even more pronounced at high
pressures where the CH4 concentration is reduced by 12 -15%, CO concentration is
reduced by 5-10% while CO2 concentration is reduced by 3%
From the single cycle results shown above it can be inferred that the addition of
278
CaO sorbent significantly increases the conversion of CH4 and the purity of H2 with the
simultaneous removal of CO2 at a temperature of 650 ºC.
7.4.3 Effect of Sorbent Calcination Conditions on the Extent of Steam Reforming:
The effect of two calcination conditions was tested on the extent of CH4
conversion. Figures 7.12 (a) and (b) show the effect of the two calcination conditions
on the purity of H2 produced and the conversion of CH4. The CaO sorbent used in
cycles 1, 2 and 3 was obtained by calcination conducted in a pure nitrogen atmosphere
at 950 ºC. It is observed that, the pre-breakthrough region of the curves during which
the sorbent is active, reduces with successive cycles. From Figure 7.12(a), it can be
observed that for cycles 1,2 and 3, high purity H2 was produced for 464, 432 and 234
seconds, respectively. The earlier onset of breakthrough with increasing cycles can be
attributed to the deactivation of the sorbent due to sintering. This decay in sorbent
activity reduces the overall capacity of the sorbent for CO2 capture, thus limiting highpurity H2 production by the reforming and water gas shift reactions. However, in the
post-breakthrough region (where the sorbent is completely exhausted), for the first
three cycles, H2 purity is almost constant. From 6(b) it can be seen that the
concentration of CH4 in the product stream follows the same trend as the breakthrough
curves for H2 purity. The time for which almost complete CH4 conversion is achieved
reduces with increase in the number of cycles for the first three cycles.
279
The CaO sorbent for cycle 4 was obtained by calcining the spent sorbent from
the previous cycle (cycle 3) in a 50:50 mixture of CO2 and steam at 950 ºC which is
more representative of the realistic conditions used in commercial calciners. From
Figure 7.12(a), it can be seen that for cycle 4, high-purity H2 is not produced and there
is no pre-breakthrough region observed. A similar observation is made for CH4
concentration in the product H2 stream from Figure 7.12(b). This is due to extensive
sintering of the sorbent during calcination in the presence of CO2 and steam. While the
H2 purity in the post-breakthrough region of cycle 4 was almost the same as H2 purity
in the post-breakthrough periods in the first three cycles the same is not true for the
concentration of CH4 in the product stream.
7.4.4 Calcination in N2 with Sorbent Hydration
The effect of sorbent reactivation by hydration was investigated on the cyclic
carbonation and calcination of CaO sorbent during the production of H2 from steam
methane reforming. In Figures 7.13 (a) and (b) the CaO sorbent for all the 4 cycles was
obtained by calcination in the presence of pure nitrogen at 950 ºC. No hydration was
conducted before the first three cycles while sorbent hydration at 600 ºC and 11 atms
was conducted before the 4th reforming cycle. For the first three cycles, it is observed
that the pre-breakthrough region of the curves during which the sorbent is active,
reduces with successive cycles. From Figure 7.13(a), it is observed that for cycles 1,2
and 3, high purity H2 is produced for 500, 285 and 130 seconds, respectively. The
earlier onset of breakthrough with increasing cycles can be attributed to the
280
deactivation of the sorbent due to sintering. However, in the post-breakthrough region
(where the sorbent is completely exhausted), for the first three cycles, H2 purity is
almost constant. This is because the reforming and water gas shift reactions occur in
the presence of the nickel catalyst alone and the H2 production is not enhanced by insitu CO2 capture. Figure 7.13 (b) illustrates the effect of cycling on CH4 concentration
in the product gas stream. During the first three cycles, the time of the prebreakthrough region for CH4 also decreases with the increase in cycle number.
However, in the 4th cycle, it can be observed that the conversion of CH4 and the
purity of H2 are higher than cycles 2 and 3, and high purity H2 is produced for about
470 seconds. This longer duration of the pre-breakthrough region can be attributed to
the reactivation of the sorbent due to hydration, which improves the CO2 capture
capacity of the sorbent.
Figure 7.14 illustrates the purity of H2 produced in 3 cycles of steam CH4
reforming in the presence of CaO sorbent which is obtained by calcination of the spent
sorbent from the previous cycle followed by hydration. The calcination of the sorbent
was conducted in nitrogen at 950 ºC and the hydration was conducted for every cycle
at 600 ºC and 11 atms.
Although there is still a decrease in the purity of H2 produced during the three
cycles, the decrease is small and lower than that observed in the first three cycles of
both figures 7.12(a) and 7.13(a). High purity H2 was produced for 620, 570 and 480
281
seconds in cycles 1, 2 and 3 respectively in Figure 7.14. Thus, hydration helps to
reduce the extent of sintering and arrest the rapid decline in sorbent activity.
7.4.5 Realistic Sorbent Calcination in a Steam/CO2 Atmosphere with Sorbent
Hydration
It is observed in Figure 7.12 (a) that the pre-breakthrough region decreases
from 234 sec to 0 sec due to calcination in the presence of CO2 and steam at 950 ºC
during cycle 4. The effect of hydration on the purity of H2 produced and the extent of
CH4 converted during 4 cycles in which the calcination was conducted in the presence
of a 50:50 H2O/ CO2 atmosphere is shown in Figures 7.15 (a) and (b). The sorbent for
all 4 cycles was obtained by calcination in a steam and CO2 mixture at 950 ºC followed
by hydration at 600 ºC and 11 atms. In Figure 7.15(a), although the sorbent was
calcined in CO2 and steam, a complete loss in sorbent reactivity is not observed as
hydration was conducted every cycle. Although an initial decrease in H2 purity is
observed between the first and the second cycles, the purity is maintained at almost a
constant value in the subsequent cycles. A similar observation is made for the CH4
concentration in the H2 product stream in Figure 7.15(b).
From the multicyclic investigation conducted for different calcination
conditions, it is evident that the reactivation of the sorbent by hydration aids in
reducing the extent of sorbent sintering.
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7.5 APPLICATIONS OF CLP IN HYDROCARBON REFORMING
The CLP enhanced steam reforming of hydrocarbons can be applied to the
production of H2 from natural gas and other hydrocarbon feedstock. It can also be used
for the production of electricity in a carbon constrained scenario from hydrocarbons.
Another important area for its application is in the production of synfuels with carbon
capture through the indirect coal conversion process comprising coal gasification and
the Fisher Tropsch (F-T) Process. Further description of the integration of the CLP
enhance reforming process in natural gas conversion and liquid fuel production is
described in the following sections.
7.5.1 Steam Reforming of Natural Gas and Other Hydrocarbons for H2 and
Electricity Generation:
Figure 7.16 shows the integration of the CLP in a natural gas reforming process
in which the unit operations namely, reforming, water gas shift, CO2 capture and sulfur
removal are integrated in a single reactor system. Within the H2 production reactor, the
natural gas is reformed with steam in the presence of the reforming catalyst and CaO
sorbent. The removal of CO2 removes the thermodynamic limitation of the water gas
shift and the reforming reaction and results in a high conversion of the CH4 to H2. The
H2 production reactor is almost heat neutral due to the exothermic energy from the
water gas shift and carbonation reactions being equal to the endothermic reforming
reaction heat duty. Hence the temperature of operation for the reforming reaction can
283
be reduced from over 900 ºC to 650 ºC resulting in cost savings for the reactor
material. The spent sorbent containing CaCO3 and CaO is separated from the H2 and
regenerated in a calciner at 900 ºC to produce a sequestration ready CO2 stream. The
CaO sorbent is then recycled back to the integrated H2 production reactor. To improve
the reactivity of the sorbent, a sorbent hydration reactor may be added downstream of
the calciner and a part or all of the calcined sorbent may be hydrated before it is fed
back into the H2 production reactor.
7.5.1.1 Technical analysis of the natural gas to H2 process using the CLP
A preliminary process analysis has been conducted for the production of H2 and
electricity from natural gas by the CLP as shown in Figure 7.17. Water, pressurized to
15 atm, is converted to high temperature steam at 650 ºC and fed to the H2 reactor
along with preheated natural gas. The CaO sorbent at 900 ºC from the calciner is also
sent to the H2 production reactor. Although the combination of the reforming, water
gas shift and carbonation reactions is almost heat neutral, the heat given out by the
solids that cool from 900 ºC to 650 ºC in the H2 production reactor makes it slightly
exothermic. The product gases are separated from the spent sorbent in a cyclone and
sent to a Pressure Swing Absorber (PSA) for the production of high purity H2. The
final high pressure H2 product at the exit of the PSA is cooled from 650 ºC to ambient
temperature. The tail gas from the PSA is preheated from 650 to 900 ºC and fed into
the calciner for combustion, to provide the energy required for the calcination reaction.
The tail gas is combusted in oxygen, preheated to 900 ºC, to produce a concentrated
284
CO2 stream for sequestration. Spent sorbent is completely calcined in the calciner. The
hot solids at 900 ºC are conveyed to the H2 production reactor while the CO2 is cooled
from 900 ºC to 25 ºC and compressed to 150 atms for transportation and sequestration.
The results from the mass balance are shown in Table 7.2.
7.5.1.2 Basis for process analysis:
1) Sulfur free natural gas from the pipeline at 41 atms containing 90% CH4, 5% ethane
and 5% nitrogen with a Higher Heating Value of 50.72 MJ/Kg was used for the
analysis. ( DOE, 2002)
2) All the reactions were assumed to proceed to thermodynamic equilibrium. In the H2
production reactor, the extent of various reactions used for conducting the mass and
energy balance are shown in Table 7.1.
The calcination reaction proceeds to completion at a temperature of 900 ºC and hence
all the CaCO3 is converted to CaO in the calciner.
3) A S:C ratio of 3 and a Ca:C ratio of 1.5 is used for the analysis.
4) For this preliminary study, heat loses in all the equipment were assumed to be
minimal.
5) The turbine isentropic efficiency is 72%.
285
6) A five stage compression was assumed for CO2 compression with an isentropic
efficiency of 85%. The energy required for CO2 compression was found to be 105
KWh/tonne of CO2.
7) MWth is used to quantify heat energy and MWe to quantify electrical energy. An
efficiency of 40% was applied for the conversion of heat energy to electrical energy.
8) The reforming catalyst is mostly retained in the H2 production reactor which is a
fixed fluidized bed. Hence only the calcium sorbent is transported between the
carbonator and the calciner in the process.
