Conceptual Design of Supercritical O2-Based PC Boiler
Final Report
Andrew Seltzer
Zhen Fan
Archie Robertson
November 2006
DE-FC26-04NT42207
Foster Wheeler Power Group, Inc.
12 Peach Tree Hill Road
Livingston, New Jersey 07039
1
DISCLAIMER
This report was prepared as an account of work sponsored by an agency of the
United States Government. Neither the United States Government nor any
agency thereof, nor any of their employees, makes any warranty, express or
implied, or assumes any legal liability or responsibility for the accuracy,
completeness, or usefulness of any information, apparatus, product, or process
disclosed, or represents that its use would not infringe privately owned rights.
Reference herein to any specific commercial product, process, or service by
trade name, trademark, manufacturer, or otherwise does not necessarily
constitute or imply its endorsement, recommendation, or favoring by the United
States Government or any agency thereof. The views and opinions of authors
expressed herein do not necessarily state or reflect those of the United States
Government or any agency thereof.
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ABSTRACT
The objective of the system design and analysis task of the Conceptual Design of
Oxygen-Based Supercritical PC Boiler study is to evaluate the effects of oxygen
firing on supercritical PC boiler design, operation and system performance.
Simulations of the oxygen-fired plant with CO2 sequestration were conducted
using Aspen Plus and were compared to a reference air-fired 460 MWe plant.
Flue gas recycle is used to control the flame temperature and resultant wall
temperature in the O2-fired PC. Parametric trade-off studies were made to
determine the effect of flame temperature on system efficiency. The degree of
improvement in system performance of various modifications was investigated.
The objective of the advanced oxygen separation system integration task of the
study is to evaluate the benefits, effects, and limitations of the integration of
advanced oxygen separation technologies into a supercritical O2-fired PC.
Simulations of the power generation unit, oxygen separation unit, and CO2
sequestration system were conducted using the Aspen Plus software. The
improvement of the O2-fired PC system performance incorporating the Oxygen
Ion Transport Membrane (OITM) and Ceramic Auto-thermal Recovery (CAR)
were investigated. A parametric study was conducted to determine the sensitivity
of the design and performance to various variables. Compared to the other CO2
removal and sequestration technologies, the oxygen-fired PC integrated with
OITM shows substantially less CO2 removal penalty. The CO2 removal penalty of
the oxygen-fired PC integrated with CAR is between cryogenic air separation and
OITM.
The objective of the furnace and heat recovery area design and analysis task of
the study is to optimize the location and design of the furnace, burners, over-fire
gas ports, and internal radiant surfaces. The furnace and heat recovery area
were designed and analyzed using the FW-FIRE, Siemens, and HEATEX
computer programs. The furnace is designed with opposed wall-firing burners
and over-fire air ports. Water is circulated in the furnace by forced circulation to
the waterwalls at the periphery and divisional wall panels within the furnace.
Compared to the air-fired furnace, the oxygen-fired furnace requires only 65% of
the surface area and 45% of the volume. Two oxygen-fired designs were
simulated: 1) with a cryogenic air separation unit (ASU) and 2) with an oxygen
ion transport membrane (OITM).
The maximum wall heat flux in the oxygen-fired furnace is more than double that
of the air-fired furnace due to the higher flame temperature and higher H2O and
CO2 concentrations. The coal burnout for the oxygen-fired case is 100% due to a
500°F higher furnace temperature and higher concentration of O2. Because of
the higher furnace wall temperature of the oxygen-fired case compared to the airfired case, furnace water wall material was upgraded from T2 to T92.
Compared to the air-fired heat recovery area (HRA), the oxygen-fired HRA total
heat transfer surface is 35% less for the cryogenic design and 13% less for the
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OITM design due to more heat being absorbed in the oxygen-fired furnace and
the greater molecular weight of the oxygen-fired flue gas. The HRA tube
materials and wall thicknesses are nearly the same for the air-fired and oxygenfired designs since the flue gas and water/steam temperature profiles
encountered by the heat transfer banks are similar.
The capital and operating costs of the pulverized coal-fired boilers required by
the three different plants (air-fired, O2-fired with cryogenic ASU and O2-fired with
OITM) were estimated by Foster Wheeler and the balance of plant costs were
budget priced using published data together with vendor supplied quotations.
The cost of electricity produced by each of the plants was determined and
oxygen-based plant CO2 mitigation costs were calculated and compared to each
other as well as to values published for some alternative CO2 capture
technologies. Compared to other CO2 sequestration technologies, the O2-fired
PC is substantially more cost effective than both natural gas combined cycles
and post CO2 removal PCs and it is superior, especially when a supercritical
steam cycle is applied, to integrated gasification combined cycles.
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Table of Contents
ABSTRACT....................................................................................................................... 3
1.0
INTRODUCTION................................................................................................. 9
2.0
EXECUTIVE SUMMARY ................................................................................ 11
3.0
EXPERIMENTAL.............................................................................................. 17
4.0
RESULTS AND DISCUSSION ......................................................................... 18
4.1
SYSTEM DESIGN AND ANALYSIS ........................................................................ 18
4.1.1
Reference Site and Conditions .................................................................. 18
4.1.2
Air-Fired Reference Case (case-1) ........................................................... 19
4.1.3
Oxygen-Based PC Plant – Base Case....................................................... 22
4.1.4
Parametric Cases...................................................................................... 28
4.1.5
Part Load .................................................................................................. 40
4.1.6
Comparison With Post Combustion CO2 Capture.................................... 42
4.1.7
Furnace Waterwall Temperature.............................................................. 42
4.2
ADVANCED O2 SEPARATION SYSTEM INTEGRATION .......................................... 43
4.2.1
O2-Fired PC Integrated with Cryogenic ASU (case-6)............................ 43
4.2.2
O2-Fired PC Integrated with OITM.......................................................... 45
4.2.3
O2-Fired PC Integrated with CAR............................................................ 61
4.2.4
Comparisons ............................................................................................. 67
4.3
FURNACE AND HEAT RECOVERY AREA DESIGN AND ANALYSIS ........................ 69
4.3.1
Furnace Design and Analysis ................................................................... 69
4.3.2
Boundary Conditions ................................................................................ 71
4.3.3
Furnace Waterwall and Division Wall Design....................................... 115
4.3.4
Furnace Waterwall Corrosion................................................................ 123
4.3.5
Heat Recovery Area Design and Analysis .............................................. 125
4.4
ECONOMIC ANALYSIS ...................................................................................... 131
4.4.1
Main Assumptions................................................................................... 131
4.4.2
Plant Cost Basis...................................................................................... 134
4.4.3
Total Plant Investment (TPI) .................................................................. 142
4.4.4
Total Capital Requirement (TCR)........................................................... 142
4.4.5
Operating Costs And Expenses............................................................... 143
4.4.6
Cost Of Electricity (COE)....................................................................... 144
4.4.7
Tube Weld Overlay ................................................................................. 145
4.4.8
Reduced Pipeline Pressure ..................................................................... 146
4.4.9
Comparison with Other Technologies .................................................... 155
5.0
CONCLUSION ................................................................................................. 157
6.0
REFERENCES.................................................................................................. 161
7.0
BIBLIOGRAPHY ............................................................................................. 163
8.0
LIST OF ACRONYMS AND ABBREVIATIONS ........................................ 164
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List of Figures
Figure 4.1.1 - Reference Case of Air-fired Supercritical PC................................21
Figure 4.1.2 – Oxygen purity effect on ASU specific power ................................23
Figure 4.1.3 - Carbon burnout vs. O2 concentration to boiler ..............................25
Figure 4.1.4 - Base case of O2-fired supercritical PC..........................................27
Figure 4.1.5 - O2-fired supercritical PC with liquid CO2 pump (Case-3) .............31
Figure 4.1.6 - Temperature-Quality diagram for wet-end heat exchangers.........32
Figure 4.1.7 – Equilibrium and adiabatic temperatures vs. O2 concentration......33
Figure 4.1.8 - Boiler design with HRA series pass ..............................................38
Figure 4.1.9 - Boiler design with HRA series pass (56% Flue Gas Recycle) ......39
Figure 4.1.10 - Flue gas recycle flow vs. part load operation..............................41
Figure 4.1.11 Efficiency at part load operation....................................................41
Figure 4.2.1 - Base case of O2-fired supercritical PC with Cryogenic ASU .........44
Figure 4.2.2 - A Typical Oxygen Ion Transport Membrane Schematic [4]...........47
Figure 4.2.3 – Integration of Oxygen Ion Transport Membrane into O2-PC........48
Figure 4.2.4 - O2-PC with OITM .........................................................................51
Figure 4.2.5 – Effect of OITM O2 Recovery Efficiency on Efficiency ...................54
Figure 4.2.6 – Effect of Sweep Gas Pressure on LMPD .....................................55
Figure 4.2.7 – Effect of OITM Pressure on System Efficiency and LMPD: Variable
O2 Recovery................................................................................................57
Figure 4.2.8 – Effect of O2 Recovery on System Efficiency ................................57
Figure 4.2.9 – Effect of OITM Pressure on System Efficiency and LMPD:
Constant and Variable O2 Recovery ...........................................................58
Figure 4.2.10 - O2-PC with OITM with 58% Recycle Flow..................................60
Figure 4.2.11 - CAR Process Schematic ............................................................61
Figure 4.2.12 – Integration of CAR Process into O2-PC....................................65
Figure 4.2.13 - O2-PC with CAR.........................................................................66
Figure 4.3.1 – Computational Model of Air-Fired Furnace (with right side wall
removed) .....................................................................................................76
Figure 4.3.2 – Computational Model of Oxygen-Fired Furnace With OFA..........77
Figure 4.3.3 – Air-Fired Boundary Conditions .....................................................78
Figure 4.3.4 – Oxygen-Fired Boundary Conditions (Cryogenic ASU) .................79
Figure 4.3.5 – Oxygen-Fired Boundary Conditions (OITM).................................80
Figure 4.3.6 – Air-Fired Boiler Design.................................................................81
Figure 4.3.7 – Summary of FW-FIRE Furnace Modeling Results .......................82
Figure 4.3.8 – Gas Velocity for Air-Fired Case ...................................................83
Figure 4.3.9 – Gas Temperature for Air-Fired Case............................................84
Figure 4.3.10 – O2 Mole Fraction for Air-Fired Case...........................................85
Figure 4.3.11 – Wall Heat Flux for Air-Fired Case ..............................................86
Figure 4.3.12 – Wall Temperature for Air-Fired Case .........................................87
Figure 4.3.13 – Wall CO for Air-Fired Case ........................................................88
Figure 4.3.14 – Char Mass Fraction (72 microns) for Air-Fired Case..................89
Figure 4.3.15 – Char Mass Fraction (176 microns) for Air-Fired Case................90
Figure 4.3.16 – Air-Fired and Oxygen-Fired Boiler Outlines ...............................91
Figure 4.3.17 – Oxygen-Fired Boiler Design (Cryogenic ASU) ...........................92
Figure 4.3.18 – Oxygen-Fired Boiler Design (OITM)...........................................93
Figure 4.3.19 – Gas Velocity for O2-Fired Case..................................................94
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Figure 4.3.20 – Gas Temperature for O2-Fired Case..........................................95
Figure 4.3.21 – O2 Mole Fraction for O2-Fired Case ...........................................96
Figure 4.3.22 – Wall Heat Flux for O2-Fired Case...............................................97
Figure 4.3.23 – Wall Temperature for O2-Fired Case .........................................98
Figure 4.3.24 – Wall CO for O2-Fired Case.........................................................99
Figure 4.3.25 – Char Mass Fraction (69 micron) for O2-Fired Case..................100
Figure 4.3.26 – Char Mass Fraction (169 micron) for O2-Fired Case................101
Figure 4.3.27 - Flue gas recycle flow vs. part load operation...........................102
Figure 4.3.28 – Summary of O2-Fired Part Load Results, Cryogenic ASU .......103
Figure 4.3.29 – Gas Temperature for O2-Fired Part Load, Cryogenic ASU ......104
Figure 4.3.30 – Wall Heat Flux for O2-Fired Part Load, Cryogenic ASU ...........105
Figure 4.3.31 – Average and Peak Heat Flux in Waterwalls.............................106
Figure 4.3.32 – Gas Velocity for O2-Fired with OITM........................................107
Figure 4.3.33 – Gas Temperature for O2-Fired Case With OITM......................108
Figure 4.3.34 – O2 Mole Fraction for O2-Fired Case With OITM .......................109
Figure 4.3.35 – Wall Heat Flux for O2-Fired Case With OITM...........................110
Figure 4.3.36 – Wall Temperature for O2-Fired Case With OITM .....................111
Figure 4.3.37 – Wall CO for O2-Fired Case With OITM ....................................112
Figure 4.3.38 – Char Mass Fraction (69 micron) for O2-Fired Case, OITM.......113
Figure 4.3.39 – Char Mass Fraction (169 micron) for O2-Fired Case, OITM.....114
Figure 4.3.40 – Advantage of Low Mass Flux Design.......................................117
Figure 4.3.41 – Tube Wall Temperature with Smooth, Standard Rifled, and
Optimized Rifled Tubes .............................................................................118
Figure 4.3.42 – Rifled Tube Design ..................................................................119
Figure 4.3.43 – Outside Tube Wall Temperature with Peak Heat Flux .............120
Figure 4.3.44 – Tube Inlet Mass Flow With a 10% Heat Flux Step Increase (No
Pressure Equalization Header) ..................................................................121
Figure 4.3.45 – Tube Inlet Mass Flow With a 10% Heat Flux Step Increase
(With Pressure Equalization Header at 80’) ...............................................122
Figure 4.3.46 – Predicted Wall Corrosion (mil/yr) in Air-Fired and O2-Fired
Furnaces....................................................................................................124
Figure 4.3.47 – HRA Tube Bank Design...........................................................128
Figure 4.3.48 – HRA Tube Bank Performance .................................................130
Figure 4.4.1 – Increase in COE of O2 Plant Above Air-Fired Reference Plant ..152
Figure 4.4.2 – COE Breakdown ........................................................................153
Figure 4.4.3 - Comparison of Levelized Cost of Electricity Among Alternative
Technologies .............................................................................................156
Figure 4.4.4 - Comparison of Mitigation Costs Among Alternative Technologies
...................................................................................................................156
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List of Tables
Table 4.1.1 - Summary of Plant Performance and Economics ...........................15
Table 4.1.1 - Site Conditions...............................................................................18
Table 4.1.2 - Setup and Assumptions .................................................................19
Table 4.1.3 - Case Summary .............................................................................29
Table 4.1.4 - Flame adiabatic temperature effect on performance .....................34
Table 4.1.5 – Flame adiabatic temperature effect on boiler flue gas flow ...........34
Table 4.2.1 – Comparison of O2PC with Cryogenic and OITM ASU ..................50
Table 4.2.2 – Parametric Case Summary ...........................................................52
Table 4.2.3 - Effect of OITM O2 Recovery Efficiency on LMPD ..........................54
Table 4.2.4 – Effect of Flame Temperature on Performance ..............................59
Table 4.2.5 - Comparison for CAR with Cryogenic ASU .....................................62
Table 4.2.6 – Comparison for CAR Using Different Sweep Gases .....................63
Table 4.2.7 – O2 PC Efficiency with Different O2 Separation Techniques ...........64
Table 4.2.8 - Gains from OITM and CAR Compared to Cryogenic ASU.............67
Table 4.2.9 – Comparison of CO2 Removal Technologies .................................68
Table 4.4.1 - Coal Properties ............................................................................132
Table 4.4.2 - CO2 Effluent Purity Design Conditions.........................................132
Table 4.4.3 – Economic Study Assumptions.....................................................133
Table 4.4.4 - Cost of 430.2 MWe Air-Fired Supercritical PC Plant ($1000 Yr
2006) .........................................................................................................138
Table 4.4.5 - Cost of 347.0 MWe ASU Based Supercritical PC Plant ($1000 Yr
2006) .........................................................................................................139
Table 4.4.6 - Cost of 463.3 MWe OITM Based Supercritical PC Plant ($1000 Yr
2006) .........................................................................................................140
Table 4.4.7 - Comparison of Total Plant Costs ($1000 Yr 2006) ......................141
Table 4.4.8 - Air Fired PC Plant Capital Investment & Revenue Requirement .147
Table 4.4.9 - ASU Based PC Plant Capital Investment & Revenue Requirement
...................................................................................................................148
Table 4.4.10 - OITM Based PC Plant Capital Investment & Revenue
Requirement ..............................................................................................149
Table 4.4.11 - Summary of Plant Economics and CO2 Mitigation Costs...........150
Table 4.4.12 - Plant Daily Consumable Requirements .....................................151
Table 4.4.13 – COE and CO2 MC with Weld Overlay and 2200 psi Pressure ..154
8
1.0 Introduction
The objective of the Conceptual Design of Oxygen-Based supercritical PC Boiler
study is to develop a conceptual pulverized coal (PC)-fired power plant, which
facilitates the practical capture of carbon dioxide capture for subsequent
sequestration. The system design and analysis task, which was performed using
the Aspen Plus computer program, is aimed at optimizing the PC boiler plant
operating parameters to minimize the overall power plant heat rate. The flow
rates and other properties of individual streams of the power plant were
calculated as the results of the Aspen Plus simulations. The required
performance characteristics of such operating components as pulverized coalfired furnace, heat recovery area, flue gas recuperator, and economizer were
determined.
Two plant configurations were simulated: 1) a conventional air-fired PC power
plant and 2) the proposed oxygen-based supercritical PC plant. In order to
compare the performance of the oxygen-based plant with that of the conventional
plant, the main steam generation rate in the both plants was kept constant.
The objective of the advanced oxygen separation system integration task is to
identify promising, low cost advanced options of oxygen separation, which can
be integrated into the O2-fired PC power plant. Currently, a number of new
oxygen separation technologies are in development. They are categorized as
high temperature ion membranes, such as oxygen ion transport membrane
(OITM), and high temperature sorption, such as ceramic auto-thermal recovery
(CAR). The former relies on oxygen transport through a ceramic membrane, and
the latter on oxygen storage in perovskite type materials. Integration of these
advanced oxygen separations into the O2-fired PC has the potential to
substantially reduce the cost of CO2 removal. This task deals with the systemlevel evaluation of the O2-fired PC integrated with these advanced oxygen
separation methods.
The advanced oxygen separation system integration task, which was performed
using the Aspen Plus computer program, is aimed at a system level optimization
to minimize the overall heat rate and maximize system performance. Two types
of advanced oxygen separation systems and related configurations were
simulated: 1) high temperature membrane technology (OITM) and 2) high
temperature oxygen sorbent technology (CAR). Determined are the required
performance characteristics of the operating components such as the boiler (with
air heater for OITM), GT expander, wet-end economizer for low-grade heat
recovery, and air separation equipment.
The furnace and heat recovery area design and analysis task, which was
performed using the FW-FIRE, Siemens and HEATEX computer programs, is
aimed at optimizing the location and design of the furnace, burners, over-fire gas
ports, and internal radiant surfaces. Three furnace and HRA designs were
developed: 1) a conventional air-fired PC power plant and 2) an oxygen-based
9
PC plant with cryogenic ASU and 3) an oxygen-based PC plant with oxygen ion
transport membrane.
The objective of the economic analysis is to prepare a budgetary estimate of the
capital and operating costs of the O2-fired PC power plants to permit comparison
to an equivalent, conventional, air-fired power plant (e.g. the reference plant) as
well as other CO2 capture technologies.
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2.0 Executive Summary
The objective of the Conceptual Design of Oxygen-Based supercritical PC Boiler
study is to develop and design a conceptual pulverized coal-fired power plant,
which facilitates the practical capture of carbon dioxide capture for subsequent
sequestration.
The reference plant employs a supercritical steam turbine with conditions,
4035psia/1076ºF/1112ºF/2.0”Hg, fires high-volatile bituminous Illinois 6 coal, and
produces 460 MWe at the generator. A conventional air-fired case was
simulated as the comparison basis. The air-fired plant has a boiler efficiency of
88.2% and a net plant efficiency of 39.5%.
The system design and analysis task (Task 1), which was performed using the
Aspen Plus computer program, is aimed at optimizing the PC boiler plant
operating parameters to minimize the overall power plant heat rate. The oxygenbased plant model contains all the components in the conventional plant model
(with the exception of the FGD and SCR) plus the addition of an air separation
unit and a flue gas cooler. Flue gas is recycled to control the flame temperature
inside the PC boiler to minimize NOx formation, to minimize ash slagging in the
furnace combustion zone, and to avoid the application of exotic materials.
Equipment to compress and liquefy the CO2 effluent to 3000 psia and to reduce
the moisture to 50 ppm (to avoid transport pipe corrosion) was included in the O2PC system model. Compression to 3000 psia is conservative compared to the
pressure of 2200 psia specified in Reference 14.
The supercritical O2-PC power plant simulated with the same flame temperature
as the air-fired PC, shows a plant net efficiency drop from 39.5 to 30.8%, and a
power reduction from 430 to 332 MWe. The flue gas flowing through the boiler
changes from 3555 to 3363 klb/hr (179 to 125 Mcf/hr, a 30% reduction in
volumetric flow). As the result of the air separation unit (ASU) power and CO2
compressions, the specific power penalty for CO2 removal is 129 kWh/klbCO2.
Parametric trade-off runs were made by varying the amount of recycled flue gas
(which directly affects the flame temperature) while maintaining the same boiler
outlet O2 concentration (3%, vol.) as the air-fired case, and by varying the
temperature of the recycled flue gas. The results show that by reducing the
recycled flue gas flow rate by 38% (by volume), and by raising the temperature of
the recycled flue gas from 95ºF to 260ºF, the equilibrium temperature increases
from 3480ºF to 4160ºF, the system efficiency increases from 30.8% to 32.9%,
and the specific penalty for CO2 removal reduces from 129 to 99 kWh/klbCO2.
The objective of Task 2 is to develop a system design of a conceptual pulverized
coal-fired oxygen combustion power plant, integrated with advanced oxygen
separation technology. The baseline oxygen-based PC boiler incorporates
cryogenic O2 separation, which can produce oxygen with high purity; but it
requires substantial capital and operating costs. Membrane separation of O2 has
11
been demonstrated at small scale employing very thin membrane fibers, which
preferentially allow O2 to permeate, but not N2. Membrane separation has the
potential to use less power at a lower capital cost.
The advanced oxygen separation system integration task (Task 2), which was
performed using the Aspen Plus computer program, is aimed at a system level
optimization to minimize the overall heat rate and maximize system performance.
Two types of advanced oxygen separation systems and related configurations
were simulated: 1) high temperature membrane technology (OITM) and 2) high
temperature oxygen sorbent technology (CAR).
Oxygen separation by oxygen ion transport membrane is driven by the difference
in oxygen partial pressure across a membrane. To produce this pressure
difference, the air is pressurized by a compressor and heated to a high
temperature. The hot pressurized air is fed to an oxygen ion transport membrane
(OITM), where about 85% of its O2 is separated through membrane, and the
remaining O2 is carried by hot vitiated air to a gas expander. Power generated
from the expander is used to drive the air compressor. The OITM does not
consume electrical power; instead, it absorbs heat, generates power, as it
separates O2 from air.
The O2-fired reference plant with cryogenic ASU has a net plant efficiency of
31.9%, a net power generation of 338 MWe, and a CO2 removal penalty of 114
kWh/klbCO2. The O2-fired reference plant with OITM has a net plant efficiency of
36.1%, a net power generation of 463 MWe, and a CO2 removal penalty of 42
kWh/klbCO2.
Parametric trade-off runs were conducted by varying the O2 recovery efficiency,
the pressure difference across the membrane, the OITM operating pressure, the
compressor discharge temperature, and the furnace flame temperature. OITM
faces significant challenges with respect to the manufacture and stability of
membranes, and scale up and design of large plants.
The ceramic auto-thermal recovery (CAR) process is based on sorption and
storage of oxygen in a fixed bed containing ionic and electronic conductor
materials. For the CAR process utilizing extracted steam as the sweep gas, net
system efficiency is increased by only by 0.7% point compared to the cryogenic
ASU process. But if the CAR process uses recycled flue gas as the sweep gas,
system efficiency can be increased by 2.6% points, compared to the gain of 3.2%
points of the OITM process. At the same equilibrium temperature (4160ºF), the
efficiency reduction of the O2-PC compared to the air-fired PC is 6.5% points with
cryogenic air separation, 3.3% points with OITM, and 4.3% points with CAR.
The furnace and heat recovery area design and analysis task (Task 3), which
was performed using the FW-FIRE, Siemens, and HEATEX computer programs,
is aimed at optimizing the location and design of the furnace, burners, over-fire
gas ports, and internal radiant surfaces.
12
A simulation was made for both the reference air-fired case and for the oxygenfired case. Two oxygen-fired models were constructed: one for the cryogenic
ASU design (with radiant superheater partial division walls) and the second for
the OITM design with a furnace wall radiant superheater and a high temperature
air heater. Boundary conditions are based on ASPEN simulations of the power
plant.
The furnace is designed with opposed wall-firing burners and over-fire air ports
located at one burner pitch above the top burner row. The O2-PC supercritical
boiler incorporates the BENSON vertical technology, which uses low fluid mass
flow rates in combination with optimized rifled tubing. Water is circulated in the
furnace by forced circulation to the waterwalls at the periphery and divisional wall
panels within the furnace.
For the air-fired furnace simulation, the maximum flue gas temperature is
approximately 3350oF. The maximum heat flux is approximately 70,000 Btu/hr-ft2
and is located on the side wall at the top of the burner zone. The total heat
absorbed by the furnace walls before the furnace exit is 1770 MM Btu/hr. The
maximum temperature of the waterwalls is approximately 870oF and of the
division walls is approximately 1000oF. Total burnout of all particle sizes is
99.6%. Average NOx concentration at the furnace outlet is 276 ppmvw (0.38
lb/MMBtu).
Compared to the air-fired furnace, the oxygen furnace requires only 65% of the
surface area and 45% of the volume. Two oxygen-fired designs were simulated:
1) with cryogenic ASU and 2) with OITM. The mixed primary/secondary gas O2
content (before combustion) is approximately 40%.
In the oxygen-fired furnace, the maximum flue gas temperature is approximately
3900oF for cryogenic ASU and 3850oF for OITM. The maximum heat flux is
171,000 Btu/hr-ft2 for cryogenic ASU and 180,000 Btu/hr-ft2 for OITM. The
maximum wall heat flux in the oxygen-fired furnace is more than double that of
the air-fired furnace due to the higher flame temperature and higher H2O and
CO2 concentrations. The total heat absorbed by the furnace walls before the
furnace exit is approximately 2287 (cryogenic) and 2029 (OITM) MM Btu/hr. The
coal burnout for the oxygen-fired case is 100% due to the high furnace
temperature and high concentration of O2. NOx is 261 ppmvw (0.18 lb/MMBtu).
The maximum temperature of the oxygen-fired furnace is approximately 1060oF
for the waterwalls, 1065oF for the division walls, and 1100oF for OITM radiant
superheater walls. Because of the higher temperature of the oxygen-fired case
compared to the air-fired case, furnace tube material was upgraded from T2 to
T92.
Since the boiler is a supercritical once-through sliding pressure unit, part load
cases (72%, 50%, and 25%) were run and evaluated with thermal/hydraulic and
structural criteria. A pressure equalization header is included at an elevation of
80’ to ensure stable operation at low loads.
13
To reduce the corrosion in the O2-PC, tube weld overlays with high Nickel and
Chromium contents (e.g. alloy 622) are applied to the waterwalls. This reduces
the predicted maximum corrosion from 45 mil/yr to 10 mil/yr.
HEATEX was used to determine the heat recovery area (HRA) design of the
convective tube banks between the furnace exit and the SCR/air heater. These
tube banks include the finishing superheater, finishing reheater, primary
superheater, primary reheater, upper economizer, and lower economizer.
For the air-fired design, total surface area of all convective banks is 335,025 ft2.
The total heat transferred to the water/steam is 1431 MM Btu/hr as 3.59 MM lb/hr
of flue gas is cooled from 2185ºF to 720ºF.
For the cryogenic ASU oxygen-fired design, convective bank total surface area is
218,693 ft2 and the total heat transferred to the water/steam is 1151 MM Btu/hr
as 2.12 MM lb/hr of flue gas is cooled from 2450ºF to 695ºF. For the OITM
oxygen-fired design, convective bank total surface area is 274,466 ft2 and the
total heat transferred to the water/steam is 1185 MM Btu/hr and to the air is 990
MM Btu/hr as 2.68 MM lb/hr of flue gas is cooled from 2950ºF to 695ºF.The total
heat transfer surface required in the oxygen-fired HRA is less than the air-fired
HRA due to more heat being absorbed in the oxygen-fired furnace and the
greater molecular weight of the oxygen-fired flue gas.
The HRA tube materials and wall thicknesses are nearly the same for the airfired and oxygen-fired designs since the flue gas and water/steam temperature
profiles encountered by the heat transfer banks are similar.
A tubular convective air heater is included in the OITM O2-PC to provide the
necessary air heating for the membrane separation process. The furnace air
heater is an Incoloy MA956 three-pass tubular design situated above the furnace
nose.
The objective of the economic analysis task (Task 4) is to prepare a budgetary
estimate of the capital and operating costs of the O2-fired PC power plants to
permit comparison to an equivalent, conventional, air-fired power plant (e.g. the
reference plant) as well as other CO2 capture technologies. The economic
analyses of the plants were carried out based on the EPRI Technical
Assessment Guide (TAG) methodology. Plant capital costs were compiled under
the Code of Accounts developed by EPRI. The estimate basis is year 2006
dollars, a 20-year life, and an 85 per cent capacity factor. Table 4.1.1
summarizes the performance and economics of the plants.
14
Table 4.1.1 - Summary of Plant Performance and Economics
Reference
Air-Fired Plant
ASU Based
Plant
OITM Based
Plant
Net Power Output, MWe
430.2
347.0
463.3
Efficiency, % (HHV)
39.5
33.0
36.1
Coal Flow, Klb/hr
319.0
308.0
375.4
633.0
1,471
723.3
2,084
953.0
2,057
50.41
66.17
63.48
20.23
16.77
Total Plant Cost
Millions of Dollars
$/kW
Levelized COE, $/MWhr
CO2 Mitigation Cost, $/tonne
Total plant costs are $633 million (1471 $/kW) for the air-fired reference plant,
$723 (2084 $/kW) million for the cryogenic ASU O2-PC, and $953 (2057 $/kW)
million for the OITM O2-PC. Even though the OITM O2-PC has a total plant cost
that was 32 per cent higher than the ASU O2-PC, its higher power output results
in a slightly lower $/kW cost of $2,057/kW versus $2,084/kW.
The levelized cost of electricity (COE) was calculated for each of the plants
assuming an 85 per cent capacity factor. The COE value is made up of
contributions from capital cost, operating and maintenance costs, consumables,
and fuel costs. The levelized COE was calculated to be $50.41/MWhr for the
reference plant, $66.17/MWhr for the cryogenic ASU O2-PC, and $63.48/MWhr
for the OITM O2-PC. Again, because of its higher output, the OITM O2-PC has a
lower levelized cost of electricity ($63.48/MWhr versus $66.17/MWhr) and a
lower CO2 mitigation cost (MC) ($16.77/tonne versus $20.23/tonne) than the
ASU O2-PC.
The addition of weld overlay increases the COE approximately 0.75%, but
reducing the pipeline pressure from 3000 psia to 2200 psia (specified in Ref. 14)
reduces the COE by 0.3%. Thus, the combined effect of these two adjustments
on the COE is relatively small (+0.4% for cryogenic O2-PC and +0.1% for the
OITM O2-PC).
Compared to the COE of the supercritical cryogenic O2 PC, the COE for the other
technologies is 52% higher for Air PC, 35% higher for NGCC, 15% higher for
IGCC, and 5% higher for the subcritical O2PC, and 4% lower for the supercritical
O2PC with OITM. Compared to the MC of the supercritical cryogenic O2 PC, the
15
MC for the other technologies is 238% higher for NGCC, 192% higher for Air PC,
25% higher for IGCC, 5% higher for the subcritical O2PC, and 17% lower for the
supercritical O2PC with OITM.
16
3.0 Experimental
This work performed for this report was performed utilizing computer program
simulations. No experimental equipment was used.
17
4.0 Results and Discussion
4.1
System Design and Analysis
4.1.1 Reference Site and Conditions
In December 2000, Parsons published a study of the cost of electricity of several
case studies of CO2 sequestration from a PC boiler by post-combustion capture
(Ref. 1). In September 2005, Foster Wheeler released a report of conceptual
design of O2-fired PC boiler (Ref. 3). To provide a consistent comparison with the
cases analyzed in the previous reports, the same site conditions (59ºF, 14.7 psia,
60% RH) and the same fuel (Illinois #6) were used. Site Conditions and fuel
properties are presented in Table 4.1.1. Fuel HHV and LHV were estimated by a
DuLong’s method and the stoichiometric air ratio of 867 lbair/lbcoal was calculated
based on the fuel ultimate analysis.
