Journal of Energy Chemistry 22(2013)795–803
Effect of mass transfer limitations on catalyst performance
during reduction and carburization of iron based
Fischer-Tropsch synthesis catalysts
Akbar Zamaniyana,b∗ , Yadollah Mortazavib , Abbas Ali Khodadadib,
Ali Nakhaei Poura
a. Department of Natural Gas Conversion, Gas Research Division, Research Institute of Petroleum Industry (RIPI), Tehran 14665-137, Iran;
b. School of Chemical Engineering, University College of Engineering, University of Tehran, Tehran 11155-4563, Iran
[ Manuscript received September 2, 2012; revised October 8, 2012 ]
Abstract
Existence of intraparticle mass transfer limitations under typical Fischer-Tropsch synthesis has been reported previously, but there is no suitable
study on the existence of intraparticle diffusion limitations under pretreatment steps (reduction and activation) and their effect on catalytic
performance for iron based catalysts. In this study, Fe-Cu-La-SiO2 catalysts were prepared by co-precipitation method. To investigate the
intraparticle mass transfer limitation under reduction, activation and reaction steps, and its effect on catalytic performance, catalyst pellets with
different sizes of 6, 3, 1 and 0.5 mm have been prepared. All catalysts were calcined, pretreated and tested under similar conditions. The
catalysts were activated in hydrogen (5% H2 in N2 ) at 450 ◦ C for 3 h and exposed to syngas (H2 /CO = 1) at 270 ◦ C and atmospheric pressure
for 40 h. Afterwards, FTS reaction tests were performed for approximately 120 h to reach steady state conditions at 290 ◦ C, 17 bar and a feed
flow (syngas H2 /CO = 1) rate of 3 L/h (STP). Using small pellets resulted in higher CO conversion, FT reaction rate and C5+ productivity
as compared with larger pellets. The small pellets reached steady state conditions just 20 h after starting the reaction. Whereas for larger
pellets, CO conversion, FT reaction rate and C5+ productivity increased gradually, and reached steady state and maximum values after 120 h of
operation. The results illustrate that mass transfer limitations exist not only for FTS reaction but also for the reduction and carburization steps
which lead to various phase formation through catalyst activation. Also the results indicate that some effects of mass transfer limitations in
activation step, can be compensated in the reaction step. The results can be used for better design of iron based catalyst to improve the process
economy.
Key words
Fischer-Tropsch; mass transfer limitation; activation; catalyst; GTL
1. Introduction
In the face of increasing oil prices and decreasing oil reserves, alternative routes for production of liquid fuels attract
considerable attention in research and industrial application.
Fischer-Tropsch synthesis (FTS) offers the possibility of the
conversion of synthesis gas to a mixture of linear hydrocarbons. Fuels produced with FTS are of a high quality due to a
very low aromaticity and near zero sulfur content. New and
stringent regulations may promote replacement or blending of
conventional fuels by sulfur and aromatic-free FTS products
[1−4].
The most common Fischer-Tropsch catalysts are group
VIII metals (Co, Ru and Fe). In conventional FTS processes,
iron and cobalt based catalysts are known to be very effective
catalysts which are used commercially in the temperature
∗
range of 200–300 ◦ C and pressure range of 10–60 bar. Iron
catalysts are commonly used due to their low costs in comparison with other active metals [5].
Various types of reactors (fixed-bed, fluidized-bed,
ebulliating-bed and slurry-phase) differ with respect to the
most suitable particle size of the catalyst used. The range of
catalyst sizes reflects the tradeoff between the desires to provide high catalyst utilization and maintain a manageable pressure drop across the reactor. If a fixed-bed mode of operation
is envisaged, the FT catalyst will generally consist of particles
with a few millimeters in size for reasons of pressure drop and
heat transfer. It can be easily envisaged that in this size, intraparticle pore diffusion limitation may exist, as previously
reported [6,7].