7.5.1.3 Results from the process analysis
The results for the energy balance are shown in Table 7.3. The cold gas
efficiency of the CPL, defined as the ratio of the Higher Heating Value (HHV) of H2
produced to the Higher Heating Value (HHV) of natural gas, is 84%. In addition to the
production of H2, the process also produces 47.7 MWe of power after accounting for
all the parasitic energy required within the plant.
The detailed explanation for the energy required/produced from each of the unit
operations in the process, listed in Table 7.3 is provided below.
Table 7.4 illustrates the energy required for the production of steam for the
reforming and water gas shift reaction in the H2 production reactor. Since a total of
4515.5 kmoles/hr of carbon is fed to the H2 production reactor and a S:C ratio of 3 is
286
used, 12946.6 kmoles of steam is produced in the steam generator at 15 atms. The Cp
and latent heat values were determined based on the fact that water boils at 199C at 15
atms.
Table 7.5 illustrates the heat required for preheating the natural gas from -27C
to 650 ºC. The CaO sorbent at the exit of the calciner is at 900 ºC and hence it releases
heat in the H2 production reactor which is operated at 650 ºC. The average Cp of the
calcium sorbent over the temperature range of 650 to 900 ºC was determined and used
to calculate the heat released as shown in Table 7.6.
Table 7.7 shows the heat balance within the H2 production reactor. The heat
required for the endothermic reforming reaction is provided by the heat released from
the exothermic water gas shift and carbonation reaction and the hot solids from the
calciner. The amount of heat released by the solids is calculated in Table 7.6.
The H2 produced from the CLP is at a temperature of 650 ºC and is cooled to
ambient temperatures for transportation. The heat released is calculated from Table 7.8
to be 57.14 MWth. The tail gas from the PSA which is at 650 ºC is combusted in the
calciner to provide heat for the calcination reaction. Since the calciner operates at 900
ºC, the tail gas is preheated to 900 ºC before being fed to the calciner and the details for
the heat required is shown in Table 7.9.
287
The oxygen for the combustion of the tail gas in the calciner also needs to be
preheated to 900 ºC and the energy required for preheating is shown to be 23.39 MWth
from Table 7.10.
The spent calcium sorbent consisting of CaO and CaCO3 at the exit of the H2
production reactor is at a temperature of 650 ºC and hence absorbs heat from the
combustion of tail gas in the calciner to heat up to 900 ºC. This energy has been
calculated in Table 7.11 to be 39.23 MWth.
The tail gas is combusted in the calciner with oxygen to produce heat which is
partly used for the endothermic calcination reaction. The remaining heat is used to
produce electricity as shown in Table 7.12. The hot CO2 at 900 ºC from the calciner is
cooled down to 25 ºC and produces 280 MWth as shown in Table 7.13. The steam in
the CO2 stream is condensed out at 100 ºC and the dry CO2 is compressed for
sequestration.
7.5.2 Implementation of Carbon Capture in Liquid Fuels Production From Coal:
Crude oil satisfies majority of the transportation-based energy needs of the
United States and 60% of the crude oil requirement is achieved through imports.
However, with the increasing fluctuation in the price of crude and the desire to achieve
energy independence, there is a renewed focus on alternative technologies to satisfy the
rising demand of energy. Coal-to-Liquids (CTL) through the Fischer-Tropsch (FT)
process is one such promising technology which enables the production of high quality
288
and cleaner liquid fuels from the abundantly present fossil fuel – coal. The estimated
carbon footprint of a CTL plant is 150-175% higher than a petroleum-based plant.
Implementation of CCS in a CTL plant can help in achieving 20% lower life cycle CO2
emissions compared to petroleum based fuel.
The CLP is capable of producing a sequestration ready CO2 stream by
capturing all the CO2 emitted during the CTL process. In addition to achieving carbon
capture, the CLP improves the efficiency of the CTL process by conversion of the
Fischer Tropsch reactor’s off gases to H2. This H2 is used to adjust the H2:CO ratio,
making it suitable for the FT reaction as well as for the product upgrader.
The CTL process can be broken down into two main blocks which consist of
the gasifier block and the Fischer Trospch synthesis block. The gasifier block consists
of the gasifier and other unit operations like particulate and heavy metal removal.
Similarly the FT synthesis block consists of the FT reactor followed by the product
separation unit, hydrocracking and hydrotreating unit, etc. Extensive studies have been
conducted on the gasifier block leading to several demonstration and pilot plant studies
and improvements are being made to their design and operation currently. A few of the
coal gasification projects for electricity, H2, or liquid fuels production are the 313 MW
Tampa Electric’s IGCC Plant in Florida, USA; the 292 MW Wabash River
Gasification Repowering IGCC Project in Indiana, USA; the 253 MW Nuon
Buggenum IGCC Power Plant in Buggenum, the Netherlands; the Shenhua Group
Corp’s 70,000+ barrel liquid fuel/day CTL project which is under construction, China.
289
Similarly the Fischer Tropsch synthesis block has been commercially operated since
World War 2. 9 CTL plants were set up in the Germany at the end of world war II
which produced 4 MMT/year of liquid fuels. In South Africa, Sasol is operating
150,000 BPD CTL plants and currently China is working with Sasol in building 2
plants which are estimated to produced 30 MMtons if liquid fuel. Although liquid fuels
are being produced commercially, no large scale demonstration exists with advanced
technology. The route for the conversion of coal to the FT feed consists of various
stages and consumes excessive parasitic energy. The CLP could reduce the parasitic
energy consumption by achieving contaminant capture at high temperature and by
integrating various operations like reforming, water gas shift, CO2 and sulfur removal
in a single stage reactor system.
7.5.2.1 Description of the Processes
Currently, the production of coal derived liquid fuels is though the coal
gasification and the Fisher Tropsch (F-T) process illustrated in Figure 7.18. A
conventional CTL plant consists of a gasifier which produces the syngas. The H2 to
carbon monoxide (H2/CO) ratio of the syngas is around 0.63, which is much lower than
the optimal ratio of ~2, required for liquid fuel production. Hence, in order to modify
the amount of H2 in the syngas, part of the syngas is introduced to a water gas shift
reactor to be shifted to H2. Since the gas stream contains sulfur impurities, a sulfur
tolerant water gas shift catalyst is used. The rest of the syngas stream passes through a
hydrolysis unit where the COS is converted into H2S.
290
The gas streams from the water gas shift reactor and the hydrolysis reactor are
mixed together and passed through several gas cleanup units that consist of a mercury
removal bed, bulk sulfur removal units, sulfur polishing unit, and CO2 removal units.
After the pollutants are removed, a portion of the syngas is sent to a PSA to separate H2
for use in the product upgrader. The bulk of the clean syngas stream with a H2/CO ratio
of around 2 is sent to the F-T reactor for the production of liquid fuel. The F-T reactor
is capable of converting more than 70% syngas into a wide range of hydrocarbons
ranging from CH4 to wax. The products from the F-T reactor are sent to a product
upgrader where the high molecular weight hydrocarbons are refined into liquid fuel or
naphtha while the low molecular weight offgas stream is sent to a power generation
block to generate electricity for the ASU and other parasitic energy consumption (
Mayer, 2005, Choi et al, 1997). The addition of CO2 capture to the CTL process will
add units to reform the C1-C4 hydrocarbons present in the offgases from the F-T
reactor, water gas shift reactors to shift the CO to H2, and CO2 capture units, making
the over all process very energy intensive. The addition of CO2 capture to a CTL plant
can be simplified by the CLP. The CLP can be integrated in a CTL plant in two
configurations as shown in Figure 7.19.
In configuration 1, the CLP is placed downstream of the F-T reactor and
converts the offgases from the F-T reactor to H2. The F-T reactor offgas contains a
mixture of C1-C4 hydrocarbons and unconverted syngas. The CLP integrates the
reforming of hydrocarbons and the conversion of unconverted syngas to H2 with the
291
capture of CO2 in a single reactor leading to the production of a pure H2 stream. The H2
is separated from the spent sorbent. A part of the H2 can be added to the syngas feed
entering the F-T reactor to improve the H2/CO ratio of the F-T reactor feed. A part of
the H2 can also be fed to the product upgrader to refine the liquid fuel product. Figure
7.21 illustrates a detailed schematic of this proposed process.
In configuration 2, the CLP is placed upstream of the F-T reactor and the feed
to the carbonation reactor consists of the syngas from the gasifier and the offgas from
the F-T reactor. The CLP achieves the following objectives:
a) Converts the C1-C4 hydrocarbons and unconverted syngas from the FT
process, and syngas from the gasifier, into a 2:1 H2:CO stream by shifting the
equilibrium of the water gas shift and reforming reaction in the forward direction by
removing the CO2 product insitu,
b) Achieves simultaneous CO2 and H2S capture at high temperatures,
c) Produces a sequestration ready CO2 stream in the calcination stage
d) Reduces the excess steam requirement which aids in higher levels of H2S
removal
As shown in Figure 7.21 the unreacted syngas and light hydrocarbons from the
FT reactor are mixed with the syngas from the gasifier and sent into the single reactor
system which adjusts the ratio of the H2:CO in the syngas stream by reforming the
292
hydrocarbons and shifting the syngas in the presence of CaO. The concomitant
carbonation of the metal oxide leading to the formation of the metal carbonate
incessantly drives the equilibrium-limited water gas shift and the reforming reaction
forward by removing the CO2 product from the gas mixture. The metal carbonate can
then be regenerated by heating, to give back the metal oxide and a pure CO2 stream. By
improving the equilibrium conversion of the reforming and water gas shift reaction,
steam addition can be greatly reduced. The reduction in steam consumption not only
reduces energy consumption but also aids in the removal of H2S to ppb levels by the
CaO)as steam poses an equilibrium constrain to the removal of H2S. Various reactions
occurring in this system are as follows
xCO + (y/2+x) H2
(7.1)
Æ
H2 + CO2
(7.3)
CaO + CO2
Æ
CaCO3
(7.4)
Sulfidation:
CaO + H2 S
Æ
CaS + H2O
(7.12)
Calcination:
CaCO3
Æ
CaO + CO2
(7.13)
Reforming:
CxHy + xH2O Æ
Water gas shift:
CO + H2O
Carbonation:
7.6 CONCLUSIONS
Single cycle tests have shown that the conversion of CH4 is improved to a large
extent by the addition of CaO sorbent at 650 ºC. High purity H2 is obtained at low S:C
ratios of 3:1 for various pressures ranging from 1 – 11 atms. The purity of H2 was
found to be higher at 650 ºC than at 700 ºC due to the favorable thermodynamics of the
293
carbonation of CaO. Although the conversion of CH4 in the conventional steam
methane reforming process decreases with the increase in pressure, the removal of CaO
during steam methane reforming reduces this effect and results in almost a constant
amount of CH4 conversion with the increase in pressure. The effect of calcination
conditions on the extent of CH4 reforming was determined. The reactivity of the
sorbent is found to decrease over multiple cycles due to calcination in both pure
nitrogen and in a mixture of steam and CO2. This reduction in reactivity of the sorbent
results in a decrease in both CH4 conversion and H2 purity. In order to improve the
recyclability of the sorbent over multiple cycles, sorbent reactivation step by hydration
is found to be effective. By the addition of the hydration step during every cycle, the
extent of sorbent sintering is reduced. This reactivation will aid in the production of a
constant purity of H2 during the steam reforming of CH4. The CLP for reforming of
hydrocarbons can be directly applied to the production of H2 and electricity from
natural gas. This integration would benefit the steam reforming process by providing a
method of internal heat integration. In the CLP, although the endothermic calciner is
operated at a high temperatures(800 -1000 ºC) similar to the steam methane reforming
it is at atmospheric pressure while the heat neutral reformer is at a high pressure and a
relatively low temperature of 650 ºC. The energy required for the reformer in the
conventional process is supplied to the calciner in the CLP and an additional benefit of
producing a sequestration ready CO2 stream is obtained by integrating steam reforming
with the CLP. The CLP can also be integrated in a CLT plant for the conversion of F-T
offgases to H2 with CO2 capture.