Table 4.1.1 - Site Conditions
Standard site:
elevation, ft
amb p, psia
amb T, F
amb T, wet, F
RH, %
P-H2O, psia
Y-H2O, %v
condenser P, "Hg
0
14.70
59
51.5
60
0.247
1.010
2.00
air, %v
N2
O2
Ar
CO2
H2O
sum
dry
78.085
20.947
0.935
0.033
0.000
100.000
wet
77.297
20.735
0.926
0.033
1.010
100.000
coal, Ill#6
C
H
O
N
S
A
M
V
F
sum
fuel HHV
given
aspen
%w
63.75
4.5
6.88
1.25
2.51
9.99
11.12
34.99
44.19
100.0
sorb
CaCO3
%w
100
btu/lb
11666
11631
The CO2 fluid produced from the oxygen-based PC power plant is not chemically
pure, but can be readily sequestered in geologic formations (depleted oil and gas
reservoirs, unmineable coal seams, saline formations, and shale formations) or in
oceans. The liquid CO2 exits the plant at over 2000 psia. The CO2 fluid inside
pipeline under this pressure is in a liquid or a supercritical state. The other gases
in the delivered CO2 are limited to H2O < 50 ppm (to avoid acid corrosion), and
Ar+N2 < 3% (to avoid phase separation). The excess gases in CO2 stream either
have to be purged or recycled. However, since SO2, as an acid gas, similar to
CO2, it can be sent to pipeline directly under moisture free condition, and as
mentioned in literature, it does not need to be separated out from CO2 product.
Furthermore it is also not necessary to remove the small concentration of NOx in
the CO2 effluent since it can be sequestered along with the CO2.
18
4.1.2 Air-Fired Reference Case (case-1)
To study the effects of CO2 removal on the performance of the power plant, an
air-fired supercritical PC boiler was been simulated in detail as a reference case.
This model was used as the base, which was then extended to include the air
separation unit (ASU) and CO2 compression unit for O2-fired PC cases.
The reference plant employs a supercritical steam turbine with conditions,
4035psia/1076ºF/1112ºF/2.0”Hg, fires high-volatile bituminous coal, and
produces 460 MWe at the generator. It employs eight feed water heaters to raise
the final feed water temperature to 569ºF. It uses an auxiliary steam turbine to
directly drive the high-pressure feed water pump.
Case 1 is the reference air-fired supercritical PC boiler case, and the model and
results shown in Figure 4.1.1 and Table 4.1.2.
Table 4.1.2 - Setup and Assumptions
Setup and Result
ST
Result
main P, psia
main T, F
RH P, psia
RH T, F
FWHs
end wet, %
end P, "Hg
4035
1076
823
1112
8
9.5
2.0
F
5
10
569
FWH
TD
DC
FW T
DeSuperheat
SH, %
water T, F
5
569
Boiler
PA, %
UBC, %
radiation/margin, %
EXA, %
flame T, F
stack T, F
blowdown, %
miller exit T, F
20
1.0
0.59
17.9
3685
289
0
219
net power, MWe
net eff, %
gross @ST, MW
aux power, MW
as %
HHV in, mmbtu
Q to st, mmbtu
Q, cond, mmbtu
boiler eff, %
ST cycle eff, %
Generator eff, %
Flow
air, klb
coal, klb
sorb, klb
flue gas, klb
O2, %
H2O, %
CO, ppmv
NOx, ppmv
SOx, ppmv
after FGD
Ash, klb
C, %
main st, klb
RH st, klb
end st, klb
Aux power
430
39.5
477
46.0
9.7
3721
3281
1696
88.2
47.7
98.3
condensed water pump
HP feed water pump
FGD pump
CT pump
PA Fan
SA Fan
ID Fan
FGD Fan
cooling tower Fan
coal handling
sorb handling
ash handling + ESP
others (=1%)
3270
319
26
3556
3.0
8.6
14
22
1979
37
33
6.0
2950
2406
1573
total
MWe
0.6
17.5
0.8
3.4
1.1
2.0
5.4
4.1
1.9
2.0
0.8
1.8
4.6
46.0
FGD & SCR
L/G
Ca/S
Excess air, %
NH3/NOx
DeSOx, %
DeNOx, %
10
1.05
85
1.0
98
90
sum
"H2O
4.0
6.0
10.0
20.0
FSH
FRH
RH
PSH
UECO
ECO
AAHX
Damper
BHG
FGD
sum
"H2O
0.5
0.5
1.0
0.7
0.3
2.0
1.6
4.3
5.5
12.0
28.4
dP-air
AAHX
duct
nozzle
dp-gas
dP-Fan
PAFan
IDFan
SAFan
eff
FDFan
IDFan
CWPump
BFPump
Motor/mech
The Aspen Plus model includes coal mills, flue gas heater, pulverized coal-fired
furnace, steam generator, superheater, reheater, economizer, ash-removal unit,
nitrogen oxides (NOx) selective catalytic reactor (SCR), flue gas de-sulfurization
reactor (FGD), air blower, induced draft (ID) fan, feed water pump, cooling water
19
"H2O
60
28
20
%
75
70
80
80
95
pump, feed water heaters, and a single reheat steam turbine. A coal drying
function has been modeled and added into mill module to produce the correct
mill exit gas temperature. The furnace was simulated by a zero dimensional
model for heat and mass balances. However, all key tube banks of the heat
recovery area (HRA) were individually modeled. The furnace roof heat absorption
was also simulated. The high-pressure steam temperature is controlled by water
spray for de-superheat. The simulation also included heat losses from the boiler
and HRA enclosure, as well as from the steam pipes. Some user-defined models
were included to perform emission calculations. User built-in calculations have
been added to determine boiler efficiency, system net efficiency, and net power.
The heat carried by exhaust streams was automatically calculated by the
program.
For a given steam turbine output and a fuel, Aspen Plus iterates to determine the
feed rates of air, coal, etc., based on specified temperature approaches and
excess air requirement.
The system configuration, detailed setup parameters and summary of results for
the case 1 reference case are shown in Figure 4.1.1 and Table 4.1.2. The
system has a steam turbine cycle efficiency (generator power divided by heat
transferred to the steam cycle) of 47.6%, a boiler efficiency (heat to steam cycle
divided by heat input from fuel to boiler) of 88.2%, an unburned carbon loss
(UBC) of 1.0%, and a net plant efficiency of 39.46% (net plant heat rate of 8647
Btu/kWh). It has a gross power of 460 MWe at the generator, 18 MWe from
auxiliary steam expander, an auxiliary power of 46 MWe, and a net power of 430
MWe. Total heat input from the fuel is 3720 MM Btu/hr.
The temperature of the flue gas exhausted to the stack is 289ºF. The flue gas
exiting the boiler contains 3.0%, vol., wet O2 (18% excess air) and contains 739
klb/hr (1.72 lb/kWh) of CO2. This 3.0% O2 level is kept constant for all of the O2fired cases. A SCR is applied to control NOx with NH3/NOx=1.0, and an FGD is
used to control SOx by lime solution with Ca/S=1.05, L/G=10, and 85% excess
air for aeration.
The breakdown of auxiliary power for case 1 is listed in Table 4.1.2. Most of
these power consumptions were simulated directly by the Aspen module. Some
required user Fortran for those processes lacking Aspen modules, such as solids
handling. The power consumption was based on stream flows and design data.
Fan power consumption was simulated based on the pressure drops from both
air side and gas side. The total auxiliary power consumption, including FGD, for
case 1 is approximately 9.7% of the gross power, while it was about 9.2% for a
subcritical case.
Because of high temperature and high CO2 concentration, the CO concentration
is high, producing a flame temperature, which is lower than the adiabatic
combustion temperature. To represent this effect on boiler design, an estimation
of the equilibrium flame temperature was modeled.
20
Figure 4.1.1 - Reference Case of Air-fired Supercritical PC
Case: O2-PC-SC-01,
06/03/2005
1083
1076
4269.0
4035.0
2950
48
PC Boiler
1113
1112
1112
849.0
823.2
823.0
2406
2406
2406
2950
LS1
1076
823.0
4035.0
2389
67
43
835
PRPB
Q=36
Q=137
2950
1585
1567
14.7
14.7
3588
3588
3588
G5
G6
1869
FRH
14.7
44
4595.0
890.0
2803
2406
72
1904
653
871
1305.3
890.0
353.2
3
2878
5
2679
SL6
G8
SL2
21
45
1076
2803
V7
11
890.0
435
323
863.3
6
522
89
80
V6
713
88
194.8
56
114
63
199
569
Q=335
60
V5
522
65
4929.0
2950
FWH8
148
382
218
38.9
36.9
127
127
99
226
92
25
10.0
3.9
5
79
SSR
9
FSPLIT
10.0
15.2
15.8
59
644
155
234
20
348
122
1
V4
VD
425
522
4954.0
15
610
114
4979.0
2950
FWH7
Q=0
181.2
386
66
V3
94
191.6
2950
FWH6
317
2226
429
V1
V2
712
430
342.6
863.3
3.7
77
78
262
545
36.9
93.8
257
127
191.6
FWH4
69
2226
214
208
226
15.2
206.6
70
2226
FWH3
145
71
150
COND
3.7
14
429
221.6
348
Q=-1696
2226
FWH2
FWH1
87
3588
G11
838
868
1651
81
G13
85
74
Q=311
396
1758
851
33
LECO
100
109
STK1
569
34
60
14.7
N2
46
60
15.0
14.7
SORB
AIR
0
60
20.0
14.7
G15
3588
2803
99
Duty (MMBtu/hr)
109
L2
90
67.0
FGD
PREC
3555
Mass Flow Rate (Mlb/hr)
0
W=0
POND
15.1
0
Temperature (F)
Pressure (psi)
67.0
60
DIST
Power(MW)
0
AC
14.7
DENOX
LIQ
NH3
EQ
35383
G20
99
L1
308
289
724
G16
3556
14.1
3589
3589
ESP
AAHX
0
G18
15.1
G21
3588
W=5
580
0
GBY
80
14.9
14.8
58.4
231.6
0
0
0
0
49
AV
15.3
Q=-407
2720
NOZ
60
14.7
A4
A2
550
89
289
16.9
14.1
ASH
550
C3
308
C1
33
COL1
60
14.7
16.6
60
A5
16.7
869
MILL
319
580
A3
Q=0
14.7
3270
2720
W=1
0
A1
AIRB
70
308
15.4
15.1
BYP
2720
380
SEP1
C5
915.2
COMP3
367
W=0
0
C2
0
1200.0
PIPE
926.2
C10
0
0
CCT
14.8
0
0
Q=-0
14.9
D1
0
46
15.0
0
GRH
RY1
O2
21
COL3
COL2
Q=-0
90
D2
80
58.4
231.6
D3
0
231.6
D4
50
925.2
80
80
0
0
368
W=0
233.6
C7
COL-4
14.8
RY2
50
DEM
Q=-0
Q=0
W=2
59.4
W=0
0
90
257
SA
COMP2
COMP1
90
Q=-0
SAFAN
SPA
16.7
550
COAL
60
80
231.6
0
14.9
PA
CO2
W=-0
C6
C4
15.1
VANT
PAFAN
A6
Q=0
0
EXPD
RECY
3589
Power(MW)
67.0
80
90
W
-146
308
90
Duty (MMBtu/hr)
-183
IDFAN
G17
14.7
ASU
Q=-0
W=1
SPF
3556
G19
3685
Q
Q=-0
35578
15.1
289
14.3
14.3
E1
A0
14.7
SLV
RGIBBS
SP1
53
W=1
401
COMP
SLV
35578
308
0
219
CP
0
26
29.7
720
14.4
50
14.7
35552
3750
Q=128
41
G2
2226
64
BYFW
FGD
14.7
478
WETECO
68
4822.0
2803
1.0
0
0
2950
W=18
2226
102
236.6
206.6
181.2
7
Q=95
102
Q=0
2950
SPQ
35
4770.0
14.7
Q=141
4839.0
606
4787.0
638
Q=-0
86
2803
Q=-1751
BFP1
52
569
14.5
374
610
FWV
3588
16
Q=110
269
DEA
342.6
179
G14
61
Q=135
Q=118
2950
DAMP
HEATER
73
102
236.6
Q=162
32
22
BOILER
93
4904.0
14.6
14.7
QB
83
569
864
14.6
RSTOIC
267
93.8
152
57
Q=117
Q
FLAME
545
93.8
95
527
1266.2
14.7
G1
AUX
76
353.2
18
746
SPRAY
4822.0
E2
644
93.8
871
199
1758
UECO
24
3552
40
17
654
112
RHB
2185
EL4
14.6
G12
2406
2803
532
1266.2
1.0
1573
219
EL3
EL2
SL3
658
199
4510.0
1651
857.0
46
750
1305.3
PW
Q=315
2406
102
1652
108
W=-18
879
Q=-36
883
30
1775
EI3
657
1567
4665.0
G3
219
877.0
Q=210
3.9
1.2
96
G9
854
2950
29
152
1874
LS4
23
14.7
810
EI2
EI1
28
234
15.8
93.8
EH2
PSH
4500.0
545
2187
DPCRH
G4
Q=389
8
1758
SPG
Q=-39
Q=310
38.9
93.8
2301
SC1
3588
DIVW
6
2389
LP4
LP3
382
540
194.8
22
14.7
G10
LP2
LP1
IP3
713
750
1487
Q=-42
G7
Q=-19
14.7
IP2
IP1
HP2
658
EH1
SC23
FSH
Q=1304
WEG
MULT
88
SC4
47
EVAP
W
WE
84
HP1
42
17
2803
ROOF
964
4445.0
459
823.0
SL5
2878
27
19
26
1112
1112
MAIN
DPHRH
830
4630.0
ST=4035/1076/1112/2.0"Hg=460 MW
0
915.2
Q=0
0
Q=-0
0
C11
D5
PCO2
0
4.1.3 Oxygen-Based PC Plant – Base Case
4.1.3.1
Boiler Plant Modifications
The oxygen-based (or oxygen-fired) plant model contains essentially all the
components in the conventional plant model. In addition, it also includes an air
separation unit (ASU) and a flue gas cooler. In the O2-fired plant, the FGD is not
needed because the SO2 is acid gas similar to CO2 and can thus be sent to
pipeline together with the CO2. A substantial portion of the SOx will be removed
as the flue gas is cooled down in the CO2 cooling and compression equipment.
The steam side components remain very similar to the air-fired case with only
some changes in heat bundle duties in the heat recover area (HRA).
In O2-fired cases, flue gas is recycled to control the flame temperature inside the
PC-fired boiler to minimize NOx formation, minimize ash slagging in the furnace
combustion zone, and avoid the application of exotic materials.
Before the flue gas is separated into a recycled and effluent stream (to the
pipeline), it is cooled to 90ºF. Since this is below the acid/moisture dew point, a
heat exchanger containing acid-resistant materials must be used. The recycled
gas is then reheated, before the forced draft (FD) fan, by mixing it with a
bypassed hot gas to avoid reaching the dew point. After the O2 from ASU plant is
mixed with recycled flue gas, it is heated by the flue gas exiting the boiler in a
gas-gas heat exchanger, which acts as a recuperator to improve cycle efficiency
and reduce fan power requirements.
It is assumed in this study that there is no tramp air ingress through the sealed
boiler.
4.1.3.2
Air Separation Unit
For an O2-fired PC, O2 purity is a key parameter for system performance and
economics. A high purity O2 will produce a high purity of product CO2 gas, which
will reduce CO2 purification and compression power. However, producing high
purity O2 requires high ASU plant operational and equipment costs. Furthermore,
too high O2 purity is not necessary because the fuel combustion itself will
generate some gases, such as N2 and some excess O2 is required for complete
combustion, in additional to CO2 as flue gas. Therefore there is a balance point
to give an optimum.
The method of air separation chosen for this study is the commercially available
large-scale cryogenic air separation technique. A traditional cryogenic ASU plant
was simplified in the simulation to include the power consumption, but without
details of distillation columns and cold heat exchangers. The Aspen model does
not include the air purifier, which removes moisture, hydrocarbons, CO2, and
NOx in an adsorber and is located between the cold box and air compressor.
22
Although the separated N2/Ar gases could potentially be sold as byproducts, no
economic credit for this is taken in this study. No heat recovery from the ASU air
compressor inter-stage coolers is included, because recovery of this low grade
heat is inefficient.
Power consumption for different O2 purity is plotted in Figure 4.1.2. For the O2
purity of 99.5% used in this study, a power consumption of 24.5 kWh/klbair is
applied. For a 460 MWe steam turbine generation, the ASU plant consumes
about 70 MWe, or 15% of generated power, which is a large penalty for CO2
removal.
Figure 4.1.2 – Oxygen purity effect on ASU specific power
Power Demand vs Oxygen Purity
0.30
0.29
0.28
0.27
Pow e r, kWh/k gO2
0.26
0.25
0.24
0.23
0.22
0.21
0.20
0.19
0.18
0.17
0.16
0.15
70
75
80
85
90
95
100
Oxyge n Purit y, %v
Advanced air separation methods such as high temperature membrane method
will be incorporated into the cycle in task 2 (Section 4.2) to reduce both operation
and capital costs.
23
4.1.3.3
CO2 Compression Unit
The flue gas effluent stream (mainly CO2) has to be compressed to the high
pipeline pressure of 1200 to 3000 psia depending upon end user specification.
The CO2 sequestration equipment is added to the system and the effluent is
conservatively compressed to 3000 psia. The dominant moisture in flue gas is
condensed out first during flue gas cooling before the first stage compression.
The condensed water contains acid gases and has to be treated before recycle
or discharge.
The flue gas composition after cooling but before the first stage CO2 compressor
from this base case (case-2) is:
CO2
90.4
O2
3.3
N2+Ar SOx
1.3
1.1
H2O
3.6
In literature, residual O2 as low as 1.3% was used. Reducing O2 content, such as
from 3.0% to 2.0% by reducing excess air, is helpful in reducing CO2
compression power, but it is judged that and oxygen content of approximately
3.0% is required for good combustion efficiency. Both CO2 and SOX are acid
gases. They combine with moisture to form acid, which causes a corrosion
problem along CO2 pipeline. Therefore, after the 2nd stage, a chemical method of
active dehydration with TEG (Triethyleneglycol) is applied to remove the rest of
moisture to a very low level of less than 50 ppm, where the TEG can be
regenerated by heating. In the model, the TEG dehydration was simulated, but
the TEG itself was not simulated.
A four-stage compression with inter-stage cooling was applied. To reduce power,
an equal compression pressure ratio of approximately 4.0 was applied.
4.1.3.4
O2-fired PC: Base Case (case-2)
Figure 4.1.4 presents the base case system model. The key parameter for the
base case is that the adiabatic temperature in the boiler was kept the same as
the reference air case (at about 3690ºF) by adjusting the recycle gas flow. Both
fuel and air feed rates were iterated to match the heat duty and 3.0%v exit O2
level. Compared to reference air-fired case, the case 2 air flow rate is reduced by
13.4% (from 3270 klb/hr to 2832 klb/hr), O2 concentration to the boiler is 28.3%
(compared to 20.7% for the air-fired case) yielding an increased combustion
efficiency and a lower UBC (unburned carbon). The relation of UBC and O2% has
been included in modeling as shown by Figure 4.1.3 (based on FW-FIRE
modeling results).
24
Figure 4.1.3 - Carbon burnout vs. O2 concentration to boiler
Carbon Burnout Function
100.0
99.8
Carbon burnout, %
99.6
99.4
99.2
99.0
98.8
0.0
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
O 2 fraction
The oxidant from the ASU was 99.5% O2 purity. ASU and CO2 plant power
consumptions were 68.7 and 38.4 MWe respectively. This resulted in a total
auxiliary power of 144.5 MWe, or about 30.4% of gross power. All CO2 generated
from fuel combustion was 100% removed with a final CO2 purity of 93.8%v. The
efficiency penalty for CO2 removal was 8.7% in points with efficiency drop from
39.46 to 30.81%, and the related electric power penalty was 129 kWh/klbCO2. The
extra cooling duties for the CO2 stream were added into cooling tower calculation
for auxiliary power.
The volumetric flow in the O2-fired PC case was reduced from 179 to 125
MMft3/hr, or to 70%, due to a high gas molecular weight of CO2 instead of N2.
The gas mass flow was reduced from 3555 to 3363 klb/hr, or 95% in comparison
to the air-fired case. The recycled gas flow was 2425 klb/hr, at a temperature of
95ºF, which has been heated 5ºF by a bypass flue gas stream as shown in
Figure 4.1.4. It is noted that compared to the air-fired case, the coal feed rate
was reduced from 319 to 315 klb/hr as the result of low UBC, and low excess air.
25
Because of high CO2 content in flue gas at high flame temperature, part of CO2
dissociated and more CO slipped, the equilibrium flame temperature differed
from adiabatic temperature.
26
Figure 4.1.4 - Base case of O2-fired supercritical PC
ST=4035/1076/1112/2.0"Hg=460 MW
Case: O2-PC-SC-02, 06/03/2005
Temperature (F)
Pressure (psi)
Mass Flow Rate (Mlb/hr)
1083
1076
4269.0
4035.0
2950
2950
48
PC Boiler
1113
1112
1112
849.0
823.2
823.0
2406
2406
2406
Duty (MMBtu/hr)
1112
MAIN
LS1
DPHRH
830
1076
823.0
4035.0
2389
67
43
PRPB
Q=36
Q=137
1867
FRH
14.7
47
1580
1562
14.7
14.7
3394
3394
G5
G6
750
4595.0
890.0
1305.3
2803
2406
3394
3
2878
5
2679
2389
382
93.8
38.9
8
545
2187
EI1
G8
219
Q=315
2803
6
890.0
435
323
863.3
522
89
80
152
10.0
3.9
5
57
79
SSR
76
9
FSPLIT
353.2
713
88
194.8
18
56
114
92
545
267
382
93.8
93.8
38.9
36.9
127
127
99
226
25
218
V5
746
199
522
234
15.8
59
155
644
15.2
10.0
20
348
122
1
569
4822.0
Q=333
60
65
4929.0
2950
FWH8
148
425
522
4954.0
66
386
610
15
114
4979.0
2950
FWH7
V3
712
Q=0
181.2
342.6
863.3
191.6
2950
FWH6
2226
150
36.9
93.8
257
127
191.6
94
429
V1
V2
262
545
317
3.7
77
78
V4
VD
430
527
1266.2
63
14.7
95
V6
V7
SPRAY
24
69
226
2226
FWH4
214
208
15.2
206.6
70
2226
FWH3
3.7
145
71
COND
14
429
221.6
348
FWH2
Q=-1696
2226
FWH1
87
Q=109
3394
83
848
849
1561
QB
85
32
2950
1663
Q=135
Q=118
396
610
14.5
G14
33
LECO
2803
60
SORB
G15
700
60
14.4
20.0
3394
0
60
14.7
AIR
2177
400
67.0
COMP
2832
W=69
0
317
60
120
L2
90
67.0
POND
FGD
PREC
0
LIQ
NH3
2832
DIST
AC
14.3
14.3
E1
G16
3394
3394
302
14.1
3363
3363
ESP
AAHX
15.1
IDFAN
G19
3690
G18
15.1
G21
580
3363
2530
Q=-348
NOZ
95
14.7
A4
A2
550
119
302
16.9
14.1
ASH
550
90
90
90
14.9
14.8
58.4
788
779
RECY
A5
60
95
32
16.6
14.7
14.7
865
16.7
580
0
COL1
SEP1
550
SPA
AIRB
A1
317
15.1
C5
BYP
2530
SA
C2
95
90
14.8
2375
D1
150
46
15.0
Q=0
RY2
RY1
O2
27
D2
323
3000.0
W=7
926.2
C10
90
90
90
58.4
231.6
9
D3
776
776
PIPE
Q=-43
Q=-49
90
14.9
GRH
W=1
W=10
779
COL3
COL2
Q=0
655
A3
C7
Q=-1
50
COMP4
281
233.6
W=11
788
Q=-341
14.8
2425
COAL
COMP3
313
59.4
W=11
3312
90
CO2
DEM
777
310
90
1875
SAFAN
16.7
MILL
315
W=1
12
COMP2
COMP1
15.4
776
90
231.6
3312
14.9
60
916.2
C6
C4
C3
317
PA
217
90
0
ASU
15.1
C1
RSTOIC
Q=0
-183
AV
90
231.6
776
GBY
PAFAN
A6
Q=-250
Q=-3
3163
15.3
473
14.7
3395
W=4
A0
0
317
G17
14.7
L1
W=0
SPF
3394
98
14.7
SLV
317
302
700
0
G20
EQ
RGIBBS
SP1
53
2832
14.7
DENOX
CP
W=1
29.7
15.1
0
N2
46
15.0
0
SLV
97
FGD
Q=135
41
G2
14.7
50
14.7
0
0
4822.0
34
2226
64
Q=0
14.7
14.7
STK1
569
35
2803
0
1.0
WETECO
68
100
2
COL4
231.6
D4
2226
236.6
102
BYFW
SPQ
637
0
2950
W=18
120
4770.0
Q=95
102
206.6
181.2
Q=-0
2950
2803
Q=-1755
Q=141
4839.0
4787.0
3394
374
86
52
569
608
BFP1
Q=110
269
DEA
342.6
843
HEATER
102
16
236.6
Q=311
FWV
DAMP
61
7
Q=162
170
23
BOILER
73
93
4904.0
14.6
G13
14.7
81
74
569
838
14.6
G11
Q
FLAME
AUX
112
UECO
2185
93.8
871
199
1663
RHB
2803
11
1266.2
40
644
14.6
G12
2406
46
532
199
4510.0
45
1065
Q=-36
857.0
4665.0
EL4
17
654
SL3
658
1305.3
21
1.0
1573
219
EL3
EL2
2406
879
Q=210
102
1652
W=-18
750
PW
30
108
1.2
LS4
SL2
657
1561
882
1775
1874
96
1562
854
29
3.9
877.0
23
2950
28
152
EI3
EH2
PSH
4500.0
G1
W EG
MULT
LP4
LP3
234
15.8
93.8
DPCRH
1663
SPG
14.7
E2
W
LP2
LP1
540
2301
14.7
G10
Q=311
3481
6
194.8
EI2
22
1482
Q=-42
G7
G9
810
353.2
890.0
SL6
Q=-39
Q=1306
Power(MW)
713
871
EH1
G4
Q=389
72
IP3
IP2
IP1
653
SC1
3394
DIVW
658
SC4
Q=-19
14.7
G3
Duty (MMBtu/hr)
WE
HP2
835
44
SC23
1902
FSH
EVAP
459
Q
W
88
84
HP1
42
17
2803
ROOF
964
4445.0
2950
Power(MW)
823.0
SL5
2878
27
19
26
4630.0
1112
1
Q=-55
D5
916.2
0
4.1.4 Parametric Cases
Various cases were simulated to evaluate the effects on O2-PC performance of
different designs and operating conditions, as shown by Table 4.1.3, in which the
red highlight shows the key parameter changed in comparing with previous
cases. All these cases, including part load cases (100% to 25%), will be
explained and discussed in details in the following sections.
In Table 4.1.3, the efficiency drop was directly calculated from the difference
between cases with and without CO2 removal. But in general, the specific power
penalty for the CO2 removal cannot be calculated by difference directly, because
extra power may be produced by firing more when integrated with different O2
separation techniques. In some cases, the extra power by added firing may be
greater than the power required for CO2 removal. Therefore a new definition,
here, is introduced as
kWh/lbCO2 = (eff drop)*(power/efficiency)air /lbCO2
(1)
Equation (1) is the way to compare penalty for CO2 removal for different systems
without cost estimation, especially for a complex system such as O2-PC
integrated with a high temperature oxygen separation membrane method, where
even more net power is produced from with CO2 removal than without.
28
Table 4.1.3 - Case Summary
case
Waste heat economizer
HRA Arrangement
CO2 Condensation
Vent Gas Recycle
Coal flow
Oxidant flow
Air flow
O2 purity
01
no
parallel
klb/hr
319.0
klb/hr
3270
klb/hr
3270
%
klb/hr
0
Recycle gas flow
%
0.0
Recycle gas temperature
F
Boiler Inlet O2
%, v
20.7
Boiler Outlet O2
%, v
3.0
MM cf/hr
179
Boiler outlet flue gas flow
klb/hr
3555
Adiabatic Temperature
F
3685
Equilibrium Temperature
F
3552
Gas Temp. to FSH
F
2185
Water temp. to evap.
F
638
UBC
%
1.00
Boiler Efficiency
%
88.16
Pipeline pressure
psia
0
Generated CO2 flow
klb/hr
739
CO2 purity
%,v wet
14.0
Removed CO2 flow
klb/hr
0
CO2 removal efficiency
%
0
CO2 purity
%
Gross Power
MW
476.2
Auxiliary Power
MW
46.0
ASU power
MW
0.0
CO2 compression power
MW
0.0
Net Power
MW
430.2
Net Efficiency
%
39.46
Efficiency Drop
% pts.
CO2 removal energy
kWh/klbCO2
-
02
no
parallel
no
315.0
655
2832
99.5
2425
72.1
95
28.3
3.0
125
3363
3690
3481
2185
638
0.49
89.28
3000
732
81.3
732
100
93.8
476.2
37.4
68.7
38.4
331.7
30.81
8.7
129
03
no
parallel
yes
315.0
655
2832
99.5
2425
72.1
95
28.3
3.0
125
3363
3690
3481
2185
638
0.49
89.28
3000
732
81.3
732
100
93.8
476.2
37.6
68.7
34.7
335.2
31.14
8.3
124
04
no
parallel
yes
no
312.0
649
2806
99.5
2429
72.4
138
27.8
3.0
126
3357
3678
3474
2185
638
0.51
90.14
3000
725
80.0
725
100
93.8
476.2
37.6
68.1
34.4
336.1
31.53
7.9
119
05
no
parallel
yes
no
310.7
647
2798
99.5
2046
68.8
142
30.8
3.0
109
2972
4010
3669
2185
615
0.39
90.52
3000
723
78.7
723
100
93.7
476.2
37.1
67.9
34.3
336.9
31.73
7.7
117
06
no
parallel
yes
no
309.4
645
2790
99.5
1751
65.5
146
33.8
3.0
96
2675
4321
3830
2185
597
0.30
90.90
3000
720
77.5
720
100
93.7
476.2
36.6
67.7
34.2
337.7
31.94
7.5
114
07
no
parallel
yes
no
308.0
644
2784
99.5
1162
55.8
163
42.2
3.0
75
2083
5178
4182
2185
592
0.15
91.31
3000
718
74.2
718
100
93.6
476.2
36.9
67.6
34.2
337.5
32.07
7.4
112
29
08
yes
parallel
yes
no
308.0
644
2784
99.5
1162
55.8
163
42.2
3.0
75
2083
5178
4182
2185
592
0.15
91.31
3000
718
74.2
718
100
93.6
483.1
37.5
67.6
34.2
343.9
32.67
6.8
103
09
yes
parallel
yes
no
308.0
644
2784
99.5
1166
55.9
260
41.1
3.0
77
2087
5104
4161
2185
596
0.16
91.31
3000
718
71.3
718
100
93.6
486.4
38.1
67.6
34.2
346.6
32.93
6.5
99
10
yes
parallel
yes
no
308.0
644
2784
99.5
1166
55.9
260
41.1
3.0
77
2087
5104
4161
2185
596
0.16
91.31
2000
718
74.2
718
100
93.6
486.4
38.1
67.6
32.9
347.9
33.05
6.4
97
p100
yes
series
yes
no
308.0
644
2784
99.5
1166
55.9
259
41.1
3.0
75
2087
5104
4161
2450
593
0.16
91.31
3000
718
71.3
718
100
93.6
486.5
37.7
67.6
34.2
347.0
32.97
6.5
99
p72
yes
series
yes
no
229.4
480
2074
99.5
1151
62.7
225
35.5
3.0
61
1837
4509
3919
2146
574
0.27
3000
535
73.9
535
100
93.5
349.0
26.1
50.3
25.5
247.1
31.52
7.9
0
p50
yes
series
yes
no
168.0
350
1513
99.5
1072
68.2
205
30.9
3.0
49
1572
4001
3672
1918
553
0.37
3000
391
76.2
391
100
93.7
239.2
16.5
36.7
18.6
167.4
29.16
10.3
0
p25
yes
series
yes
no
98.4
206
889
99.5
697
70.3
206
29.1
3.0
26
991
3723
3514
1458
495
0.43
3000
230
77
230
100
93.6
119.3
7.2
21.6
10.9
79.6
23.67
15.8
0
6A
no
series
yes
no
309.4
645
2790
99.5
1751
65.5
146
33.8
3.0
96
2675
4321
3830
2450
666
0.30
90.90
3000
720
77.5
720
100
93.7
476.2
36.6
67.7
34.2
337.7
31.94
7.5
114
6B
no
series
yes
no
309.4
645
2790
99.5
1751
65.5
146
33.8
3.0
96
2675
4321
3830
2185
603
0.30
90.90
3000
720
77.5
720
100
93.7
476.2
36.6
67.7
34.2
337.7
31.94
7.5
114
9A
yes
series
yes
no
308.0
644
2784
99.5
1166
55.9
260
41.1
3.0
77
2087
5104
4161
2450
591
0.16
91.31
3000
718
71.3
718
100
93.6
486.4
38.1
67.6
34.2
346.6
32.93
6.5
99
9B
yes
series
yes
no
308.0
644
2784
99.5
1166
55.9
259
41.1
3.0
75
2087
5104
4161
2450
594
0.16
91.31
3000
718
71.3
718
100
93.6
486.5
37.7
67.6
34.2
347.0
32.97
6.5
99
4.1.4.1
O2-fired PC with Liquid CO2 Pump (case-3)
Case-3, as shown in Figure 4.1.5, was formed from case-2 by adding the CO2
condensing process into CO2 plant. The CO2 gas is compressed first by a threestage compressor to over 900 psia, and then is cooled down by cooling water to
its dew point. Most of the CO2 is condensed out as liquid, some of the CO2 stays
in gas by equilibrium partial pressure, where the gas dew point reduces from 76
to 55ºF during condensation. Both the gas stream and the liquid stream from the
condenser are boosted to the pipeline pressure by a small gas compressor and a
liquid pump, respectively. Thus a potentially large gas compressor is replaced by
a small compressor and a liquid pump. The gas stream is cooled down to nearly
same temperature as the liquid, and then mixed with the liquid CO2 stream. The
final CO2 stream therefore consists of the same liquid compositions as that of the
case-2.