Using a model that assumed first-order kinetics with
respect to hydrogen, Zimmerman et al. [8] strongly suggested
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Copyright©2013, Dalian Institute of Chemical Physics, Chinese Academy of Sciences. All rights reserved.
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Akbar Zamaniyan et al./ Journal of Energy Chemistry Vol. 22 No. 5 2013
that transport limitations of H2 occur at particle with diameters larger than 0.2 mm at temperature higher than 235 ◦ C with
a fused iron ammonia synthesis catalyst. Post et al. [9] also
used first-order behavior and observed transport limitations of
hydrogen at high temperatures (T >220 ◦ C, dP >0.4 mm) with
a number of iron and cobalt based catalysts in a fixed-bed micro reactor. Robert Becker et al. [10] showed computationally that under typical Fischer-Tropsch conditions (200 ◦ C,
2.1 MPa), CO can also become diffusion limited. Intrinsic FT kinetic measurements should be performed with small
(dP <0.2 mm) catalyst particles in order to eliminate diffusion
limitations of one of the reactants. Some of the most pertinent
information can be found in the results obtained by Post et al.
[9], who have demonstrated that the synthesis rate over a fused
iron catalyst which is essentially nonporous in the oxidic state,
increases with the decreasing particle size. This could be due
to the diffusion limitations during synthesis reaction or to the
activation of the nonporous oxidic catalyst particles limited to
a superficial layer only [9].
There are number of studies on mass transfer limitations
and their effects during FTS, but to the best our knowledge
there is no suitable study previously published on the existence of intraparticle pore diffusion limitations under pretreatment steps (reduction and carburization) and specially their
effects on catalyst performance for iron based catalyst. The
diffusion limitations result in an incomplete utilization of the
catalyst particles and lead to changes in reactivity and selectivity. It is of practical importance in the design of iron based
FT catalysts to obtain knowledge of various phases formed
during different pretreatments and under industrial FTS conditions. In addition of FT chemistry, an optimal operation with
respect to product yield and selectivity requires a deep understanding of catalyst behavior in reduction, carburization and
reaction steps.
In this study, using different catalyst sizes, the existence
of intraparticle pore diffusion limitations in porous iron based
FT catalysts under reduction, carburization (activation) and
reaction steps and their effects on catalyst performance, have
been examined and described.
2. Experimental
2.1. Catalyst synthesis
Most early FT catalysts were prepared with precipitation techniques [5]. Alkali-promoted iron catalysts have been
applied industrially for Fischer-Tropsch synthesis for many
years that have high water-gas shift activity, high selectivity to
olefins and appear to be stable with a high H2 /CO ratio synthesis gas [3,11]. Here, Fe-Cu-La-SiO2 catalysts were prepared
by co-precipitation of Fe and Cu nitrates at constant pH of
7.5 to form porous Fe-Cu oxyhydroxide powder, which were
promoted by La(NO3 )3 precursor and silica sol solution after
treatment in air [12]. The nominal composition of the catalyst
by weight ratio was tended to be 100Fe : 6Cu : 6La : 20SiO2 .
To study the effect of intraparticle mass transfer limitations during reduction, activation and reaction steps and
their effects on catalyst performance, the catalyst pellets with
different sizes of 6, 3, 1 and 0.5 mm have been prepared as
shown in Figure 1. All samples were calcined at 450 ◦ C in air
for 3 h.
Figure 1. Appearance photographs of the prepared catalyst pellets with different sizes. (a) 6 mm, (b) 3 mm, (c) 1 mm, (d) 0.5 mm
Journal of Energy Chemistry Vol. 22 No. 5 2013
2.2. Catalyst characterization
Atomic absorption spectroscopy (Analyst 200 Perkin
Elmer) was used to determine the iron and copper content of the catalyst.