294
Reaction
Extent of Reaction
@ 650 ºC and 15
atms
Methane Reforming
CH4 + H2O = CO + 3H2
Ethane Reforming
C2H6 + 2H2O = 2CO + 5H2
Water Gas Shift Reaction
CO +H2O = CO2 + H2
Carbonation Reaction
CaO + CO2 = CaCO3
70.5%
99.997%
99.31%
99.38%
Table 7.1: Thermodynamic extent of the various reactions occurring in the carbonator
295
296
Stream
1
25.0
T ºC
1.0
P bar
Mass Flow kg/hr
CH4
0.0
C2H6
0.0
N2
0.0
H2O
233236.3
CO
0.0
H2
0.0
CO2
0.0
O2
0.0
CaCO3
0.0
CaO
0.0
Mole Flow kmol/hr
CH4
0.0
C2H6
0.0
N2
0.0
H2O
12946.6
CO
0.0
H2
0.0
CO2
0.0
O2
0.0
CaCO3
0.0
CaO
0.0
2
25.2
15.2
3
650.0
15.2
4
25.0
41.4
5
-27.0
15.2
6
650.0
15.2
7
650.0
15.2
8
650.0
15.2
9
650.0
15.2
10
25.0
15.2
0.0
0.0
0.0
0.0
0.0
0.0
233236.3 233236.3
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
62309.8
6488.3
6044.7
0.0
0.0
0.0
0.0
0.0
0.0
0.0
62309.8
6488.3
6044.7
0.0
0.0
0.0
0.0
0.0
0.0
0.0
62309.8
6488.3
6044.7
0.0
0.0
0.0
0.0
0.0
0.0
0.0
18373.1
0.3
6044.7
119365.8
553.1
25088.7
944.9
0.0
313177.2
187536.7
18373.1
0.3
6044.7
119365.8
553.1
25088.7
944.9
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
22579.8
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
22579.8
0.0
0.0
0.0
0.0
0.0
0.0
0.0
12946.6
0.0
0.0
0.0
0.0
0.0
0.0
3884.0
215.8
215.8
0.0
0.0
0.0
0.0
0.0
0.0
0.0
3884.0
215.8
215.8
0.0
0.0
0.0
0.0
0.0
0.0
0.0
3884.0
215.8
215.8
0.0
0.0
0.0
0.0
0.0
0.0
0.0
1145.3
0.0
215.8
6625.8
19.7
12445.5
21.5
0.0
3129.0
3344.2
1145.3
0.0
215.8
6625.8
19.7
12445.5
21.5
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
11201.0 11201.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
Continued
0.0
0.0
0.0
12946.6
0.0
0.0
0.0
0.0
0.0
0.0
Table 7.2: Stream data for the integration of the CLP in a steam methane reforming process
296
297
Table 7.2 continued
Stream
11
900.0
T ºC
1.0
P bar
Mass Flow kg/hr
CH4
0.0
C2H6
0.0
N2
0.0
H2O
0.0
CO
0.0
H2
0.0
CO2
0.0
O2
0.0
CaCO3
0.0
CaO
363005.3
Mole Flow kmol/hr
CH4
0.0
C2H6
0.0
N2
0.0
H2O
0.0
CO
0.0
H2
0.0
CO2
0.0
O2
0.0
CaCO3
0.0
CaO
6473.3
12.0
650
1.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
313177.2
187536.7
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
3129.0
3344.2
13.0
650
1.0
14.0
900.0
1.0
15.0
25.0
1.0
16.0
900.0
1.0
17.0
900.0
1.0
18.0
900.0
1.0
19.0
25.0
1.0
20.0
25.0
151.7
18373.1 18373.1
0.0
0.0
0.0
0.0
0.0
0.0
0.3
0.3
0.0
0.0
0.0
0.0
0.0
0.0
6044.7
6044.7
0.0
0.0
6044.7
6044.7
6044.7
6044.7
119365.8 119365.8
0.0
0.0
183051.4 183051.4
0.0
0.0
553.1
553.1
0.0
0.0
0.0
0.0
0.0
0.0
2508.9
2508.9
0.0
0.0
0.0
0.0
0.0
0.0
944.9
944.9
0.0
0.0
189925.7 189925.7 189925.7 189925.7
0.0
0.0
93756.5 93756.5
233.9
233.9
233.9
233.9
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
363005.3
0.0
0.0
0.0
1145.3
0.0
215.8
6625.8
19.7
1244.6
21.5
0.0
0.0
0.0
1145.3
0.0
215.8
6625.8
19.7
1244.6
21.5
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
0.0
2930.0
0.0
0.0
297
0.0
0.0
0.0
0.0
0.0
0.0
0.0
2930.0
0.0
0.0
0.0
0.0
215.8
10160.9
0.0
0.0
4315.5
7.3
0.0
6473.3
0.0
0.0
215.8
10160.9
0.0
0.0
4315.5
7.3
0.0
0.0
0.0
0.0
215.8
0.0
0.0
0.0
4315.5
7.3
0.0
0.0
0.0
0.0
215.8
0.0
0.0
0.0
4315.5
7.3
0.0
0.0
Inputs
Outputs
Natural Gas feed rate
Steam boiler duty
NG heater duty
O2 preheater duty
Tail gas preheater duty
Pumps
Net energy Input
Mass Flow
Kg/hr
Heat
MWth
74842.7401
1054.41
250.85
42.41
23.39
28.44
H2 production rate
22579.8107
H2 Reactor heat duty
Calciner heat duty
Hot H2 cooler duty
Hot CO2 cooler duty
NG turbine
Net energy out
Cold gas efficiency
(HHV)
Total
Net thermal heat available
Net Power available
CO2 Compression
Export power from plant
1399.51
Power
MWe
0.17
0.17
889.29
10.68
159.84
57.14
280.57
1397.79
2.45
2.45
0.84
163.41
65.36
2.28
19.94
47.70
Table 7.3: Energy balance for the production of H2 and electricity from natural gas
using the CLP.
298
Mole
flow
kmol/hr
Steam
Boiler
Duty
12946.6
T IN
T out
(C)
(C)
For H2O Cp
Values
25
650
Cp (l)
kJ/kg-K
25-199C
5.3
Cp (v)
kJ/kgK
199650 ºC
2.01
Latent
heat
Q
(MWth)
kJ/kg
at 199C
2040
250.85
Table 7.4: Heat required for the production of steam at 650 ºC from water at 15 atms
299
Natural Gas
Preheating
CH4
C2H6
N2
Total
Mass
flow
kg/hr
Mole
flow
kmol/hr
62244.49
6484.08
6041.74
74770.33
3883.0
215.8
215.8
4314.6
T IN
(C)
T out
(C)
Energy
Reqd.
kJ/kmole
-27
-27
-27
650
650
650
34912.01
58703.02
20649.53
Table 7.5: Heat required for preheating the natural gas at 15 atms to 650 ºC
300
Q
(MWth)
37.66
3.52
1.24
42.41
Mass
flow
kg/hr
Solid from
calciner
CaO
Mole
flow
kmol/hr
T IN
(C)
T out
(C)
Cp
kJ/kg-K
363009.16 6473.29
900
650
0.96
Q
(MWth)
-24.20
Table 7.6: Heat released by the solids from the calciner in the H2 production reactor.
301
Hydrogen Production
Reactor
Reforming of CH4
Reforming of C2H6
Water gas shift reaction
Carbonation
Sensible heat from hot CaO
solids
Net Reactor heat duty
Heat of
Reaction
KJ/Mole
Moles
Reacted
kmoles/hr
Heat
Duty
MWth
Heat duty
224
376
-35.7
-170.5
2737.74
215.771
3147.539
3128.07
170.35
22.54
-31.21
-148.15
Endothermic
Endothermic
Exothermic
Exothermic
-24.20
-10.68
Exothermic
Exothermic
From Table 7.5
Table 7.7: Heat generated from the H2 production reactor
302
H2
outlet
cooler
Mass flow
kg/hr
Mole flow
kmol/hr
T IN
(C)
T out
(C)
Energy
Reqd.
kJ/kmole
22581.15
11200.97
650
25
18366.03
Q
(MWth)
-57.14
Table 7.8: Heat released on cooled the H2 from 650 ºC to ambient temperature
303
Tail Gas
Preheating
CH4
C2H6
N2
H2O
CO
H2
CO2
Total
Mass flow Mole flow
kg/hr
kmol/hr
T IN
(C)
T out
(C)
Energy
Reqd.