In this way, the last stage of compression is accomplished by liquid pumping.
The power saving is about 3.7 MWe, or 1.1% to the net power, while the
efficiency increases from 30.81 to 31.14%. The electric power penalty is reduced
from 129 in case-2 to 124 kWh/klbCO2 in case-3. Because of this power saving,
this option will be applied to most of CO2 compression processes for the other
cases.
30
Figure 4.1.5 - O2-fired supercritical PC with liquid CO2 pump (Case-3)
Case: O2-PC-SC-03,
ST=4035/1076/1112/2.0"Hg=460 MW
06/03/2005
Temperature (F)
Pressure (psi)
Mass Flow Rate (Mlb/hr)
1083
1076
4269.0
4035.0
2950
1112
1112
849.0
48
PC Boiler
Duty (MMBtu/hr)
2950
1113
2406
823.2
823.0
2406
2406
26
SL5
2878
27
19
Power(MW)
17
2389
4035.0
67
LS1
459
823.0
823.0
1076
MAIN
DPHRH
830
1112
1112
88
84
Q
Duty (MMBtu/hr)
W
Power(MW)
W
WE
EG
W MULT
4630.0
2803
ROOF
HP2
HP1
43
964
4445.0
835
PRPB
Q=36
Q=137
4595.0
2950
44
1867
1580
1562
14.7
14.7
3394
3394
3394
G5
G6
FRH
14.7
47
1902
3
2878
5
2679
SC4
LP1
6
93.8
2301
8
38.9
2187
EI2
23
219
G8
42
21
1266.2
863.3
199
522
89
80
713
88
194.8
92
93.8
127
218
38.9
25
127
36.9
99
59
226
234
155
15.8
15.2
122
348
9
10.0
20
1
863.3
4822.0
Q=333
60
65
425
610
2950
FWH7
15
386
Q=0
114
429
V1
V2
2226
2950
FWH6
36.9
93.8
94
150
262
545
317
191.6
4979.0
4954.0
2950
FWH8
148
66
V3
712
181.2
342.6
522
3.7
77
78
V4
VD
430
527
199
4929.0
257
127
69
208
226
FWH4
2226
70
14
429
221.6
348
2226
FWH3
71
145
15.2
206.6
191.6
COND
3.7
214
Q=-1696
2226
FWH2
FWH1
87
Q=109
G13
14.7
81
236.6
7
4904.0
14.6
849
1561
32
Q=118
Q=311
Q=162
2950
1663
170
QB
23
102
206.6
236.6
0
608
4787.0
WETECO
LECO
STK1
569
120
100
14.7
14.7
0
0
60
N2
46
60
15.0
14.7
2803
34
AIR
2177
2832
14.7
SORB
0
400
FGD
2803
Q=135
700
SLV
67.0
2832
W=69
0
317
20.0
3394
COMP
97
29.7
60
14.4
G15
60
POND
15.1
0
PREC
FGD
0
0
L2
90
67.0
120
2832
DIST
AC
14.7
DENOX
53
CP
W=1
4822.0
41
50
64
BYFW
2950
SPQ
35
637
G2
2226
68
4839.0
33
4770.0
14.7
1.0
0
2950
LIQ
NH3
0
98
G20
EQ
A0
14.7
SLV
RGIBBS
L1
0
317
700
SP1
E1
G16
302
14.3
14.3
3394
3394
302
Q=-3
W=0
SPF
IDFAN
G19
317
G17
14.7
G18
G21
W=4
3363
90
90
90
14.9
14.8
58.4
231.6
916.2
236
3163
788
779
776
20
3000.0
C4
C6
C9
C11
AV
55
0
20
RECY
C3
317
2530
14.7
90
ASU
GBY
15.3
473
3395
-183
15.1
3394
580
FLAME
3363
3363
ESP
AAHX
3690
G1
Q=-250
15.1
14.1
302
119
NOZ
95
14.1
16.9
14.7
A4
A2
550
A6
ASH
550
C1
3312
COL1
A5
217
60
95
16.6
14.7
14.7
865
60
580
16.7
16.7
MILL
315
550
W=1
A1
317
W=11
15.4
15.1
2530
BYP
SA
COL2
Q=0
926.2
C10
W=0
776
2375
D1
COL3
Q=-49
90
14.9
46
150
15.0
90
Q=-43
90
655
A3
323
W=10
779
97
COL4
14.8
14.8
2425
COAL
C7
C2
50
90
95
C8
233.6
W=11
788
Q=-1
SAFAN
SPA
313
59.4
C5
Q=-341
1875
0
AIRB
310
SEP1
COMP3
COMP4
90
14.9
3312
90
58.4
D2
9
231.6
90
231.6
D3
D4
Q=-115
55
3000.0
916.2
755
755
1
CPUMP
PIPEL
2
D5
GRH
W=1
RY2
CO2
DEM
777
COMP2
COMP1
PA
PIPEG
231.6
32
PAFAN
RSTOIC
Q=0
90
15.1
RY1
O2
Q=0
31
2226
102
Q=0
14.5
3394
2803
Q=-1755
269
W=18
569
843
G14
181.2
86
52
DAMP
Q=-0
610
FWV
HEATER
374
BFP1
Q=95
Q=141
Q=110
Q=135
DEA
396
342.6
BOILER
102
16
61
73
569
838
14.6
G11
Q
93
85
74
83
848
3481
SSR
644
382
267
95
V5
1266.2
522
14.7
E2
545
93.8
114
746
63
UECO
3394
Q=1306
5
79
FSPLIT
56
V6
V7
2185
57
76
353.2
18
10.0
3.9
AUX
112
569
G3
435
323
2803
SPRAY
24
EVAP
11
4510.0
1663
RHB
2803
890.0
532
199
2406
46
4665.0
152
6
14.6
G12
857.0
Q=315
810
45
1065
PW
40
644
93.8
871
658
1305.3
Q=-36
882
EL4
17
W=-18
879
1561
854
219
EL3
SL3
SL2
1562
4500.0
2950
EL2
2406
750
Q=210
1.0
1573
1.2
654
G9
14.7
102
1652
108
877.0
G4
Q=311
30
1775
96
657
PSH
3.9
1874
LS4
SPG
Q=-39
152
EI3
EH2
SC1
3394
DIVW
Q=389
29
234
15.8
93.8
DPCRH
1663
28
545
LP4
LP3
382
14.7
G10
LP2
540
EI1
22
1482
Q=-42
IP3
713
194.8
2389
SL6
G7
Q=-19
14.7
72
2406
IP2
871
353.2
EH1
SC23
FSH
1305.3
890.0
2803
IP1
653
890.0
750
658
W=3
4.1.4.2
O2-fired PC with Hot Gas Recycle (case-4)
A case with hot gas recycle (case 4) was run to evaluate its effect on the system
performance. A hot gas recycle brings more energy back into boiler, reduces fuel
and O2 feed rates, and reduces ASU duty, but it requires more power to the fan
because of the increased recycle gas volumetric flow.
As result, increasing the recycle gas temperature from 95ºF to 138ºF increases
the boiler efficiency from 89.3% to 90.1% and net efficiency from 31.14% to
31.53%. The resultant fuel saving is approximately 3.0 klb/hr or 1.0%. The size of
the gas-gas heat exchanger increases due to the reduction in LMTD (fluid
temperature difference is reduced from 212ºF to 166ºF for the hot end, and from
120ºF to 78ºF for the cold end).
There is a limit to increasing the recycle gas temperature without increasing the
stack gas temperature, which will reduce efficiency and increase cooling duty.
One option mentioned in literature is to raise both stack gas and recycle gas
temperatures, and then recover heat from stack gas to replace part of the
feedwater heaters.
Figure 4.1.6 - Temperature-Quality diagram for wet-end heat exchangers
T-Q Diagram
400
Col-1, HT
350
Temperature, F
300
250
Col-1, LT
200
150
100
50
0
-300
Col-4
-250
-200
-150
Col-2&3
-100
Duty, MMBtu/hr
32
-50
0
50
Figure 4.1.6 shows the cooling curves for flue gas cooling before each stage of
CO2 compressor. Part of the heat can be recovered to preheat the condensate. If
90 MMBtu/hr of heat is recovered (exit temperature = 142ºF), the steam saved
from extraction generates an additional 2.2 MWe, or an efficiency increase of
0.2% point. If a total 150 MMBtu/hr of heat is recovered, the steam saved from
extraction generates an additional 3.4 MWe, or an efficiency increase of 0.3%
points.
4.1.4.3
O2-fired PC with High Flame Temperature (case 4 to 7)
In cases of 4 to 7 the amount of flue gas recycle was reduced, to increase the
O2% level to the boiler. This increase in boiler O2% creates a higher adiabatic
temperature and less flue gas flow, which reduces the size of the furnace and
increases the boiler and overall cycle efficiency. The adiabatic temperature is the
maximum theoretical temperature that can be reached by the combustion with no
loss of heat and no dissociation. Actual flame temperature is lower than the
adiabatic temperature especially at adiabatic temperatures greater than 3600ºF
due to flue gas dissociation of CO2 to CO and O2 (Figure 4.1.7).
Figure 4.1.7 – Equilibrium and adiabatic temperatures vs. O2 concentration
Adiabatic and Equilibrium Temp. Vs. Inlet O2
5500
5000
Temperature, F
Adiabatic
4500
4000
Equilibrium
3500
3000
2500
15
20
25
30
O2 Percentage
33
35
40
45
The effect of adiabatic temperature on cycle efficiency is shown by Table 4.1.4,
where case-4 has nearly the same adiabatic temperature as air-fired reference
case, while cases 5 to 7 have higher adiabatic temperatures. Although there is
little change in gas exhaust flow to CO2 compressor among cases 4 to 7, the
decreasing recycle gas flow results in reduced auxiliary power consumption for
both the FD and ID fans. It is noted that both the total power and the auxiliary
power are relatively insensitive to the adiabatic temperature when it is greater
than 4320ºF. However the efficiency related indices such as net efficiency, coal
flow rate, and CO2 removal penalty improve with increased flame temperature
even when it is beyond 4320ºF.
Table 4.1.4 - Flame adiabatic temperature effect on performance
case
Adiabatic Temperature
F
Boiler Inlet O2
%, v
Coal flow
klb/hr
Boiler Efficiency
%
Total Aux Power, MW
MW
Net Power
MW
Net Efficiency
%
Generated CO2 flow
klb/hr
Efficiency Drop
% pts.
CO2 removal energy
kWh/klbCO2
01
3685
20.7
319.0
88.16
46.0
430.2
39.46
739
-
04
3678
27.8
312.0
90.14
140.1
336.1
31.53
725
7.9
119
05
4010
30.8
310.7
90.52
139.3
336.9
31.73
722.7
7.7
117
06
4321
33.8
309.4
90.9
138.5
337.7
31.94
720
7.5
114
07
5178
42.2
308.0
91.31
138.7
337.5
32.07
718
7.4
112
Table 4.1.5 shows the relationship of flue gas flow both in mass and in volume to
boiler adiabatic temperature. The air-fired data is shown for comparison. It can
be observed that the O2-fired PC has a lower volumetric flow rate than does the
air-fired PC due to the higher molecular weight of the flue gas (i.e. CO2 versus
N2).
Table 4.1.5 – Flame adiabatic temperature effect on boiler flue gas flow
case
Boiler outlet flue gas flow
Adiabatic Temperature
Equilibrium Temperature
01
179
3555
3685
3552
MM cf/hr
klb/hr
F
F
04
126
3357
3678
3474
05
109
2972
4010
3669
06
96
2675
4321
3830
07
75
2083
5178
4182
From case 4 to case 7 the ratio of the O2-fired PC volumetric flow rate to the airfired PC volumetric flow rate drops from 70% to 42%, which means for a constant
flue gas velocity, boiler size is reduced.
34
Another advantage of decreasing the quantity of recycle gas is the increase in O2
content in the boiler (from 27% to 42% by vol. from case 4 to case 7), which
improves the fuel combustion and reduces the required height of the furnace.
This credit from the reduction of UBC has been approximated in this system
study, while a more thorough treatment will be modeled in the 3-D CFD boiler
simulation study (Section 4.3).
4.1.4.4
O2-fired PC with Wet-End Economizer (case-8, 9)
One option to increase system efficiency is to use an economizer to recover lowgrade heat from compressor inter-stage coolers, and from the flue gas cooler
before the CO2 plant. Case-8 is such a case for integration of low grade heat
recovery (based on case-7), where the heat recovered was used for low pressure
feed water heating. A method of condensate split stream was applied. The steam
extractions from ST at different LP ports were adjusted to match the duties. This
option incurs no cost increase because all these heat exchangers were already
installed to cool the flue gas for the CO2 plant. The only difference is in that the
cold side cooling water was replaced by condensate from the ST. In this way, the
cooling water duty was reduced too.
The temperature-quality diagrams for these low-grade heat coolers are plotted in
Figure 4.1.6, where 150ºF, just above the flue gas dew point, was used as hot
exhaust temperature for heat recovery. This leads to a temperature approach at
the hot side of about 50ºF, which is sufficient for economizer design and sizing.
Because the tube wall temperature is below the flue gas temperature, part of the
moisture is condensed out on the wall. A wet-end heat exchanger design
therefore has to be considered for this application.
Additional cooling is required to reduce gas temperature from 150 to 90ºF by
cooling water before gas compression, and the condensate from gas moisture
needs to be drained out before the compressor.
The dominant change of this option is to produce more power from the generator
as the result of less steam extraction. The total power generated increased
substantially from 476 to 483 MWe, or 2% increase in net power. This leads to
better system performance as the electrical power penalty is reduced from 112 to
103 kWh/klbCO2, an 8% drop. The efficiency drop changed from 7.4% to 6.8% in
points. The net efficiency increased from 32.07% to as 32.67%, a net 0.6% point
increase.
The application of the wet-end heat exchanger to recover low-grade heat brings
the benefit that the temperature of the hot recycle gas can be further boosted.
This will cause a high gas exit temperature from gas-gas heat exchanger to ID
fan. The heat from this high temperature exit stream can be recovered by an
economizer, which means more steam condensate will be extracted and sent to
economizer to pick up more low-grade heat. Because of temperature approach
limitation at cold side of gas-gas heat exchanger, this option will not bring any
35
more energy back to the boiler. Instead, it will discharge more heat to the
economizer and to the feed water. Therefore both the system efficiency and total
power will be increased.
Since this option does not bring more energy to boiler, the fuel feed rate, and the
boiler efficiency were not changed. The main change is power generation, from
483 to 486 MWe in total as compared with case-8. As a result, the power penalty
reduced to 99 kWh/klbCO2, and the efficiency drop to 6.5% in points. The net
system efficiency changed to 32.93%.
This hot gas recycle brings back more moisture to the boiler, as 15%v, and flue
gas moisture went up to 23.8%v. This moisture slightly changes the boiler
performance, such as recycle gas mass and volumetric flows, adiabatic
temperature, and O2%v to boiler. The high moisture content in gas tends to
reduce the difference between equilibrium and adiabatic flame temperatures,
which is similar to water-gas shift equilibrium, where CO is shifted by water vapor
to as H2, and the generated H2 is much easier to be burned than the CO.
4.1.4.5
O2-fired PC with Reduced Pipeline pressure (case-10)
This pressure change only affects the CO2 plant power. When this pressure
changed from 3000 to 2000 psia, the CO2 compression power changed from 34.2
to 32.9 MWe with liquid CO2 pumping as the last stage. While the power for the
last stage compression, if only gas compressor is used without CO2 liquid
pumping, would change from 6.5 to 3.9 MWe. Therefore even without the
requirement of CO2 purification, the CO2 condensation before last stage is
important in power saving, especially for the high pipeline pressure. At the Ref.
14 specified pressure of 2200 psia, overall cycle efficiency is increased by 0.1%
points versus a pressure of 3000 psia.
4.1.4.6
O2-fired PC with HRA Series Pass (case-6A/6B, 9A/9B)
Because of the reduced flue gas flow and lower HRA duty for high O2
concentration operating cases, designing the HRA in series instead of parallel
will produce a more compact design The modified HRA arrangement will not
affect the heat balance, but will affect the mechanical design and the heat duty
arrangement. Due to less heat picked up by the pre-superheater, the water
spraying for steam superheat temperature control was setup before the finishing
superheater.
Case-6B (Figure 4.1.8) applies nearly the same HRA inlet temperature (1451 vs.
1460ºF), as the case-6, while the case-6A applies a higher HRA inlet
temperature, 1732ºF, to increase the energy discharge to HRA to maintain a
better temperature profile. In case-6a, part of heat duty was shifted from boiler to
HRA, which was absorbed by a larger economizer resulting in a higher feed
water temperature (666 vs. 603ºF), while the furnace water wall heat duty was
reduced from 1441 to 1198 MMBtu/hr.
36
Because of less flue gas recycle required for the higher adiabatic temperature of
case-9, there was not enough heat to maintain the same HRA temperature
profile. Consequently, either the economizer water temperature can be reduced
or the HRA inlet temperature can be increased. A high HRA inlet temperature
was applied for case-9A to keep the water temperature from economizer at about
590ºF. The higher HRA inlet temperature also helped to shift heat duty from
division wall to the finish superheater.
The exit steam temperature from the division wall was about 1021ºF in case-9A,
which could require expensive high-grade materials to be used. In order to
reduce this temperature, a different approach was used in case-9B, where the
PSH (primary superheater) was removed and its heat duty was shifted to the
finishing superheater. Kept constant were the flue gas temperature to the
economizer, the water temperature to the evaporator, and the LMTD of lower
reheater. Consequently, the finishing superheater size increased since its duty
changed from 289 to 416 MMBtu/hr.
Case-9B (Figure 4.1.9) has a modification, where the heat picked up by roof was
reduced from 134 to 50 MMBtu/hr, and the difference was shifted to water wall,
which leads to an increased steam exit temperature from water wall. This does
not affect the heat and material balances, but only the design.
37
Figure 4.1.8 - Boiler design with HRA series pass
ST=4035/1076/1112/2.0"Hg=460 MW
Case: O2-PC-SC-06B, 09/28/2005
Temperature (F)
Pressure (psi)
Mass Flow Rate (Mlb/hr)
1076
PC Boiler
Duty (MMBtu/hr)
4035.0
569
4822.0
148
1083
4286.0
2803
48
1112
1112
823.2
823.0
2406
2406
2406
LS1
67
823.0
2389
Q=134
1451
1409
14.7
14.7
2705
2705
2705
G5
G6
1829
FRH
14.7
47
837
4612.0
44
1864
1021
14.7
4472.0
2705
871
890.0
353.2
72
8
38.9
545
2193
EI1
28
29
SL2
219
199
877.0
55
1283
361
6
AUX
G11
2705
863.3
199
522
89
80
18
5
47
SSR
9
FSPLIT
713
194.8
88
56
108
545
267
382
218
234
93.8
93.8
38.9
36.9
15.8
78
78
59
137
92
25
59
644
155
10.0
15.2
20
210
73
23
B1
V7
2406
V6
V5
522
2950
FWH8
522
4954.0
65
4929.0
2950
425
2950
FWH7
66
V3
386
610
15
108
Q=0
191.6
4979.0
257
150
262
545
36.9
93.8
257
78
191.6
94
1332
2950
FWH6
318
V1
V2
712
181.2
342.6
863.3
3.7
77
78
V4
VD
430
527
199
4904.0
Q=138
95
746
1266.2
63
569
45
1
112
877.0
PSH
2803
435
323
10.0
3.9
57
76
353.2
890.0
152
657
14.6
4527.0
11
1266.2
40
644
93.8
871
658
532
863
EL4
17
654
SL3
1305.3
21
1.0
1751
W=-18
657
102
1797
108
EI3
750
B2
30
1870
EL3
EL2
LS4
Q=106
3.9
1.2
219
DPCRH
49
152
1929
24
2406
234
15.8
93.8
96
RHB
869
Q=289
93.8
2301
EI2
EH2
DIVW
Q=479
6
2389
LP4
LP3
382
22
910
857.0
Q=329
5
2679
EH1
2045
SC1
G4
3
2878
540
194.8
857.0
G8
Q=-36
Q=-31
2803
653
890.0
2406
LP2
713
750
1305.3
2803
SC23
81
WEG
MULT
WE
Power(MW)
LP1
IP3
IP2
IP1
658
SL6
SPRAY
FSH
W
W
88
HP2
HP1
PRPB
Q=36
2950
Duty (MMBtu/hr)
2950
43
4462.0
Q
17
84
60
974
469
823.0
SL5
2878
27
19
26
1076
4035.0
Power(MW)
1112
1112
MAIN
DPHRH
832
4647.0
ROOF
2950
1113
849.0
69
137
1332
FWH4
145
15.2
206.6
70
1332
FWH3
3.7
214
208
71
COND
14
257
221.6
210
Q=-1860
1332
FWH2
FWH1
87
812
2185
4682.0
42
32
G3
Q
41
G12
2045
Q=162
657
2705
Q=311
RHA
QB
396
2045
374
BFP1
Q=-0
610
52
4839.0
BOILER
G13
4822.0
100
122
SPQ
STK1
50
CP
695
60
14.5
20.0
2705
0
SORB
N2
14.7
AIR
2145
400
COMP
SLV
67.0
2790
W=68
60
POND
FGD
PREC
90
67.0
122
L2
0
2790
0
0
309
15.1
0
2790
DIST
AC
14.7
DENOX
LIQ
NH3
EQ
0
G20
98
L1
309
SP1
296
695
14.4
14.4
E1
G16
2705
2705
ESP
AAHX
296
15.1
14.2
2674
309
G18
15.1
2705
G21
W=3
625
Q=-246
Q=-3
IDFAN
G17
14.7
0
W=0
SPF
2674
G19
4321
A0
14.7
SLV
RGIBBS
2674
-183
90
90
90
90
14.9
14.8
58.4
231.6
2126
776
767
764
C4
C6
GBY
AV
55
0
236
916.2
3000.0
29
29
RECY
15.3
1846
A2
550
ASH
550
C1
249
Q=0
PA
16.6
A5
60
625
16.7
MILL
550
60
145
14.7
14.7
0
1201
A1
Q=0
323
926.2
767
C10
W=0
764
98
COL4
1350
COL3
90
14.9
D1
147
46
15.0
90
90
231.6
58.4
D2
9
D3
3000.0
55
90
Q=-42
Q=-48
Q=0
645
A3
233.6
C7
COL2
14.8
14.8
1751
COAL
776
W=10
Q=-1
C2
401
90
145
SA
C5
W=11
15.1
BYP
1846
SAFAN
SPA
AIRB
15.4
W=1
313
59.4
W=11
2273
309 Q=-281
119
859
16.7
309
SEP1
14.9
COMP3
COMP4
311
90
DEM
765
COMP2
COMP1
COL1
PIPEG
231.6
2273
31
PAFAN
A6
C11
90
15.1
14.2
16.9
14.7
A4
296
171
145
2706
C3
309
Q=-282
NOZ
RSTOIC
53
W=1
14.7
0
0
60
29.7
G15
489
2232
64
15.0
60
14.7
14.7
2803
G2
14.7
1.0
WETECO
BYFW
2705
231.6
D4
Q=-113
735
916.2
735
CPUMP
1
PIPEL
2
D5
GRH
W=1
RY2
RY1
O2
38
1332
102
900
900
2950
W=18
97
FLAME
236.6
Q=200
33
34
G1
206.6
46
569
LECO
14.7
7
Q=57
102
850
14.6
Q=118
3826
Q=84
322
68
35
E2
102
16
2950
Q=210
603
4787.0
2803
181.2
86
FWV
569
12
Q=-2042
61
Q=65
DEA
342.6
20
Q=1440
HEATER
73
Q=83
Q=118
877.0
EVAP
93
85
236.6
14.6
867.0
2705
74
1115
820
14.7
2803
83
W=3
Figure 4.1.9 - Boiler design with HRA series pass (56% Flue Gas Recycle)
Case: O2-PC-SC-09B,
ST=4035/1076/1112/2.0"Hg=460 MW
09/28/2005
Temperature (F)
Pressure (psi)
Mass Flow Rate (Mlb/hr)
1076
PC Boiler
Duty (MMBtu/hr)
4035.0
569
4822.0
148
832
1083
4647.0
4286.0
48
2803
ROOF
2950
1113
1112
1112
849.0
823.2
823.0
2406
2406
2406
MAIN
19
1076
823.0
4035.0
2389
67
LS1
DPHRH
26
Q=50
1850
1379
1326
14.7
14.7
2117
2117
2117
G5
G6
FRH
14.7
47
658
837
81
974
14.7
4542.0
2117
72
2406
871
1305.3
353.2
8
EI2
EI1
2406
21
B2
55
219
EL2
EL3
G11
23
B1
93.8
6
AUX
152
10.0
3.9
5
47
SSR
435
323
863.3
522
89
80
871
713
353.2
194.8
18
56
88
108
545
267
382
93.8
93.8
38.9
36.9
78
78
59
137
92
25
218
95
527
199
863.3
234
15.8
59
155
644
15.2
10.0
20
210
73
425
522
4954.0
522
65
2950
FWH8
712
181.2
66
15
386
610
108
318
93.8
257
78
191.6
FWH4
1332
94
257
V1
V2
150
262
545
1332
2950
FWH6
V3
Q=0
191.6
4979.0
2950
FWH7
430
342.6
36.9
69
208
137
70
1332
FWH3
145
15.2
206.6
820
71
93
85
74
83
210
FWH2
Q=-1860
1332
FWH1
102
16
61
73
236.6
7
Q=118
Q=311
Q=162
657
2117
396
2045
342.6
DEA
BFP1
Q=-0
610
86
52
4839.0
BOILER
322
102
374
206.6
236.6
181.2
900
1.0
2232
WETECO
68
W=18
50
64
833
BYFW
14.6
G13
33
SPQ
2803
100
126
4822.0
LECO
60
46
569
2117
2803
34
SORB
AIR
2140
2784
0
400
SLV
97
Q=84
N2
14.7
0
0
67.0
COMP
2784
29.7
G2
G15
695
60
14.5
20.0
2117
0
W=68
0
383
60
15.1
0
L2
90
67.0
POND
PREC
FGD
0
126
2784
DIST
AC
14.7
DENOX
LIQ
NH3
98
0
G20
EQ
W=1
14.7
15.0
60
14.7
14.7
STK1
35
14.7
A0
14.7
SLV
383
RGIBBS
368
695
SP1
E1
G16
14.4
14.4
2117
2117
14.2
2087
2087
ESP
AAHX
368
15.1
G19
15.1
G21
2118
Q=-246
Q=-3
383
G18
2117
625
FLAME
0
IDFAN
G17
14.7
L1
W=0
SPF
5104
1260
W=2
2087
-183
90
90
90
90
14.9
14.8
58.4
231.6
1314
774
766
763
C4
C6
AV
55
0
916.2
236
37
3000.0
GBY
37
RECY
Q=-188
NOZ
259
ASH
550
C1
30
257
RSTOIC
A5
PA
60
625
16.7
MILL
550
60
259
14.7
14.7
A1
SAFAN
BYP
A3
SPA
SA
Q=0
COMP4
59.4
C5
323
233.6
W=11
774
C7
Q=-269
W=10
766
926.2
C10
W=0
763
Q=-1
COL2
259
90
14.8
14.8
1166
540
98
COL4
C2
626
Q=0
14.9
D1
146
COL3
Q=-48
90
46
90
58.4
D2
90
Q=-42
90
15.0
644
COAL
COMP3
313
15.1
1260
616
0
AIRB
383
15.4
W=1
W=11
1461
163
858
16.7
308
SEP1
14.9
16.6
Q=0
DEM
763
311
90
COL1
PIPEG
COMP2
COMP1
PAFAN
C11
231.6
1461
550
A6
90
15.1
14.2
16.9
A2
C3
383
368
288
14.7
A4
53
CP
Q=200
2950
Q=210
594
4787.0
9
231.6
D3
231.6
D4
Q=-112
55
3000.0
916.2
725
725
CPUMP
1
PIPEL
2
D5
GRH
W=1
RY2
RY1
O2
39
1332
102
900
2950
FWV
569
12
Q=57
Q=84
Q=65
Q=83
877.0
19
14.7
14
257
221.6
14.6
G12
2045
Q=1558
15.3
COND
3.7
214
87
867.0
457
1
3.7
77
78
V4
VD
V5
746
32
1161
RHA
G1
9
FSPLIT
890.0
199
4929.0
2950
EVAP
11
1266.2
63
Q=4
2450
41
57
76
V6
V7
2406
4904.0
14.7
40
644
112
PSH
2117
EL4
657
877.0
2117
1.0
1751
17
654
658
199
657
14.6
569
E2
219
102
1797
108
W=-18
1266.2
361
30
1870
1.2
LS4
Q=106
3.9
1929
SL3
SL2
45
4161
29
152
EI3
1305.3
837
Q=-2089
28
234
15.8
93.8
750
2803
HEATER
38.9
545
2193
LP4
LP3
24
4597.0
QB
93.8
2301
DPCRH
RHB
Q
LP2
382
96
877.0
G3
LP1
540
EH2
1167
2803
6
2389
22
910
532
825
5
2679
IP3
2045
49
42
3
2878
DIVW
4682.0
WEG
MULT
Power(MW)
713
194.8
857.0
G8
869
Q=416
IP2
653
890.0
2803
857.0
Q=329
IP1
750
SL6
SC1
G4
Q=487
W
88
EH1
Q=-36
Q=-25
2803
890.0
4612.0
44
SC23
1885
FSH
W
WE
HP2
HP1
SPRAY
469
Duty (MMBtu/hr)
2950
PRPB
Q=36
2950
Q
17
84
43
4532.0
Power(MW)
823.0
SL5
2878
27
60
934
1112
1112
W=2
4.1.5 Part Load
For a sliding pressure supercritical system design, part load performance has to
be considered and evaluated. For this purpose, the Aspen model can be used by
converting its design model to a simulation model, where all of the equipment
sizes are fixed from the design case. Based on the 100% load condition, the heat
transfer coefficients are functions of operating conditions, such as temperatures
and Reynolds number. At full load, the heat transfer coefficients vary from 14.4
for the finishing superheater to 9.7 Btu/ft2-hr-F for economizer. In order to
simulate the air-air heat exchanger, the original lumped module was replaced by
two modules for both the primary air and the secondary air. The steam turbine
was also converted in the simulation model such that the turbine exit parameters
are floated as loading changes. The model for feedwater heaters was adjusted
so that the extraction pressures were automatically calculated.
The water wall absorption of boiler is difficult to be simulated by the zero
dimensional Aspen model. Therefore its performance such as the gas exit
temperature was calculated separately using the 3-D FW-FIRE model and
iterated with Aspen.
Note that in conventional air fired PC boiler, the excess air has to be increased
during part load to maintain steam exit temperatures. Increasing excess air is not
necessary for the oxygen-fired boiler since the quantity of flue recycle can be
increased at part load to maintain steam temperatures. Increasing flue gas
recycle flow rather than excess air has less adverse effect on plant performance
and emissions.
The 72% part load (72% generator power) is shown as Case p72 in Table 4.1.3.
The amount of flue gas recycled is adjusted to produce sufficient SH and RH
steam temperatures. The water spray to SH was adjusted from 5% to 4% to
boost SH steam temperature. In this way, the efficiency loss is kept at a minimum
when boiler is operated in part load.
The loading is further reduced to 50% as shown by p50 in Table 4.1.3. At this
load, the main steam pressure is 2125 psia. At this pressure, the steam turbine
can be operated at a reasonably high temperature to keep the system efficiency
high. Therefore the boiler was adjusted to maintain both the SH and RH steam
conditions as near as possible to those at 100% load. Note that the flue gas
recycle flow changes only slightly as the loading changes from 100% to 50%.
The lowest part load case for this study was 25% as shown by p25. In
corresponding to a sliding pressure, both SH and RH steam temperatures were
adjusted through the combination of firing rate and gas circulation rate to make
sure the end steam exhausted from the steam turbine at right conditions.
Because of lower steam conditions, the system efficiency reduced substantially.
40
Figure 4.1.10 and Figure 4.1.11 are plots, which show the effect of part load on
performance. At some loadings, the flue gas absolute recycling flow does not
change much, but its ratio to boiler flow changes because of less total boiler gas
flow as shown by Figure 4.1.10. Figure 4.1.11 shows the relation between part
load and system efficiency.
Figure 4.1.10 - Flue gas recycle flow vs. part load operation
Part Load Performance
100
1400
90
1200
70
60
800
50
600
40
30
400
gas recycling, %
Gas recycling, klb/hr
80
1000
20
200
10
0
120
0
0
20
40
60
80
100
Load, %
Figure 4.1.11 Efficiency at part load operation
Part Load Performance
40
35
Efficiency, %
30
25
20
15
10
0
20
40
60
Loading, %
41
80
100
120
4.1.6 Comparison With Post Combustion CO2 Capture
CO2 cannot be captured and sequestrated without reducing both the plant power
and efficiency because of the potential energy stored in the pressurized liquid
CO2. A minimum of 40 kw/klbCO2 additional auxiliary power is required for CO2
compression. The difference between technologies lies in the difference in power
requirements of the different CO2 or O2 separation techniques.