The contents of lanthanum and
silica were determined by ICP (Wear Metal Analyzer
400 Perkin Elmer) and gravimetric methods, respectively.
The composition of the catalyst was determined to be
100.0Fe : 5.6Cu : 5.6La : 19.0SiO2.
The surface area was calculated from the BrunauerEmmett-Teller (BET) equation, and pore volume and average pore diameter of the catalyst were determined by N2 physisorption using a Micromeritics ASAP 2010 automated system. A 0.5 g catalyst sample was degassed at 373 K for 1 h
and then at 573 K for 2 h prior to analysis. The analysis was
done using N2 adsorption at 77 K.
XRD was used to determine the phase composition of
the synthesized catalysts. The XRD pattern of the catalysts were collected using an X-ray diffractometer (Philips
PW1840 X-ray diffractometer) by monochromatized Cu/Kα
radiation (40 kV, 40 mA) using a step scan mode at a scan rate
of 0.02o/s (2θ) from 10o −80o. XRD peak identification was
performed by comparison with the JCPDS database software.
2.3. Catalyst reduction and activation
The catalyst is synthesized in the form of a metal oxide, so
it must be subjected to an activation treatment to become active for Fischer Tropsch synthesis. The pretreatment for iron
based catalysts is not straightforward. The pretreatment environment has been reported to influence the initial catalytic
activity significantly [13,14]. The common reported activa-
797
tion treatments for iron catalysts are H2 , CO, or synthesis gas
reduction [5].
Here, the catalyst was activated in hydrogen (5% H2 in
N2 ) at 450 ◦ C for 3 h. The reactor temperature was then
cooled down to 270 ◦ C under flowing hydrogen. Then, the
catalyst was exposed to syngas (H2 /CO = 1) and remained under this condition for about 40 h that is required in order to
form carbide phases. Afterwards, the reactor was pressurized
to 17 bar and the temperature was increased to 290 ◦ C for FTS
reaction tests. The pretreatment and reaction conditions were
selected based on previous works [3,11,15].
2.4. Experimental setup and test procedure
All prepared catalysts were tested in an experimental setup as shown in Figure 2. The feed stream was supplied from
high-pressure carbon monoxide and hydrogen cylinders using
individual mass flow controllers. The experiments were carried out using a stainless steel reactor (14.3 mm ID) placed in
a tubular triple zones furnace. The catalyst, without any dilution, was located between two layers of ceramic beads to
homogenize and preheat the feed before it entered the catalyst bed. For measuring and controlling the reaction temperature, three J-type TICs (temperature indicators and controllers) were installed. The first thermocouple was located
before the catalyst bed in the preheater section whilst the second and the third ones were inside and after the catalyst bed.
No temperature runaway was observed during the reaction
tests. The system pressure was controlled by a backpressure
regulator.
The performances of all the catalysts were measured
under similar conditions. For this purpose, 2.7 g catalyst was
Figure 2. Experimental setup
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Akbar Zamaniyan et al./ Journal of Energy Chemistry Vol. 22 No. 5 2013
loaded in the reactor. After catalyst reduction and activation the feed flow (syngas H2 /CO = 1) equal to 3 L/h (STP)
that provided the space velocity of 750 h−1 was conducted at
17 bar and 290 ◦ C to start FTS reactions. The FTS test was
took time about 120 h to reach steady state conditions.
The reactor effluent was passed through a hot separator operated at 80 ◦ C and cold separator operated at 0 ◦ C to
remove heavy and light liquid products, respectively. Noncondensable gases and feed were analyzed using an on-line
gas chromatograph (Varian CP 3800) equipped with a packed
column (3 m long) and a capillary column (Petrocol DH
100 m long) to determine CO conversion and the formation
of light hydrocarbon products. The liquid products were collected and weighed from hot and cold separators every 24 h.
9 nm, as determined by N2 adsorption (Micrometrics ASAP
2010).
Figure 3 shows XRD analysis for the catalyst samples.