KJ/kmole
18358.52
0.27
6041.75
119264.6
553
2489.1
944.63
147651.9
650
650
650
650
650
650
650
900
900
900
900
900
900
900
18640.69
31353.93
8233.88
10463.18
8352.67
7599.04
13675.83
1145.3
0.0
215.8
6625.8
19.8
1244.6
21.5
9272.6
Q
(MWth)
5.93
0.00
0.49
19.26
0.05
2.63
0.08
28.44
Table 7.9: Heat required for preheating the PSA tail gas from 650 ºC to 900 ºC
304
Mass flow
kg/hr
Oxygen
Preheating
46880
Mole flow T IN
kmol/hr
(C)
2930
25
T out
(C)
Energy
Reqd.
kJ/kmole
Q
(MWth)
900
28740.76
23.39
Table 7.10: Heat required for preheating the oxygen from ambient temperature to 900
ºC
305
Mass
flow
kg/hr
Solids to calciner
CaCO3
312900
CaO
187615.23
500515.23
Total
Mole
flow
kmol/hr
3129
3344.3
6473.3
T IN
(C)
T out
(C)
Cp
kJ/kg-K
650
650
900
900
1.23
0.96
Q
(MWth)
26.72
12.50
39.23
Table 7.11: Heat absorbed by the solids from the H2 production reactor in the calciner
306
Calcination Reactor
Calcination of CaCO3
Heat needed to heat the
solids
Tail gas combustion
Combustion of CH4
Combustion of C2H6
Combustion of H2
Combustion of CO
Net Reactor heat duty
Heat of
Reaction
KJ/Mole
Moles
Reacted
kmoles/hr
Heat
Duty
MWth
Heat duty
165.5
3129.043
143.85
Endothermic
From Table 7.10
-802.5
-1429
-249
-282
Table 7.12: Heat released from the calciner
307
1145.256
0.009
1244.552
19.747
39.234557 Endothermic
-255.30
~0.00
-86.08
-1.55
-159.84
Exothermic
Exothermic
Exothermic
Exothermic
Exothermic
Mass flow
kg/hr
CO2
cooler
N2
H2O
CO2
O2
Total
6041.747
183048.6
189948.1
233.8909
379272.3
Mole
flow
kmol/hr
215.77
10160.89
4315.53
7.30
T
T
IN
out
(C) (C)
For H2O
Cp Values
900
25
900
25
900
25
900
25
Latent
ht
kJ/kg
Cp (l)
Cp (v)
kJ/kg-K
kJ/kg-K
25-100
ºC
100-900 ºC at 100 ºC
1.2
4.85
2.2
2259.36
1.21
1.35
Table 7.13: Heat released from cooling the CO2 from 900 ºC to ambient temperature
308
308
Q
(MWth)
1.76
222.87
55.86
0.08
-280.57
Sorbent Makeup
Sorbent Purge
Reaction
Regeneration
H2
Pure CO 2 gas
Integrated
reactor
Net Heat
Output
Heat
Input
Hydrocarbon Feed
Dehydration :
Reforming :
WGSR
:
CO2 removal :
Sulfur
:
Halide
:
Ca(OH) 2 Æ CaO + H 2O
CxHy +H2O Æ CO + H 2
CO + H2O Æ CO2 + H2
CaO + CO2 Æ CaCO3
CaO + H 2 S Æ CaS + H 2O
CaO + 2HX ÆCaX 2 + H2O
Calciner
Calcination: CaCO3 Æ CaO + CO2
Reactivation
Heat
Output
Hydrator
H 2O
Hydration : CaO + H 2O Æ Ca(OH) 2
Figure 7.1: Schematic of the CLP for the conversion of hydrocarbons to H2
309
K, Equilibrium Constant
800
Reforming
Reforming + WGS
Reforming + WGS + Carbonation
600
400
200
0
200
400
600
800
Temperature (C)
Figure 7.2: Thermodynamic data illustrating the equilibrium constants of the steam
reforming of CH4, water gas shift and carbonation reaction
310
Thermocouple
And
Pressure Guage
Steam Generator
Steam &
Gas Mixture
Catalyst
Powder
+ Sorbent
Water In
MFC
Back Pressure
Regulator
Analyzers (CO,
CO2, H2, H2S)
MFC
N2
Heated Steel
Tube Reactor
Hydrocarbon
Analyzer
Gas
Gas
Mixture
Mixture
MFC
CH4
Water Syringe
Pump
Heat
Exchanger
Water Trap
Figure 7.3: Simplified schematic of the bench scale experimental setup
311
MFC
C2H6
C3H8
90
S:C = 3
S:C = 5
H2 Purity (%)
85
80
75
70
600
650
700
750
800
850
900
Temperature (C)
(a)
30
CH4 (3:1)
CH4 (5:1)
CO (3:1)
CO (5:1)
CO2 (3:1)
CO2 (5:1)
Gas Composition (%)
25
20
15
10
5
0
600
650
700
750
800
850
900
Temperature (C)
(b)
Figure 7.4: Effect of temperature and S:C ratio on (a)H2 purity and (b) the amount of
CO, CO2 and CH4 remaining in the product gas for the steam methane
reforming reaction in the presence of Ni-based catalyst ( P = 1 atm)
312
100
Gas composition (%)
H2
CH4
80
CO
CO2
60
40
20
0
0
500
1000
1500
2000
2500
3000
Time (sec)
Figure 7.5: Breakthrough curve in the composition of the product gases obtained
during the simultaneous reforming, water gas shift and carbonation
reaction. (T = 650 ºC, P = 1 atm)
313
1
0.9
Methane Conversion
0.8
0.7
0.6
0.5
0.4
0.3
0.2
0.1
0
560
1060
1560
2060
2560
3060
3560
4060
4560
5060
Time (sec)
Figure 7.6: CH4 conversion obtained during the simultaneous reforming, water gas
shift and carbonation reaction. (T = 650 ºC, P = 1 atm)
314
(a)
(b)
Figure 7.7: Effect of temperature and S:C ratio on (a) H2 purity (b) conversion of CH4
(P = 1atm)
315
(a)
(b)
Figure 7.8: Effect of temperature and S:C ratio on the amount of (a) CO and (b) CO2
remaining in the product gas for H2 production from methane with/without
sorbent. ( P = 1 atm)
316
100
H2 Purity (%)
90
80
1 atm
3 atm
4.5 atm
11 atm
70
60
0
500
1000
1500
2000
Time (sec)
(a)
CH4 in the Product Stream (%)
40
1 atm
3 atm
4.5 atm
11 atm
30
20
10
0
0
500
1000
1500
2000
2500
3000
Time (sec)
(b)
Figure 7.9: Effect of pressure on (a) H2 purity and (b) CH4 concentration in the product
stream. (T = 650 ºC, S:C ratio = 3)
317
CO2 in the Product Stream (%)
6
1 atm
3 atm
4.5 atm
5
4
3
2
1
0
0
500
1000
1500
2000
2500
3000
Time (sec)
(a)
CO in the Product Stream (%)
40
1 atm
3 atm
4.5 atm
11 atm
30
20
10
0
0
500
1000
1500
2000
2500
3000
Time (sec)
(b)
Figure 7.10: Effect of pressure on (a) CO2 and (b) CO concentration in the product
stream. (T = 650 ºC, S:C ratio = 3)
318
Gas Composition in the Product Stream (%)
20
CH4-Post
18
16
14
CH4
12
10
8
CO-Post
6
CO
CH4-Pre
4
2
CO2
0
CO2-Post
CO-Pre
CO2-Pre
0
2
4
6
8
10
12
Pressure (atms)
Figure 7.11: Effect of pressure on the pre-breakthrough and post-breakthrough
concentration of CH4, CO and CO2 in the product stream. (T = 650 ºC, S:C
ratio = 3)
319
100
Cycle 1
Cycle 2
Cycle 3
Cycle 4
H2 Purity (%)
95
90
85
80
75
0
500
1000
1500
2000
2500
Time (sec)
CH4 in the Product Stream (%)
(a)
8
6
4
Cycle 1
Cycle 2
Cycle 3
Cycle 4
2
0
0
500
1000
1500
2000
2500
Time (sec)
(b)
Figure 7.12: Effect of calcination conditions on (a) H2 purity and (b) CH4 composition
in the product gas for cycles 1,2,3 and 4. [(Reforming reaction conditions
:T= 650 ºC, P = 1 atm, S:C ratio = 3) (Calcination conditions: sorbent for
cycles 1, 2 and 3 are calcined in pure N2 at 950 and the sorbent for cycle 4
is calcined in a 50:50 CO2/H2O atmosphere at 950 ºC.)]