Parsons (Ref. 1) performed studies on CO2 removal by a post capture method for
a conventional PC boiler. In their study for post combustion CO2 removal, the
plant efficiency drops from 40.5% to 28.9% for a supercritical (3500
psia/1050ºF/1050ºF/1050ºF/2.0”Hg) boiler, and from 42.7% to 31.0% for an ultra
supercritical (5000 psia/1200ºF/1200ºF/1200ºF /2.0”Hg) boiler. In the study
presented herein, the CO2 removal using an O2-fired PC is used, which relies on
an ASU. At a pipeline pressure of 2000 psia (as applied in the Parsons study) the
efficiency
drops
from
36.7%
to
30.4%
for
the
subcritical
(2415psia/1000ºF/1000ºF/2.5”Hg) boiler [3], and from 39.5% to 33.1% for the
supercritical (4035psia/1076ºF/1112ºF/2.0”Hg) boiler. The net efficiency drops for
these cases are
11.7% points for supercritical, post combustion CO2 removal
11.7% points for ultra supercritical, post combustion CO2 removal
6.3% points for subcritical, O2 fired
6.4% points for supercritical, O2 fired
Another comparison basis is the kWh per klb CO2 removal, where the kW is
power generation difference between cases with and without CO2 removal.
Comparing the post CO2 capture to the O2-fired case yields:
187 kWh/lbCO2 for supercritical, 90% post combustion removal
188 kWh/lbCO2 for ultra supercritical, 90% post combustion removal
99 kWh/lbCO2 for subcritical, O2 fired 100% removal
97 kWh/lbCO2 for supercritical, O2 fired 100% removal
Thus the efficiency drop and change in power penalty for CO2 removal appears
independent of steam cycle. From the above data, it is clear that the O2-PC has
significant advantages over the post combustion CO2 capture.
4.1.7 Furnace Waterwall Temperature
The level of radiation in the O2-fired boiler is significantly higher than an air-fired
boiler due to greater concentrations of radiating gas species (CO2 and H2O) and
higher flame temperature. Consequently, it is important to select the proper
amount of recycled flue gas to limit the water wall temperature such that a
reasonable waterwall material can be used. The maximum waterwall temperature
and furnace heat flux is determined in Section 4.3.
42
4.2
Advanced O2 Separation System Integration
4.2.1 O2-Fired PC Integrated with Cryogenic ASU (case-6)
The model of the oxygen-fired plant with cryogenic ASU was described in
Section 4.1.3.
Figure 4.2.1 presents the cryogenic ASU base case system model (case-6). Flue
gas recycle flow rate is 65.5% resulting in a boiler equilibrium temperature of
3830ºF and an overall cycle efficiency of 31.94%. The oxidant from the ASU is
99.5% O2 pure. ASU and CO2 plant power consumptions is 67.7 and 34.2 MWe,
respectively, resulting in a total auxiliary power of 138.5 MWe, or about 29.1% of
gross power. The efficiency drop compared to the air-fired case is 7.5% in points
(39.46% to 31.94%), and the associated power penalty is 114 kWh/klbCO2.
43
Figure 4.2.1 - Base case of O2-fired supercritical PC with Cryogenic ASU
Case: O2-PC-SC-06,
ST=4035/1076/1112/2.0"Hg=460 MW
06/03/2005
1083
Mass Flow Rate (Mlb/hr)
4035.0
2950
48
1113
1112
1112
849.0
823.2
823.0
2406
2406
2406
830
2950
1112
MAIN
67
LS1
19
ROOF
HP2
HP1
Q=137
47
1818
1460
1442
14.7
14.7
14.7
FRH
2705
2705
G5
G6
PRPB
Q=36
2705
44
890.0
2803
2406
G10
G8
W
WE
EG
W
MULT
PW
G12
EI2
EI1
532
199
1266.2
219
EL2
435
323
863.3
Q=333
194.8
56
114
545
267
382
218
93.8
93.8
38.9
36.9
127
127
99
226
92
25
60
863.3
425
522
4954.0
65
527
81
386
610
4979.0
G13
15
114
2950
69
36.9
208
226
206.6
DEA
BFP1
2226
86
695
60
14.4
20.0
0
DENOX
122
100
14.7
14.7
0
0
STK1
N2
AIR
2145
PREC
FGD
400
67.0
2790
60
POND
122
L2
LIQ
14.3
14.1
2674
98
0
A0
14.7
0
Q=-246
15.1
Q=-3
W=0
SPF
311
G18
G21
90
90
15.1
2705
90
14.9
2674
90
58.4
14.8
2126
625
W=3
NOZ
A4
550
C1
ASH
31
COL1
PA
Q=0
120
16.7
309
60
14.7
14.7
0
1201
859 625
16.7
MILL
COAL
550
SAFAN
AIRB
SPA
SA
A1
C5
776
W=11
COMP3
C7
146
767
C10
W=0
764
98
COL4
COL2
90
14.8
14.8
1350
GRH
926.2
Q=-1
C2
401
1751
RY2
CO2
COMP4
323
W=10
Q=-282
BYP
A3
W=1
W=11
DEM
765
233.6
15.1
1846
29
C11
PIPEG
313
59.4
SEP1
311
15.4
W=1
311
2273
146
249
16.6
236
3000.0
29
231.6
COMP2
COMP1
14.9
RSTOIC
916.2
90
2273
90
PAFAN
55
0
ASU
C6
C3
15.1
14.1
16.9
A2
550
-183
764
RECY
311
296
171
A6
GBY
AV
231.6
767
776
146
Q=-282
14.7
90
AC
2674
IDFAN
14.7
W=68
2790
DIST
SLV
311
ESP
AAHX
4321
W=1
2790
0
COMP
G20
2705
53
CP
14.7
15.0
60
14.7
67.0
0
14.3
2226
50
60
0
311
2705
1.0
0
Q=0
SORB
RY1
Q=0
44
Q=0
90
14.9
D1
147
46
Q=-48
O2
90
Q=-42
58.4
15.0
645
COL3
D2
9
90
90
231.6
D3
2
231.6
D4
Q=-113
55
3000.0
916.2
735
735
CPUMP
1
2226
102
64
15.1
296
Q=95
WETECO
97
2803
NH3
G16
Q=-1696
102
BYFW
FGD
296
COND
14
429
FWH1
102
0
46
EQ
SP1
71
236.6
2950
14.7
E1
2226
Q=141
W=18
569
2705
695
145
221.6
348
FWH2
68
0
RGIBBS
3.7
214
15.2
16
Q=110
29.7
G15
150
206.6
181.2
Q=-0
4822.0
Q=79
429
V1
61
269
374
52
SPQ
34
G2
70
FWH3
73
2950
35
A5
1
112
77
4839.0
2803
60
20
348
122
236.6
592
2803
4770.0
1846
644
10.0
V2
2226
FWH4
4787.0
LECO
597
257
191.6
127
Q=135
396
569
799
2705
9
155
15.2
7
Q=118
Q=311
Q=162
342.6
14.5
G14
14.7
5
79
SSR
4904.0
32
811
135
DAMP
545
93
85
610
HEATER
2706
3.9
57
262
93.8
94
2226
2950
74
FWV
15.3
10.0
15.8
59
V3
Q=0
317
191.6
FWH6
FWH7
83
19
489
152
87
14.6
14.7
181.2
342.6
66
2950
2950
FWH8
569
818
14.6
41
234
78
V4
712
430
527
522
199
4929.0
148
95
VD
V5
Q=17
14.7
40
644
93.8
FSPLIT
713
88
80
1266.2
63
2185
1758
1.0
1573
EL4
17
6
353.2
18
V6
569
BOILER
219
EL3
654
76
522
746
4822.0
G11
890.0
11
89
UECO
811
102
1652
3.7
811
V7
2705
30
108
1775
1.2
871
199
2803
RHB
Q=-2042
29
1874
93.8
SL3
658
1305.3
21
SPRAY
24
QB
545
2187
3.9
AUX
750
2406
46
Q
8
152
14.6
Q=-36
14.7
2301
LP4
LP3
234
15.8
W=-18
2406
4510.0
45
893
2803
LP2
382
38.9
EI3
SL2
855
Q=114
1758
882
857.0
4665.0
LP1
540
93.8
96
1442
4500.0
FLAME
Power(MW)
LS4
14.7
Q=299
2389
SL6
657
23
Q=311
G1
W
877.0
PSH
G9
810
2679
6
DPCRH
811
SPG
2950
IP3
194.8
EH2
SC1
Q=-31
835
3
2878
14.7
G4
3826
Duty (MMBtu/hr)
713
353.2
22
1362
DIVW
Q=501
72
IP2
871
890.0
1305.3
SC4
Q=-20
G7
Q=-15
2705
4595.0
IP1
653
750
EH1
SC23
14.7
658
835
1853
FSH
Q=1453
88
84
43
G3
Q
SL5
2878
27
970
EVAP
Power(MW)
17
2803
4445.0
42
459
823.0
2389
4035.0
Duty (MMBtu/hr)
1112
823.0
1076
DPHRH
26
4630.0
2950
Pressure (psi)
1076
4269.0
PC Boiler
Temperature (F)
PIPEL
D5
W=3
4.2.2
O2-Fired PC Integrated with OITM
Advanced oxygen separation through high temperature membranes such as
OITM is currently under development [2, 4, 5, 7]. Oxygen ion transport
membranes have the potential to provide a major reduction in oxygen separation
capital and energy consumption.
Oxygen ion transport is driven by the difference in oxygen partial pressure across
a membrane. Oxygen atoms adsorb on the cathode (high oxygen partial
pressure side of the membrane) and dissociate into ions as they pick electrons.
These ions travel from cathode to anode (the low oxygen partial pressure) by
jumping through lattice sites and vacancies until they reach the anode side of the
membrane. On the anode side, the oxygen ions yield their electrons to become
atoms/molecules, which are then desorbed into the gas phase. Electrons from
the anode side are carried through the membrane to the cathode side to
complete the circuit. The rate of oxygen transport through such membranes is
temperature sensitive, and can be very fast at high temperatures. These
membranes have infinite selectivity for oxygen over other gases, because only
oxygen ions can occupy the lattice positions. A typical schematic of the oxygen
ion transport membrane process is presented in Figure 4.2.2.
The OITM process and design data are based on Reference 7. A schematic of
the OITM process incorporated within the O2-PC plant is shown in Figure 4.2.3.
To integrate the OITM into the O2-fired PC, air is pressurized by a compressor,
and then heated to a high temperature. High pressure air provides a high oxygen
partial pressure on the airside of the OITM to reduce the size and cost of the
OITM. Air enters the compressor at 60ºF and 14.7 psia and is compressed at
85% efficiency to 215 psia and 743ºF. The compressed air is heated to 1652ºF
by a tubular air heater inside the PC furnace. This hot pressurized air is fed to the
OITM, where about 85% of its O2 is separated through the membrane to an exit
pressure of 16 psia, and the rest of the vitiated air or O2-depleted air is sent to a
gas expander to generate power at 86% turbine efficiency. Power generated
from the expander is used to drive the air compressor. Since the power
generated in the gas expander is greater than the air compressor power the
OITM system produces a net power output.
The separated O2 from the membrane is carried by a heated sweep gas. The use
of a sweep gas reduces the oxygen partial pressure on the low-pressure side of
the membrane and consequently reduces the size and cost of the OITM. A
recuperator is applied between inlet and outlet sweep gas flows to reduce the
heat requirement. Heat transfer from the sweep gas to the membrane to the air is
neglected as it is expected to be relatively small (it was also neglected in
Reference 7). The mixture of sweep gas and O2 is then injected into the furnace.
In this design, the OITM does not consume electrical power; instead, it absorbs
heat, generates power, and separates O2 from air. Therefore it is expected to
reduce ASU operating and capital costs.
45
From a heat and mass balance point of view, if there is no heat transfer between
the air and sweep gas, the model can be simplified by directly mixing the
separated O2 with the sweep gas. Thus, to simplify the model, the recuperator
and OITM are combined into a single module in the Aspen model. The
operational details of the OITM, such as the O2 flux through the membrane as a
function of OITM temperature and pressure, were not included in the Aspen
model. These details are necessary only for OITM size and cost estimations.
Moreover, the Aspen model does not directly model the recuperator, but the
recuperator design and configuration is required for the economic analysis.
Without performing a detailed economic study, it is difficult to determine the
optimum OITM operating pressure. However, Reference 8 specifies that for an
economic design, the ratio of oxygen partial pressures of the feed gas (air) to the
permeate stream (sweep gas) should be approximately 7. The base case OITM
design has an oxygen partial pressure ratio of 4.9 at the air inlet and 10.2 at the
air outlet, yielding an approximate average ratio of 7.5. Reference 8 specifies that
an 85% O2 recovery and an operating temperature of 1652ºF is within the
operating range of the OITM. The effect of O2 recovery and OITM operating
pressure on system performance and OITM design is examined in Section
4.2.2.3.1 and Section 4.2.2.3.2, respectively.
Several cases incorporating an OITM ASU into the O2-PC have been simulated
to evaluate the effect of different conceptual designs and operating conditions on
power plant performance.
46
Figure 4.2.2 - A Typical Oxygen Ion Transport Membrane Schematic [4]
47
Figure 4.2.3 – Integration of Oxygen Ion Transport Membrane into O2-PC
48
4.2.2.1
Process Gains
The inclusion of a gas turbine expander in the O2-PC power plant utilizing the
OITM increases the overall system efficiency by allowing work to be done in the
gas turbine at a higher temperature than can be achieved by the steam from the
boiler. This principle has been applied by FW in its 1st generation Pressurized
Fluidized Bed Combustor (PFBC) design. A further enhancement is the 2nd
generation PFBC design (similar to IGCC) in which the turbine entrance
temperature is increased by syngas combustion. This concept was also applied
in the FW High Performance Power System (HIPPS) design, which includes an infurnace high temperature air heater similar to the one required for the O2-PC
OITM concept.
4.2.2.2
Base Case (case-11)
Figure 4.2.3 shows a schematic of the OITM ASU integrated with the O2-fired PC
power plant. The OITM system includes a sweep gas system, and an air supply
system. The pressurized vent gas from the CO2 plant is recycled back to the air
separation unit, where it is mixed with the compressed air. In this way the rich O2
in the vent gas can be recovered, and the air to compressor can be reduced.
Therefore this vent gas recycling increases system efficiency and reduces the
operating cost.
Figure 4.2.4 (case-11) is a process flow diagram generated by Aspen Plus for the
OITM ASU integration. Air is compressed to about 200 psia by a compressor,
and then heated within the boiler to about 1650ºF. This hot pressurized air is fed
to OITM, where about 85% of its O2 is separated through a membrane, and the
rest of vitiated air is sent to an expander. The separated O2 from the membrane
is carried by a heated recycled flue gas after gas-to-gas heat exchange. A
recuperator is applied between inlet and exit sweep gas flows (not shown in
Figure 4.2.4 since it lumped inside the OITM module). The mixture of sweep gas
and O2 is fed directly to the boiler. The exhaust gas from the expander passes
through an economizer to release its heat for feedwater heating. Power
generated from the expander is used to drive the air compressor.
The O2 obtained from the OITM is swept with recycled flue gas. Since the
compressed air is heated to 1650ºF inside the boiler, a special heat exchanger
and boiler design has to be used for the OITM application (Sections 4.3.2.2.3 and
4.3.5.4). The boiler air heater duty is 974 MMBtu/hr. The coal feed rate is 377
klb/hr as compared with 319 klb/hr for the air-fired case-1, and 309 klb/hr for the
O2-fired (cryogenic ASU) case-6. The corresponding flue gas flow increased to
3497 klb/hr, nearly approaching the air-fired flow of 3552 klb/hr. The boiler O2
concentration is 31%v, which is nearly the same as case-6.
Because of increased flue gas flow created by greater coal-firing, more heat is
carried to the boiler HRA in case-11 than in case-6. As result, distribution of heat
49
duty shifts: the heat duty of the division wall reduces from 501 MMBtu/hr in case6 to 396 MMBtu/hr in case-11, while the sum of the primary superheater and
upper economizer duties increases from 131 MMBtu/hr in case-6 to 411
MMBtu/hr in case-11. Consequently, the inlet furnace feedwater temperature
increases from 597ºF in case-6 to 666ºF in case-11. The total furnace duty for
case-11 increases because of the 974 MMBtu/hr air heater duty, although the
heat to waterwalls is reduced from 2042 to 1583 MMBtu/hr.
Because more low-grade heat is released from the gas turbine (GT) exhaust and
from boiler, all low-pressure feedwater heaters are shut off. The low-pressure
feedwater heating is provided in parallel by the GT economizer and a HRA flue
gas exhaust economizer (in parallel to flue gas heat recuperator), as well as by
part of the compressor inter-stage coolers.
The OITM O2-PC incorporates the recycling of vent gas from the CO2
compression system back to the OITM. The high-pressure vent gas from the CO2
plant is heated with GT exhaust gas before it is expanded for power recovery.
After the expansion to the OITM operation pressure, the O2-rich vent gas is
mixed with compressed air and sent to the OITM. Consequently, the compressor
air flow and power is reduced. The emission control equipment treats the vent
gas prior to the OITM requiring smaller equipment sizes due to the high pressure
of the vent gas.
The replacement of the cryogenic ASU with the OITM greatly reduces the
efficiency loss penalty. This is caused by lower equivalent ASU power, better
system integration, increased power from the OITM GT, and more low-grade
heat recovery from cooling. As shown in Table 4.2.1 the OITM reduces the
efficiency drop by 52% and the CO2 removal power penalty by 61%.
Table 4.2.1 – Comparison of O2PC with Cryogenic and OITM ASU
Air-fired
ASU
Main steam flow, klb/hr
2950
Coal flow, klb/hr
319
Net power, MW
430
Net efficiency, %
39.5
ASU Power, MW
0
CO2 Compression Power, MW
0
CO2 removal flow, klb/hr
0
Efficiency penalty, % points
0
CO2 removal penalty, kWh/klbCO2
0
50
O2-fired O2-fired
cryogenic OITM
2950
2950
309.4
377
338
462
31.9
35.8
67.7
-26.7
34.2
41.3
720
874
7.5
3.6
114
45
Figure 4.2.4 - O2-PC with OITM
Case: O2-PC-SC-11,
ST=4035/1076/1112/2.0"Hg=460 MW
06/15/2005
Temperature (F)
Pressure (psi)
Mass Flow Rate (Mlb/hr)
1083
1076
4294.0
4035.0
2873
48
PC Boiler
1113
1112
1112
849.0
823.2
823.0
2483
2483
2406
Duty (MMBtu/hr)
2950
19
26
823.0
4035.0
2389
67
LS1
DPHRH
830
1076
MAIN
HP2
HP1
PRPB
Q=36
1907
1632
FRH
14.7
47
14.7
14.7
3534
3534
3534
G5
G6
1942
658
750
890.0
1305.3
2730
2483
72
3
2878
1534
Q=-50
LP1
540
382
2355
93.8
38.9
8
2355
G8
46
199
G12
545
93.8
93.8
0
0
92
382
545
398
38.9
25
36.9
0
746
60
65
2873
871
445
15
848
798
81
32
639
0
485.0
69
33
4812.0
3534
HR-B
1436
W=1
127
60
14.7
AIR
14.7
0
SLV
SORB
3945
782
0
15.3
29.7
68
695
GC
H3
GT
3179
W=195
W=-222
-974
POND
14.4
G15
PREC
FGD
3534
0
470.0
DENOX
958
12
20.0
695
AAHX
14.3
4199
304
14.1
14.3
3497
3534
3534
G16
G17
LIQ
3497
216
C1
600
ASH
90
90
90
14.8
58.4
231.6
940
923
915
A3
37
14.7
14.7
MILL
220.0
220.0
18
18
C15
C14
319
W=1
233.6
59.4
C5
W=13
940
C7
A1
SA
C13
93
C11
56
926.2
C8
W=13
C10
915
Q=-6
923
916.2
75
Q=-308
Q=-1
C2
1774
90
216
14.8
14.8
2374
1150
COL4
Q=0
Q=-61
90
90
D2
17
D4
Q=-115
1
D6
GRH
RY1
RY2
Q=0
51
CO2
63
56
3000.0
916.2
897
822
231.6
58.4
183
231.6
Q=-48
90
14.9
D1
90
COL3
D3
PIPEL
7
D5
COAL
W=1
W=-0
Q=3
TSA
COMP3
308
W=13
COMP4
18
RH
C12
DEM
308
309
SEP1
14.9
2273
700
896.2
56
COMP2
COL2
SPA
AIRB
1224
56
916.2
18
916.2
15.4
SAFAN
16.7
377
393
231.6
1774
0
60
393
C9
90
2273
90
15.1
BYP
225
216
60
C6
C4
916
PA
600
C3
COMP1
COL1
16.7
A5
Q=0
H4
Q=-0
90
14.1
PAFAN
3179
Q=974
14.9
304
600
625
14.7
3179
HX
784
15.1
16.9
A2
14.7
A4
200
191.6
GBY
319
977
Q=-468
H1
319
243
Q=-367
1774
16.6
3963
1652
O2
15.1
G21
625
A6
3945
OTM
RECY
15.3
231
199.6
H2
FGD
W=0
G19
214.6
16.0
2090
705
A0
1652
0
SPF
G18
W=4
L1
743
1652
3497
IDFAN
ESP
14.7
3534
SLV
0
0
15.1
304
0
G20
319
14.7
14.7
319
15.1
60
100
127
EQ
RGIBBS
HR-A
L2
QOTM
400
RSTOIC
ASU/OTM
60
0
14.7
STK1
Q=161
100
2730
53
CP
Q=152
41
G2
14.7
NOZ
2394
50
BYFW
0
3535
1.0
14.7
470.0
0
102
Q=465
WETECO
958
7
Q=0
102
W=17
4847.0
34
102
Q=0
Q=0
Q=0
1436
100
567
2730
FWH1
500.0
2873
SPQ
35
4795.0
Q=-2251
0
16
52
270
LECO
666
0
FWH2
450.0
86
BFP1
64
2730
Q=-2557
1
485.0
61
4864.0
612
COND
14
500.0
419
DEA
2873
Q=-0
958
G14
70
0
FWH3
71
430
567
846
1
3.7
639
15.2
485.0
73
342.6
80
500.0
14.5
1
3.7
641
66
FLAME
20
V1
376
639
0
0
FWH4
93
102
FWV
14.7
0
V2
36.9
639
87
Q=232
Q=162
2873
1943
141
DAMP
NH3
15.2
0
863.3
74
445
16
SP1
15.8
398
545
93.8
94
485.0
0
455
4929.0
14.6
G13
34
2873
FWH7
568
848
14.7
9
644
10.0
77
V3
Q=0
342.6
4979.0
83
14.6
1449
V4
VD
445
144
Q=205
E1
376
78
527
863.3
519
199
FWH8
4847.0
Q=344
HEATER
234
59
0
95
4954.0
14.7
3534
BOILER
SSR
FSPLIT
18
1266.2
63
567
G11
5
0
0
34
V6
V7
SPRAY
Q
57
579
UECO
2185
56
10.0
3.9
AUX
W=-18
89
1943
RHB
24
2730
644
152
76
247
199
2483
4690.0
11
1266.2
2730
40
17
654
14.6
Q=-36
857.0
Q=292
21
4535.0
45
1199
PW
EL4
93.8
194.8
353.2
890.0
529
219
EL3
6
658
1305.3
1.0
2169
713
871
750
1449
4525.0
882
EL2
SL3
SL2
657
879
102
2169
EI3
2483
Q=206
30
2169
108
96
1614
854
29
3.9
1.2
219
877.0
23
LP4
152
2169
LS4
14.7
G1
WEG
MULT
LP3
234
15.8
93.8
EI2
EI1
EH2
PSH
2873
28
545
DPCRH
1943
SPG
Q=321
E2
WE
LP2
194.8
14.7
G10
G9
QB
IP3
2389
SL6
Q=-42
810
5
2679
EH1
G4
3776
Power(MW)
713
6
SC1
3534
Q=396
IP2
871
353.2
22
G7
DIVW
Q=1149
W
W
IP1
653
890.0
SC4
Q=-21
14.7
835
4620.0
44
1614
SC23
FSH
G3
Duty (MMBtu/hr)
88
84
43
Q=135
EVAP
Q
17
2730
ROOF
970
4470.0
42
502
823.0
SL5
2878
27
4655.0
2873
Power(MW)
1112
1112
CPUMP
W=3
4.2.2.3
OITM Parametric Studies
Several OITM parametric cases were run as described in the following sections.
The results are summarized in Table 4.2.2.
Table 4.2.2 – Parametric Case Summary
case
Air separation method
Waste heat economizer
HRA Arrangement
CO2 Condensation
Vent Gas Recycle
Coal flow
Natural Gas Flow
Oxidant flow
Air flow
O2 purity
01
06
09
10
11
12
13
14
15
16
17
18
19
None Cryo Cryo Cryo OITM OITM OITM OITM OITM OITM OITM OITM CAR
no
no
yes
yes
yes
yes
yes
yes
yes
yes
yes
yes
yes
parallel parallel parallel parallel parallel parallel parallel parallel parallel parallel parallel parallel series
yes
yes
yes
yes
yes
yes
yes
yes
yes
yes
yes
yes
no
no
no
yes
yes
yes
yes
yes
yes
yes
yes
no
klb/hr
319.0 309.4 308.0 308.0 377.7 376.0 403.8 377.7 377.7 377.7 389.4 375.4 307.7
klb/hr
4.0
klb/hr
3270
645
644
644
784
780
838
784
784
784
808
781
1849
klb/hr
3270 2790 2784 2784 3945 3685 4945 4284 3810 3875 4067 3930 3165
%
99.5
99.5
99.5
100
100
100
100
100
100
100
100
43.2
klb/hr
0
1751 1166 1166 2374 2356 2523 2374 2374 2374 2392 1525 1190
Recycle gas flow
%
0.0
65.5
55.9
55.9
67.9
67.8
67.8
67.9
67.9
67.9
67.4
57.7
56.0
Recycle gas temperature
F
146
260
260
216
217
223
214
217
216
220
249
150
Boiler Inlet O2
%, v
20.7
33.8
41.1
41.1
31
31.1
31.1
31
31
31
31.4
39.3
41.8
Boiler Outlet O2
%, v
3.0
3.0
3.0
3.0
3.0
3.0
3.0
3.0
3.0
3.0
3.0
3.0
3.0
MM cf/hr
179
96
77
77
139
138
150
139
139
139
142
100
81
Boiler outlet flue gas flow
klb/hr
3555 2675 2087 2087 3497 3474 3723 3497 3497 3497 3550 2644 2126
Adiabatic Temperature
F
3685 4321 5104 5104 4196 4202 4208 4196 4196 4196 4243 5120 5120
Equilibrium Temperature
F
3552 3830 4161 4161 3770 3773 3775 3770 3771 3771 3792 4169 4155
Gas Temp. to FSH
F
2185 2185 2185 2185 2185 2185 2185 2185 2185 2185 2185 2185 2450
Water temp. to evap.
F
638
597
596
596
666
664
690
666
666
666
672
599
591
UBC
%
1.00
0.30
0.16
0.16
0.38
0.38
0.38
0.38
0.38
0.38
0.38
0.16
0.15
Boiler Efficiency
%
88.16 90.90 91.31 91.31
Pipeline pressure
psia
0
3000 3000 2000 3000 3000 3000 3000 3000 3000 3000 3000 3000
Generated CO2 flow
klb/hr
739
720
718
718
874
869
934
874
874
874
900
866
729
CO2 purity
%,v wet
14.0
77.5
71.3
74.2
75
74.9
74.9
75
75
75
74.8
70.8
76.5
Removed CO2 flow
klb/hr
0
720
718
718
874
869
934
874
874
874
900
866
729
CO2 removal efficiency
%
0
100
100
100
100
100
100
100
100
100
100
100
100
CO2 purity
%
93.7
93.6
93.6
96.9
96.9
96.9
96.9
96.9
96.9
96.9
96.9
94.2
Gross Power
MW
476.2 476.2 486.4 486.4 519.7 515.8 520.3 519.3 520.0 519.8 521.8 519.5 488.7
Auxiliary Power
MW
46.0
36.6
38.1
38.1
43.0
42.8
44.0
43.1
43.0
43.0
43.4
41.6
42.5
ASU power
MW
0.0
67.7
67.6
67.6 -26.7 -22.0 -42.2 -24.3 -27.3 -27.1 -35.9 -26.6 32.0
CO2 compression power
MW
0.0
34.2
34.2
32.9
41.3
41.1
44.1
41.3
41.4
41.4
42.6
41.2
34.3
Net Power
MW
430.2 337.7 346.6 347.9 462.1 453.9 474.4 459.2 462.9 462.5 471.7 463.3 379.9
Net Efficiency
%
39.46 31.94 32.93 33.05 35.80 35.33 34.38 35.58 35.87 35.83 35.45 36.11 35.19
Efficiency Drop
% pts.
7.5
6.5
6.4
3.7
4.1
5.1
3.9
3.6
3.6
4.0
3.3
4.3
CO2 removal energy
kWh/klbCO2
114
99
97
46
52
59
48
45
45
49
42
64
52
4.2.2.3.1 Effect of O2 Recovery Efficiency (case 11 to 13)
As described in Section 4.2.2.1, shifting more duty to the GT will increase system
efficiency. Increasing OITM O2 recovery efficiency reduces the required air flow
rate and decreases the GT mass flow and power generated.
In Case-12 the O2 recovery is increased from 85% to 90.5%. As a result of high
O2 recovery, less air is required to be fed to the system, and so less power is
generated in the GT. This leads to a reduction of system efficiency from 35.8% to
35.3%.
In Case-11 and Case-12 the low-grade heat released from GT exhaust, flue gas
cooling before CO2 compression, and CO2 compressor inter-stage cooling is
recovered to heat the low-pressure feedwater. In Case-13, when the O2 recovery
is reduced from 85% to 73%, more air has to be fed to OITM (4945 vs. 3945
klb/hr) and so more heat is required by the OITM. The system fires more coal
(404 vs. 378 klb/hr) and as result, the system generates more low-grade heat
than can be recovered, which results in an increase of GT exhaust temperature
from 200ºF to 330ºF. This reduces the system efficiency even with increased
extra power from the GT.
It is clear that the two opposing effects, more GT power and more low grade heat
from increased air flow, form a system with an optimum performance for a given
air side pressure as shown by Figure 4.2.5, where the system efficiency is
maximum when the O2 recovery is approximately 85%. Note that recovering
additional low-grade heat as through the use of a high-pressure economizer will
shift the optimum O2 recovery efficiency. As more of the low grade heat is
recovered by the use of more complex heat integration schemes or by cogeneration heat export, then the optimum O2 recovery efficiency is reduced and
the maximum system efficiency is increased. Such a reduction in O2 recovery
efficiency reduces the size of the OITM because of increased logarithm mean
pressure difference (LMPD).
Table 4.2.3 shows a comparison of different cases under the same airside
pressure, including a case published by Alstom [7]. It is obvious that the higher is
the O2 recovery efficiency, the lower is the LMPD, and the larger is the OITM
size.
53
Figure 4.2.5 – Effect of OITM O2 Recovery Efficiency on Efficiency
Efficiency, %
36.0
35.5
35.0
34.5
34.0
70
75
80
85
90
95
O2 recovery, %
Table 4.2.3 - Effect of OITM O2 Recovery Efficiency on LMPD
Ca s e
O2 r e c o v e r y , %
O2 t o b o i l e r , %
L MPD
Al s t o m
85
70
14. 7
Ca s e - 1 1
85
31
16. 5
54
Ca s e - 1 2
90
31
13. 9
Ca s e - 1 3
73
31
20. 9
4.2.2.3.2 Effect of LMPD Across the Membrane (cases 14-16)
The performance of the OITM relies on the O2 partial pressure difference across
the OITM membrane. Similar to the LMTD used in a heat transfer process, a
LMPD is a key design parameter for a mass transfer process. The LMPD can be
used to compare the design size of different options.
An increase in the O2 level to the boiler increases the sweep gas outlet O2 partial
pressure resulting in a reduced LMPD as shown in Table 4.2.3 (compare Alstom
to case-11). Similarly an increase in sweep gas total pressure reduces the LMPD
as shown in Figure 4.2.6. The effect is fairly small because of the limited
operating pressure variation of the sweep gas.
LMPD, psia
Figure 4.2.6 – Effect of Sweep Gas Pressure on LMPD
20
18
16
14
12
10
8
6
4
2
0
0
5
10
15
20
25
Sweep gas pressure, psia
30
The air side pressure directly affects the OITM performance. Increasing the air
side pressure (OITM operating pressure) will:
(1) Increase LMPD, and so reduce OITM size
(2) Increase compressor discharge temperature (CDT), and require less heat
from the boiler
(3) Reduce turbine exhaust temperature (TET), and so release less low grade
heat and reduce the optimum O2 recovery efficiency
(4) Increase equipment thickness
55
In Case-14 the OITM operating pressure is raised from 200 psia to 250 psia with
constant coal feed rate and furnace air heater duty. This 250 psia OITM pressure
requires the compressor discharge pressure to be increased from 214 to 265
psia, which raises the compressor discharge temperature. Thus, for the same
operating temperature the OITM needs less heat per unit mass of air, and so the
air to the OITM is increased to balance the heat released from the boiler.