The diffraction lines belong to magnetite (Fe3 O4 ) should be
at about 2θ = 31o , 37o , 59o and 65o , and the diffraction lines
belong to hematite (Fe2 O3 ) should be at about 2θ = 24o , 33o ,
35o , 49o and 54o .
Due to the existence and distribution of SiO2 , the intensity of the hematite peaks was low. Figure 3 illustrates clearly
that the synthesized catalysts contained mainly magnetite and
hematite phases.
2.5. Reaction system and calculations
The overall reactions of Fischer-Tropsch synthesis can be
simplified as a combination of FT reaction and water-gas shift
(WGS) reaction as follows [16]:
CO+(1+m/2n)H2−→(1/n)Cn Hm +H2 O, rFTS
CO+H2 O⇐⇒CO2 +H2 , rWGS
(1)
Figure 3. XRD patterns of synthesized catalyst samples
(2)
where, n is the average carbon number and m is the average
number of hydrogen atoms in hydrocarbon products. Water is
a primary product of FT reaction, and CO2 can only be produced by the water-gas shift reaction, so it can be concluded
that:
RWGS = Rate of water-gas shift reaction (mol/(gFe ·h))
= Rate of CO2 production (3)
Thus the rate of FT reaction can be obtained by the following formula:
RFTS = Rate of FTS (mol/(gFe ·h)) = RCO –RWGS
= RCO –RCO2 (4)
RCO = Rate of CO consumption (mol/(gFe·h))
(5)
CO conversion, CH4 selectivity and catalyst productivity
are required to evaluate the catalyst performance which have
been calculated from the following formulas:
CO conversion = Mole of CO converted/Total CO inlet
(6)
CH4 selectivity = Mole of CH4 produced/CO converted
(7)
Productivity = W desired product (g)/[WFe (g)×time (h)]
3.2. FTS reaction tests
The results for CO conversion versus time during 120 h of
FTS reaction tests are presented in Figure 4. It can be seen that
CO conversions for smaller catalyst size (0.5 mm and 1 mm)
were higher than those of larger ones (3 mm and 6 mm). Also,
as demonstrated in Figure 4, the approach to steady state was
markedly different for different catalyst sizes. It should be
emphasized that the steady-state activity (and selectivity) of
iron catalysts does not reached in minutes but rather over periods of several hours. The time required to reach steady state
conditions and maximum attainable CO conversion for these
catalyst sizes are also different. These results clearly indicated the existence of mass transfer limitations in FTS for the
applied catalyst under operation conditions.
(8)
To ensure the reproducibility of the results, at least two
sets of data under similar conditions were obtained. The statistical analysis on the data was done and error lower than
±2.0% was obtained.
3. Results
3.1. Catalyst characterization
BET surface area of the catalyst was 107 m2 /g, the pore
volume was 0.24 cm3 /g and the average pore diameter was
Figure 4. CO conversion versus time during FTS reaction for various catalyst
sizes
Journal of Energy Chemistry Vol. 22 No. 5 2013
The time required to reach steady state and maximum
value for CO conversion was about 20 h for the 0.5 and 1 mm
catalysts. At this time, CO conversion was about 93% and
82% for the 0.5 and 1 mm catalysts, respectively, whereas at
the same point in time CO conversion for the 3 and 6 mm catalyst pellets were 38% and 60%, respectively. It is surprising
and unexpected that CO conversion for the 6 mm catalyst was
higher than that for the 3 mm one. The following figures and
discussion may help to interpret this behavior.
As Figure 4 indicates, the time required for CO conversion to reach steady state and maximum value was about 100 h
for the 6 mm catalyst. CO conversion for the 3 mm catalyst
increased continuously even after 120 h. Therefore, for the
same catalyst composition, the catalyst with smaller size requires less activation time and shows higher CO conversion
which in turn affects the economy of the process.