320
100
Cycle 1
Cycle 2
Cycle 3
Cycle 4
H2 Purity (%)
95
90
85
80
75
70
0
500
1000
1500
2000
2500
3000
Time (sec)
(a)
Continued
Figure 7.13: Effect of hydration on (a) H2 purity and (b) CH4 composition in the
product gas for cycles 1, 2, 3 and 4. [(Reforming reaction conditions :T=
650 ºC, P = 1 atm, S:C ratio = 3) (Calcination conditions: sorbent for
cycles 1, 2, 3 and 4 are calcined in pure N2, T = 950, P = 1 atm)(Hydration
conditions: hydration of calcined sorbent from the 3rd cycle in a 80:20
H2O/N2 atmosphere, T = 600, P = 11 atm)]
321
Table 7.13 continued
CH4 in the Product Stream (%)
8
6
4
Cycle 1
Cycle 2
Cycle 3
Cycle 4
2
0
0
200
400
600
800
Time (sec)
(b)
322
1000
1200
1400
100
Cycle 1
Cycle 2
Cycle 3
H2 Purity (%)
95
90
85
80
75
0
500
1000
1500
2000
2500
3000
Time (sec)
Figure 7.14: Effect of hydration on H2 purity for cycles 1,2,3 and 4. [(Reforming
reaction Conditions: T= 650 ºC, P = 1 atm, S:C ratio = 3) (Calcination
conditions: sorbent for cycles 1, 2, 3 and 4 are calcined in pure N2, T =
950, P = 1 atm)(Hydration conditions: hydration for cycles 1, 2, 3 and 4 in
a 80:20 H2O/N2 atmosphere, T = 600, P = 11 atm)]
323
100
Cycle 1
Cycle 2
Cycle 3
Cycle 4
H2 Purity (%)
95
90
85
80
75
70
0
500
1000
1500
2000
2500
3000
Time (sec)
(a)
Continued
Figure 7.15: Effect of hydration on (a) H2 purity and (b) CH4 content in the product gas
for cycles 1,2,3 and 4. [(Reforming reaction Conditions: T= 650 ºC, P = 1
atm, S:C ratio = 3) (Calcination conditions: sorbent for cycles 1, 2, 3 and 4
are calcined in a 50:50 CO2/H2O atmosphere, T = 950, P = 1 atm)
(Hydration conditions: hydration for cycles 1, 2, 3 and 4 in a 80:20 H2O/N2
atmosphere, T = 600, P = 11 atm)]
324
Figure 7.15 continued
Cycle 1
Cycle 2
Cycle 3
Cycle 4
CH4 in the Product Stream (%)
14
12
10
8
6
4
2
0
0
500
1000
1500
2000
Time (sec)
(b)
325
2500
3000
Steam
Electricity
Natural Integrated Reforming, WGSR,
Gas
CO2 and Sulfur Capture
Gas Solid
Separation
Hydrogen
Chemicals and
Liquid Fuels
CaO
Calciner
Spent Sorbent
CO2
CO2 Compression
and Sequestration
Figure 7.16: Integration of the CLP in a natural gas reforming system
326
250.85 MWth
0.17 MWe
Water
1
2
Pump
57.14 MWth
Steam
Generator
10.68 MWth
8
42.41 MWth
2.45 MWe
Natural Gas 4
from Pipeline
3
5
6
1054.41 MWth
Hydrogen
Production
Reactor
7
9
PSA
Pure
Hydrogen
10
889.29 MWth
H2 Cooling
Cyclone
13
NG Preheating
159.84 MWth 12
327
11
Tailgas
Preheating
28.4 MWth
14
Cyclone
17
23.39 MWth
Calciner
16
15
O2 Preheating
19.94 MWe
18
20
19
280.57 MWth
CO2 Compression
CO2 Cooling
Figure 7.17: Detailed schematic for H2 production from natural gas
327
CO2 to
Sequestration
Oxygen
328
Figure 7.18: Conventional CTL plant
328
Flue Gas
Conventional
Gas
Turbine
COS
Hydrolysis
Sulfur
and CO2
Reformer
Gases
Selexol
H2
Recovery
WGSR
CLP -1
Gasifier
H2
Sulfur
Removal
Separator
/Upgrader
FT Reactor
H2
CLP
FT Reactor
H2
Gases
Separator
/Upgrader
Fuel
CLP -2
Gases
CLP
FT Reactor
H2
Recovery
Separator
/Upgrader
Fuel
H2
Figure 7.19: Integration of the CLP in a CTL plant in two configurations
329
Fuel
Raw Syngas
H2/CO = 0.5
Steam
HC
Reforming
+ WGSR
+CO2
removal
Mercury
Removal
Hg
Coal
Fly
Ash
Pretreatment
Sulfur
Byproduct
Reactivator
Steam
Calciner
High
Pressure
BFW
O2
CO2
330
Gasifier
Generator
C1-C4
Slag
Unconverted
Syngas
Steam
Air
Product
Separation
N2
Separator II
Gasoline
C5 – C14
N2
JP-8 Fuel
Water
Warm
Clean Syngas
F-T Reactor
H2/CO = 2:1
Hydrocraking
Hydrotreatment
C15 and above
Figure 7.20: Integration of the CLP in a CTL plant – configuration 1
330
H2
High
Pressure
Calciner
To Sequestration
CO2
One Step Process
•
WGSR
•
HC Reforming
•
CO2
•
Sulfur
Raw Syngas
H2/CO = 0.5
Steam
Coal
WGSR+ HC Reforming +
sulfur
Pretreatment
BFW
Candle
CaCO3 to Filter
Fly
Calciner
Steam
Ash
O2
331
Gasifier
Generator
C1-C4
Unconverted
Syngas
Steam
Slag
Air
Product
Separation
N2
Separator II
Gasoline
C5 – C14
N2
JP-8 Fuel
Water
Warm
Clean Syngas
F-T Reactor
H2/CO = 2:1
Hydrocraking
Hydrotreatment
C15 and above
Figure 7.21: Integration of the CLP in a CTL plant – configuration 2
331
H2
CHAPTER 8
SUBPILOT SCALE TESTING AND RECOMMENDATIONS FOR FUTURE
WORK
8.1 INTRODUCTION
The CLP has been studied in the lab and bench scale for H2 production from
syngas and hydrocarbons. The process is being scaled up to a 25 KWth subpilot unit
demonstration at the Ohio State University. The design for the subpilot scale unit is
based on the thermodynamic, kinetic and sorbent reactivity studies detailed in the
previous chapters. Based on the design, cold flow models were constructed and sorbent
flow testing was conducted. The subpilot scale unit consists of the gas delivery system
and the gas regulating panel, steam generation system, sorbent feeding system, H2
production or carbonation reactor, gas cooling and particle capture system and gas
analysis. The unit is designed to operate at a temperature of up to 900 ºC and pressure
of 4.5 atms. In this chapter, the results from cold flow testing and the design of the
subpilot scale unit, currently under construction are discussed.
332
8.2 COLD FLOW TESTING
In the CLP, the carbonation reactor, calciner and hydrator are fluidized or
entrained bed reactors and hence an understanding of the flow characteristics of the
sorbent through these reactors is very critical to the process. Since CaO and Ca(OH)2
are cohesive particles that are difficult to fluidize, cold flow testing was conducted
before construction of the subpilot scale unit. As discussed in Chapter 3, hydration and
the subsequent dehydration of the sorbent has been found to produce CaO with a
particle size of 2-20 micron(D50). Hence the active sorbent that captures CO2, sulfur
and halides in the carbonation reactor has a D50 of 2- 20 microns.
A cold model with a diameter of 4 inches was constructed for the carbonation
reactor with a sorbent feeding system, flow controllers for air and a particulate filter at
the exit of the reactor. Air was used as the fluidization gas. Ca(OH)2 sorbent with a
D50 of 2-20 micron was used as the active sorbent. Since Ca(OH)2 is difficult to
fluidize, it was mixed with a fluidizing aid, ground lime which has a D50 of 600
microns. Other powders like sand and other metal oxides could also be used as
fluidizing aids. The gas velocity in the carbonation reactor was designed such that the
active Ca(OH)2 sorbent is entrained while the fluidizing aid is in the turbulent regime.
Hence the gas velocity in the carbonation reaction should be above the entrainment
velocity of the Ca(OH)2 particles and between the critical and entrainment velocities of
the fluidizing aid.
333
Cold model tests were conducted to determine the critical velocity for the
fluidizing aid particles in the 4 inch cold model of the carbonation reactor. Critical
velocity is the minimum velocity for turbulent fluidization and is determined from the
standard deviation of pressure change across the bed. The standard deviation of
pressure change is a maximum at the critical velocity of the particles. Pressure taps
were placed along with 2 micron inline filter along the height of the column and were
connected to a U-tube manometer to measure the pressure change. The manometer was
connected to a transducer and the voltage signal generated by the transducer was
recorded using a data logging system. The collapsed bed height of the sorbent was ~10
cms. The air flowrate was varied between 1-1600 scfh and the change in pressure drop
across the bed was determined at each flowrate. Figure 8.1 shows the standard
deviation of pressure change(STD(Pa)) for different gas velocities(Ug). The standard
deviation is a maximum at velocities greater than 1.3 m/s and hence turbulent
fluidization of the fluidizing aid would occur above a velocity of 1.3 m/s. Since the
active Ca(OH)2 particles are entrained at velocities above 1.3 m/s, the carbonation
reactor can be operated at a velocity of 1.3m/s and greater.
After determining the suitable gas velocity for the carbonation reactor, a
continuous fluidization test was conducted with a mixture of 10% fluidizing aid
(ground lime power with a D50 of 500 to 600 microns) and 90% active sorbent
(Ca(OH)2 with a D50 of 2- 20 microns). The mixture of sorbent was fed to the
carbonation reactor from a sorbent storage hopper by a screw feeder. The air flow rate
334
was set at 1400 SCFH. When the sorbent was conveyed at a flow rate of 15 g/min, ~
80% of the total sorbent that was conveyed into the reactor was entrained at the end of
2 hours. At a higher sorbent flowrate of 60g/min, ~70% of the sorbent was entrained.
From these test it can be seen that not all the active Ca(OH)2 sorbent was entrained by
the gas. At the end of the tests, agglomerates were observed in the bed material which
might have been formed by the cohesive Ca(OH)2 sorbent at room temperature. The
moisture in the fluidizing air might also cause the aggregates to form in the sorbent
mixture. Hence fluidization test need to be conducted at the reaction temperature of
600 ºC in the subpilot scale with air prior to testing the system with the simulated
syngas stream.
A cold model of the entire CLP with the carbonation reactor, calciner and
hydrator was also constructed as shown in Figure 8.2(a). Figure 8.2(b) is a picture of
the cold model unit. The carbonation reactor and hydrator were fluidized beds while
the calciner was a rotary bed. As shown in Figure 8.2(a), the cold flow model consists
of a riser (H2 production or carbonation reactor), cyclones, a rotary calciner, a U-valve
and a hydrator. Air is introduced into the riser to carry the sorbent through the riser
reactor. At the top of the riser, two cyclones separate the solids from the air. The
recovered solids enter the rotary calciner via a U-valve. At the outlet of the calciner,
the sorbent enters the hydrator. Cold flow tests with air and a mixture of active
Ca(OH)2 sorbent and fluidization aid have shown that the sorbent can be fluidized and
made to flow smoothly in the continuous reactor system. Figure 8.3 is a picture of the
335
cold model for the hydrator. Air was used instead of steam to fluidize a mixture of CaO
and Ca(OH)2. The sorbent mixture fluidizes well with minimum gas channeling and no
slugging was observed.
8.3 DESIGN OF THE SUBPILOT SCALE UNIT
The 25 KWth sub-pilot scale reactor system being constructed at OSU is shown
in Figure 8.4. The calcium sorbent will be continuously fed using a sorbent hopper and
a motor-driven screw feeder. This sorbent is entrained using the reactant gas mixture
entering from the bottom. The product gas is analyzed using a micro GC. The subpilot
scale unit consists of the gas delivery system and the gas regulating panel, steam
generation system, sorbent feeding system, H2 production or carbonation reactor, gas
cooling and particle capture system and gas analysis.