Because the amount of the O2 required is fixed, the O2 recovery is adjusted for
the increased air flow through the OITM. Figure 4.2.7 shows that raising the air
side pressure from 200 psia to 250 psia results in a small reduction in system
efficiency (from 35.8 to 35.6%) and an attendant small increase in the CO2
removal specific power penalty (from 46 to 48 kWh/klbCO2). The system efficiency
decreases with increased OITM pressure because of less heat being carried to
the OITM per unit air. However, due to the increased air flow to the OITM, the
LMPD is increased from 16.5 to 24.9 psia, which results in a decrease in OITM
size of 34%.
Opposite to the case-14, cases 15 and 16 were run with reduced OITM
pressures of 180 psia and 190 psia, respectively. As shown in Figure 4.2.7, lower
OITM operating pressure (for a given operation temperature) increases system
efficiency because more heat is transferred from the boiler to the OITM cycle per
unit mass of air. However, as the OTM pressure is reduced the LMPD decreases
requiring a larger OITM size. In this scenario, the oxygen recovery is adjusted to
changes on the airside pressure, which results in an indirect relationship between
system efficiency and the O2 recovery as shown in Figure 4.2.8. Figure 4.2.8
presents a linear correlation for which each percent change in O2 recovery
causes a 0.03 percent point change in efficiency.
Figure 4.2.9 shows the effect of OITM pressure on LMPD when the O2 recovery
efficiency is constant and O2 flow rate is variable (as in Figure 4.2.7). This is
compared in Figure 4.2.9 with the effect of OITM on LMPD when the O2 flow rate
is constant and the O2 recovery efficiency is variable (as in Figure 4.2.7).
The selection of the OITM operating pressure is a trade-off between the cost of
OITM and the system efficiency. If the OITM cost is not too high in future
commercial application, the optimum OITM operating pressure will be relatively
low. Furthermore, a higher OITM operation temperature will be better for the
system efficiency. However, the magnitude of this temperature is constrained by
material limitations.
56
Figure 4.2.7 – Effect of OITM Pressure on System Efficiency and LMPD:
Variable O2 Recovery
38
30
eff, %
20
36
Eff
15
35
10
34
LMPD, psia
25
LMPD
37
5
33
100
150
200
0
300
250
OTM pressure, psia
Figure 4.2.8 – Effect of O2 Recovery on System Efficiency
efficiency, %
35.9
35.8
35.7
35.6
35.5
76
78
80
82
84
O2 recovery, %
57
86
88
90
LMPD, psia
Figure 4.2.9 – Effect of OITM Pressure on System Efficiency and LMPD:
Constant and Variable O2 Recovery
35
30
25
20
15
10
5
0
100
150
200
250
300
air pressure, psia
350
400
O2 recovery constant, O2 flow variable
O2 recovery variable, O2 flow constant
4.2.2.3.3 Effect Of Compressor Discharge Temperature (case 17)
Another way to boost the system efficiency is to transfer more heat to the gas
turbine by increasing the temperature difference between the compressor
discharge temperature (CDT) and the turbine inlet temperature (TIT). Increasing
the TIT results in an increase in OITM operating temperature, which may be
restricted due to material limitations. Without increasing the pressure, the
increase of TIT will increase turbine exhaust temperature, which results in more
low-grade heat that can be recovered in the HRSG. Another approach is to
reduce CDT by the use of a compressor with inter-stage cooling. This will reduce
compressor power, and let more heat be transferred from the boiler to the
turbine, but it releases more low-grade heat from inter-stage cooling. If this heat
cannot be recovered, system efficiency would be reduced. As discussed before,
for the present configuration and integration, there was no margin to recover
more low-grade heat. Therefore the inter-stage cooler will be applicable only for
co-generation of heat and power, where low-grade heat could be recovered by
low-pressure steam export.
Case-17 employs an alternative method to reduce compressor power by interstage water quench to avoid the need for low-grade heat recovery. This quench
(20 klb/hr of water) reduced the CDT from 743 to 698ºF, and therefore more heat
flowed from the boiler to the gas turbine. As a result, the coal to boiler increased
from 377.7 to 389.4 klb/hr, and the corresponding air to the OITM increased from
3945 to 4067 klb/hr (for the same OITM O2 recovery efficiency). The power from
58
the GT increased from 26.7 to 35.9 MWe, and the net power increased from 462
to 472 MWe. However, the system net efficiency was reduced from 35.80 to
35.45%, and the corresponding net penalty for CO2 removal increased from 46 to
49 kWh/klbCO2 because of efficiency loss. The great benefit from this option was
the 10 MWe net power gain. Note also that the furnace air heat duty increased
from 974 to 1060 MMBtu/hr.
4.2.2.3.4 Effect of Furnace Flame Temperature (case 18)
Increased furnace flame temperature increases heat transfer, especially for
radiant transfer, and so it will reduce the furnace size for both for the waterwalls
and air heater. Similar to the effect of excess air in the boiler, higher flame
temperature slightly increases system efficiency because of less flue gas flow out
of the system.
Higher flame temperature cases have been evaluated for the cryogenic ASU O2PC, as reported in Section 4.1.4.3. For the cryogenic ASU O2-PC raising the
equilibrium temperature from 3830ºF (case 6) to 4182ºF (case 7) increased
system efficiency about 0.15% in point, and reduced specific power penalty for
CO2 removal from 114 to 112 kWh/klbCO2 as shown in Table 4.2.4.
Case 18 (Figure 4.2.10) was generated from case 11 by raising the equilibrium
temperature to nearly the same equilibrium temperature as case 7. Table 4.2.4
shows that for the OITM ASU O2-PC raising the equilibrium temperature to
4169ºF increases system efficiency by about 0.3% in point, and reduces the CO2
removal specific power penalty from 46 to 42 kWh/klbCO2. It is clear that
increased flame temperature can reduce equipment size and slightly improve
system efficiency, but could be limited by material cost in the furnace.
Table 4.2.4 – Effect of Flame Temperature on Performance
Case
6
4321
3830
146
31.94
114
Adiabatic Flame T, F
Equilibrium flame T, F
Recycle flue gas T, F
system eff, %
kWh/klbCO2
59
7
5178
4182
163
32.07
112
11
4196
3770
216
35.80
46
18
5120
4169
249
36.11
42
Figure 4.2.10 - O2-PC with OITM with 58% Recycle Flow
Case: O2-PC-SC-18,
ST=4035/1076/1112/2.0"Hg=460 MW
06/19/2006
P ressur e ( p si)
1083
Mass Flo w Rate (Mlb/h r )
1076
4344.0
4035.0
2873
1113
1112
1112
849.0
823.2
823.0
2483
48
PC Boiler
2483
2950
2406
MAIN
DPHRH
19
2878
27
HP2
HP1
Q=135
925
47
1831
2873
1472
PRPB
Q=36
14.7
2873
43
14.7
FRH
2681
2681
G5
G6
14.7
835
658
750
4655.0
890.0
1305.3
2730
2483
72
14.7
751
23
836
G8
SL2
Q=296
830
G12
199
Q=344
519
199
4954.0
56
35
219
EL3
654
EL4
40
644
17
93.8
152
10.0
3.9
5
57
0
SSR
60
445
445
4979.0
2873
92
545
545
382
398
93.8
93.8
38.9
36.9
0
0
0
0
25
234
376
644
15.2
10.0
15.8
59
20
0
0
4929.0
14.6
G13
32
Q=162
0
20
639
641
0
485.0
69
0
70
0
376
639
15.2
485.0
0
FWH3
3.7
71
1
0
Q=-2250
FWH2
73
COND
14
FWH1
61
102
16
418
2873
450.0
DEA
Q=0
Q=0
Q=0
7
Q=0
102
1436
Q=-0
1.0
1436
HR-B
2393
66
FWV
BOILER
DAMP
HEA TER
G14
592
14.5
4864.0
4812.0
2681
2873
2730
34
Q=82
G2
0
68
G15
AAHX
14.3
5120
321
321
14.1
14.3
2644
2681
2681
G16
G17
726
14.7
2682
FLAME
IDFAN
NOZ
Q=-272
RSTOIC
249
337
A2
A5
A3
60
C1
ASH
249
14.7
0
925
920
W=1
A1
BYP
936
FGD
RECY
COL1
925
393
220.0
220.0
912
18
19
C15
C14
59.4
W=13
1718
C5
Q=-287
W=13
936
233.6
C7
C12
C13
Q=3
COMP3
309
W=13
C10
912
920
56
916.2
Q=-6
78
63
Q=-1
925
C2
249
90
14.8
Q=0
14.8
1525
600
90
Q=-60
90
14.9
D1
182
Q=-48
17
231.6
90
D4
1
COL4
56
D6
916.2
Q=-114
894
D3
PIPEL
6
D5
COAL
W=1
Q=0
GRH
RY2
RY1
60
3000.0
816
231.6
58.4
D2
90
COL3
W=-0
TSA
308
C11
926.2
C8
COMP4
RH
DEM
COMP2
310
SEP1
19
56
96
913
14.9
700
896.2
916.2
231.6
1718
56
916.2
19
90
C6
C3
COL2
SA
H4
393
231.6
15.4
SPA
AIRB
58.4
90
15.1
258
60
14.7
SAFAN
MILL
337
37
3167
H1
90
90
14.8
COMP1
PA
600
16.7
375
GBY
14.1
16.7
Q=0
1536
321
600
PAFAN
14.7
3167
Q=970
90
14.9
2644
15.1
600
200
Q=-466
191.6
HX
781
90
15.1
16.9
14.7
625
A4
H2
O2
337
975
1652
Q=0
278
925
16.6
A6
3948
16.0
SPF
W=3
15.3
236
OTM
3930
W=0
G21
625
199.6
1652
0
LIQ
G18
2681
A0
14.7
0
15.1
HR-A
1652
SLV
0
3167
W=-221
ASU/OTM
214.6
100
2644
ESP
14.7
G1
G20
337
12
695
0
782
15.3
H3
GT
W=194
L2
14.7
337
20.0
NH3
3930
743
957
60
SP1
AIR
0
GC
15.1
E1
14.7
14.7
SORB
130
EQ
RGI BBS
W=1
POND
PREC
FGD
QOTM
470.0
DENOX
60
0
100
-970
400
2681
0
53
BYFW
60
0
14.4
14.7
CP
29.7
2730
695
41
4169
STK1
Q=161
4847.0
2730
50
52
14.7
130
14.7
957
567
35
WETECO
470.0
SPQ
604
4795.0
Q=463
100
270
LECO
Q=-2931
W=17
64
567
801
0
102
500.0
86
BFP1
36.9
500.0
500.0
957
485.0
FWH4
342.6
80
102
0
93
430
Q=232
2873
751
639
94
1
V1
V2
398
93.8
87
863.3
74
445
161
485.0
455
83
727
14.7
35
15
568
81
639
77
V3
545
Q=0
342.6
2873
144
748
14.6
1
579
78
V4
871
FWH7
Q=41
1769
9
FSPLI T
0
VD
527
863.3
65
FWH8
829
QB
EL2
95
V6
746
567
G11
1.0
2168
1.2
AUX
18
1266.2
63
4847.0
Q
102
2168
W=-18
89
751
2185
Q=1508
93.8
30
108
3.7
V7
14.7
2168
76
247
199
SPRAY
2681
3.9
6
890.0
11
1266.2
2730
29
EG
M ULT
LP4
152
2168
SL3
658
529
1305.3
UECO
24
G3
28
219
713
RHB
EVAP
EI1
38.9
545
2354
EI2
194.8
2483
46
2730
8
871
21
W
LP3
234
15.8
14.6
Q=-36
4690.0
42
93.8
2354
353.2
4620.0
45
934
PW
6
2483
750
1769
882
857.0
2389
877.0
855
Q=103
WE
LP2
382
96
1454
4610.0
P ow er ( MW )
W
EI3
657
PSH
2873
2679
SL6
G9
Q=366
W
LS4
SPG
14.7
Du ty ( MMBtu/h r)
LP1
540
DPCRH
SC1
Q=-32
IP3
713
194.8
EH2
14.7
G10
G4
Q=321
3
2878
IP2
871
353.2
22
1374
Q=-19
G7
IP1
653
890.0
EH1
SC4
2681
Q=-16
2681
DIVW
44
SC23
1866
FSH
1454
88
84
970
Q
SL5
ROOF
4520.0
502
P ow er ( MW )
17
2389
4035.0
67
823.0
823.0
1076
Du ty ( MMBtu/h r)
1112
1112
LS1
26
4555.0
Tem per atu r e ( F )
CPUMP
W=3
4.2.3
O2-Fired PC Integrated with CAR
4.2.3.1
Overview
The ceramic auto-thermal recovery (CAR) process [5, 6] is based on sorption
and storage of oxygen in a fixed bed containing ionic and electronic conductor
materials operated at high temperature and increased pressure. The stored
oxygen is then released by pressure reduction using sweeping gas, such as
recycled flue gas, or steam extracted from low-pressure section of steam turbine
as shown in Figure 4.2.11. The continuous operation is obtained by employing
multiple beds in a cyclic way, which is similar to the Pressure Swing Absorption
(PSA) process. A large vessel is provided to provide a five second buffer time to
smooth out any fluctuations in either flow and/or composition caused by a batch
adsorption-desorption operation cycle. An important feature of the CAR process
is in that it can be tailored to produce low-pressure oxygen at the concentration
required for O2-fired combustion by using recycled cleaned flue gas as a sweep
gas. The CAR process is based on conventional sorbent bed adsorption that is
easy to fabricate and readily available. The scaling up for such a process is
similar to the PSA process, and has fewer challenges than the OITM technology.
Figure 4.2.11 - CAR Process Schematic
steam/CO 2
N2
regenerator layer
perows kite pellets
lay er
regenerator layer
air
s team /C O 2 and O 2
In general during adsorption, heat is released and system temperature is raised.
The CAR adsorption process has to be operated at high temperature, therefore
the air is preheated to a certain temperature before being fed to the adsorption
bed. The heat generated during adsorption can be recuperated to heat up the
61
fresh air to reduce heat loading. For this purpose, a recuperator is employed to
transfer heat between exhaust oxygen depleted air and the fresh air, where the
heat generation during adsorption raises the exhaust air temperature. As result,
the fresh air needs less preheating (i.e. 1020ºF) for the CAR process, compared
to 1650ºF for the OITM process. On the other hand, the desorption process is
associated with heat absorption, where heat has to be transferred to the bed
during oxygen release. This is done by direct injection of natural gas into bed to
combust with oxygen and by pre-stored heat in the sorbent bed during
adsorption. The CAR process is so designed that it also stores the energy in the
bed during the heat release in the adsorption cycle and releases this heat during
the stripping cycle, by means of installation of inert layers on both ends of
sorbent bed for heat storage. In this so called “auto-thermal recovery”, more heat
is stored in bed, less heat is carried out by oxygen-depleted air, and less heat is
required from fuel combustion to strip oxygen.
High oxygen concentration can be obtained by steam striping providing that the
steam is condensed out downstream. Oxygen concentration is limited to about
30-40% if the flue gas is applied as a sweep gas. As reported in the BOC study
[5] with steam as sweeping gas, the air-to-steam molar ratio is about 2.66 with O2
recovery over 90%, which leads to an oxygen concentration in the sweeping gas
of about 33-36%v before steam condensing.
Simulated integration of the CAR process to the O2-PC has been reported by
BOC to determine the technical and economical feasibility [5]. The air-fired
reference plant is an existing ultra-supercritical lignite-fired 865 MWe Lippendorf
power plant near Leipzig, Germany. A simulation of an oxyfuel power plant with
cryogenic air separation was employed for comparison. The same plant was then
applied for integration study with CAR process, where low-pressure steam
extracted from steam turbine was applied as a sweep gas. Table 4.2.5
summarizes the key results.
Table 4.2.5 - Comparison for CAR with Cryogenic ASU
Case
Net Power, MWe
Net efficiency, %(LHV)
Efficiency drop, % point
Air
865
42.6
-
ASU
687
33.3
9.3
CAR
726
34.0
8.6
Comparing the CAR to the cryogenic process, the efficiency drop reduced from
9.3 to 8.6% points, and the net power increased from 687 to 726 MWe. As
reported, it is clear from the results of study that the steam consumption is a
critical variable for this option. The steam extraction required for oxygen stripping
is about 200 kg/s in comparison with the main steam flow as 692 kg/s.
The reason recycled flue gas was not used as sweep gas was that the flue gas
has to be cleaned up to avoid any contaminates to the sorbent bed, and the
62
effects of contaminates have not yet been studied in detail. In general, to be
economic, recycle gas clean up should be avoided because the clean gas will be
sent back to boiler, where it mixes with combustion gases to become dirty gas
again. Optimally, the gas clean up process should be applied to the flue gas
exiting from system and flowing to the CO2 plant. Moreover, this stream has
much less flow to be treated than does the recycled flue gas.
Since the CAR process is based on swings in the partial pressure of oxygen, it
will be affected by pressure, and LMPD (logarithmic mean pressure difference).
Increased adsorption pressure will enhance oxygen adsorption from air, and a
low desorption pressure will favor adsorbed oxygen release. However too high a
pressure will require more power for air compression. If steam is used as a
sweep gas, heat carried by the sweep gas can be recovered through feedwater
heating during steam cooling and condensing. If flue gas is used as a sweep gas,
hot sweep gas can be directly fed to boiler.
In order to explore the advantage of integration with CAR process for O2-fired
combustion, an Aspen simulation has been made for the CAR process operated
with recycled flue gas sweep gas. The difference in principle between steam and
flue gas sweep gases is their pressure ratios, where steam extracted at 1.6 bar
could continuously expand to a pressure as low as 0.038 bar to generate more
power with a pressure ratio of 1.6/0.038, while the flue gas has to be compressed
from about 1.0 bar to 1.6 bar to consume power with pressure ratio as 1.6/1.0.
Thus, the substitution of flue gas by extracted steam reduces power because the
low pressure ratio of the flue gas is replaced by the high pressure ratio of steam
for the same amount of gas volumetric flow. The reduced power can be
calculated by difference between power from steam expansion and power for
recycle gas compression, without including the changes in auxiliary power and
low grade heat integration. The net result is shown by Table 4.2.6.
Table 4.2.6 – Comparison for CAR Using Different Sweep Gases
Oxidant
Air Separation Method
Sweep Gas
Net Power, MWe
Net Efficiency, % (LHV)
Efficiency Drop, % point
Air
865
42.6
-
Cryo
687
33.3
9.3
O2
CAR
steam
726
34.0
8.6
CAR
flue gas
767
35.9
6.7
As can be seen in Table 4.2.6, the system efficiency increases about 1.9% points
when steam is replaced by recycled flue gas as the sweep gas. As compared
with the cryogenic ASU, the CAR process with gas recycle sweep gas has an
increased system efficiency of 2.6% in points, which is close to 3.2% points
achieved by the OITM. It is obvious from the standpoint of system efficiency that
the future of the CAR process is to use recycle flue gas as a sweep gas.
63
Note that in the CAR process, natural gas is fired to provide a portion of the heat
for stripping out the adsorbed oxygen from sorbent bed, which is similar to the
gas absorption-regeneration cycle. Because of auto-thermal recovery process,
the ratio of energy input from the natural gas to coal is only about 3.2%, which is
much less than the heat requirement by the OITM process.
4.2.3.2
Optimized Integration of CAR in O2-PC
Advanced oxygen separation CAR process has been integrated into the oxygen
fired PC as shown in the Figure 4.2.12 schematic. A multi-bed adsorptiondesorption batch operation cycle is simulated by a quasi-continuous operation of
moving bed system, where the oxygen sorbent is circulated between two beds.
The Aspen model of the O2-PC CAR power plant is presented in Figure 4.2.13.
Flue gas, instead of steam extracted from steam turbine, is used as a sweep gas
to strip out oxygen from desorption bed. Note that flue gas contains many
contaminants such as fly ash, SOx, NOx, etc. that may be harmful to the
perovskite sorbents used in the CAR process and future experimental studies will
be required to quantify this effect.
The flue gas from the boiler transfers heat to the oxygenated flue gas in an air-toair heat exchanger and then is further cooled down to 150ºF by feedwater
economizers. After cooling, the flue gas is boosted to a pressure to overcome the
pressure drop across the CAR oxygen desorption equipment. Exit gas from
desorption bed, containing about 25-45% of oxygen, is then cooled down to
280ºF through the sweep gas recuperator. This oxygen containing flue gas is
then sent to air-to-air heat exchanger to be heated up to 625ºF, and then fed to
boiler.
Air is compressed to a pressure of 30 psia and then fed to the adsorption bed
after pre-heating in the air recuperator. About 90% of O2 in the air is adsorbed by
solids in the bed. The heat released from adsorption increases the system
temperature. Adsorption bed exit gas or vitiated air is cooled down to 280ºF in
the air recuperator before being discharged to the stack. The solids circulated
between two the moving beds adsorb and release oxygen nearly reversibly.
The system net efficiency for the CAR integrated with the O2-PC power plant is
35.2%. Heat input from natural gas is about 2.7% of the total heat input. Table
4.2.7 lists a comparison among different air separation technologies in efficiency
drop and CO2 energy removal penalty. OITM has less energy penalty than CAR
due to the process gain associated with the gas turbine expander (see 4.2.2.1).
Table 4.2.7 – O2 PC Efficiency with Different O2 Separation Techniques
Air separation method
Efficiency Drop
% pts.
CO2 removal energy
kWh/klbCO2
64
Cryo
6.5
99
OITM
3.3
42
CAR
4.3
64
Figure 4.2.12 – Integration of CAR Process into O2-PC
65
Figure 4.2.13 - O2-PC with CAR
Case: O2-PC-SC-19, 07/19/2006
1083
48
4822.0
832
147
4647.0
Mass Flow Rate (Mlb/hr)
4035.0
2950
569
1113
1112
1112
849.0
823.2
823.0
2406
2406
2406
19
27
Duty (MMBtu/hr)
2950
1112
LS1
67
SL5
2878
HP2
HP1
974
43
4462.0
Q=134
2022
1559
1508
14.7
14.7
2156
2156
2156
G5
G6
14.7
FRH
837
PRPB
Q=36
4612.0
44
658
750
890.0
1305.3
2406
2156
IP3
2389
SL6
22
LP1
382
2301
38.9
8
545
2190
B2
G11
532
21
657
EI2
EI1
219
1893
EL2
199
QB
871
713
353.2
194.8
18
56
88
2156
92
545
268
382
93.8
38.9
36.9
62
62
49
111
199
65
4929.0
522
4954.0
181.2
342.6
386
610
4979.0
66
15
FWH6
94
569
569
4822.0
68
695
60
14.5
20.0
86
17.5
1115
29.7
AIR
14.7
SORB
97
2156
0
G36
A8
1022
30.0
29.0
3165
3165
HX
A1
123
98
14.7
14.7
SLV
290
15.1
15.1
2126
1284
COOL
G19
G23
90
90
1190
14.9
110
W=7
C6
290
ESP
14.4
2156
Q=0
918.5
918.5
0
0
90
303
90
15.1
14.9
842
842
C11
232.1
94
Q=75
SEP1
W=11
PIPEG
DEM
COMP2
COMP1
C5
COMP3
310
312
59.5
234.1
W=11
781
4
ASU/CAR
90
770
SPA
2510
232.1
NOZ
A4
300.0
PAFAN
772
C3
Q=-1
NG
150
Q=-142
W=2
RSTOIC
G24
769
1849
A6
Q=-1
21.5
VA
58.5
G22
781
2157
986
1190
59
280
Q=883
14.9
14.9
G21
G18
DES
22.0
1190
SPF
IDFAN
AAHX
PV3
A2
PV1
150
W=0
2126
14.7
PV4
PV5
0
LIQ
303
SPO
SPVA
ADS
POND
L2
G20
14.2
2156
FLAME
BYFW
203
W=32
PREC
FGD
53
CP
Q=170
22.5
23.0
2156
1.0
2229
50
64
W=1
3165
0
303
5120
102
2510
0
COMP
15.1
1115
Q=49
102
1124
1849
60
242
G29
Q=72
14.7
60
0
0
0
1294
102
16
1115
695
14.7
Q=-1897
FWH1
FWH2
61
2950
303
14.4
16.5
1115
14
215
WETECO
14.7
14.7
STK1
254
EQ
16.5
173
70
1115
71
236.6
100
123
206.6
NH3
457
221.6
236.6
W=18
FGD
1115
SPQ
DENOX
625
15.2
Q=53
320
156.6
0
625
206.6
FWH3
COND
3.7
111
374
DEA
0
G16
150
146
995
34
SP1
215
V1
214
BFP1
108
33
2802
RGIBBS
77
V2
211
73
Q=66
FWV
2950
14.6
2156
LECO
G15
1
112
181.2
35
G2
20
173
62
7
Q=118
Q=311
52
816
14.7
644
10.0
610
BOILER
Q=74
9
156
15.2
36.9
1115
FWH4
93
Q=-0
Q=-2088
69
87
342.6
4839.0
G13
15.8
59
78
191.6
62
1115
396
85
74
Q=162
877.0
12
Q=210
5
40
262
258
93.8
110
191.6
16
591
3.9
657
L1
C9
PV2
5
E2
Q=1485
2802
10.0
SSR
234
V3
545
Q=0
316
2950
2950
FWH7
83
2045
4787.0
152
57
W=-18
221
95
425
2950
32
RHA
HEATER
25
V4
712
430
863.3
14.6
G12
110
VD
V5
527
522
FWH8
1145
2045
40
644
93.8
93.8
80
V6
746
63
Q=138
867.0
EL4
17
FSPLIT
435
863.3
522
1266.2
2950
EVAP
219
EL3
6
323
89
V7
2406
4904.0
41
1.0
1791
3.7
877.0
23
45
Q
102
1831
657
B1
2156
30
108
654
76
890.0
199
361
2156
11
1266.2
877.0
55
PSH
820
3.9
1.2
658
750
14.6
4527.0
152
AUX
569
G1
29
SL3
SL2
Q=106
2802
14.7
MULT
LP4
234
15.8
1942
LS4
2406
Q=289
2450
EG
EI3
DPCRH
1353
812
W
LP3
24
49
E1
28
96
1305.3
4682.0
WE
LP2
540
93.8
EH2
RHB
869
863
4155
Power(MW)
857.0
Q=329
G3
W
W
93.8
DIVW
2802
Duty (MMBtu/hr)
194.8
2045
SC1
Q=-25
G4
42
Q
713
6
Q=-36
4472.0
2802
IP2
871
353.2
2679
2878
857.0
G8
14.7
1021
3
72
2802
910
SC23
2057
FSH
IP1
653
890.0
EH1
47
SPRAY
Q=479
88
84
60
81
471
17
2389
4035.0
Power(MW)
823.0
823.0
1076
MAIN
DPHRH
26
1112
2802
ROOF
2950
Pressure (psi)
1076
4286.0
PC Boiler
Temperature (F)
ST=4035/1076/1112/2.0"Hg=460 MW
C7
C8
323
COMP4
928.5
W=10
772
C10
769
247
16.4
A5
COL1
G17
290
625
308
14.2
16.5
A3
555
ASH
30
90
Q=-111
280
280
16.5
16.5
1849
1849
MILL
ECOW
G28
G27
COAL
Q=0
90
COL2
60
16.7
Q=0
66
98
Q=-1
C2
C1
862
Q=0
D1
90
Q=-48
14.9
61
COL3
Q=-42
9
90
D4
1
Q=-116
55
3000.0
918.5
769
769
232.1
58.5
D2
COL4
232.1
D3
CPUMP
PIPEL
2
D5
CO2
W=3
4.2.4
Comparisons
CO2 cannot be captured and sequestrated without incurring an energy penalty
because of the potential energy stored in the pressurized liquid CO2. A minimum
of 40 kWh/klbCO2 additional auxiliary power is required for CO2 compression. The
difference between technologies lies in the difference in power requirements of
the different CO2 or O2 separation techniques, and the process gain when
advanced power generation is integrated, such as power generation from OITM
through hot gas expansion.
Table 4.2.8 shows that compared to CAR, the OITM results in higher system
efficiency and significantly more in power because of the process gain of the hot
compressed air expanding through the gas turbine (gains are in reference to the
cryogenic ASU O2-PC). When the OITM technology is integrated with O2-fired
combustion, a conventional Rankine cycle power plant is upgraded to a
combined cycle power plant. This improvement makes the OITM technology
attractive for economic CO2 removal.
Table 4.2.8 - Gains from OITM and CAR Compared to Cryogenic ASU
Efficiency gain Power gain
(% points)
(MW)
3.2
117
0.7
20
2.6
50
OITM
CAR/steam sweep gas
CAR/flue gas sweep gas
The OITM faces more technical challenges than does the CAR process because
it operates under higher pressure and temperature. In addition to the OITM
development itself, the integration and design of the boiler air heater presents a
challenge (see Section 4.3.2.2.3). An alternative to the furnace air heater is to
provide the air heating by a natural gas duct burner, although the high heat
required (the ratio of heat absorbed by OITM to heat input by the coal is 22%)
may preclude duct firing. For the CAR process, the heat input from natural gas is
only 2.7% to the total energy input.
In general, CO2 can be removed by means of:
Post-combustion capture - Amine process or other
Pre-combustion capture - IGCC
Oxygen-combustion – cryogenic ASU, OITM, CAR
Table 4.2.9 compares the efficiency drop and specific power requirement of the
various CO2 removal technologies. Similar to the O2-PC designs employing
67
cryogenic ASU and OITM, the O2-PC CAR power plant is designed to include
liquid CO2 pumping, wet-end heat recovery, increased flame temperature, and
hot gas recycle.
Table 4.2.9 – Comparison of CO2 Removal Technologies
PC
Boiler Type
Removal Technique
Efficiency drop
CO2 removal penalty
post comb.
% ponts
kWh/klbCO2
11.6
188
O2 fired
cryo. ASU OITM
6.5
3.3
99
42
IGCC
pre comb.
CAR
4.3
64
From Table 4.2.9, it is clear that the O2-fired PC integrated with advanced oxygen
separation technology has significant advantages over both the post combustion
and pre-combustion CO2 techniques since the separation of oxygen from air is a
physical process and involves less energy than the chemical separation of CO2
from flue gas or syngas.
68
7.2
116
4.3
Furnace and Heat Recovery Area Design and Analysis
4.3.1 Furnace Design and Analysis
4.3.1.1
FW-FIRE Computer Program Description
FW-FIRE (Fossil fuel, Water-walled Furnace Integrated Reaction and Emission
Simulation) simulates furnace combustion, heat transfer and pollutant formation
based on fundamental principles of mass, momentum, and energy conservation
[9]. FW-FIRE is an extended and enhanced version of PCGC-3, which was
developed over a period of ten years by the Advanced Combustion Engineering
Research Center (ACERC), operated jointly by Brigham Young University and
the University of Utah. The FW-FIRE computer program incorporates the latest
state-of-art coal combustion/gasification, pollutant formation, and physical
analysis techniques based on extensive empirical research.
The FW-FIRE code performs general three dimensional multiphase gas
combustion steady state analysis of reactive fluid flows. The program is fully
capable of analyzing gas-fired and coal-fired boilers although FW-FIRE was
initially tailored for pulverized coal combustion and gasification.
The FW-FIRE program models the gas flow field as a three dimensional
(Cartesian or cylindrical) turbulent reacting continuum that is described locally by
the Newtonian form of the Navier-Stokes equations coupled with the energy
equation and other appropriate constitutive equations. These equations are
solved in Eulerian framework to predict gas properties such as pressure,
temperature, velocity, and pollutants and other species concentrations.
The Reynolds stress terms, which result from Favre-averaging of the
conservation equations, are approximated using the Boussinesq assumption and
effective eddy viscosity. The value of the eddy viscosity and subsequent closure
of the turbulence equations is made using either a linear or non-linear k-ε twoequation model. The effects of turbulence of the flow field on the combustion
reactions are included.
The turbulent flow field is also coupled with the combustion chemistry. Since
gaseous reactions are limited by mixing rates and not reaction kinetics, the
process chemistry is calculated using locally instantaneous equilibrium based on
the degree of mixing of the species. Rate constants for processes such as
devolatilization (two step process) and char oxidation are built-in to the program
based on empirical testing.
A Lagrangian model of the particle conservation equations is used to predict
particle transport by characterizing the particle field as a series of discrete
particle trajectories through the gas continuum. Particles interact with the gas
69
field via mass, momentum, and energy exchange. Particle properties such as
burnout, velocities, temperatures, and particle component compositions are
obtained by integrating the governing equations along the trajectory paths. The
program possesses the capability to input a particle size distribution and
chemical composition.
In a pulverized coal flame, the radiation field is a multi-component, non-uniform,
emitting, absorbing, gas-particle system. The coal particles cause anisotropic
and multiple scattering, and the flame is surrounded by non-uniform, emitting,
reflecting, absorbing surfaces. The radiation field calculations are based on an
energy balance on a beam of radiation passing through a volume element
containing the absorbing-reflecting-emitting medium. An Eulerian framework
using a discrete-ordinates approach is used to model this process. Heat transfer
via radiation and convection to waterwall and tube banks is determined by
specifying a local wall temperature and emissivity.
The set of non-linear differential equations is discretized and combined by a
upwind and weighted central-differencing scheme. The resulting gas flow field
finite difference equations are solved using variations of the SIMPLE/SIMPER
algorithm.
FW-FIRE contains a sub-model for the prediction of nitrogen pollutant emissions.