The results for FTS rate of reaction versus time are presented in Figure 5. The behavior of various catalyst sizes, in
regard to the required time to achieve steady state, was relatively similar to Figure 4 only with a few differences. Similar to Figure 4, smaller catalysts showed higher FTS rates of
reaction as compared with the larger catalysts. In contrary
to Figure 4, the 0.5 and 1 mm catalysts indicated the same
FTS rate of reaction which in turn implies that they have similar mass transfer limitation characteristics under the reaction
condition. Thus it can be concluded that mass transfer limitations can be eliminated by catalysts with sizes in the range
of 0.5 mm around under the applied condition. The rate of
FTS reaction increased continuously even after 120 h for the
3 and 6 mm catalysts whereas it approached steady state and
maximum value for the 0.5 and 1 mm catalysts faster.
Figure 6 shows the results for WGS reaction rate. The
6 mm and 0.5 mm catalysts displayed the highest WGS reaction rates. Similar to CO conversion and FTS rate of reaction, the rate of WGS reaction for the 3 mm catalyst increased
continuously even after 120 h. On the other hand, WGS rate
reached steady state conditions more rapidly for the others.
The values of major parameters obtained from an analysis
of Figures 4 to 6 are summarized in Table 1.
Figure 5. FT reaction rate versus time during FTS for various catalyst sizes
799
Figure 6. WGS reaction rate versus time during FTS for various catalyst
sizes
The results for CO2 and CH4 selectivities for various catalyst sizes versus reaction time are presented in Figures 7 and
8, respectively.
WGS reaction is the likely source of CO2 production, so
similar results to Figure 6 can be seen here. The 6 mm catalyst showed the highest CO2 and CH4 selectivities. All the
catalysts showed a maximum CH4 selectivity at around 20 h
Figure 7. CO2 selectivity versus time during FTS for various catalyst sizes
Figure 8. CH4 selectivity versus time during FTS for various catalyst sizes
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Akbar Zamaniyan et al./ Journal of Energy Chemistry Vol. 22 No. 5 2013
Table 1. Major reaction parameters for various catalyst sizes
Catalyst size
(nm)
0.5
1
3
6
Maximum CO
conversion (%)
94.3
82.7
70.5
73.3
Time to reach
steady state (h)
20
20
−
100
Initial activity (CO
conversion at 20 h) (%)
93.1
77.3
47.3
57.2
of FTS reaction time. CH4 selectivities for 6 mm and 3 mm
catalysts decreased slowly after 20 h (i.e., the maximum
value) running of FTS reaction.
Figure 9 shows C5+ productivity for various catalyst
sizes. C5+ productivity increased versus time for all the catalysts. The 0.5 mm and 1 mm catalysts exhibited about double
C5+ productivity as compared with the 6 mm and 3 mm catalysts.
Maximum FTS rate of
reaction (mol/(h·gFe ))
0.0215
0.0206
0.0199
0.0164
Maximum WGS rate of
reaction (mol/(h·gFe ))
0.0167
0.0158
0.0135
0.0167
in catalyst performance versus time can be explained by the
differences in distribution of iron phases formed during the
pretreatment which can be attributed to the differences in mass
transfer limitations resulted from differences in catalyst size.
4. Discussion
The significant differences in CO conversion, the required
time to reach steady state, FTS reaction rate and WGS reaction
rate (Figures 4−9) imply the existence of mass transfer limitations. The results can be interpreted by the effect of mass
transfer limitations on various steps, including reduction, carburization and reaction.