The gas delivered to the reactor is a simulated syngas mixture. Gas cylinders
are used as the source of H2, CO, CO2 and N2 (carrier gas). To ensure a continuous
supply of gas to the reactor during the test a change over system is installed for each
gas. 2 cylinders are connected to the changeover system and when one cylinder
becomes empty, the changeover system automatically switches over to the other
cylinder. An audible alarm is used to alert the operator to replace the empty cylinder
while the other cylinder is in use. In this manner cylinders can be changed even when
the system is in operation. The switch over valves and the cylinders are placed in gas
cabinets that are provided with adequate ventilation. Gas leak alarms are also placed in
336
the gas delivery area for H2 and CO2. Safety features like flash arrestors, check valves
to prevent back flow of the gases, automatic shutoff valves to stop gas flow in the event
of a leak, pressure relief valve to prevent an increase in pressure and flow limit shut off
valves to prevent an increase in the gas flow rates beyond the preset maximum value
are included in the gas delivery system.
The gases from the gas delivery room are sent to a gas mixing panel consisting
of digital mass flow controllers. Safety features like flash arrestors, check valves,
maximum flow limit shut-off valves and pneumatically controlled diaphragm valves
have been incorporated into the gas delivery system. The mass flow controllers and the
diaphragm valves are controlled using a computer interface. A nitrogen line is provided
to flush the system and a line is provided from the gas panel to the vent for safety. In
addition to these safety features, the gas panel is connected to safety features in the
other sections of the subpilot scale unit. The flow of gas through the gas panel will be
automatically shut off if the temperature exceeds a preset value in the reactor, the
pressure increases beyond 4.5 atms or the pressure blower in the vent malfunctions.
The gas mixture from the gas panel is passed through a preheater section consisting of
a helical coil surrounded by a ceramic heater before being sent to the carbonation
reactor. At the outlet of the heater, the gas composition and temperature of the gas
mixture is monitored. H2 and CO gas alarms are also connected to the gas panel to shut
off the gas flow in the case of a gas leak.
337
Steam is generated separately using a high precision water pump which
conveys water to a heated section similar to the gas preheater. The pressure and
temperature of the steam is constantly monitored. The steam and preheated gas mixture
are mixed just before they enter the reactor.
The sorbent is stored in a hopper and is metered by a screw feeder into the
reactor as shown in Figure 8.5. An airlock and two valves are provided in the sorbent
transport line to prevent gas flow from the reactor to the sorbent hopper.
The reactor is divided into 4 flanged parts and is surrounded by ceramic heaters
to maintain the desired temperature during operation. Provision has been made for 8
temperature ports, 4 gas sampling ports and 4 pressure ports throughout the length of
the reactor. The gas sampling ports and pressure ports are protected by in-line filters to
avoid the solids from entering the lines. The gas sampling ports are connected to the
GC for continuous gas analysis. Pressure transducers are connected to the pressure
ports and thermocouples are inserted into the temperature measurement ports. The
temperature ports of the reactor are integrated with the reactor heaters through a
feedback loop process control system. Gas leak detectors have been placed near the
reactor to alert the operators.
The H2-rich product gas containing the calcium sorbent, from the outlet of the
reactor, is then passed through a high temperature particle capture device followed by a
water cooled heat exchanger shown in Figure 8.6. Before the gas enters the baghouse,
338
it is further diluted to reduce the temperature and prevent the formation of an explosive
mixture. From the baghouse, the gas is vented out of the facility. The continuous
production of H2 by the CLP will be tested in this subpilot scale facility.
Figure 8.7 is a schematic of the reactor system along with the support structure.
The support structure has two platform levels and the sorbent is fed at the top level.
8.4 CONCLUSIONS
The CLP has been shown to enhance H2 yield and purity from syngas and
hydrocarbons in lab and bench scale tests. The process is being scaled up to a 25 KWth
subpilot unit demonstration at the Ohio State University. Cold flow test were
conducted to determine the flow characteristics of the sorbent and parameters like
fluidization velocity. The subpilot scale unit design is based on the thermodynamic,
kinetic and sorbent reactivity studies and cold flow tests. This unit will be used to
conduct continuous testing for the production of H2 from a simulated syngas stream
and a mixture of hydrocarbons.
8.5 RECOMMENDATIONS FOR FUTURE WORK
The CLP has been successfully demonstrated at the lab and bench scale for
carbon capture during the production of H2 and electricity. A subpilot scale unit is
being constructed for continuous testing of the concept at a higher scale. System
339
analysis and economic analysis have shown that the CLP has good potential to improve
the efficiency and economics of H2 production.
After construction of the subpilot scale unit, shake down and start up testing
will be conducted. Following this, leak testing and fluidization tests with hot air will be
conducted to ensure safety and to determine the temperature and pressure profile and
the flow characteristics of the sorbent at 600 ºC in the system. H2 production tests will
then be conducted with a simulated syngas stream to determine and optimize important
process parameters like the residence time and Ca:C ratios. The effect of process
parameters including temperature, pressure and S:C ratio will also be determined and a
comparison will be made with the data obtained at the bench scale. On successful
testing of the carbonation reactor for H2 production, a fluidized bed calciner and
hydrator should be integrated in the subpilot scale unit to test continuous sorbent flow
through the process during H2 production. The purity of H2 produced from the
carbonator and that of the CO2 from the calciner should be monitored over long range
tests lasting from 1 to 5 days. The performance of particle capture devices like high
temperature cyclones and metallic filter for the micron sized calcium sorbent should be
evaluated. At the end of this testing the subpilot scale unit should be moved to a
location where a slip stream testing can be conducted from a real gasifier. The effect of
other impurities in the syngas and flyash will be determined from the slip stream
testing. On successful testing of the CLP concept in the 25 KW subpilot scale facility,
340
the process should be scale up to 250 to 500 KW to move the process to
commercialization.
System analysis studies and economic analysis should be updated during every
scaleup testing to confirm the feasibility of the process with new information obtained
from the testing. Sensitivity analysis on different coal and coal surfur contents,
limestones from different locations, and process operation parameters should be
conducted to determine the advantages that certain parameters or location of the plant
could offer. Since the CLP produces a lot of high quality heat in the carbonation
reactor and hydrator, reactor design and effective methods of heat extraction should be
evaluated so that this heat can be used to produce additional electricity. Complete life
cycle analysis should be conducted to determine the impact of this process on the
environment and ecology and these lifecycle and economic studies should be used to
guide the development of the process.
Scientific studies to understand the mechanism of sintering and deactivation of
the sorbent and reactivation by hydration should be conducted. These studies will help
in further improving the reactivity of the sorbent. With complete understanding of the
sorbent performance, scale up characteristics, system analysis, economics and life
cycle analysis the CLP can be brought closer to commercialization for production of
H2, electricity, chemicals and liquid fuels.
341
40
STD (Pa)
30
20
10
0
0
0.2
0.4
0.6
0.8
1
Ug (m/s)
Figure 8.1: Standard deviation of pressure in the fluidized bed
342
1.2
1.4
1.6
Figure 8.2 (a): Schematic diagram of the cold flow model for the CLP
343
Figure 8.2 (b): Snapshot of the cold flow model for the CLP
344
Figure 8.3: Cold flow model for the hydrator
345
Sorbent Hopper
+ Motor driven
screw feeder
Particle
Capture
Device
Baghouse
To vent
Water-cooled
Heat
Exchanger
TI
TI
N2
H2
CO
TI
PI
CO2
TI
F
TI
Computer
Setup
PI
F
F
TI
F
346
TI
PI
GC
TI
Gas Delivery
System
PI
Computer
TI
PI
Steam Generator
PI
TI
TI
PI
Gas Preheater
Figure 8.4: Schematic of the subpilot scale unit being constructed at OSU for testing the Calcium Looping Concept for H2
production
346
Figure 8.5: Sorbent hopper and screw feeder
347
348
Figure 8.6: Water cooled heat exchanger
348
349
Figure 8.7: Schematic of the subpilot scale unit with the support structure
349
APPENDIX - A
LCA ANALYSIS - COMPARISON OF THE CONVENTIONAL COAL TO H2
PROCESS WITH THE CLP PROCESS
The inputs and outputs have been quantified in physical and monetary terms for
the conventional and novel process for facilities coproducing H2 and electricity from
coal. The conventional process includes coal gasification system followed by the sour
or the raw gas shift with CO2 capture using a dual stage selexol unit. The novel process
consists of coal gasification system integrated with the CLP. Since both electricity and
H2 are products of these plants in order to make a fair comparison between the
conventional and novel process, the same coal feed rate and gasifier type have been
assumed. In addition, the same assumptions have been made for all the unit operations
that are common to the two processes to compare the two processes on the same basis.
Both plants produce H2 with a purity of >99.9% suitable for applications like hydro
treating, chemical and fuel synthesis and for fuel cells. For the conventional case all the
pertinent information for the plant details and cost estimates have been obtained from
DOE/NETL reports for CO2 capture ready gasification plants (DOE, 2007). The mass
and energy balance for the CLP has been obtained from ASPEN simulations conducted
at thermodynamic equilibrium conditions.
350
Conventional coal to hydrogen process
Basis for analysis:
1) Analysis has been conducted for a production rate of 547 tonnes/day of H2 and
30MW of electricity
2) Bituminous Illinois #6 coal has been used at a feed rate of 5891 tonnes/day.
3) A GE oxygen blown gasifier with radiant cooling is used to generate the syngas
at a temperature of 2500F and pressure of 65.7 atms.
4) Sour shift is conducted in 2 sets of reactors in series with a STC and a
minimum S:C ratio of 2
5) Mercury removal is conducted using an activated carbon bed.
6) A dual stage selexol process is used for capture of H2S and CO2 in series
7) H2 purity of greater than 99.9% is obtained by using a PSA.
8) Product H2 is delivered at a minimum pressure of 21 atms while the
sequestration ready CO2 is at a pressure of 150 atms.
9) The process achieves 90% CO2 removal.
10) A part of the solid wastes which includes fly ash, bottom ash, gasifier slag, etc
is reused in the plant while the rest of it is disposed in the coal mine at a fee.
11) Waste water is treated before discharge to meet effluent disposal guidelines.
351
Figure A.1 illustrates the conventional gasification process for the coproduction
of H2 and electricity. 5891 tonnes/day of Illinois # 6 coal is ground in a rod mill and
mixed with slurry water. The coal slurry is then injected into the GE gasifier along with
95% pure oxygen from the ASU. Two GE gasifier trains have been used each operating
at a 50% capacity at a temperature of 2500F and pressure of 65.7 atms. The hot syngas
produced at the gasifier temperature and pressure is cooled in a radiant heater followed
by water quench cooler. The molten solids solidify on cooling and collect in a water
sump at the bottom of the gasifier. In addition to removing the residual solids in the
syngas, the counter current quench also removes chlorides. Following the quench, the
syngas is fed into a sour gas stripper to remove chloride, ammonia and other trace
impurities.