This sub-model has the capability of predicting both fuel and thermal NOx
formation. Fuel NO formation can proceed directly to N2 and NO (such as in the
case of char) or through HCN and NH3 which are oxidized to form NO and
reduced to N2 (such as in the case of volatiles). Global reaction rates are based
on work by de Soete and Bose. Thermal NO formation is governed by the
extended Zel=dovich mechanism.
4.3.1.2
Model Geometry
A simulation was made for both the reference air-fired case and for the oxygenfired case. The FW-FIRE model simulates the furnace, in height from the bottom
of the hopper to the roof, in depth from the front wall to the rear wall, and in width
from the left side wall to the right side wall. Furnace partial division walls are also
included in the model. Finer meshes are used to model the burners and over-fire
air (OFA) ports. The air-fired model contains 528,840 (117x113x40) nodes and is
shown in Figure 4.3.1. The oxygen-fired model contains 484,160 (136x89x40)
nodes and is shown in Figure 4.3.2.
70
4.3.2 Boundary Conditions
Boundary conditions are based on ASPEN simulations of the power plant [Sect.
4.1 and Sect. 4.2]. The air-fired, oxygen-fired with ASU and oxygen-fired with
OITM ASPEN reference cycle diagrams are presented in Figure 4.1.1, Figure
4.1.9, and Figure 4.2.10, respectively The input data required by FW-FIRE
include fuel analysis, coal particle size distribution (mass percentage for each
size bin), waterwall fluid temperatures, and the velocities, flow rates and
temperatures of primary and secondary gas streams. Boundary conditions are
detailed in Figure 4.3.3, Figure 4.3.4, and Figure 4.3.5
The waterwalls of the furnace are assumed to be gray and diffusive. The wall
temperature at each location is calculated based on the fluid temperature and the
heat flux at the wall cell.
For coal devolatilization kinetic properties, Ubhayakar rate parameters were
employed for bituminous coal. For bituminous char oxidation, Sandia kinetic and
burning mode parameters were applied for Illinois #6 coal.
The selected quantity of flue gas recycle produces a 600ºF higher equilibrium
temperature in O2-firing than air-firing. This may increase the potential for
slagging in the furnace depending on the ash fusion characteristics.
Consequently, for a dry bottom furnace design a minimum flue gas recycling flow
may be required to avoid slagging depending on fuel type. Alternatively a wetbottom or slag type furnace could be used to resolve ash deposition and removal
problems.
4.3.2.1
Air-Fired Reference Case
The general layout drawing of the air-fired reference case is shown in Figure
4.3.6. The furnace has a total 24 opposed wall-fired burners (3 vertical x 4
horizontal x 2 walls) and 10 overfire air ports. 30% of the total combustion air is
injected through the over-fire air ports located at one burner pitch above the top
burner row. The radiant heat transfer surface consists of 2.75” OD tube
waterwalls and five 2.0” OD tube partial divisional wall panels. Water is circulated
in the furnace by forced circulation.
The boundary conditions were applied to the computational model and FW-FIRE
was run until steady state conditions were achieved. The modeling results are
summarized in Figure 4.3.7. The coal burnout shown in the table is the
percentage of dry ash-free based coal burned. The furnace exit gas temperature
(FEGT) shown in the table is the average temperature of flue gas before the
platen superheater. The energy absorption listed is the total energy absorbed by
water walls and partial division walls prior to the platen superheater. Total
furnace absorption and FEGT predicted by FW-FIRE and ASPEN match closely.
71
Figure 4.3.8 is a plot of the flue gas velocity magnitude in a vertical plane through
the second burner column. It can be seen that the gas velocity near the burners
accelerates to greater than 130 ft/s due to the reduced gas density after particle
ignition. Figure 4.3.9 presents a plot of gas temperature in a vertical plane
through the second burner column. The maximum flue gas temperature is
approximately 3350oF. The mole fraction of O2 through the second burner
column is presented in Figure 4.3.10.
The heat flux at the furnace water wall is shown in Figure 4.3.11. The maximum
heat flux is approximately 70,000 Btu/hr-ft2 and is located on the side wall at the
top of the burner zone. The total heat absorbed by the furnace walls before the
furnace exit is 1770 MM Btu/hr. Figure 4.3.12 displays temperatures of the
furnace walls and roof. The maximum temperature of the waterwalls is
approximately 870oF and of the division walls is approximately 1000oF. Figure
4.3.13 presents the CO concentration at the wall, peaks at approximately 10%
due to the sub-stoichiometric conditions of the lower furnace (without overfire air
the wall CO would be below 2%).
The trajectories of the 72-micron particles are plotted in Figure 4.3.14 with colors
in each trajectory representing the mass fraction of char in the particle. Char is
formed from devolatilization and consumed by oxidation. The maximum char
mass fraction is usually less than the mass fraction of fixed carbon in a proximate
analysis. Figure 4.3.14 shows that all of the 72-micron particles are completely
burned before the furnace exit. The trajectories of 176-micron particles are
plotted in Figure 4.3.15. It can be observed from Figure 4.3.15 that some
particles are not completely burned at the exit of the furnace, causing unburned
carbon in the fly ash. Total burnout of all particle sizes is 99.66% (2.61% LOI).
Average NOx concentration at the furnace outlet is 276 ppmvw (0.38 lb/MMBtu).
4.3.2.2
Oxygen-Fired Design Case
The preliminary size of the furnace heat transfer area was based on a calculation
of average wall heat flux using the Foster Wheeler computer program, EMISS
[11]. The EMISS computer program calculates radiative heat flux of CO2 and H2O
gases. A three dimensional CFD run was then made using FW-FIRE to more
accurately determine the total heat absorption. Based on the CFD results, the
height of the furnace model was adjusted until the total heat absorption
approximately matched that required in the ASPEN oxygen-fired design case.
Figure 4.3.16 shows a comparison between the sizes of the resultant oxygenfired furnace and the air-fired reference furnace. Compared to the air-fired
furnace, the oxygen furnace has only approximately 65% of the surface area and
approximately 45% of the volume. Figure 4.3.17 presents the oxygen-fired
design general layout drawing for the cryogenic ASU design and Figure 4.3.18
for the OITM design.
72
The oxygen-fired furnace has a total 24 opposed wall-fired burners (4 vertical x 3
horizontal x 2 walls) and 8 overfire gas ports. The burner designs (including 0.5
primary air swirl) are based on the subcritical O2-PC burner design [3]. The
radiant heat transfer surface consists of 2.75” OD tube waterwalls and ten 2.0”
OD tube partial divisional wall panels. Water is circulated in the furnace by forced
circulation.
Two designs were simulated: 1) with cryogenic ASU and 2) with OITM. The
designs differ as follows:
Radiant Superheater: In the cryogenic ASU design, radiant superheat
(downstream of the primary superheater) is provided by the partial division
walls, whereas in the OITM design radiant superheat is provided by the
waterwalls above 100’ and the furnace roof. No division wall surface is
required in the OITM design due to the increased coal-firing such that less
evaporator surface (below 100’) is required (in addition, in the OITM
design the overall furnace height is reduced by 5’).
Air Heater: A tubular convective air heater is included in the OITM to
provide the necessary air heating for the membrane separation process.
The modeling results are summarized in Figure 4.3.7. The coal burnout for the
oxygen-fired cases is 100% due to the high furnace temperature and high
concentration of O2. Total furnace absorption and FEGT predicted by FW-FIRE
and ASPEN match well. NOx is reduced by oxygen firing (compared to air-firing)
by about a factor of two from 0.38 lb/MMBtu to 0.18 lb/MMBtu.
4.3.2.2.1 Cryogenic ASU – Full Load
Figure 4.3.19 is a plot of the flue gas velocity magnitude in a vertical plane
through the middle burner column. It can be seen that the gas velocity near the
burners accelerates to nearly 150 ft/s due to the reduced gas density after
particle ignition. Figure 4.3.20 presents a plot of gas temperature in a vertical
plane through the middle burner column. The maximum flue gas temperature is
approximately 3900oF. The mole fraction of O2 through the middle burner column
is presented in Figure 4.3.21. The mixed primary/secondary gas O2 content
(before combustion) is 41%.
The heat flux at the furnace water wall is shown in Figure 4.3.22. The maximum
heat flux is approximately 171,000 Btu/hr-ft2 and is located on the side wall at the
top of the burner zone. This maximum heat flux is approximately 2.5 times the
air-fired case due to the higher flame temperature and higher H2O and CO2
concentrations. The total heat absorbed by the furnace walls before the furnace
exit is 2287 MM Btu/hr. Figure 4.3.23 displays temperatures of the furnace walls
and roof. The maximum temperature of the waterwalls is approximately 1035oF
73
and of the division walls is approximately 1050oF. Because of the higher
temperature of the oxygen-fired case compared to the air-fired case, furnace
water wall material was upgraded from 0.22” thick SA-213-T2 to 0.20” thick SA213-T92. Figure 4.3.24 presents the CO concentration at the wall, which is
significantly greater than for the air-fired case (Figure 4.3.13) and its effects on
corrosion are modeled analytically in Section 4.3.4, but will also need to be
measured empirically in future work.
The trajectories of the 69-micron particles are plotted in Figure 4.3.25 with colors
in each trajectory representing the mass fraction of char in the particle. Figure
4.3.25 shows that all of the 69-micron particles are completely burned before the
furnace exit. The trajectories of 169-micron particles are plotted in Figure 4.3.26.
Note that due to the higher temperature and O2 concentration all the 169-micron
particles are completely burned as compared to the air-fired case (Figure 4.3.15)
where there is some residual unburned char at the outlet.
4.3.2.2.2 Cryogenic ASU – Part Load
Since the boiler is a supercritical once-through sliding pressure unit, part load
operation must be evaluated with thermal/hydraulic and structural criteria. Three
part loads were selected for evaluation: 72%, 50%, and 25% (minimum BENSON
load). Boundary conditions were based on the ASPEN System Analysis. Figure
4.3.27 presents the corresponding recycle flow versus load. The modeling results
are summarized in Figure 4.3.28. Figure 4.3.29 presents the gas temperature in
a vertical plane through the middle burner column for the part load cases. The
heat flux at the furnace water wall is shown for the part load cases in Figure
4.3.30. The average and peak heat flux versus height is presented in Figure
4.3.31.
4.3.2.2.3 Integration of Oxygen Ion Transport Membrane
Figure 4.3.32 is a plot of the flue gas velocity magnitude in a vertical plane
through the middle burner column. It can be seen that the gas velocity near the
burners accelerates to nearly 150 ft/s due to the reduced gas density after
particle ignition. Figure 4.3.33 presents a plot of gas temperature in a vertical
plane through the middle burner column. The maximum flue gas temperature is
approximately 3850oF. The mole fraction of O2 through the middle burner column
is presented in Figure 4.3.34. The mixed primary/secondary gas O2 content
(before combustion) is 39%.
The heat flux at the furnace water wall is shown in Figure 4.3.35. The maximum
heat flux is approximately 180,000 Btu/hr-ft2 and is located on the side wall at the
top of the burner zone. This maximum heat flux is approximately 2.5 times the
air-fired case due to the higher flame temperature and higher H2O and CO2
concentrations. The total heat absorbed by the furnace walls before the furnace
74
exit is 2029 MM Btu/hr. Figure 4.3.36 displays temperatures of the furnace walls
and roof.
The maximum temperature of the evaporator waterwalls is
approximately 1060oF, of the radiant superheater is approximately 1100oF, and
of the air heater is approximately 1800 oF. Because of the higher temperature of
the oxygen-fired case compared to the air-fired case, furnace water wall material
was upgraded from 0.22” thick SA-213-T2 to 0.23” thick SA-213-T92. The air
heater, which constructed from Incoloy MA956 material, is described in detail in
Section 4.3.5.4. Figure 4.3.37 presents the CO concentration at the wall, which is
significantly greater than for the air-fired case (Figure 4.3.13) and its effects on
corrosion will need to be studied in future work.
The trajectories of the 69-micron particles are plotted in Figure 4.3.38 with colors
in each trajectory representing the mass fraction of char in the particle. Figure
4.3.38 shows that all of the 69-micron particles are completely burned before the
furnace exit. The trajectories of 169-micron particles are plotted in Figure 4.3.39.
Note that due to the higher temperature and O2 concentration all the 169-micron
particles are completely burned as compared to the air-fired case (Figure 4.3.15)
where there is some residual unburned char at the outlet.
75
Figure 4.3.1 – Computational Model of Air-Fired Furnace (with right side
wall removed)
Partial division walls
OFA ports
Burners
76
Figure 4.3.2 – Computational Model of Oxygen-Fired Furnace With OFA
Superheater
Partial
division walls
Air Heater
OFA ports
Evaporator
Burners
OITM
Cryogenic ASU
77
Figure 4.3.3 – Air-Fired Boundary Conditions
Coal
Ultimate Analysis
Ash
S
H
C
H2O
N
O
Total
Volatile Matter (daf)
Density
HHV, as received
TCA
Excess O2
OFA
%
%
%
%
%
%
%
%
%
lb/ft3
Btu/lb
lb/hr
%
%
Primary
Inner Sec. Air
Outer Sec. Air
Tertiary Air
Overfire Air - Inner
Overfire Air - Outer
Size (micron)
Distribution
9.99%
9.3
2.51%
33.7
4.50%
71.7
63.75%
121.8
11.12%
175.5
1.25%
Total
6.88%
100.00% < 75 micron
< 150 micron
44.35%
80.0
11,666
3,299,687
11.0
%
%
%
%
%
%
Mass Percent
28.9
27.3 Coal Flow
23.9 Moisture in Coal
10.7
9.2
100.0
%
%
319,000
35,473
70
99
Heat Input
3721.45 MM Btu/hr
Divwall
Inlet
Outlet
Temp (F) Temp (K)
854
730
964
791
WW
Inlet
Outlet
Temp (F) Temp (K)
638
610
810
705
Density
Inner Diam. Outer Diam. Area per Port No. of Ports Axial Velocity Tan./Axial Rad./Axial Coal Flow
in
in
ft2
ft/sec
lb/hr
Velocity
Velocity
20.0%
Flow Rate
lb/hr
%
550,000
16.5
425,044
12.7
1,700,178
51.0
0
0.0
335,664
10.1
324,274
9.7
3,335,159
100.0
Temperature
F
219
580
580
580
580
580
lb/ft3
0.059
0.039
0.039
0.039
0.039
0.039
13.100
23.750
33.875
0.000
0.000
19.500
78
22.750
33.375
49.000
12.100
19.000
27.000
1.887
2.999
6.837
0.799
1.969
1.902
24
24
24
24
10
10
57.1
42.5
74.6
0.0
122.7
122.7
0.00
0.00
0.41
0.00
0.00
0.00
0.00 283,527
0.21
0.41
0.00
0.00
0.21
Figure 4.3.4 – Oxygen-Fired Boundary Conditions (Cryogenic ASU)
Coal
Ultimate Analysis
Ash
S
H
C
H2O
N
O
Total
Volatile Matter (daf)
Density
HHV, as received
TCA
Excess O2
OFA
%
%
%
%
%
%
%
%
%
lb/ft3
Btu/lb
lb/hr
%
%
Primary
Inner Sec. Air
Outer Sec. Air
Tertiary Air
Overfire Air - Inner
Overfire Air - Outer
Size (micron)
Distribution
9.99%
10.7
2.51%
33.7
4.50%
69.3
63.75%
118.2
11.12%
169.2
1.25%
Total
6.88%
100.00% < 75 micron
< 150 micron
44.35%
80.0
11,666
1,809,811
11.0
%
%
%
%
%
%
Mass Percent
29.5
38.2 Coal Flow
25.9 Moisture in Coal
5.5
1.0
100.0
%
%
308,000
34,250
85
99
Heat Input
3593.13 MM Btu/hr
Divwall
Inlet
Outlet
Temp (F) Temp (K)
837
720
974
796
WW
Inlet
Outlet
Temp (F)
593
825
Density
Inner Diam. Outer Diam. Area per Port No. of Ports Axial Velocity
Temp (K)
585
714
20.0%
Flow Rate
lb/hr
%
495,000
26.8
197,420
10.7
789,679
42.8
0
0.0
170,790
9.3
191,172
10.4
1,844,061
100.0
Temperature
F
257
625
625
625
625
625
lb/ft3
0.073
0.046
0.046
0.046
0.046
0.046
in
9.500
16.000
21.750
0.000
0.000
14.000
79
in
15.000
20.750
29.000
8.500
13.500
20.000
ft2
0.735
0.952
2.007
0.394
0.994
1.113
24
24
24
24
8
8
ft/sec
106.3
52.6
99.8
0.0
130.8
130.8
Tan./Axial Rad./Axial Coal Flow
Velocity
Velocity
lb/hr
0.50
0.00
0.41
0.00
0.00
0.00
0.00 273,750
0.21
0.41
0.00
0.00
0.21
Figure 4.3.5 – Oxygen-Fired Boundary Conditions (OITM)
Coal
Ultimate Analysis
Ash
S
H
C
H2O
N
O
Total
Volatile Matter (daf)
Density
HHV, as received
TCA
Excess O2
Drum Pressure
OFA
Primary
Inner Sec. Air
Outer Sec. Air
Tertiary Air
Overfire Air - Inner
Overfire Air - Outer
%
%
%
%
%
%
%
%
%
lb/ft3
Btu/lb
lb/hr
%
psia
%
Size (micron)
Distribution
9.99%
10.7
2.51%
33.7
4.50%
69.3
63.75%
118.2
11.12%
169.2
1.25%
Total
6.88%
100.00% < 75 micron
< 150 micron
44.35%
80.0
11,666
2,306,255
11.0
%
%
%
%
%
%
Mass Percent
29.5
38.2 Coal Flow
25.9 Moisture in Coal
5.5
1.0
100.0
%
%
375,400
41,744
85
99
Heat Input
4379.42 MM Btu/hr
Divwall
Inlet
Outlet
Temp (F) Temp (K)
837
721
970
794
WW
Inlet
Outlet
Temp (F)
599
830
Density
Inner Diam. Outer Diam. Area per Port No. of Ports Axial Velocity
Temp (K)
588
716
20.0%
Flow Rate
lb/hr
%
600,000
25.6
257,350
11.0
1,029,399
43.8
0
0.0
217,639
9.3
243,612
10.4
2,347,999
100.0
Temperature
F
236
625
625
625
625
625
lb/ft3
0.076
0.046
0.046
0.046
0.046
0.046
in
9.500
16.000
21.750
0.000
0.000
14.000
80
in
15.000
20.750
29.000
8.500
13.500
20.000
ft2
0.735
0.952
2.007
0.394
0.994
1.113
24
24
24
24
8
8
ft/sec
125.1
68.6
130.1
0.0
166.6
166.6
Tan./Axial Rad./Axial Coal Flow
Velocity
Velocity
lb/hr
0.50
0.00
0.41
0.00
0.00
0.00
0.00 333,656
0.21
0.41
0.00
0.00
0.21
Figure 4.3.6 – Air-Fired Boiler Design
81
Figure 4.3.7 – Summary of FW-FIRE Furnace Modeling Results
Air-Fired
Burnout
LOI
Total Furnace Absorption
Division Wall Absorption
FEGT
NOx
%
%
M Btu/hr
M Btu/hr
F
ppmvw
lb/MMBtu
FW-FIRE
99.6
2.76
1770
445
2107
276
0.38
ASPEN
99.0
7.32
1751
451
2185
%
%
M Btu/hr
M Btu/hr
F
ppmvw
lb/MMBtu
FW-FIRE
100.0
0.00
2096
548
2266
261
0.18
ASPEN
99.9
0.78
2089
485
2450
%
%
M Btu/hr
M Btu/hr
M Btu/hr
F
ppmvw
lb/MMBtu
FW-FIRE
100.0
0.00
2029
458
935
1822
273
0.20
ASPEN
99.9
0.78
1961
502
970
2185
Oxygen-Fired (Cryogenic ASU)
Burnout
LOI
Total Furnace Absorption
Division Wall Absorption
FEGT
NOx
Oxygen-Fired (OITM)
Burnout
LOI
Total Furnace Absorption
Radiant SH Absorption
Air Heater Absorption
FEGT
NOx
82
Figure 4.3.8 – Gas Velocity for Air-Fired Case
83
Figure 4.3.9 – Gas Temperature for Air-Fired Case
84
Figure 4.3.10 – O2 Mole Fraction for Air-Fired Case
85
Figure 4.3.11 – Wall Heat Flux for Air-Fired Case
86
Figure 4.3.12 – Wall Temperature for Air-Fired Case
87
Figure 4.3.13 – Wall CO for Air-Fired Case
88
Figure 4.3.14 – Char Mass Fraction (72 microns) for Air-Fired Case
89
Figure 4.3.15 – Char Mass Fraction (176 microns) for Air-Fired Case
90
Figure 4.3.16 – Air-Fired and Oxygen-Fired Boiler Outlines
(Black = Air-Fired, Red = Oxygen Fired)
91
Figure 4.3.17 – Oxygen-Fired Boiler Design (Cryogenic ASU)
92
Figure 4.3.18 – Oxygen-Fired Boiler Design (OITM)
93
Figure 4.3.19 – Gas Velocity for O2-Fired Case
94
Figure 4.3.20 – Gas Temperature for O2-Fired Case
95
Figure 4.3.21 – O2 Mole Fraction for O2-Fired Case
96
Figure 4.3.22 – Wall Heat Flux for O2-Fired Case
97
Figure 4.3.23 – Wall Temperature for O2-Fired Case
98
Figure 4.3.24 – Wall CO for O2-Fired Case
99
Figure 4.3.25 – Char Mass Fraction (69 micron) for O2-Fired Case
100
Figure 4.3.26 – Char Mass Fraction (169 micron) for O2-Fired Case
101
Figure 4.3.27 - Flue gas recycle flow vs. part load operation
Part Load Performance
1400
100
90
1200
1000
70
60
800
50
600
40
30
400
20
200
10
0
0
20
40
60
80
Load, %
102
100
0
120
gas recycling, %
Gas recycling, klb/hr
80
Figure 4.3.28 – Summary of O2-Fired Part Load Results, Cryogenic ASU
100% Load 72% Load 50% Load 25% Load
Burnout
%
100.00
100.00
100.00
99.81
LOI
%
0.00
0.00
0.00
1.48
Total Furnace Absorption MM Btu/hr
2096
1572
1073
747
Division Wall Absorption MM Btu/hr
548
380
248
193
FEGT
F
2266
2069
1912
1658
103
Figure 4.3.29 – Gas Temperature for O2-Fired Part Load, Cryogenic ASU
72 %
50 %
25 %
104
Figure 4.3.30 – Wall Heat Flux for O2-Fired Part Load, Cryogenic ASU
72 %
50 %
25 %
105
Figure 4.3.31 – Average and Peak Heat Flux in Waterwalls
Average Flux Vs. Height
160000
Heat Flux (Btu/hr-ft2)
140000
120000
100000
80000
60000
40000
20000
0
0
20
40
60
80
100
120
140
160
Height (ft)
no ofa, 100%
ofa, 100%
ofa, 25%
ofa, 72%
ofa, 50%
OTM, 100%
air-fired
Peak Flux Vs. Height
200000
Heat Flux (Btu/hr-ft2)
180000
160000
140000
120000
100000
80000
60000
40000
20000
0
0
20
40
60
80
100
120
140
160
Height (ft)
no ofa, 100%
ofa, 100%
ofa, 25%
ofa, 72%
106
ofa, 50%
OTM, 100%
air-fired
Figure 4.3.32 – Gas Velocity for O2-Fired with OITM
107
Figure 4.3.33 – Gas Temperature for O2-Fired Case With OITM
108
Figure 4.3.34 – O2 Mole Fraction for O2-Fired Case With OITM
109
Figure 4.3.35 – Wall Heat Flux for O2-Fired Case With OITM
110
Figure 4.3.36 – Wall Temperature for O2-Fired Case With OITM
Air heater
Furnace walls
111
Figure 4.3.37 – Wall CO for O2-Fired Case With OITM
112
Figure 4.3.38 – Char Mass Fraction (69 micron) for O2-Fired Case, OITM
113
Figure 4.3.39 – Char Mass Fraction (169 micron) for O2-Fired Case, OITM
114
4.3.3
Furnace Waterwall and Division Wall Design
The O2-PC supercritical boiler incorporates the state-of-the-art once-through utility
(OTU) BENSON Vertical technology, which uses low fluid mass flow rates in
combination with optimized rifled tubing and offers the following advantages:
“Natural Circulation” Flow Characteristic: By designing for low mass flow
rates that minimize frictional pressure losses, high heat flux tubes receive more
flow to minimize temperature unbalances (see Figure 4.3.40).
Improved Heat Transfer Coefficient: By using optimized rifled tubing,
departure from nucleate boiling (DNB) when operating near the critical pressure
can be suppressed, even with relatively low fluid mass flow rates, and dryout can
be prevented from occurring until steam qualities of greater than 90% are
achieved. This lowers the wall temperature permitting the use of thinner wall less
expensive materials (see Figure 4.3.41).
Simple Configuration: A standard, simple support system can be used to
support the vertical tube panels so that interconnecting piping and headers
between multiple passes are not required.
Low Pressure Losses: The low mass flow rates significantly reduce pressure
loss, which reduces auxiliary power and therefore increases cycle efficiency. The
boiler design pressure can also be lowered.
The furnace heat transfer tube design utilizing optimized rifled tubes is summarized in
Figure 4.3.42.
4.3.3.1
Tube Wall Temperature and Pressure Loss
Thermal/hydraulic modeling of the waterwalls and division walls was performed using
the Siemens computer program, Stade2 [12]. Stade2 creates a one-dimensional model
of the tube panels to determine inside heat transfer, wall temperatures, and pressure
loss and to perform linear dynamic and static stability analyses.
Stade2 models of the waterwalls and division walls were created and the average heat
flux (Figure 4.3.31) and average mass flow were applied to establish the net pressure
difference between the inlet and outlet headers. Models of the peak heat flux tubes
were then created and Stade2 was used to determine the increased mass flow rate
corresponding to the average pressure drop (due to the natural circulation characteristic
of the optimized rifled tube geometry). For example, for the cryogenic O2-PC design, the
average heat flux and average mass flux (832 kg/m2-sec) produced a pressure loss of
33.7 psi. To match this pressure loss for the peak heat flux tube requires a mass flux of
966 kg/m2-sec (0.71 MM lb/ft2-hr). Due to the increased mass flow rate the maximum
tube wall temperature of the peak heat flux tube is reduced. Figure 4.3.43 presents the
115
outside wall temperature of the peak heat flux waterwalls and division walls for the airfired and O2-fired designs.
Based on the maximum outside wall temperatures of Figure 4.3.43 the minimum tube
wall thickness is computed using stress allowables from the ASME Boiler and Pressure
Vessel Code as follows (design pressure = 5000 psi):
Air-Fired:
O2-Fired Cryo.:
O2-Fired OITM:
4.3.3.2
Material = SA-213-T2, min. wall thickness = 0.22”
Material = SA-213-T92, min. wall thickness = 0.20”
Material = SA-213-T92, min. wall thickness = 0.23”
Tube Panel Stability
The waterwall and division wall designs were determined by Stade2 to be statically and
linearly dynamically stable at full and part loads (although the 25% load case appears to
be near the dynamically unstable region). A more thorough treatment of dynamic
stability was performed using the Siemens program, Dynastab [13]. Dynastab performs
calculations for a single tube within a greater number of parallel tubes. The calculation
starts from steady state and compares the flow behavior of a single tube, with slightly
changed conditions (e.g. heat input, tube geometry), with the mean tubes of the heating
surface. The mass flow of the single tube is stimulated by a distinct disturbance (e. g.
global or local heating factor, inlet enthalpy), while pressure drop is kept constant. If this
disturbance causes continuous oscillations, the tube is dynamically unstable; whereas if
the disturbance results in a (new) steady state condition, the tube is dynamically stable.
A 10% step increase in heat flux was applied to the single tube model. Figure 4.3.44
presents the transient results of inlet mass flow versus time (although a single tube was
modeled the flow per entire furnace is presented). The 100% and 72% loads are stable,
the 50% load is marginally stable, and the 25% load is unstable. To ensure stable
operation at low loads, a pressure equalization header is added to the design at an
elevation of 80’ and the resultant dynamic stability response is shown in Figure 4.3.45.
With the pressure equalization header all loads are dynamically stable.
116
Figure 4.3.40 – Advantage of Low Mass Flux Design
117
Figure 4.3.41 – Tube Wall Temperature with Smooth, Standard Rifled, and
Optimized Rifled Tubes
118
Figure 4.3.42 – Rifled Tube Design
Material
OD
twall
ID
Pitch
Ligament
Ligament Width
Number of waterwall tubes
Number of radiant superheater tubes
Number of ribs
Rib Height
in
in
in
in
in
in
in
Air-Fired
O2 PC SC
Cryogenic
O2 PC SC
OITM
SA-213-T2
1.40
0.22
0.96
2.00
0.60
0.25
1218
618
6
0.048
SA-213-T92
1.40
0.20
1.00
2.00
0.60
0.25
888
720
6
0.050
SA-213-T92
1.40
0.23
0.94
2.00
0.60
0.25
888
666
6
0.048
119
Figure 4.3.43 – Outside Tube Wall Temperature with Peak Heat Flux
Outside Tube Wall Temperature Vs. Height
1150
1100
1050
Temperature (F)
1000
950
900
850
800
750
700
650
600
0
20
40
60
80
100
120
140
160
Height (ft)
100% no OFA
Div. Wall 100%
100% load
Div. Wall 72%
72% load
Div. Wall air
120
50% load
100% oitm
100% air
Figure 4.3.44 – Tube Inlet Mass Flow With a 10% Heat Flux Step Increase (No
Pressure Equalization Header)
Flow Vs. Time: No Pressure Equilization Header
6
5
Flow Rate (MM lb/hr)
4
3
2
1
0
0
20
40
60
80
100
120
140
-1
-2
Time (sec)
100%
72%
50%
121
25%
160
180
200
Figure 4.3.45 – Tube Inlet Mass Flow With a 10% Heat Flux Step Increase (With
Pressure Equalization Header at 80’)
Flow Vs. Time: With Pressure Equilization Header at 80'
3.5
3
Flow Rate (MM lb/hr)
2.5
2
1.5
1
0.5
0
0
20
40
60
80
100
120
140
Time (sec)
100%
72%
122
50%
25%
160
180
200
4.3.4 Furnace Waterwall Corrosion
Waterwall corrosion is caused by sulfidation from a sub-stoichiometric gas containing
H2S and from deposits containing carbon and iron sulfide. Corrosion is especially
significant at the high waterwall surface temperatures of a supercritical and ultrasupercritical boiler. In the O2-PC without FGD, recycling of the flue gas will increase the
H2S concentration of the furnace by a factor of 1/(1 – recycle fraction). For example for
a recycle flow rate of 55%, the H2S concentration is increased by a factor of 2.2. An
empirical EPRI formula was developed to predict the magnitude of waterwall corrosion
due to the presence of H2S as follows:
1
⎛ 15818 ⎞
0.574
CR = 3.2 x10 5 exp⎜ −
⎟[H s S ]
(Cr % + 10.5)1.234
⎝ 1.987T ⎠
(Cr wt% < 10)
1
⎛ 19230 ⎞
0.29
CR = 1.04 x10 7 exp⎜ −
⎟[H s S ]
(Cr % − 1.40)1.37
⎝ 1.987T ⎠
(Cr wt% > 16)
where,
CR = corrosion rate (mil/year)
T = metal temperature, K
H2S = H2S flue gas concentration, ppm
Cr% = Cr concentration of the metal
Figure 4.3.46 shows the predicted corrosion rate for the air-fired and O2-fired cases. For
the air-fired case the maximum corrosion on the waterwalls is 12 mil/yr and on the
division walls is 18 mil/yr (maximum H2S concentration is 1600 ppm). For the oxygenfired case maximum corrosion on the waterwalls is 35 mil/yr and on the division walls is
45 mil/yr (maximum H2S concentration is 5600 ppm). Note that for the O2-fired case the
corrosion is increased by the higher water wall temperature, higher concentration of
H2S, but is reduced by the higher chromium content of T92 vs. T2. To reduce the
corrosion in the O2-PC, weld overlays with high Nickel and Chromium contents are
proposed. This will be much more cost-effective than adding an FGD system, which will
be substantially more costly and reduce the corrosion by only a factor of 1.6 (by
reducing the H2S by a factor of 2.2). Applying alloy 622 (20% Cr) as a weld overlay
significantly reduces the corrosion rate as shown in Figure 4.3.46 to less than 10 mil/yr.
Note that other alternatives to weld overlays include upgrading the base tube metal and
thermal spray coatings.
123
Figure 4.3.46 – Predicted Wall Corrosion (mil/yr) in Air-Fired and O2-Fired
Furnaces
Air-Fired
O2-Fired
O2-Fired
(Without Tube Overlay)
(With Tube Overlay)
124
4.3.5
Heat Recovery Area Design and Analysis
4.3.5.1
HEATEX Program Description
HEATEX [10] is a Foster Wheeler general-purpose program for thermal/hydraulic
analysis of tube banks. The program performs heat transfer calculations on a local basis
by dividing the tube bundle into a number of small heat transfer elements.