It has been reported that cobalt, nickel and ruthenium
remain in metallic state under FTS conditions [16], but the
composition of iron-based catalysts changes during FischerTropsch synthesis. Several phases of iron are known in ironbased catalysts subjected to FTS conditions. These include
metallic iron (α-Fe), iron oxides (hematite, α-Fe2 O3 ; magnetite Fe3 O4 , and Fex O), and five different forms of iron carbides. The formation and composition of these iron phases depend on the process conditions, catalyst activation and deactivation, and catalyst composition [17−21]. There are a number
of studies on iron catalysts concerning the role of iron phases
in FT synthesis that indicate the carbide phases result in high
FT activity whereas the magnetite phase has negligible catalytic activity toward FT reactions but may be active for WGS
reaction [5,16,17,22]. Thus, it is expected that the differences
Figure 9. C5+ productivity versus time during FTS for various catalyst sizes
As mentioned earlier, the reduction and carburization processes are performed by hydrogen and syngas (H2 , CO), respectively. Bulk and surface changes in the speciation of iron
during the pretreatment of FTS reaction on iron catalysts have
been studied and sequential bulk phase modifications during
the activation process from Fe2 O3 to Fe3 O4 to FeCx (iron carbides) have been reported [19−21]. The sequence of transformation for iron phases during the pretreatment and FTS can
be illustrated schematically in Figure 10. It also addresses the
changes that occur during the exposure to H2 and syngas (H2
and CO) environments.
Figure 10. Transformation sequence for iron phase during pretreatment and FTS [20]
The studies indicate that during the pretreatment and reaction steps, Fe2 O3 crystallites are rapidly converted to Fe3 O4
crystallites, but the conversion of Fe3 O4 crystallites to iron
metal and carbides is relatively slow and requires a number of
steps [19,20]. On the other hand, the diffusivities of H2 and
CO are different, with hydrogen diffusivity being greater than
that of CO. Also, different catalyst sizes result in different
diffusion lengths. These parameters influence the H2 and CO
concentration profiles in reduction, carburization and reaction
steps. Thus, the combinations of various phases (hematite,
magnetite, carbides and Fe metal) are different for various catalyst sizes. Figure 11 indicates that this is the case indeed. In
this figure, the sharp peaks observed clearly at 2θ equal to 35o ,
57o and 63o assign to Fe3 O4 , and the peaks at 2θ equal to 39o ,
41o , 43o , 44o , 70o and 78o assign to various types of iron carbides. Also the peaks at 2θ equal to 24.3o, 33.3o, 35.8o, 40.8o,
Journal of Energy Chemistry Vol. 22 No. 5 2013
49.6o, 54.1o, 57.6o and 64.1o assign to Fe2 O3 .
Figure 12 traces the iron phase transformations of iron
catalyst from calcined catalyst through the pretreatments (reduction and carburization) and reaction steps. In this figure,
801
the color contrast, albeit in a qualitative manner depicts, the
amount of each phase present with black color denoting high
content, gray denoting less content and white meaning the
lowest content.
Figure 11. XRD patterns of the catalysts after reaction
Figure 12. Qualitative trace of iron phase transformations of iron catalyst from the calcined catalyst through pretreatments and reaction steps
After 3 h pretreatment of the calcined catalyst in the presence of H2 , hematite phase (Fe2 O3 ) was reduced partially to
Fe3 O4 and totally to metal Fe. Thus, a mixture of Fe2 O3 ,
Fe3 O4 and Fe metal formed. In this step, the content of various phases in the catalyst increased in the order of Fe metal,
Fe2 O3 and Fe3 O4 , as shown in Figure 12. With regard to the
high diffusivity of H2 , it is expected that a similar combination of the aforementioned phases formed for all the catalyst
sizes during this step.
At carburization step after exposure to syngas for 40 h,
Fe3 O4 , Fe2 O3 and metal Fe were converted partially to carbide phases. The reaction itself may promote the carburization. Thus during the reaction step, some of other remaining
phases may be converted to carbide phases and lead to an increase of carbide phase content as shown in Figure 12. It must
be noted that the mass transfer limitation is highlight and important here, due to low diffusivity of CO, number of slow
chain steps, and different diffusion lengths. Thus the degree of
carburization process and the amount of various phases will be
differed for different catalyst sizes. Therefore, it is expected
that the extent of carburization (carbide phase formation) in
terms of catalyst size decreases in the order of 0.5, 1, 3 and
6 mm. In other words, the amount of magnetite phase in terms
of catalyst size increases in the order of 0.5, 1, 3 and 6 mm.