The syngas is then shifted in two sets of parallel reactors arranged in series with
interstage cooling to remove the exothermic heat of the water gas shift reaction. STC
from Haldor Topsoe is used for the sour shift with a S:C ratio of 2:1 resulting in a
conversion of 98% of the CO to H2. At the exit of the shift reactors, the syngas is
cooled to 100F and sent through activated carbon beds developed by Eastman
Chemical Company for the removal of 90 to 95% of the mercury and other heavy
metals.
The syngas is then fed to a dual stage selexol scrubber unit for the removal of
H2S and CO2. H2S is removed in the first stage using a CO2 saturated selexol solution
which is regenerated in a stripper to produce the acid gas used for sulfur production in
352
a claus unit. In the second stage of the selexol unit, 95% of the CO2 is removed using
an unsaturated selexol solution. The H2 product stream thus produced is purified to
99.9% purity in a PSA. The regeneration gas from the PSA is combusted to produce
power in a steam turbine.
The CO2 produced from the regeneration of the selexol solution is dried and
compressed to 150 atms suitable for pipeline transportation in a five stage compressor
with inter-cooling at an adiabatic efficiency of 75%. The CO2 is then transported to a
geological site for sequestration.
Based on Figure A.1, an input-output diagram has been shown in Figure A.2 for
the conventional gasification process with CO2 capture. In order to quantify the input
output values, a mass and energy balance was conducted.
Table A.1 shows the energy balance for the process. 172.5 MWe is generated
from the steam turbine of which 142.21 is used for the parasitic energy requirement
within the plant. Hence 30.29 MWe is produced for sale as a result from the process.
A preliminary water balance was conducted as shown in Table A.2 to determine
the water make up. 1,819,200 Kg/hr of water is required in the plant for various unit
operations and 940,200 kg/hr is recycled within the plant. Hence 879,000 kg/hr is
required as make up water.
353
Tables A.3 and A.4 give the input and output from the process in physical and
monetary terms. As shown in Table A.4 around 10% of the carbon in coal is emitted as
CO2 due to the combustion of the PSA tail gas with air. The H2S captured in the
selexol unit is converted to elemental sulfur and sold as a byproduct.
CLP integrated with coal gasification
Basis for analysis:
1) Analysis has been conducted for a production rate of 440.8 tonnes/day of H2
and 144.4 MWe of electricity
2) Bituminous Illinois #6 coal has been used at a feed rate of 5891 tonnes/day.
3) A GE oxygen blown gasifier with radiant cooling is used to generate the syngas
at a temperature of 2500F and pressure of 65.7 atms
4) Syngas enters the H2 production reactor at a temperature of 600 ºC.
5) A Ca:C ratio of 1.1 is used since the sorbent undergoes a conversion of 90% in
the H2 production reactor. The S:C ratio is also 1:1 for H2 production. A solids
purge and makeup rate of 6% is used for the calcium based sorbent.
6) 100% calcination occurs in the calciner and the heat is supplied by the
combustion of syngas and PSA tail gas with oxygen.
7) Mercury removal is conducted using an activated carbon bed.
8) H2 purity of greater than 99.9% is obtained by using a PSA.
354
9) Product H2 is delivered at a minimum pressure of 21 atms while the
sequestration ready CO2 is at a pressure of 150 atms.
10) The process achieves nearly 100% CO2 removal since the PSA tail gas is burnt
in oxygen and not air.
11) A part of the solid wastes which includes fly ash, bottom ash, gasifier slag, etc
is reused in the plant while the rest of it is disposed in the coal mine at a fee.
12) Waste water is treated before discharge to meet effluent guidelines.
Figure A.3 illustrates the gasification process integrated with the CLP for the
coproduction of H2 and electricity. 5891 tonnes/day of Illinois # 6 coal is injected as a
slurry into the GE gasifier along with 95% pure oxygen from the ASU. Two GE
gasifier trains have been used each operating at a 50% capacity at a temperature of
2500F and pressure of 65.7 atms. The hot syngas produced at the gasifier temperature
and pressure is cooled in a radiant heater to 600 ºC. The molten solids solidify on
cooling and collect in a water sump at the bottom of the gasifier. The hot syngas is then
fed into the H2 production reactor of the CLP along with the Ca(OH)2 sorbent from the
hydrator. In this reactor, the water gas shift reaction, carbonation, sulfur and chloride
removal occur simultaneously to produce a H2 stream with a purity of >95%. The spent
sorbent is then separated from the H2 stream and sent to a calciner. The H2 stream is
sent through an activated carbon bed to remove mercury and purified in a PSA to
99.9% purity. The sorbent is calcined to produce a sequestration ready CO2 stream
which is compressed to 150 atms in a 5 stage compressor similar to the conventional
355
process. The calcined sorbent is then hydrated with stoichiometric steam and the
Ca(OH)2 is conveyed back to the H2 production reactor. The mass and energy balance
for this process was conducted using ASPEN plus under equilibrium conditions as
shown in Figure A.4. A 6% purge and make up rate is used for the calcium sorbent and
the heat for calcination is obtained by the combustion of the syngas and PSA tail gas.
Exothermic energy from the carbonator and hydrator is used to make electricity.
Figure A.5 shows the input output diagram for the CLP integrated with a coal
gasifier. In this process, there are no emissions as the PSA tail gas is not combusted in
air and almost 100% of the CO2 is captured and sent for sequestration.
The CLP produces a total of 292.7 MWe of electricity approximately half of
which is used for the parasitic energy requirement and the other half is sold along with
the H2 as shown in Table A.5.
Table A.6 gives a water balance in the plant and shows that 710,715 Kg/hr of
makeup water is required as an input.
The inputs to the process are shown in Table A.7. The amount of makeup
limestone required for a purge rate of 6% of the spent solids is 1699.2 T/d. This
amount of sorbent is sufficient to achieve almost 100% CO2 removal at high pressures
in the H2 production reactor at a solids conversion of 90%.
356
As shown in Table A.8, along with the H2 and electricity, the spent sorbent is also a
salable byproduct which is used for construction.
The life cycle carbon footprint is the total amount of green house gas emissions
over the entire lifecycle of the product. In order to determine the lifecycle carbon foot
print of a process it is necessary to determine the carbon footprint of all the inputs, the
total green house gas (GHG) emissions during the process and the carbon foot print of
all the waste streams and emissions generated. For the coal to H2/electricity systems
described earlier, this analysis includes all the GHG emissions during the mining,
transportation of coal and limestone, production of solvents, catalysts and sorbents,
production of electricity for internal plant demands (ex .CO2 compression), emissions
from plant, treatment of the wastes, and manufacture of all the equipment required for
the entire lifecycle, mining, transportation, etc of the steel required for the equipment,
etc. It is very complicated to manually determine the various paths required to conduct
this analysis. In order to simplify this analysis the EIOLCA/ECO-LCA ( Matthews et
al, 2008, DOE, 2005) software is used as it contains comprehensive information on the
GHG emissions for various products and processes. Depending on the economic value
of an activity or product, the GHG emissions are determined using this software.
In order to determine and compare the lifecycle carbon footprint of the
conventional and novel process, the monetary values calculated for the inputs and
357
outputs in Tables A.3, A.4 and Tables A.7, A.8 are used respectively. The boundary
considered for this analysis does not include the process equipment and scrap from the
equipment after its useful life.
For the conventional process, the lifecycle GWP for the annual costs of inputs
like coal, water, shift catalyst, activated carbon for mercury removal, selexol solution,
water treatment chemicals and Claus catalyst are determined using the EIOLCA/ECOLCA software. Similarly the lifecycle GWP for the treatment of waste water, solid
waste disposal are also determined. Finally, the GHG emissions during the lifecycle of
CO2 sequestration (Khoo and Tan, 2006) are determined. The lifecycle GWP related to
all the inputs, outputs, CO2 sequestration and the emission of CO2, NOx, particulates,
SO2 are added together. Since elemental sulfur is a byproduct in this process, its GWP
is subtracted from the above mentioned total to give the lifecycle carbon footprint of
the conventional process.
For the CLP, the cumulative lifecycle GWP of the inputs (coal, water, activated
carbon for mercury removal, water treatment chemicals and limestone), outputs
(treatment of waste water, solid waste disposal, disposal of the particulates, mercury
and heavy metals captured from the syngas), CO2 sequestration (Khoo and Tan, 2006)
are determined. Since the spent sorbent is a saleable product, its GWP is subtracted
from the above calculated total to give the lifecycle carbon foot print. This process
does not give rise to any emissions and all the CO2 produced is sequestered.
358
Since the conventional and CLP are optimized for the production of different
amounts of H2 and electricity, the carbon foot prints mentioned above cannot be
directly used for comparison. In order to compare the two processes, the two products,
H2 and electricity are lumped together by adding their energy contents or HHV’s and
the carbon footprint per MWth of the total product is determined and compared for the
conventional and CLP process.
Another method of comparison is to treat the electricity as a byproduct and
subtract its GWP from the total carbon footprints calculated earlier. The comparison
then is made on the basis of GHG emissions / ton of H2 product.
The life cycle GWP for each of the inputs and outputs was calculated as shown
in Tables A.9 and A.10.
The total carbon footprint for the conventional process is calculated as:
(a) When H2 and electricity are the main products:
Lifecycle carbon footprint = 353,929 + 6,536 + 40,215 +452,130 +72.5 – 8070
= 844,813 MTCO2E
=880.02 MTCO2E/MWth of total product (H2 and Electricity)
(b) When H2 is the main product:
359
Lifecycle carbon footprint = 353,929 + 6,536 + 40,215 +452,130 +72.5 – 8070 201000
= 643,813 MTCO2E
= 3.362824758 MTCO2E/tonnes of H2
The total carbon footprint for the CLP is calculated as:
(c) When H2 and electricity are the main products:
Lifecycle carbon footprint = 361,738 + 6536 + 72.5 – 7850
= 360,497 MTCO2E
= 350.427662 MTCO2E/MWth of total product (H2 and Electricity)
(d) When H2 is the main product:
Lifecycle carbon footprint = 361,738 + 6536 + 72.5 – 7850 - 957000
= -596,504 MTCO2E
= -4 MTCO2E/tonnes of H2
From the life cycle analysis it was found that the CLP has a smaller carbon foot
print when compared to the conventional process. The CLP has a smaller footprint
360
because it captures all the carbon for sequestration while the conventional process
emits about 10% of the carbon to the atmosphere.