4.3.5.2
Air-Fired Reference Case
HEATEX was used to determine the heat recovery area (HRA) design of the convective
tube banks between the furnace exit and the SCR/air heater. These tube banks include
the finishing superheater, finishing reheater, primary superheater, primary reheater,
upper economizer, and lower economizer. Flue gas exits the furnace at 2185ºF and
flows over the finishing superheater and finishing reheater tube bundles where it heats
the main steam and reheat steam to 1083ºF and 1113ºF, respectively. The gas flow is
then split into two parallel flows: one passing over the primary reheater and the other
passing over the primary superheater and upper economizer. The gas split is controlled
by dampers to achieve the proper reheater outlet temperature. Attemperating spray is
used to control superheat temperature. The flue gas is combined downstream of the
dampers and flows over the lower economizer, which receives water from the last
feedwater heater stage. Flue gas exits the lower economizer at 720ºF and is sent to the
SCR and then to the air heater. Figure 4.1.1 presents the heat transfer requirements of
the HRA banks. Figure 4.3.47 presents the corresponding design of the HRA banks.
Total surface area of all convective banks is 335,025 ft2. The performance of HRA tube
banks is shown in Figure 4.3.48. The total heat transferred to the water/steam is 1431
MM Btu/hr as 3.59 MM lb/hr of flue gas is cooled from 2185ºF to 720ºF.
4.3.5.3
Oxygen-Fired Case, Cryogenic ASU
Due to the 40% lower flue gas flow rate of the cryogenic ASU oxygen-fired case versus
the air-fired case, the cross sectional area of the HRA is reduced to maintain the same
gas side velocity (and pressure drop). HEATEX was used to determine the heat
recovery area design of the convective tube banks between the furnace exit and the gas
recuperator. These tube banks include the finishing superheater, finishing reheater,
primary reheater, and lower economizer. Flue gas exits the furnace at 2450ºF and flows
over the finishing superheater and finishing reheater tube bundles where it heats the
main steam and reheat steam to 1083ºF and 1113ºF, respectively. Because of the
reduced flue gas flow and lower HRA duty, the HRA is designed in series instead of
parallel to produce a more compact design After exiting the finishing reheater the flue
gas flows over the primary reheater and then the lower economizer, which receives
water from the last feedwater heater stage. Flue gas exits the lower economizer at
695ºF and is sent to the gas recuperator.
125
Figure 4.1.9 presents the heat transfer requirements of the HRA banks. Figure 4.3.47
presents the corresponding design of the HRA banks. Total surface area of all
convective banks is 218,693 ft2. The performance of HRA tube banks is shown in Figure
4.3.48. The total heat transferred to the water/steam is 1151 MM Btu/hr as 2.12 MM
lb/hr of flue gas is cooled from 2450ºF to 695ºF. The total heat transfer surface required
in the oxygen-fired HRA is 35% less than the air-fired HRA due to the following main
reasons:
1. More heat is absorbed in the oxygen-fired furnace (2089 MM Btu/hr) than the
air-fired furnace (1751 MM Btu/hr) due to the higher adiabatic temperature
and the greater specific heat of the oxygen-fired furnace flue gas. This
requires less heat transfer duty in the HRA (as a consequence the upper
economizer is not needed)
2. A higher heat transfer coefficient can be achieved in the oxygen-fired HRA
than the air-fired HRA for the same flue gas pressure loss due to greater
molecular weight (38 mol/lb-mol vs. 29 mol/lb-mol) of the oxygen-fired flue
gas.
The HRA tube materials and wall thicknesses are nearly the same for the air-fired and
oxygen-fired design (except for the finishing superheater where the 0.42” wall thickness
of the air-fired case is increased to 0.46” for the oxygen-fired case) since the flue gas
and water/steam temperature profiles encountered by the heat transfer banks are very
similar.
4.3.5.4
Oxygen-Fired Case, OITM
Due to the 25% lower flue gas flow rate of the OITM oxygen-fired case versus the airfired case, the cross sectional area of the HRA is reduced to maintain the same gas
side velocity (and pressure drop). HEATEX was used to determine the heat recovery
area design of the convective tube banks between the furnace exit and the gas
recuperator. These tube banks include the finishing superheater, finishing reheater,
primary superheater, primary reheater, upper economizer, and lower economizer. Flue
gas exits the furnace at 2185ºF and flows over the finishing superheater and finishing
reheater tube bundles where it heats the main steam and reheat steam to 1083ºF and
1113ºF, respectively. The flue gas is then used to heat the air used for the OITM
process. The gas flow is then split into two parallel flows: one passing over the primary
reheater and the other passing over the primary superheater and upper economizer.
The gas split is controlled by dampers to achieve the proper reheater outlet
temperature. Attemperating spray is used to control superheat temperature. The flue
gas is combined downstream of the dampers and flows over the lower economizer,
which receives water from the last feedwater heater stage. Flue gas exits the lower
economizer at 695ºF and is sent to the gas recuperator.
Figure 4.2.10 presents the heat transfer requirements of the HRA banks. Figure 4.3.47
presents the corresponding design of the HRA banks. Total surface area of all
126
convective banks is 291,150 ft2 (242,510 ft2 without the air heater). The total heat
transfer surface required in the oxygen-fired OITM HRA is 13% less than the air-fired
HRA (28% less not including the air heater). The performance of HRA tube banks is
shown in Figure 4.3.48. The total heat transferred to the water/steam is 1181 MM Btu/hr
and to the air is 990 MMBtu/hr as 2.68 MM lb/hr of flue gas is cooled to 695ºF. The total
furnace + HRA heat transfer surface area for the OITM O2-PC is 309,799 ft2 as
compared to 382,574 ft2 for the air-fired PC and 252,002 ft2 for the cryogenic ASU O2PC.
The HRA tube materials and wall thicknesses are nearly the same for the air-fired and
OITM oxygen-fired design (except for the finishing superheater where the 0.42” wall
thickness of the air-fired case is increased to 0.46” for the oxygen-fired case) since the
flue gas and water/steam temperature profiles encountered by the heat transfer banks
are very similar.
The air heater is constructed Incoloy MA956 which is an iron-chromium-aluminum alloy.
Incoloy MA956 has been used in advanced gas turbine engines and is resistant to
creep, oxidation, and corrosion at temperatures up to 2200ºF. The furnace air heater is
a three pass tubular design situated above the furnace nose to reduce radiation and
maximum metal temperature. Air is heated from 745 to 1650ºF as the flue gas is cooled
from 2950 to 2185ºF. Maximum metal temperature is approximately 1900ºF.
127
Figure 4.3.47 – HRA Tube Bank Design
Air-Fired
O2 PC SC
Cryogenic
O2 PC SC
OITM
Air Heater
Length
No. of Tubes Deep
No. of Tubes Wide
Total Number of Tubes
Tube Outside Diameter
Tube Thickness
Tube Material
Design Pressure
Design Temperature
Stress Allowable
Min. Wall
Total Surface Area
Finishing Superheater
Length
No. of Tubes Deep
No. of Tubes Wide
Total Number of Tubes
Tube Outside Diameter
Tube Thickness
Tube Material
Design Pressure
Design Temperature
Stress Allowable
Min. Wall
Total Surface Area
ft
28.0
42
79
3,318
2.000
0.06
MA956
250
2000
8200
0.040
48,640
in
in
psi
F
psi
in
ft2
ft
in
in
psi
F
psi
in
ft2
36.0
44
30
1,320
2.000
0.42
SA-213-T92
5000
1150
10200
0.414
26,539
24.0
56
32
1,792
2.000
0.47
SA-213-T92
5000
1180
8379
0.470
24,021
24.0
40
32
1,280
2.000
0.47
SA-213-T92
5000
1180
8379
0.470
17,152
30.5
40
63
2,520
2.250
0.165
SA-213-T92
1000
1200
7990
0.173
45,270
20.0
72
51
3,672
2.250
0.165
SA-213-T92
1000
1200
7330
0.173
43,260
20.0
60
51
3,060
2.250
0.165
SA-213-T92
1000
1200
6730
0.173
36,050
Vertical Reheater
Length
No. of Tubes Deep
No. of Tubes Wide
Total Number of Tubes
Tube Outside Diameter
Tube Thickness
Tube Material
Design Pressure
Design Temperature
Stress Allowable
Min. Wall
Total Surface Area
ft
in
in
psi
F
psi
in
ft2
Primary Superheater
Length
No. of Tubes Deep
No. of Tubes Wide
Total Number of Tubes
Tube Outside Diameter
Tube Thickness
Tube Material
Design Pressure
Design Temperature
Stress Allowable
Min. Wall
Total Surface Area
ft
0
0
0
in
in
17.0
96
63
6,048
2.250
0.42
SA-213-T2
5000
925
11550
0.412
63,252
psi
F
psi
in
ft2
128
9.5
64
90
5,760
2.250
0.42
SA-213-T2
5000
925
11550
0.412
33,662
Figure 4.49 – HRA Tube Bank Design (Continued)
Air-Fired
O2 PC SC
Cryogenic
O2 PC SC
OITM
17.0
70
125
8,750
2.250
0.165
SA-213-T2
1000
950
9200
0.127
93,461
24.0
75
90
6,750
2.250
0.165
SA-213-T2
1000
950
9200
0.127
101,792
21.0
75
90
6,750
2.250
0.165
SA-213-T2
1000
950
9200
0.127
89,067
17.0
30
126
3,780
2.250
0.34
SA-210-A1
5000
700
15600
0.322
40,373
0.0
0
0
0
0.000
0
0
0
0
0
0
0
9.5
36
90
3,240
2.250
0.32
SA-210-A1
5000
650
17100
0.298
19,339
24.0
39
90
3,510
2.250
0.32
SA-210-A1
5000
650
17100
0.298
49,620
27.0
33
90
2,970
2.250
0.32
SA-210-A1
5000
650
17100
0.298
47,240
218,693
252,002
291,150
309,799
Horizontal Reheater
Length
No. of Tubes Deep
No. of Tubes Wide
Total Number of Tubes
Tube Outside Diameter
Tube Thickness
Tube Material
Design Pressure
Design Temperature
Stress Allowable
Min. Wall
Total Surface Area
ft
in
in
psi
F
psi
in
ft2
Upper Economizer
Length
No. of Tubes Deep
No. of Tubes Wide
Total Number of Tubes
Tube Outside Diameter
Tube Thickness
Tube Material
Design Pressure
Design Temperature
Stress Allowable
Min. Wall
Total Surface Area
ft
in
in
psi
F
psi
in
ft2
Lower Economizer
Length
No. of Tubes Deep
No. of Tubes Wide
Total Number of Tubes
Tube Outside Diameter
Tube Thickness
Tube Material
Design Pressure
Design Temperature
Stress Allowable
Min. Wall
Total Surface Area
ft
psi
F
psi
in
ft2
27.0
33
126
4,158
2.250
0.32
SA-210-A1
5000
650
17100
0.298
66,130
Total HRA Surface Area
Total Furnace + HRA Surface Area
ft2
ft2
335,025
382,574
in
in
129
Figure 4.3.48 – HRA Tube Bank Performance
Bank
O2 PC OITM Air Heater
Surface Heat Trans. Mean Temp.
Heat
Gas
Area
Coeff.
Diff.
Transfer
Press. Drop
(ft2)
(Btu/hr-ft2-F)
(F)
(MM Btu/hr) (in H2O)
48,640
17.1
1189
990
0.38
Air PC
Finishing Superheater
O2 PC
Finishing Superheater
O2 PC OITM Finishing Superheater
26,539
24,021
17,157
11.4
14.7
16.0
1038
1173
1026
315
414
281
0.12
0.19
0.20
Air PC
Primary Superheater
O2 PC
Primary Superheater
O2 PC OITM Primary Superheater
63,252
0
33,662
8.6
356
194
11.2
268
101
0.29
0.00
0.44
Air PC
Finishing Reheater
O2 PC
Finishing Reheater
O2 PC OITM Finishing Reheater
45,270
43,260
36,050
10.2
12.2
14.1
722
629
658
333
331
334
0.18
0.35
0.47
Air PC
Primary Reheater
O2 PC
Primary Reheater
O2 PC OITM Primary Reheater
93,461
101,792
89,067
9.1
10.8
10.9
402
290
353
341
320
342
0.47
0.58
0.55
Air PC
Upper Economizer
O2 PC
Upper Economizer
O2 PC OITM Upper Economizer
40,373
0
19,339
7.5
367
111
15.6
367
111
0.20
0.00
0.19
Air PC
Lower Economizer
O2 PC
Lower Economizer
O2 PC OITM Lower Economizer
66,130
49,620
47,240
9.9
9.8
10.4
196
177
165
128
86
81
0.35
0.27
0.28
1431
1151
2171
1.61
1.39
2.51
Air PC
Total
O2 PC
Total
O2 PC OITM Total
335,025
218,693
291,155
130
4.4
Economic Analysis
Economic analysis was performed for three cases: air-fired (reference), O2-fired with
cryogenic ASU, and O2-fired with OITM. Economic analysis was not performed for the
O2-fired PC with CAR since detailed cost information on the CAR process was not
available in the literature nor was it provided by the vendor (BOC).
4.4.1 Main Assumptions
The economic analysis was performed based on the DOE/NETL guidelines [14] using
the EPRI Technical Assessment Guide (TAG) methodology. Plant capital costs were
compiled under the Code of Accounts developed by EPRI.
The estimate basis and major assumptions are listed below:
•
Total plant costs were estimated in January 2006 dollars.
•
Plant book life was assumed to be 20 years.
•
The net power output of the reference air-fired plant is 430.2 MWe versus 347.0
MWe for the ASU oxygen-based plant and 463.3 MWe for the OITM oxygenbased plant.
•
The plants operate with a capacity factor of 85 per cent (Plant operates at 100
per cent load 85 per cent of the time).
•
Cost of electricity (COE) was determined on a levelized constant dollar basis.
•
Average annual ambient air conditions for material balances, thermal efficiencies
and other performance related parameters are at a dry bulb temperature of 60ºF
and an air pressure of 14.7 psia.
•
The coal is 2.5 per cent sulfur Illinois #6 coal (see Table 4.4.1 for analysis).
•
Design CO2 effluent purity is presented in Table 4.4.2
•
Terms used are consistent with the EPRI TAG.
Economic study assumptions are detailed in Table 4.4.3.
131
Table 4.4.1 - Coal Properties
Illinois No. 6 Coal
C
H
O
N
Cl
S
Ash
H2O
Total
%
%
%
%
%
%
%
%
%
LHV Btu/lb
HHV Btu/lb
63.75%
4.50%
6.88%
1.25%
0.29%
2.51%
9.70%
11.12%
100.00%
11,283
11,631
Table 4.4.2 - CO2 Effluent Purity Design Conditions
Constituent
N2
H2O
O2
Units
vppm
vppm
vppm
132
Value
< 300
< 20
< 50
Table 4.4.3 – Economic Study Assumptions
GENERAL DATA/CHARACTERISTICS
Levelized Capacity Factor / Preproduction (equivalent months):
85%
Capital Cost Year Dollars (Reference Year Dollars):
2006 (January)
Design/ Construction Period:
4 years
Plant Start-up Date (1st year Dollars):
2010 (January)
Land Area/Unit Cost:
100 acres
$1,600 / Acre
FINANCIAL CRITERIA
Project Book Life:
Book Salvage Value:
Project Tax Life:
Tax Depreciation Method:
Inflation Rate
Property Tax Rate:
Insurance Tax Rate:
Federal Income Tax Rate:
State Income Tax Rate:
Investment Tax Credit/% Eligible
20 years
0%
20 years
Accel. Based on ACRS Class
3.0 %
1.0 %
1.0 %
35.0 %
4.0 %
0%
Economic Basis:
Over Book Constant Dollars
Capital Structure
Common Equity
Preferred Stock
Debt
Weighted Cost of Capital: (after tax)
% of Total
20
0
80
Cost (%)
12
6.5
5.57%
Coal Price Escalation Rate
3.0%
Total Capital Requirement
Initial Chemical Inventory
(same as general escalation)
30 days
Startup Costs
2% TPI
30 days of fuel and chemicals
labor and miscellaneous items
Spare Parts
0.5% TPC
Working Capital
30 days fuel and consumables
30 days direct expenses
Consumable Costs
Coal, $/MMBtu
Limestone, $/ton
Water, $/kgal
Water Treatment Chemicals, $/kgal
Ash/Slag Disposal, $/ton
$1.34
$15.00
$1.00
$0.50
$10.00
Plant Labor
Operating labor
Labor Rate
42.25 $/hr
(includes labor burden)
personnel
14 per shift
Supervisory/clerical 30% of operating + maintenance labor cost
Maintenance Costs
Labor 0.88% TPC
Materials 1.32% TPC
133
4.4.2 Plant Cost Basis
For each of the plants, heat and material balances (Aspen simulations) were prepared
that identified the flow rates and operating conditions of all major flow streams (Sect. 4.1
and Sect. 4.2). These balances also identified each boiler’s operating requirements and
enabled design calculations to be performed that established the overall boiler
dimensions, tube surface areas, materials of construction, weights, and auxiliary
equipment requirements. With this information defined, the cost of each boiler, which
together with its auxiliary equipment constitutes the plant “boiler island”, was determined
from Foster Wheeler’s cost estimating database.
4.4.2.1
Air-Fired Reference Plant Cost
In [15] Parsons presents a conceptual design of a supercritical pressure PC plant with a
net power output of 550.2 MWe. Similar to the air-fired reference plant of this study, the
Parsons plant burned 2.5 per cent sulfur Illinois No. 6 coal with air and, to control
emissions, the boiler was provided with low NOx burners, SCR, wet flue gas
desulfurization, and a baghouse filter. A detailed EPRI TAG cost estimate of the plant
totaled $745.7 million or $1355/kW in year 2006 dollars and was broken down into 14
accounts, one of which (Account 4 entitled, “PC Boiler and Accessories”) included the
cost of the PC boiler and its auxiliaries. Since the primary difference between the
Parsons and the FW air-fired reference plant is size, the Parsons balance of plant costs
(excluding the boiler island) were scaled down on an account-by-account basis to obtain
the reference plant balance of plant costs. As recommended in [16] a scaling exponent
of 0.65 applied to flow rate/output was used and a small adjustment to the PC plant’s
feedwater and steam turbine accounts was made for the slightly higher operating
pressure. Boiler island costs were estimated directly by FW based on designs
generated in Task 3 (Sect. 4.3). To validate the scaling process, Parsons Account 4
boiler costs were scaled down and determined to be within 6 per cent of Foster
Wheeler’s directly estimated costs. The results of the scaling, together with Foster
Wheeler’s determined Account 4 (boiler island) costs, yielded a total plant cost of
$633.0 million or $1471/kW for the 430.2 MWe reference plant. The account-by-account
costs of the air-fired reference plant (case-1, Figure 4.1.1) are presented in Table 4.4.4.
The total plant cost (TPC), also referred to as the plant capital cost is comprised of the
following elements:
1. Bare erected plant cost (includes equipment supply and erection)
2. Architect engineering, construction management, and fee
3. Project and process contingencies
The reference plant estimate uses the same erection factors, fees, and contingencies
as the Parsons plant estimate (i.e. boiler erection at 80 per cent of equipment supply
costs, architect engineering/construction management/home office/fees at 10 per cent
134
of bare erected costs, and contingency, totaling 10 per cent, applied to the sum of items
1 and 2).
4.4.2.2
Oxygen Based PC Plant Costs
In [7] Parsons presents conceptual designs of two plants that burned 2.5 per cent sulfur
Illinois No. 6 coal with oxygen to facilitate CO2 capture/removal for pipeline transport to
a sequestering site. Both plants used flue gas recirculation to control their boiler
combustion temperatures and their CO2 rich exhaust gases were dried and compressed
for pipeline transport.
The first plant had a net power output of 132.2 MWe; its oxygen was supplied by a
conventional, cryogenic ASU, and the gas flow to the CO2 processing unit totaled 422.3
Klb/hr. The oxygen was 99 per cent pure and it was delivered to an “air heater” at the
boiler where 660ºF flue gas heated the oxygen to 610ºF for delivery to the boiler at a
rate of 327.1 Klb/hr. In the FW study’s ASU based plant, the oxygen from the ASU is
also heated by an “air heater” at the boiler but to 625ºF with 695ºF flue gas; aside from
some slight temperature differences, the plant arrangements are similar and their ASUs
differ primarily in size/through put.
The second Parsons plant had a net output of 197.4 MWe; its oxygen was supplied at a
rate of 398.8 Klb/hr by an ion transport membrane, and the gas flow to CO2 processing
was 515.8 Klb/hr. The OITM operated with 200 psia air that was heated to 1652ºF via
heat transfer surface placed in the boiler; the hot air was then delivered to the OITM for
separation of the oxygen and nitrogen. The oxygen, with a purity of 100 per cent, was
then delivered to the boiler, whereas, the nitrogen was passed through a hot gas
expander followed by a heat recovery steam generator for power recovery. Excepting
for differences in flow rates, the operating conditions of the Parsons OITM plant are
essentially identical to the OITM based plant of the FW study.
Praxair Inc, a developer of oxygen transport membranes, participated in the detailed
Alstom/Parsons study [7] and, per the Acknowledgement section of their study report,
provided detailed design, performance, and cost information on the OITM system. Aside
from a difference in flow rate, the FW OITM operates at essentially the same pressure
and temperature as that used in the Alstom/Parsons study and, hence, the FW system
cost estimate was obtained by scaling their system capital costs. Consequently,
individual OITM component designs were not developed in the FW study. It is assumed
that Parsons/Praxair selected the operating conditions of the OITM to minimize the
overall system cost. Such a sensitivity/optimization cost evaluation of the OITM plant
design is beyond the scope of the FW study. Note that there are several variables,
which have direct influence on the OITM plant cost and performance (see Sect. 4.2.2.3
for more details). These variables include
O2 Recovery Percentage: CO2 removal power penalty is minimum at an O2
recovery of 85% (The Alstom/Parsons study uses a O2 recovery of 85%).
135
Air Pressure: Increasing air pressure reduces OITM size, but also reduces system
efficiency. Optimum pressure depends on the relative cost of the OITM.
OITM Temperature: Increasing OITM operating temperature will increase system
efficiency, but will increase boiler air heater cost and presumably OITM cost.
A detailed 14 account EPRI TAG cost estimate in year 2003 dollars was prepared by
Parsons for each of the two plants; in these estimates the oxygen (ASU or OITM) and
CO2 processing system costs appear, respectively, as separate sub-accounts under the
Boiler and Accessories and Flue Gas Clean Up Accounts. Since the arrangement and
scope of supply of these systems is essentially identical to that of the FW study’s O2fired plants, the Parsons costs were used to estimate the costs of Foster Wheeler’s two
plants. The Parsons ASU system costs were scaled up based on oxygen flow rate
raised to the 0.65 exponent and escalated to year 2006 dollars; comparison of the
scaled up costs with a vendor supplied budget price yielded good agreement further
validating the scale up exponent. Comparison of the gas processing system costs given
for Parsons’ two plants, however, revealed a greater sensitivity to flow rate and resulted
in a scaling exponent of 0.87; this exponent was then used to calculate the CO2 gas
processing system costs of Foster Wheeler’s plants.
The Table 4.4.4 air-fired reference plant cost estimate served as a starting point for the
determination of O2 based plant costs. The cost of the air-fired PC boiler package was
removed from the Parsons cost estimate and the new value, determined by Foster
Wheeler for the particular plant configuration under study, was inserted. Then each
balance of plant account and or component was individually scaled based on flow
rate/output and adjusted, where necessary, to reflect design differences. Some
examples of the changes that were made are removal of SCR systems, limestone
systems, flue gas desulfurization systems, etc. and elimination of the stack from the
ASU based plant.
Table 4.4.5 and Table 4.4.6 present a detailed cost breakdown of the ASU based O2
plant (case-9B, Figure 4.1.9) and the OITM based O2-PC plant (case-18, Figure 4.2.10)
and Table 4.4.7 compares the total costs of all three plants. Although the oxygenbased plants use flue gas recirculation to control the boiler combustion temperature,
they operate with higher oxygen concentrations than the air-fired reference plant boiler.
As a result they produce a higher combustion temperature and a lower flue gas flow
rate, which reduce the size of the boiler and its downstream flue gas components. With
the SCR eliminated and the flue gas flow rate reduced, the PC boiler cost of the ASU
based plant is about $32 million less than that of the air-fired case. Despite additional
savings associated with elimination of some systems/components (e.g. SCR, limestone,
FGD, and stack plus a reduction in size of other components, i.e. baghouse filter,
ducting, foundations, etc.), the total cost of the cryogenic ASU based plant is
approximately $91 million higher than the air-fired plant because of the high cost of the
ASU ($115.0 million) and CO2 processing systems ($111.5 million).
136
In the OITM based plant, the PC boiler not only meets the steam generation and
heating needs of the steam cycle, but it also contains tubing that heats air to 1600ºF for
delivery to the OITM. Although the boiler flue gas flow rate of the OITM based plant is
about 25 per cent less than the air-fired plant, the cost savings it provides is more than
negated by the high cost of the air heater tubing; as a result, the OITM based PC boiler
costs $18 million more than that of the air-fired plant. The compressor required to
pressurize OITM air to 200 psia, the OITM, the OITM associated heat exchangers and
piping, the OITM power recovery system (hot gas expander and HRSG), and a larger
CO2 gas processing system add additional costs to the plant; as a result, the OITM
plant costs approximately 50 per cent more than the air-fired reference plant ($953.0
million versus $633.0 million) and approximately 32 per cent more than the cryogenic
ASU based plant ($953.0 million versus $723.3 million). Since the OITM based plant
operates with a higher electrical output than the cryogenic ASU based plant (463.3
MWe versus 347.0 MWe), its total plant costs on a dollar per kilowatt basis are slightly
lower ($2.057/kW versus $2,084/kW) but much higher than the air-fired plant
($2,057/kW versus $1,471/kW).
137
Table 4.4.4 - Cost of 430.2 MWe Air-Fired Supercritical PC Plant ($1000 Yr
2006)
Account #
1
2
Bare Erected
Costs
Account Title
Contingency
at 10%
Total Plant
Costs
Coal & Sorbent Handling
Coal
Limestone
12,985
3,114
1,299
311
1,428
343
15,712
3,768
Coal
Limestone
2,599
6,934
260
693
286
763
3,145
8,390
47,879
4,788
5,267
57,933
155,621
15,562
17,118
188,301
0
17,111
58,705
0
1,711
5,871
0
1,882
6,458
0
20,705
71,033
0
Coal and Sorbent Prep & Feed
3
Feedwater & Misc BOP Systems
4
Boiler and Accessories
Boiler with SCR, Air Heater, Fans, Ducts, etc
Oxygen Supply; None
5
Engr, C.M.
H.O. & Feee
at 10%
Flue Gas Clean Up
Baghouse & Accessories
FGD
CO2 Processing
6
Combustion Turbine & Accessories
7
HRSG, Ducting, & Stack
0
Duct Work
Stack
Foundations
9,094
9,774
1,456
909
977
146
1,000
1,075
160
11,004
11,827
1,762
8
Steam Turbine Generator
88,706
8,871
9,758
107,334
9
Cooling Water System
24,540
2,454
2,699
29,693
10
Slag/Ash Handling Systems
7,746
775
852
9,373
11
Accessory Electric Plant
27,500
2,750
3,025
33,275
12
Instrumentation & Control
12,237
1,224
1,346
14,807
13
Improvements to Site
7,773
777
855
9,405
14
Buildings & Structures
29,353
2,935
3,229
35,517
523,126
52,313
57,544
632,984
Totals
$/kW
138
1,471
Table 4.4.5 - Cost of 347.0 MWe ASU Based Supercritical PC Plant ($1000 Yr
2006)
Account #
Bare Erected
Costs
Account Title
1
Coal Handling
2
Coal Prep & Feed
3
Feedwater & Misc BOP Systems
4
Boiler and Accessories
Boiler with Air Heater, Fans, Ducts, etc
Oxygen Supply: ASU
5
Engr, C.M.
H.O. & Feee
at 10%
Contingency
at 10%
Total Plant
Costs
12,692
1,269
1,396
15,357
2,540
254
279
3,074
47,879
4,788
5,267
57,933
129,052
12,905
14,196
156,153
115,005 *
12,102
1,210
1,331
14,643
0
111,493 *
Flue Gas Clean Up
Baghouse & Accessories
FGD
CO2 Processing
6
Combustion Turbine & Accessories
0
7
HRSG, Ducting, & Stack
Duct Work and Foundations
Stack
HRSG
6,432
643
707
7,782
0
0
8
Steam Turbine Generator
88,706
8,871
9,758
107,334
9
Cooling Water System
25,966
2,597
2,856
31,419
10
Slag/Ash Handling Systems
7,281
728
801
8,810
11
Accessory Electric Plant
27,885
2,789
3,067
33,741
12
Instrumentation & Control
12,408
1,241
1,365
15,014
13
Improvements to Site
7,882
788
867
9,537
14
Buildings & Structures
29,764
2,976
3,274
36,014
41,059
45,165
723,310
Totals
$/kW
2,084
*Values Scaled from Parsons /Alstom Year 2003 Study and Escalated to 2006 at 5% per Year
139
Table 4.4.6 - Cost of 463.3 MWe OITM Based Supercritical PC Plant ($1000
Yr 2006)
Account #
Bare Erected
Costs
Account Title
1
Coal Handling
2
Coal Prep & Feed
3
Feedwater & Misc BOP Systems
4
Boiler and Accessories
Boiler, Air Heater, Fans, Ducts, etc
Oxygen Supply: OTM
5
Engr, C.M.
H.O. & Feee
at 10%
Contingency
at 10%
Total Plant
Costs
14,434
1,443
1,588
17,465
2,889
289
318
3,496
47,879
4,788
5,267
57,933
170,902
17,090
18,799
206,791
188,031 *
14,113
1,411
1,552
17,077
0
131,662 *
Flue Gas Clean Up
Baghouse & Accessories
FGD
CO2 Processing
6
Combustion Turbine & Accessories
7
HRSG, Ducting, & Stack
48,627 *
Duct Work
Stack
Foundations
HRSG
7,501
750
825
9,076
2,680 *
205 *
18,420 *
8
Steam Turbine Generator
88,706
8,871
9,758
107,334
9
Cooling Water System
29,489
2,949
3,244
35,682
10
Slag/Ash Handling Systems
8,344
834
918
10,097
11
Accessory Electric Plant
29,101
2,910
3,201
35,212
12
Instrumentation & Control
12,949
1,295
1,424
15,668
13
Improvements to Site
8,225
823
905
9,953
14
Buildings & Structures
31,061
3,106
3,417
37,584
46,559
51,215
952,993
Totals
$/kW
*Values Scaled from Parsons /Alstom Year 2003 Study and Escalated to 2006 at 5% per Year
140
2,057
Table 4.4.7 - Comparison of Total Plant Costs ($1000 Yr 2006)
Reference
Air-Fired Plant
Net Power Output, MWe
ASU
Based Plant
OITM
Based Plant
430.2
347.0
463.3
319.0
26.0
33.0
3556.0
0.0
308.0
0.0
30.0
2087.0
774.0
375.4
0.0
37.0
2644.0
936.0
1696.0
1850.0
2250.0
Flow Rates, Klb/hr
Coal
Limestone
Ash
Boiler Flue Gas
Gas to CO2 Processing
Condenser Duty, MMBtu/hr
Account #
1
2
Account Title
Plant Costs by Account
Coal and Limestone Handling
Coal
Limestone
15,712
3,768
15,357
0
17,465
0
Coal
Limestone
3,145
8,390
3,074
0
3,496
0
57,933
57,933
57,933
188,301
156,153
115,005
206,791
188,031
20,705
71,033
0
14,643
0
111,493
17,077
0
131,662
0
0
48,627
11,004
11,827
1,762
0
7,782
0
in duct work
0
9,076
2,680
205
18,420
107,334
107,334
107,334
29,693
31,419
35,682
9,373
8,810
10,097
Coal and Limestone Prep & Feed
3
Feedwater & Misc BOP Systems
4
Boiler and Accessories
Boiler, SCR*, Air Heater, Fans, Ducts, etc
Oxygen Supply
5
Flue Gas Clean Up
Baghouse & Accessories
FGD
CO2 Processing
6
Combustion Turbine & Accessories
7
HRSG, Ducting, & Stack
Duct Work
Stack
Foundations
HRSG
8
Steam Turbine Generator
9
Cooling Water System
10
Slag/Ash Handling Systems
11
Accessory Electric Plant
33,275
33,741
35,212
12
Instrumentation & Control
14,807
15,014
15,668
13
Improvements to Site
9,405
9,537
9,953
14
Buildings & Structures
35,517
36,014
37,584
632,984
723,310
952,993
Totals
$/kW
141
1,471
2,084
2,057
4.4.3 Total Plant Investment (TPI)
The TPI at date of start-up includes escalation of construction costs and
allowance for funds used during construction (AFUDC). AFUDC includes interest
during construction as well as a similar concept for timing of equity funds over the
construction period. TPI is computed from the TPC based on a linear draw down
schedule and the compounded interest (or implied equity rate) in the percentages
of debt and equity. Draw down was over the assumed 48-month construction
schedule for all three plants. As the analysis is done in constant 2006 dollars, no
escalation was applied. The full AFUDC is used in calculating returns on debt
and equity, but only the interest during construction is included in the
depreciation base.
4.4.4 Total Capital Requirement (TCR)
The TCR includes all capital necessary to complete the entire project. TCR
consists of TPI, prepaid royalties, pre-production (or start-up) costs, inventory
capital, initial chemical and catalyst charge, and land cost:
•
•
•
•
Royalty Costs have been assumed to be zero, as none apply.