Briefly, it is expected that the catalysts with larger diffusion
length contain mainly magnetite and hematite phases, and less
carbide and active metal phases.
Figure 4 shows that the 0.5 and 1 mm catalyst pellets activate rapidly and reach to maximum CO conversion
value about 20 h after reaction starts. This indicates that these
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Akbar Zamaniyan et al./ Journal of Energy Chemistry Vol. 22 No. 5 2013
catalysts have been reduced, carburized and the active phases
have been formed well. On the other hand, CO conversion
for the 3 mm pellet increases with time gradually. CO is the
key to form carbide phases. As mentioned earlier, there are
more mass transfer limitations under carburization step than
those under reduction one. As the diffusional lengths of the
3 mm catalysts are, on average, larger than those of the 0.5
and 1 mm pellets, mass transfer limitation of CO can affect the
formation rate of carbide phases during carburization. It may
be concluded that the 3 mm catalyst has likely been reduced
relatively well, whereas this is not the case for carburization.
Due to similarity of the reaction and carburization steps, the
carburization process can be continued in parallel to the reaction, and the remaining iron oxides and metal sites can be
converted to carbides phases in the presence of CO. Figure 11
shows that this is the case as the XRD patterns of the catalysts
are relatively similar to each other after reaction.
CO conversion for the 6 mm catalyst shows similar behavior, as mentioned in previous discussion. The increase rate
of CO conversion for the 6 mm catalyst was less than that for
the 3 mm catalyst. Longer diffusion lengths for the 6 mm catalyst relative to the 3 mm one lead to more CO mass transfer
limitations and less formation of carbide phases during the reaction testing. It is surprising that CO conversion for the 6 mm
catalyst is more than that of the 3 mm one.
In spite of Figure 4, the 6 mm catalyst showed a lower
FT reaction rate than the 3 mm one (Figure 5). As mentioned
above, it is expected that the 6 mm catalyst contains mainly
hematite and magnetite phases as compared with others. Referring to prior discussion, due to CO mass transfer limitations, longer diffusion lengths lead to incomplete carburization and active phase formation, not only in the carburization
step but also during the reaction. Thus, for this catalyst, a low
content of active phase (carbide and active metal) and more
magnetite phase (active phase for WGS) are the possible reasons for a lower FT reaction rate and higher WGS reaction
rate, respectively (Figures 5 and 6).
On the other hand, due to a lower FT reaction rate, less
heat is released by the reaction, causing a lower bed temperature for the 6 mm catalyst relative to the other catalysts. As
the water gas shift reaction is an exothermic and equilibrium
limited reaction, low temperature leads to an increase in WGS
reaction rate too. Thus for the 6 mm catalyst, higher WGS
reaction rate relative to the other catalysts and higher CO conversion relative to the 3 mm catalyst are expected. The 0.5 mm
catalyst leads to better filling of the catalyst bed, so the reactor
has better radial heat transfer performance, a smaller temperature gradient, and less hot spot that cause high WGS reaction
rate.
Figure 6 indicated that WGS reaction rate increases with
time for all the catalysts. This is not the reason for increase
of the amount of magnetite phase with time; because, as mentioned before, the magnetite phase decreases with time due to
the continuation of carburization process in the reaction step.
The water concentration increases due to the increase of FT
reaction rate for all the catalyst. This leads to an increase of
WGS reaction rate that is thermodynamically controlled.
Typically, steady-state activity reached only after several
hours of reaction as a result of the slow transformations of
iron and iron oxides species to carbide phase at the surface
and in the bulk of iron crystallites [19,20]. As shown in Figure 4, the 0.5 and 1 mm catalysts reach steady-state activity only after 20 h of the reaction. It demonstrates that due
to low mass transfer limitations, transformations to carbide
phase take place rapidly. Whereas, this not the case for the
3 mm and 6 mm catalysts, which need much more time to
reach steady-state activity (Figures 4 and 5).