361
Table A.1: Energy balance for the conventional process
MWe
172.5
Electricity output
Electricity needed
142.21
Coal handling, milling and coal slurry pumps
Slag handling
ASU air compressor and oxygen compressor
CO2 compressor
Feed water pumps
Condensate Pumps
Quench water pump
Shift pumps
Circulating water pump
Cooling tower fans
Scrubber pumps
Acid gas removal
Stem turbine auxilaries
Claus process
Balance of plant
Transformer losses
3.19
1.12
77.98
27.3
2.27
0.19
2.38
0.34
2.31
1.51
0.03
19.26
0.1
2.4
1.5
0.33
Net electricity produced
30.29
362
Table A.2: Water balance for the conventional process
Water Required Recycled Water Water Make-up
(Kg/hr)
(Kg/hr)
(Kg/hr)
Slag Handling
Quench/Wash
Slurry Water
Venturi Scrubber Water
Condenser Makeup
Shift Steam
Cooling Tower
SWS Blowdown
Total
30600
735600
100800
34200
166200
166200
539400
46200
1819200
363
0
735600
78000
34200
0
0
46200
46200
940200
30600
0
22800
0
166200
166200
493200
0
879000
Table A.3: Quantification of the inputs for the conventional process
Inputs
Coal
Air
Water
Shift Catalyst
Hg Removal
Selexol
Water treatment chemicals
Claus catalyst
Total flow
5981.0
28769.3
21038.4
4.0
103.0
95.0
13150.0
2.0
tpd
tpd
tpd
lb/day
lb/day
lb/day
lb/day
MWh
364
Unit cost
41.5
0.1
494.0
1.0
13.4
0.2
130.0
$/tonne
$/tonne
$/lb
$/lb
$/lb
$/lb
$/MWh
Annual Cost
86,769,358
875,344
691,600
37,492
446,548
782,425
91,000
Table A.4: Quantification of outputs for the conventional process
Outputs
Total flow
Hydrogen
547.0
Electricity
30.3
CO2 to sequestration
11265.9
Water
16560.0
Sulfur
135.7
Emissions
SO2
44.0
NOX
143.0
Particulates
148.0
Hg
0.0
CO2 from stack
1291.8
Solid waste Disposal
617.0
tpd
Mwe
tpd
tpd
tpd
tpy
tpy
tpy
tpy
tpd
lb/day
Unit cost
2040.0 $/tonne
75.0 $/MWh
Annual Cost
390,558,000
19,082,700
80.0 $/tonne
3,798,816
16.07 $/lb
365
9915.19
Table A.5: Energy balance for the CLP
Electricity output
MWe
292.68
Electricity needed
148.272
Coal handling, milling and coal slurry pumps
Slag handling
ASU air compressor and oxygen compressor
CO2 compressor
Feed water pumps
Hydration water pump
Circulating water pump
Stem turbine auxilaries
Balance of plant
Transformer losses
3.19
1.12
109
28.3
2.27
0.152
2.31
0.1
1.5
0.33
Net electricity produced
144.408
366
Table A.6: Water balance for the CLP
Slag Handling
Slurry Water
Hydration Steam
Cooling Tower
Total
Water Required Recycled Water Water Make-up
(Kg/hr)
(Kg/hr)
(Kg/hr)
30600
0
30600
100800
78000
22800
278775
114660
164115
539400
46200
493200
949575
238860
710715
367
Table A.7: Quantification of the inputs for the CLP
Inputs
Coal
Air
Water
Hg Removal
Limestone
Water treatment chemicals
Total flow
5981.0
22823.0
17057.2
103.0
1699.2
13150.0
tpd
tpd
tpd
lb/day
tpd
lb/day
368
Unit cost
41.5
0.1
1.0
20.0
0.2
$/tonne
$/tonne
$/lb
$/tonne
$/lb
Annual Cost
86,769,358
709,696
37,492
11,894,400
782,425
Table A.8: Quantification of the outputs for the CLP
Outputs
Total flow
Hydrogen
440.8
Electricity
144.4
CO2 to sequestration
12768.4
Water
12571.2
Spent Sorbent
1568.9
CaS
275.5
CaCO3
1196.4
CaCl2
24.5
Ca(OH)2
72.6
Particulates
148.0
Hg
0.0
Solid waste Disposal
617.0
tpd
Mwe
tpd
tpd
tpd
tpd
tpd
tpd
tpd
tpy
tpy
lb/day
Unit cost
2040.0 $/tonne
75.0 $/MWh
8.0 $/tonne
Annual Cost
314,765,066
90,977,670
4,393,032
16.07 $/lb
369
9915.19
Table A.9: Quantification of the Global Warming Potential(GWP) for the inputs and
outputs for the conventional process.
Inputs
Total flow
5981.0
28769.3
21038.4
4.0
103.0
95.0
13150.0
2.0
Coal
Air
Water
Shift Catalyst
Hg Removal
Selexol
Water treatment chemicals
Claus catalyst
Total
Outputs
Total flow
Hydrogen
547.0
Electricity
30.3
CO2 to sequestration
11265.9
Water
16560.0
Sulfur
135.7
Emissions
SO2
44.0
NOX
143.0
Particulates
148.0
Hg
0.0
CO2 from stack
1291.8
Solid waste Disposal
617.0
Unit cost
41.5
0.1
494.0
1.0
13.4
0.2
130.0
tpd
tpd
tpd
lb/day
lb/day
lb/day
lb/day
MWh
tpd
Mwe
tpd
tpd
tpd
tpy
tpy
tpy
tpy
tpd
lb/day
$/tonne
$/tonne
$/lb
$/lb
$/lb
$/lb
$/MWh
Unit cost
2040.0 $/tonne
75.0 $/MWh
GWP
Annual Cost MTCO2
86,769,358
349000
875,344
691,600
37,492
446,548
782,425
91,000
Annual Cost
390,558,000
19,082,700
80.0 $/tonne
3,798,816
16.07 $/lb
370
9915.19
787
2,300
58
840
642
302
353,929
GWP
MTCO2
201000
6536
8070
40215
452130
72.5
Table A.10: Quantification of the Global Warming Potential(GWP) for the inputs and
outputs for the CLP
Inputs
Total flow
5981.0
22823.0
17057.2
103.0
1699.2
13150.0
Coal
Air
Water
Hg Removal
Limestone
Water treatment chemicals
Total
Outputs
Total flow
Hydrogen
440.8
Electricity
144.4
CO2 to sequestration
12768.4
Water
12571.2
Spent Sorbent
1568.9
CaS
275.5
CaCO3
1196.4
CaCl2
24.5
Ca(OH)2
72.6
Particulates
148.0
Hg
0.0
Solid waste Disposal
617.0
Unit cost
41.5
0.1
1.0
20.0
0.2
tpd
tpd
tpd
lb/day
tpd
lb/day
tpd
Mwe
tpd
tpd
tpd
tpd
tpd
tpd
tpd
tpy
tpy
lb/day
$/tonne
$/tonne
$/lb
$/tonne
$/lb
Unit cost
2040.0 $/tonne
75.0 $/MWh
8.0 $/tonne
16.07 $/lb
371
GWP
Annual Cost MTCO2
86,769,358
349000
709,696
638
37,492
58
11,894,400
11,400
782,425
642
361,738
GWP
Annual Cost MTCO2
314,765,066 90,977,670
957000
6536
4,393,032
7850
9915.19
72.5
Claus Catalyst
H2S
Shift
Catalyst
Activated
Carbon
Regeneration
Steam
Shift
Gas
Cooling
Quench and
Syngas Scrubber
372
Coal
GE Gasifier
Final Syngas
Scrubber
Steam
Air
Steam
HRSG
Power
Generation
ASU
Water
Slag
Mercury
Removal
N2 Rich
Stream
Electricity
Selexol
Unit
PSA
Boiler
Claus
Process
CO2
Compression
CO2 to
Sequestration
Selexol
Pure H2
Air
Water +
Treatment Chemicals
Stack Gases
Figure A.1: Schematic of a conventional gasification plant for the cogeneration of H2 and electricity
372
Sulfur
Slag
Sulfur
H2 Electricity CO2 Water Emissions
Captured CO2
Coal Gasification and
CO2 Sequestration
Selexol Aided CO2 Capture
Coal
Air Water
Claus
Catalyst
Shift
Catalyst
Selexol
Activated
Water Treatment
Carbon (Hg) Chemicals
Figure A.2: Input-output diagram for the conventional coal to H2 process
373
Electricity
Limestone
Power
Generation
Steam
Hydrator
HRSG
Calciner
CO2
Compression
CO2 to
Sequestration
St e
am
Water
+ Treatment Chemicals
Coal
Radiant
Cooler
374
GE Gasifier
Water
Slag
H2 Production
Reactor
PCD
PSA
ASU
Air
Hg Removal
N2 Rich
Stream
Spent Sorbent
Figure A.3: Schematic of a CLP plant for the cogeneration of H2 and electricity
374
Pure H2
Activated
Carbon
54
B29
45
B19
Q=8
B25
B27
B7
Q=-68
Q=26
3
W=17148B24
B30
48
B3
52
Q=64
Q=31
B6
Q=11
12
43
41
B31
51
W=26659 W=21252
B8
40
15
B10
31
CO2H2S-O
19
9
W=152
CARB
CAO,CO2
39
W
B9
Q CYCCO2
18
CALC
Q=37
Q=9B13
SOLMAKUP
24
B2
4
B4
Q=46
Q
B20
B1
B15
W=-33386
Q=-39
Q=-227
TAILGAS
MIXER
1
SOLPURGE
B5
FSPLIT
Q=-9
CYCH2
PSA
CACO3
B28
44
38
2
33
Q=84
49
W=-24631
34
7
Q=-12
B18
Q=-212
Q=11
B21
B11
B22
B17
14
32
Q=15
B23
B12
W=12145 29
W=14053
27
30
CAO
B14
B26
PUREH2
B16
23
HYDROGEN
Q=-1
Q=-5
Q=20
Q
53
Figure A.4: Flow sheet developed for the CLP using ASPEN plus simulator
375
SYNGAS
Spent Sorbent
H2 Electricity
Slag
Water
Coal Gasification and
CLP
Captured CO 2
CO2 Sequestration
Water Treatment
Coal Air Water Activated
Carbon (Hg) Chemicals
Limestone
Figure A.5: Input-output diagram for the coal to H2 process using the CLP
376
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