Start-Up/Pre-Production Costs are intended to cover operator training,
equipment checkout, extra maintenance, and use of fuel and other materials
during plant start-up. They are estimated as follows:
- Hiring and phasing-in prior to and during start up of operating and
maintenance labor, administrative and support labor, variable operating
costs ramped up to full capacity (including fuel, chemicals, water, and
other consumables and waste disposal charges. These variable costs are
assumed to be compensated by electric energy payments during the start
up period.
- Costs of spare parts usage, and expected changes and modifications to
equipment that may be needed to bring the plant up to full capacity.
Inventory capital is the value of inventories of fuel, other consumables, and
by-products, which are capitalized and included in the inventory capital
account. The inventory capital is estimated as follows:
- Fuel inventory is based on full-capacity operation for 30 days.
- Inventory of other consumables (excluding water) is normally based on
full-capacity operation for the same number of days as specified for the
fuel.
- ½ percent of the TPC equipment cost is included for spare parts.
Initial catalyst and chemical charge covers the initial cost of any catalyst or
chemicals that are contained in the process equipment (but not in storage,
which is covered in inventory capital). No value is shown because costs are
assumed to have been included in the component equipment capital cost.
1. Land cost is based on 100 acres of land at $1,600 per acre.
142
4.4.5 Operating Costs And Expenses
Operating costs were expressed in terms of the following categories:
•
•
•
•
•
Operating Labor
Maintenance Cost
- Maintenance labor
- Maintenance materials
Administrative and Support Labor
Consumables
Fuel Cost
These values were calculated consistent with EPRI TAG methodology. All costs
were based on a first year basis in January 2006 dollars. The first year costs do
not include start-up expenses, which are included in the TCR.
The cost categories listed above are calculated, on a dollars per year basis, as
follows:
•
Operating labor is calculated by multiplying the number of operating
personnel with the average annual (burdened) compensation per person.
•
Maintenance costs are estimated to be 2.2% of the TPC and are divided
into maintenance labor and maintenance materials
-
Maintenance labor is estimated to be 40% of the total maintenance
cost
Maintenance materials are estimated to be 60% of the total
maintenance cost
•
Administrative and support labor is estimated to be equal to 30% of the
sum of operating and maintenance labor.
•
Consumables are feedstock and disposal costs calculated from the annual
usage at 100 per cent load and 85 per cent capacity factor. The costs is
expressed in year 2006 dollars and levelized over 20 years on a constant
dollar basis.
Fuel cost is calculated based on a coal delivered cost of $1.34/MMBtu. Fuel cost
is determined on a first year basis and levelized over 20 years on a constant
dollar basis. The calculation of first year fuel costs is done as follows:
Fuel (tons/day) = Full Load Coal Feed Rate (lb/hr) x 24 hr/day / 2000 lbs/ton
Fuel Unit Cost ($/ton) = HHV (Btu/lb) x 2000 lb/ton
143
Fuel Cost (1st year) = Fuel (tons/day) x Fuel Unit Cost ($/ton) x 365 day/yr x 0.85 (CF)
The operating and maintenance costs, excluding fuel and consumables, are
combined and divided into two components: 1) 90 per cent for Fixed O&M, which
is independent of power generation, and 2) 10 per cent for Variable O&M, which
is proportional to power generation.
4.4.6 Cost Of Electricity (COE)
The COE value is made up of contributions from the capital cost (called the
carrying charge), operating and maintenance costs, consumables, and fuel costs.
The following relationship is used to calculate COE from these cost components:
COE = LCC + LFOM x 100/(8760 x CF) + LVOM + LCM +LFC
LCC = Levelized carrying charge, ¢/kWh
LFOM = Levelized fixed O&M, $/kW-yr
LVOM = Levelized variable O&M, ¢/kWh
LCM = Levelized consumables, ¢/kWh
LFC = Levelized fuel costs, ¢/kWh
CF = plant capacity factor (0.85)
The CO2 mitigation cost (MC) shows the cost impact, in dollars per tonne of CO2
that would otherwise be emitted, of a configuration that allows CO2 capture
relative to the air-fired reference plant.
The MC is calculated as follows:
MC = COEwith removal - COEreference x 0.01 $/¢
Ereference – Ewith removal
COE = Cost of electricity in ¢/kWh
E = CO2 emission in tonnes/kWh
The capital investment and revenue requirements of the three plants are
presented in detail in Table 4.4.8 through Table 4.4.10 and summarized in Table
4.4.11.
With the oxygen based plants having higher total plant costs than the air-fired
reference plant, their interest during construction is higher and they have higher
total plant investment costs; start-up and working capital costs, which are
somewhat related to total plant costs are also higher and yield total capital
requirement costs that are 14 and 51 per cent higher than the air-fired plant.
144
The air-fired reference plant incorporates an ammonia based SCR and a
limestone based scrubber to control its NOx and SOx emissions. Although these
systems are not required with the oxygen based plants, the latter incorporates
oxygen supply and CO2 processing systems. Since the number of systems
added equaled the number deleted, it was assumed, in the absence of a detailed
staffing study, that all three plants required the same number of operating
personnel. Although operator costs are identical, maintenance and administration
costs, which key off of total plant costs, are higher for the oxygen-based plants.
The consumable requirements of the three plants are given in Table 4.4.12. With
the SCR and scrubber systems deleted, the consumable costs of the oxygenbased plants are lower (ammonia and limestone costs are eliminated), but
because of their lower efficiency, their fuel costs are higher (per net MWe). The
higher fuel and higher operating and maintenance costs exceed the lower
consumable cost savings and, as a result, the ASU and OITM plants have higher
20 year levelized production costs of $24.42/MWhr and $22.27/MWhr,
respectively, versus $20.93/MWhr for the air-fired plant. When added to their
respective higher capital carrying costs of $41.75/MWhr and $41.21/MWhr
versus $29.48/MWhr, their costs of electricity of $66.17/MWhr and $63.48/MWhr
are 31 and 26 per cent higher than the air-fired reference plant at $50.41/MWhr
(see Figure 4.4.1). The breakdown of the COE for the three plants is summarized
in Figure 4.4.2, where plant capital cost is divided into CO2 Processing, ASU,
Turbines, Boiler, and Balance of Plant (BOP).
The CO2 mitigation costs of the ASU and OITM based plants were calculated to
be $20.23 and $16.77 per tonne of CO2 sent to the pipeline for sequestering (not
including transportation cost). With the OITM based plant offering the promise of
a lower CO2 mitigation cost, this analysis has shown that reducing the costs of
both the oxygen supply and the CO2 processing systems are a key to reducing
both the cost of electricity and CO2 mitigation costs of these sequestration ready
power plants.
4.4.7 Tube Weld Overlay
As discussed in Section 4.3.4 tube wall weld overlays may be required to
mitigate furnace waterwall corrosion. Conservatively assuming weld overlay is
applied to the entire furnace adds approximately $7.5 million to the total plant
cost. This increases the COE of the cryogenic O2-PC to 66.7 $/MWhr (increase
of 0.75%) and the MC to $20.8 per tonne (increase of 3.0%). The weld overlay
increases the COE of the OITM O2-PC to 63.7 $/MWhr (increase of 0.3%) and
the MC to $17.0 per tonne (increase of 1.5%). Utilizing weld overlays would be
more cost effective in mitigating corrosion than FGD since adding an FGD
system (upstream of the recycle split) would increase the COE by approximately
9%. Note that adding an FGD system downstream of the recycle split (to remove
SOx from the pipeline) would increase the COE by approximately 6%.
145
4.4.8 Reduced Pipeline Pressure
The analysis performed herein was conservatively performed for a pipeline
pressure of 3000 psia. The Reference 14 guidelines specify a pipeline pressure
of 2200 psia. At 2200 psia system efficiency is increased by 0.1% and COE is
reduced by 0.3%. The economic effects of adding tube overlay and reducing the
pipeline pressure are shown in Table 4.4.13.
146
Table 4.4.8 - Air Fired PC Plant Capital Investment & Revenue Requirement
TITLE/DEFINITION
Case:
Plant Size:
Fuel (type):
Design/Construction:
TPC (Plant Cost) Year:
Capacity Factor:
Air Fired
430.2 MWe (net)
Illinois No 6 Coal
48 Months
Jan-06
85.0%
CAPITAL INVESTMENT:
TOTAL PLANT COST
AFUDC
TOTAL PLANT INVESTMENT
Royalty Allowance
Start Up Costs
Working Capital
Debt Service Reserve
TOTAL CAPITAL REQUIREMENT
OPERATING & MAINTENANCE COSTS (2006)
Operating Labor
Maintenance Labor
Maintenance Material
Administrative & Support Labor
TOTAL OPERATION & MAINTENANCE (2006)
FIXED O&M (2006)
VARIABLE O&M (2006)
CONSUMABLE OPERATING COSTS, LESS FUEL (2006)
Water and Treatment
Limestone
Ash Disposal
Ammonia
Other Consumables
TOTAL CONSUMABLES (2006)
Steam Turbine:
Net Efficiency:
Fuel Cost:
Book Life:
4005psig/1076F/1112F
39.5 % HHV
$1.34/MMBtu
20 Years
$x1000
632,982
105,458
732,440
$/kW
1,471
1,717
0
16,458
4,795
0
759,693
1,815
4,921
5,570
8,355
3,147
21,994
19,795
2,199
1,297
1,452
2,423
1,768
1,005
7,945
0
BY-PRODUCT CREDITS (2006)
FUEL COST (2006)
Coal
FUEL COST (2006)
37,131
st
PRODUCTION COST SUMMARY
Fixed O&M
Variable O&M
Consumables
By-Product Credit
Fuel
TOTAL PRODUCTION COST (2006)
LEVELIZED 20 YEAR CARRYING CHARGES (Capital)*
LEVELIZED 20 YEAR BUSBAR COST OF POWER
*Levelized Fixed Charge Rate = 12.5%
147
1 Year
(2010))
$/MWhr
6.18
0.69
2.48
0.00
11.59
20.94
20 Year Levelized
$/MWhr
6.18
0.69
2.48
0.00
11.59
20.93
29.48
50.41
Table 4.4.9 - ASU Based PC Plant Capital Investment & Revenue
Requirement
TITLE/DEFINITION
Case:
Plant Size:
Fuel (type):
Design/Construction:
TPC (Plant Cost) Year:
Capacity Factor:
Oxygen by ASU
347.0 MWe (net)
Illinois No 6 Coal
48 Months
Jan-06
85.0%
CAPITAL INVESTMENT:
TOTAL PLANT COST
AFUDC
TOTAL PLANT INVESTMENT
Royalty Allowance
Start Up Costs
Working Capital
Debt Service Reserve
TOTAL CAPITAL REQUIREMENT
OPERATING & MAINTENANCE COSTS (2006)
Operating Labor
Maintenance Labor
Maintenance Material
Administrative & Support Labor
TOTAL OPERATION & MAINTENANCE (2006)
FIXED O&M (2006)
VARIABLE O&M (2006)
CONSUMABLE OPERATING COSTS, LESS FUEL (2006)
Water and Treatment
Limestone
Ash Disposal
Ammonia
Other Consumables
TOTAL CONSUMABLES (2006)
Steam Turbine:
Net Efficiency:
Fuel Cost:
Book Life:
4005psig/1076F/1112F
33.0 % HHV
$1.34/MMBtu
20 Years
$x1000
723,310
120,507
843,818
$/kW
2,084
2,432
0
18,806
5,480
0
868,103
2,502
4,921
6,365
9,548
3,386
24,220
21,798
2,422
898
0
1,111
0
1,005
3,014
0
BY-PRODUCT CREDITS (2006)
FUEL COST (2006)
Coal
FUEL COST (2006)
35,851
st
PRODUCTION COST SUMMARY
Fixed O&M
Variable O&M
Consumables
By-Product Credit
Fuel
TOTAL PRODUCTION COST (2006)
LEVELIZED 20 YEAR CARRYING CHARGES (Capital)*
LEVELIZED 20 YEAR BUSBAR COST OF POWER
*Levelized Fixed Charge Rate = 12.5%
148
1 Year
(2010))
$/MWhr
8.43
0.94
1.17
0.00
13.88
24.42
20 Year Levelized
$/MWhr
8.43
0.94
1.17
0.00
13.88
24.42
41.75
66.17
Table 4.4.10 - OITM Based PC Plant Capital Investment & Revenue
Requirement
TITLE/DEFINITION
Case:
Plant Size:
Fuel (type):
Design/Construction:
TPC (Plant Cost) Year:
Capacity Factor:
Oxygen by Ion
Transport Membrane
463.3 MWe (net)
Illinois No 6 Coal
48 Months
Jan-06
85.0%
Steam Turbine:
Net Efficiency:
Fuel Cost:
Book Life:
CAPITAL INVESTMENT:
TOTAL PLANT COST
AFUDC
TOTAL PLANT INVESTMENT
Royalty Allowance
Start Up Costs
Working Capital
Debt Service Reserve
TOTAL CAPITAL REQUIREMENT
OPERATING & MAINTENANCE COSTS (2006)
Operating Labor
Maintenance Labor
Maintenance Material
Administrative & Support Labor
36.1 % HHV
$1.34/MMBtu
20 Years
$x1000
952,993
158,774
1,111,767
2,400
0
24,778
7,220
0
1,143,764
2,469
$/kW
2,057
4,921
8,386
12,579
3,992
TOTAL OPERATION & MAINTENANCE (2006)
FIXED O&M (2006)
VARIABLE O&M (2006)
29,879
26,891
2,988
CONSUMABLE OPERATING COSTS, LESS FUEL (2006)
Water and Treatment
Limestone
Ash Disposal
Ammonia
Other Consumables
TOTAL CONSUMABLES (2006)
899
0
1,356
0
1,005
3,259
0
BY-PRODUCT CREDITS (2006)
FUEL COST (2006)
Coal
4005psig/1076F/1112F
FUEL COST (2006)
43,696
st
PRODUCTION COST SUMMARY
Fixed O&M
Variable O&M
Consumables
By-Product Credit
Fuel
TOTAL PRODUCTION COST (2006)
LEVELIZED 20 YEAR CARRYING CHARGES (Capital)*
LEVELIZED 20 YEAR BUSBAR COST OF POWER
*Levelized Fixed Charge Rate = 12.5%
149
1 Year
(2010))
$/MWhr
7.79
0.86
0.94
0.00
12.67
22.27
20 Year Levelized
$/MWhr
7.79
0.86
0.94
0.00
12.67
22.27
41.21
63.48
Table 4.4.11 - Summary of Plant Economics and CO2 Mitigation Costs
Reference
Air-Fired Plant
ASU Based
Plant
OITM Based
Plant
Net Power Output, MWe
430.2
347.0
463.3
Efficiency, % (HHV)
39.5
33.0
36.1
Coal Flow, Klb/hr
319.0
308.0
375.4
718.0
0.939
866.0
0.848
633.0
1,471
723.3
2,084
953.0
2,057
Total Capital Requirement in Millions of Dollars
760.0
868.1
1143.8
Levelized Production Costs, $/MWhr
Operating & Maintenance
Consummables
Fuel
Total
6.86
2.48
11.59
20.93
9.37
1.17
13.88
24.42
8.66
0.94
12.67
22.27
Levelized Capital Carrying Charges, $/MWhr
29.48
41.75
41.21
Levelized COE, $/MWhr
50.41
66.17
63.48
20.23
16.77
CO2 to Stack
Klb/hr
Tonnes/MWhr
739.0
0.779
CO2 to Pipe Line
Klb/hr
Tonnes/MWhr
Total Plant Cost
Millions of Dollars
$/kW
CO2 Mitigation Cost, $/tonne
150
Table 4.4.12 - Plant Daily Consumable Requirements
Air-Fired
3,483
17,041
312
25
371
Cryogenic
ASU
2,414
17,041
0
0
358
Fuel, tons/day
3,828
3,696
4,505
Fuel, lbs/hr/MWe
Fuel, Btu/hr/kWe
741.5
8625
887.6
10324
810.3
9424
Water, 1000s gal/day
Water Treatment Chemicals, lbs/day
Limestone, tons/day
Ammonia (28% NH3), tons/day
Ash Disposal, tons/day
151
OITM
2,414
17,041
0
0
437
Figure 4.4.1 – Increase in COE of O2 Plant Above Air-Fired Reference Plant
35.0%
Increase in COE
30.0%
25.0%
20.0%
15.0%
10.0%
5.0%
0.0%
Cryogenic ASU
OITM
152
Figure 4.4.2 – COE Breakdown
7.00
6.00
COE (cents/kwh)
5.00
CO2 Processing
ASU
BOP
Turbines
Boiler
Consumables
O&M
Fuel
4.00
3.00
2.00
1.00
0.00
Air
Cryogenic
153
OITM
Table 4.4.13 – COE and CO2 MC with Weld Overlay and 2200 psi Pressure
Reference
Air-Fired Plant
ASU Based
Plant
OITM Based
Plant
50.4
66.2
66.7
66.5
63.5
63.7
63.6
20.2
20.8
20.6
16.8
17.1
16.8
Levelized COE, $/MWhr
Base
With tube overlay & 3000 psi pipeline pressure
With tube overlay & 2200 psi pipeline pressure
CO2 Mitigation Cost, $/tonne
Base
With tube overlay & 3000 psi pipeline pressure
With tube overlay & 2200 psi pipeline pressure
154
4.4.9
Comparison with Other Technologies
An economic comparison was performed between the O2-PC and other
competing CO2 removal technologies. For comparison the following alternate
technologies were chosen:
Air PC:
Supercritical PC plant with post-combustion CO2 mitigation (Ref.
[15] case 12).
NGCC:
Natural Gas Combined Cycle with post combustion (Ref. [15] case
14).
IGCC:
Integrated Gasification Combined Cycle with pre-combustion CO2
mitigation (Ref. [15] case 4).
SUB O2PC: Oxygen-fired subcritical PC (Ref. [3]).
The economics of these technologies were compared with the supercritical O2PC using both the levelized cost of electricity and the CO2 mitigation cost as
indexes. The CO2 mitigation cost (MC) shows the cost impact, in dollars per
tonne of CO2 that would otherwise be emitted, of a configuration that allows CO2
capture relative to the reference plant.
The COE and MC for the Air PC, NGCC, and IGCC were obtained from Ref. 15.
Since the economic analysis of Ref. 15 were made for a larger power plant (480550 MWe net power) they were scaled to a 30% smaller power plant to be
consistent with the supercritical O2-PC analyzed herein. The COE and MC for the
subcritical O2-PC were obtained from Ref. 3 and adjusted from 2004 to 2006
dollars and from a coal cost of $1.14/MMBtu to $1.34/MMBtu.
Figure 4.4.3 and Figure 4.4.4 present a comparison of the COE and MC using an
85% capacity factor. Compared to the COE of the supercritical cryogenic O2-PC,
the COE for the other technologies is 52% higher for Air PC, 35% higher for
NGCC, 15% higher for IGCC, and 5% higher for the subcritical O2-PC, and 4%
lower for the supercritical O2-PC with OITM. Compared to the MC of the
supercritical cryogenic O2-PC, the MC for the other technologies is 238% higher
for NGCC, 192% higher for Air PC, 25% higher for IGCC, 5% higher for the
subcritical O2-PC, and 17% lower for the supercritical O2-PC with OITM. Since
based on operating experience an 85% capacity factor for IGCC technology
appears too optimistic, the COE and MC with a 70% capacity factor is also
shown in Figure 4.4.3 and Figure 4.4.4 (COE is increased by 16% and MC by
18%).
155
Figure 4.4.3 - Comparison of Levelized Cost of Electricity Among
Alternative Technologies
12.0
Levelized COE, c/kWh
10.0
70%
CF
8.0
6.0
4.0
2.0
0.0
Air PC
NGCC
Post Comb. Post Comb.
IGCC
Pre Comb.
O2 PC Sub
O2-Fired
O2 PC ASU O2 PC OITM
O2-Fired
O2-Fired
Figure 4.4.4 - Comparison of Mitigation Costs Among Alternative
Technologies
Mitigation Cost, $/tonne of CO2
80.00
70.00
60.00
50.00
40.00
30.00
70%CF
20.00
10.00
0.00
NGCC
Air PC
Post Comb. Post Comb.
IGCC
O2 PC Sub
Pre Comb.
O2-Fired
156
O2 PC ASU O2 PC OITM
O2-Fired
O2-Fired
5.0 Conclusion
To assure continued U.S. power generation from its abundant domestic coal
resources, new coal combustion technologies must be developed to meet future
emissions standards, especially CO2 sequestration. Current conventional coalfired boiler plants burn coal using 15-20% excess air producing a flue gas, which
is only approximately 15% CO2. Consequently, CO2 sequestration requires noncondensable gases stripping, which is both expensive and highly powerconsumptive. Several different technologies for concentrating the CO2 by
removing the non-condensable gases have been proposed including aminebased absorption and membrane gas absorption. However, these techniques
require substantial energy, typically from low-pressure steam.
A new boiler is presented where the combustion air is separated into O2 and N2
and the boiler uses the O2, mixed with recycled flue gas, to combust the coal.
The products of combustion are thus only CO2 and water vapor. The water vapor
is readily condensed, yielding a pure CO2 stream ready for sequestration. The
CO2 effluent is in a liquid form and is piped from the plant to the sequestration
site. The combustion facility is thus truly a zero emission stackless plant.
A conceptual design of a CO2 sequestration-ready oxygen-based 460 MWe
supercritical PC boiler plant was developed. The selected O2-fired design case
has a system efficiency of 32.9% compared to the air-fired system efficiency of
39.5%.
The efficiency and cost-effectiveness of carbon sequestration in oxygen-firing
boilers can be optimized by specifically tailoring boiler design by appropriate
surface location, combustion system design, material selection, furnace layout,
and water/steam circuitry. Boiler efficiencies of near 100% can be achieved by
recovery of virtually all of the flue gas exhaust sensible heat. Boiler size can be
drastically reduced due to higher radiative properties of O2-combustion versus
air-combustion. Furthermore, a wider range of fuels can be burned due to the
high oxygen content of the combustion gas and potential for high coal preheat.
A conceptual design of a CO2 sequestration-ready oxygen-based 460 MWe
supercritical PC boiler plant was developed with integration of advanced oxygen
separation techniques, such as OITM and CAR. The optimized OITM O2-fired
design case has a CO2 removal specific power penalty of 42 kWh/klbCO2 and a
system efficiency of 36.1% compared to the air-fired system efficiency of 39.5%.
Considering that CO2 compression itself consumes 40 kWh/klbCO2, the OITM
integration into the O2-PC is a breakthrough in CO2 removal. The CAR process
efficiency loss and specific power penalty lies approximately midway between the
cryogenic ASU and OITM.
The O2-fired PC CO2 removal penalty with integration of OITM is nearly a quarter
of that from post combustion CO2 removal technologies, and only a half of IGCC.
157
OITM faces significant challenges with respect to the manufacture and stability of
membranes, and scale up and design of large plants.
A design and analysis of a reference air-fired boiler, an oxygen-fired boiler with
cryogenic ASU, and an oxygen-fired boiler with oxygen ion transport membrane
were performed. The O2-PC supercritical boiler incorporates the OTU BENSON
vertical technology, which uses low fluid mass flow rates in combination with
optimized rifled tubing. The following conclusions are made comparing the airfired furnace with the oxygen-fired furnace design and performance:
1. The oxygen furnace has only approximately 65% of the surface area and
approximately 45% of the volume of the air-fired furnace due to the higher
heat flux of the oxygen-fired furnace.
2. Due to the higher O2 concentration of the oxygen-fired furnace versus the
air-fired furnace (40% vs. 21%), the maximum flame temperature of the
oxygen-fired furnace is approximately 500°F higher.
3. Maximum wall heat flux in the oxygen-fired furnace is about 2.5 times that
of the air-fired furnace (175,000 Btu/hr-ft2 vs. 70,000 Btu/hr-ft2) due to the
higher flame temperature and higher H2O and CO2 concentrations.
4. 100% coal burnout is achieved in the oxygen-fired furnace (compared to
99.6% burnout in the air-fired furnace) due to higher furnace temperature
and higher concentration of oxygen. The burnout differential between the
oxygen-fired boiler and the air-fired boiler is expected to be significantly
greater when harder to burn fuels are fired.
5. The higher heat flux of the oxygen-fired furnace significantly increases the
maximum waterwall temperature (from 870°F for the air-fired furnace to
1060°F for the oxygen-fired furnace) requiring the material to be upgraded
from T2 to T92.
6. NOx is reduced by oxygen firing (compared to air-firing) by about a factor
of two from 0.38 lb/MMBtu to 0.18 lb/MMBtu.
7. A pressure equalization header is included at an elevation of 80’ to ensure
stable operation at low loads.
8. The total heat transfer surface required in the HRA is 35% less than the
air-fired HRA due to more heat being absorbed in the oxygen-fired furnace
and the greater molecular weight of the oxygen-fired flue gas. To minimize
the required surface area, the HRA design is in series for the cryogenic
ASU O2-PC and parallel for the OITM O2-PC.
158
9. To reduce the corrosion in the O2-PC, weld overlays with high Nickel and
Chromium contents (e.g. alloy 622) are applied to the waterwalls. This
reduces the predicted maximum corrosion from 45 mil/yr to 10 mil/yr.
10. The required HRA tube materials and wall thicknesses are nearly the
same for the air-fired and oxygen-fired design since the flue gas and
water/steam temperature profiles encountered by the heat transfer banks
are similar.
11. A tubular convective air heater is included in the OITM O2-PC to provide
the necessary air heating for the membrane separation process. The
furnace air heater is an Incoloy MA956 three-pass tubular design situated
above the furnace nose.
The levelized cost of electricity of the supercritical pressure, air-fired reference
plant was calculated to be $50.41/MWhr and it operated with an efficiency of 39.5
per cent. The oxygen supply and CO2 gas processing systems required by the
oxygen-based plants significantly increase their plant costs and parasitic power
requirements. The cryogenic ASU based plant had a cost of electricity of
$66.17/MWhr and an efficiency of 33.0 per cent.
The oxygen transport membrane of the OITM based plant operates at 200 psia
with air heated to 1600ºF via tubes placed in the boiler. The high cost of this
boiler tubing, together with piping, heat exchangers, and the oxygen transport
membrane itself, add considerable costs to the plant. Since the nitrogen
exhausting from the membrane is hot and at pressure, it can be used for power
recovery via a hot gas expander and HRSG. Although addition of these
components further increases plant costs, the added power they provide
increases the plant efficiency to 36.1 per cent and enables the OITM based plant
to operate with a cost of electricity that is less than that of the ASU based plant
e.g. $63.48/MWhr versus $66.17/MWhr. As a result the CO2 mitigation cost of the
OITM based plant was calculated to be less than that of the ASU based plant
e.g. $15.66/tonne versus $17.06/tonne, respectively.
The addition of weld overlay increases the COE approximately 0.75%, but
reducing the pipeline pressure from 3000 psia to 2200 psia (specified in Ref. 14)
reduces the COE by 0.3%. Thus, the combined effect of these two adjustments
on the COE is relatively small (+0.4% for cryogenic O2-PC and +0.1% for the
OITM O2-PC).
The oxygen supply and the CO2 gas processing systems have a major impact on
the economics and efficiency of oxygen-based plants. Additional R&D aimed at
improving these systems, especially OITMs, is necessary to improve both
electricity costs and CO2 mitigation costs.
The O2-fired PC CO2 removal penalty is nearly half that of post combustion CO2
removal technologies and 15% to 30% than that of IGCC pre-combustion
159
capture. Furthermore, of the CO2 sequestration-ready technologies, the O2-fired
PC is the simplest, requires the least modification of existing proven designs, and
requires no special chemicals for CO2 separation.
Thus CO2 sequestration with an oxygen-fired combustion plant can be performed
in a proven reliable technology while maintaining a low-cost high-efficiency power
plant. As new lower power-consuming air separation techniques, such as
membrane separation, become commercially available for large-scale operation
in O2-fired plants, the CO2 removal power consumption and efficiency reduction
will continue to decline.
160
6.0 References
1. White, J.S., Buchanan, T.L., Schoff, R.L. (Parsons); Stiegel, G. (DOE);
Holt, N.A., Booras G. (EPRI); Wolk, R. (WITS), “Evaluation of Innovative
Fossil Fuel Power Plants with CO2 Removal”, EPRI Report No. 1000316,
December 2000.
2. Wilkinson, M., Boden, John (BP), Panesar, R. (Mitsui Babcock), Allam, R.
(Air Products), “CO2 Capture Via Oxygen Firing: Optimisation of a Retrofit
Design Concept for a Refinery Power Station Boiler”, First National
Conference on Carbon Sequestration, May 15-17, 2001, Washington DC.
3. Seltzer, Andrew and Fan, Zhen, Conceptual Design of Oxygen-Based PC
Boiler, DOE # DE-FC26-03NT41736, September 2005.
4. Philip A. Armstrong, E.P. Foster, Dennis A Horazak, Harry T Morehead,
VanEric E. Stein, “Ceramic and Coal: ITM Oxygen for IGCC”, The 22nd
Pittsburgh Coal Conference, Sept.11-15, 2005, Pittsburgh, USA
5. Divyanshu Acharya, Krish R. Krishnamurthy, Michael Leison, Scott
MacAdam, Vijay K. Sethi, Marie Anheden, Kristin Jordal, Jinying Yan,
“Development of a High Temperature Oxygen Generation Process and Its
Application to Oxycombustion Power Plants with Carbon Dioxide
Capture”, ibid
6. A.A. Leontiou, A.K. Ladavos, T.V. Bakas, T.C. Vaimakis, & P.J. Pomonis,
“Reverse uptake of oxygen from La1-xSrx(Fe3+/Fe4+)O3 perovskite-type
mixed oxides”, Applied Catalyst A: General 241 (2003) p143-154.
7. “Greenhouse Gas Emissions Control By Oxygen Firing In Circulating
Fluidized Bed Boilers: Phase 1 – A Preliminary Systems Evaluation”,
Alstom Power Inc. Power Plant Laboratories, PPL Report No. PPL-03CT-09, DOE/NETL Cooperative Agreement No. DE-FC26-01NT41146,
May 15, 2003.
8. Philip A. Armstrong, “Method for Predicting Performance of an Ion
Transport Membrane Unit-Operation”, Air Products and Chemicals, Inc.
9. FW-FIRE, Fossil-fuel Water-walled Furnace Integrated Reaction Emission,
Theory and User’s Manual, Foster Wheeler Development Corp., 11/30/99.
10. HEATEX Computer Program, “A Program to Determine the
Thermal/Hydraulic Performance of Gas-Cooled Heat Exchangers”,
Revision 24, Foster Wheeler Corporation.
161
11. EMISS Computer Program, “A Program
Emissivity”, Foster Wheeler Corporation.
to
Determine
Gaseous
12. STADE2 Computer Program, ”Calculation Program for Pressure Drop and
Heat Transfer in Tubes”, Siemens, Version 4.40.
13. DYNASTAB Computer Program, “Dynamic Stability of the Flow through
Evaporator Heating Tubes in Fossil Fired Steam Generators”, Siemens,
Version 1.1.
14. “Carbon Capture and Sequestration System Analysis Guidelines”,
DOE/NETL, April 2005.
15. “2006 Cost and Performance Comparison of Fossil Energy Power Plants”,
DOE/NETL-401/053106, Draft Final May 2006.
16. “Task 1 Topical Report - IGCC Plant Cost Optimization”, Gasification Plant
Cost and Performance Optimization, DOE/NETL Contract No. DE-AC2699FT40342, May 2002.
162
7.0 Bibliography
N/A
163
8.0 List of Acronyms and Abbreviations
ASU
BOP
CAR
CDT
CF
CFD
COE
DNB
E
EPRI
FD
FEGT
FGD
FW
FW-FIRE
GT
HHV
HIPPS
HRA
HRSG
ID
IGCC
LCC
LCM
LFC
LFOM
LHV
LMPD
LMTD
LOI
LP
LVOM
MC
NGCC
NOx
O&M
OD
OFA
OITM
OTU
PC
PFBC
PSA
RH
Air Separation Unit
Balance of Plant
Ceramic auto-thermal recovery
Compressor discharge temperature
Capacity Factor
Computational fluid dynamics
Cost of Electricity
Departure from nucleate boiling
Emission of CO2
Electric Power Research Institute
Forced draft
Furnace exit gas temperature
Flue Gas Desulfurization
Foster Wheeler
Fossil fuel, Water-walled Furnace Integrated Reaction and
Emission Simulation
Gas turbine
Higher Heating Value
High performance power system
Heat recovery area
Heat Recovery Steam Generator
Induced draft
Integrated gasification combined cycle
Levelized Carrying Charge
Levelized Consumables
Levelized Fuel Costs
Levelized Fixed O&M
Lower Heating Value
Log mean pressure difference
Log mean temperature difference
Loss on ignition
Low pressure
Levelized Variable O&M
Mitigation Cost (CO2)
Natural gas combined cycle
Nitrogen Oxides
Operation and Maintenance
Outside diameter
Over-fired air
Oxygen Ion Transport Membrane
Once-through utility
Pulverized Coal
Pressurized fluidized bed combustion
Pressure swing absorption
Reheater
164
SCR
SH
SOx
ST
T
TCR
TEG
TET
TPC
TPI
UBC
Selective Catalytic Reactor
Superheater
Sulfur Oxides
Steam Turbine
Temperature
Total Capital Requirement
Triethyleneglycol
Turbine Exit temperature
Total Plant Cost
Total Plant Investment
Unburned carbon loss
165