As WGS reaction is the source of CO2 production and
with respect to previous figures and results, the behavior of
Figure 7 is known clearly and there is no need for additional
discussion, but the results of Figure 8 need some discussion
as following:
1. Low CO conversion for the 6 mm catalyst as shown in
Figure 4, leads to high CH4 selectivity based on Equation 7.
2. High methane selectivity can be explained by catalyst
surface wetting effect [23]. The methane formation is considered mainly due to methanation reaction at the dry catalyst
surface. Evolution of the liquid product along the catalytic
bed is illustrated in Figure 13. The liquid product can be accumulated in the made volume among the particles. For large
catalyst particles, the accumulated liquid may be low due to
carrying by the gas stream. Also as shown in Figure 5, the rate
of FT reaction is low for large particles due to incomplete carburization. Thus, there is more dry catalyst surface, resulting
in more methane formation. A decrease in methane selectivity for all the catalysts versus time may be explained by this
effect too. The dry catalyst surface can be covered by liquid
product with time, so methane formation decreases slightly
versus time for the catalysts with larger size. Also it can be
concluded that this not the case for the 0.5 and 1 mm sizes.
The methane selectivity reaches to steady state condition after 20 h for these catalysts. Here, the catalyst with small size
leads to good filling of the catalyst bed, resulting in that the
catalyst surface is covered rapidly with liquid product. After
20 h, there is no significant change in methane selectivity for
the 0.5 and 1 mm catalysts which can be explained by Equation 7.
Figure 13. Evolution of liquid product along the catalytic bed
3. High WGS reaction rate causes the usage of CO and
production of H2 , so H2 /CO increases. High H2 /CO leads to
an increase of light products such as methane [24].
4. It has been reported that heat and mass transfer limitations result in increases in methane formation [25]. As ex-
Journal of Energy Chemistry Vol. 22 No. 5 2013
plained before, mass transfer limitations exist for large catalyst size. Also due to a low radial heat transfer coefficient,
heat transfer limitation exists too.
With respect to previous figures and results, the behavior
of Figure 9 is known clearly and there is no need to discuss
further.
5. Conclusions
The effects of pretreatment steps (reduction and carburization) on the catalytic performance of iron based FTS catalyst have been investigated. The observations reveal that mass
transfer limitations exist during the reduction and carburization in addition to the reaction step.
Based on the results of our studies, we conclude that the
observed differences in activity-time behavior, reaction rates,
selectivity and productivity, are not only for mass transfer
limitations in the reaction step, but also some of them refer
to differences in the distributions of iron metal, carbides and
oxides (hematite and magnetite) in the catalyst resulted from
mass transfer limitations in the pretreatment step. The high
CO conversion, high FT reaction rate and short time required
to reach steady state conditions are indicative of complete reduction and carburization and a high content of active phase
for small pellets. By contrast, a low FT reaction rate, high
WGS reaction rate, and long time required to reach steady
state condition are indicative of incomplete carburization and
higher content of magnetite phase for large pellets. The catalysts with longer diffusion lengths contain mainly magnetite
and hematite phases and less carbide and active metal. In addition to higher CO conversion and FT reaction rates, the small
catalysts need much shorter activation times due to higher extents of reduction and carburization, which in turn affects the
process economy.
Due to similarities in the reaction and carburization steps,
the carburization process can be continued in parallel to the
reaction. Thus, some effects of mass transfer limitations from
the pretreatment can be compensated in the reaction step after
a long time, so that the remaining iron oxides and metal sites
can convert to carbides phases in the presence of CO.
Combining the information from this study with earlier
works on catalyst design, it may provide a new perspective
of better design of iron based catalyst to improve the process
economy.
803
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