Comparative Simulations of Cobalt-and Iron-Based Fischer-Tropsch

Apr 24, 2008 - Department of Chemical Engineering, Laval University, Québec G1V 0A6, Canada, TOTAL-CSTJF, Avenue Larribau, F-64000 Pau, France, ...
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Ind. Eng. Chem. Res. 2008, 47, 3861–3869

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Comparative Simulations of Cobalt- and Iron-Based Fischer-Tropsch Synthesis Slurry Bubble Column Reactors I. Iliuta,† F. Larachi,*,† J. Anfray,‡ N. Dromard,§ and D. Schweich| Department of Chemical Engineering, LaVal UniVersity, Que´bec G1V 0A6, Canada, TOTAL-CSTJF, AVenue Larribau, F-64000 Pau, France, TOTAL-Research Center of GonfreVille, Refining & Process, F-76700 Harfleur, France, and LGPC-CNRS, ESCPE, 43 Bd. du 11 NoVembre, BP 2077, 69616 Villeurbanne, France

The influence of the catalyst type (Fe and Co) on CO and H2 conversions, CO2 selectivity, and the composition in Fischer-Tropsch synthesis slurry bubble column reactors was simulated for representative commercialscale units (7 m i.d. and 30 m height). A nonisothermal, core-annulus multicompartment multicomponent two-bubble class model was used to account for a relatively detailed hydrodynamics. It was coupled to comprehensive Fischer-Tropsch synthesis and water-gas-shift reactions, in addition to descriptions of thermodynamics and thermal effects, variable gas flow rate due to chemical/physical contraction, and gas and slurry backmixing and (re)circulation. Two mechanistic kinetic models with consideration of olefin readsorption were employed to describe the paraffin and olefin formation with cobalt- and iron-based catalysts, in addition to relatively large activities for CO2 and oxygenate formation, mainly alcohols, for the latter catalyst. The influence of the temperature and superficial gas velocity on CO and H2 conversions was more evident for a cobalt-based catalyst. For both catalysts, the space-dependent superficial gas velocity directly affected the gas-phase mean residence time, influencing in the return reactor temperature and conversions. Reliable estimation of the gas velocity due to chemical contraction was critical for conversions exceeding 50%. For both catalysts, the nonisothermal simulations reveal that, because heat removal is well managed from the heat-exchange area, the reactor operation can be considered as nearly isothermal. 1. Introduction The current interest in Fischer-Tropsch synthesis (FTS) has grown up as a consequence of environmental demands, technological developments, and changes in fossil energy reserves. This catalytic synthesis leads to a complex multicomponent mixture of linear and branched hydrocarbons and oxygenates, whose abundance depends on the catalysts employed, as well as on operating conditions. Although all group VIII metals display some activity in the C-C coupling reaction during CO hydrogenation, the most active metals for FTS are ruthenium, nickel, iron, and cobalt, with iron- or cobalt-based catalysts being the most commonly used at an industrial scale. Cobalt-based catalysts exhibit high yields, long lifetimes, and low water-gasshift (WGS) activity while producing linear alkanes and especially heavy hydrocarbons.1 By contrast, iron-based catalysts have high WGS activity and high selectivity toward both olefins and oxygenated products and appear to be stable when syngas with a high H2/CO ratio is reacted.2 Traditionally, these processes are realized using fixed- and fluidized-bed reactors as well as slurry bubble column reactors (SBCRs). There are positive features as well as drawbacks to conducting FTS by the traditional gas-phase route or even by the more advanced liquid-phase methods. For example, a gasphase fixed-bed FTS reactor produces high product yields because of the superior catalyst inventory per reactor volume. However, these high rates in relation with the potential for poor heat removal typically lead to localized overheating of the catalyst as well as deposition of heavy waxes within catalyst * To whom correspondence should be addressed. Tel.: +1-418-6563566. Fax: 1-418-656-5993. E-mail: [email protected]. † Laval University. ‡ TOTAL-CSTJF. § TOTAL-Research Center of Gonfreville, Refining & Process. | LGPC-CNRS, ESCPE.

pores. Both factors negatively affect the catalyst activity.3 Better heat control throughout the reactor can be gained by conducting FTS in the liquid phase, because of the better heat removal capacities of the liquid. Liquid-phase FTS is typically conducted in SBCRs. SBCRs offer nearly isothermal operation and large catalyst loadings, as well as high productivity, besides an operational flexibility and low power consumption. Deactivation rates of the catalyst are lower because the liquid medium facilitates dissolution of the wax products, both internal and external to the catalyst pores. However, the liquid itself provides a resistance to the diffusional transport of gas-phase reactants to the active sites, resulting in a possible decrease of the reaction rate in comparison to gas-phase FTS. Also, the separation of the attrited catalyst from the waxy product remains a demanding task for liquid-phase FTS in comparison to the fixed-bed gasphase reactor, where wax products typically trickle down the catalyst bed. The interactions between phases in a SBCR still defy the elaboration of a reliable modeling framework for optimal reactor design capable of describing the fluid dynamics and transport phenomena coupled with the complex chemistry, thermodynamics, and thermal effects.4 Procurement of chemical reaction engineering models, embedding the observed physics, for improved SBCR design has a two-fold goal.5 First, the current representation of the complex flow pattern in SBCR by the axial dispersion model attempts to lump the description of too many physical phenomena into a single dispersion coefficient, which cannot be done in a precise manner.5 Second, the computational fluid dynamics (CFD) codes have not yet sufficiently matured to be relied on as effective tools for the design of industrial-scale slurry bubble columns.6,7 Furthermore, the inclusion of the complex FTS chemistry to simulate large-scale units using CFD is still a remote goal. Meanwhile, there is still room for ameliorating

10.1021/ie701764y CCC: $40.75  2008 American Chemical Society Published on Web 04/24/2008

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and enriching the classical chemical reaction engineering models to provide better descriptions of mixing, transport, and reaction performances. The comprehensive model developed by our group4 follows this direction and takes into account the following phenomena in SBCR: (1) catalyst sedimentation/advection/dispersion/lateral exchange, (2) coupling to gas buoyancy-driven slurry recirculation with the slurry and gas rising in the center and flowing downward alongside the wall due to gas holdup radial profile,5 (3) detailed chemical kinetics with the complete products’ distribution typical of FTS (i.e., high temperature, pressure, and industrial catalyst), (4) thermodynamic equilibria, and (5) heat transfer. Because the interminglement between the FTS chemistry and the SBCR hydrodynamics is important, it appears to the authors’ best knowledge that a valid question, not addressed in the open literature, concerns the behavior of commercialscale SBCR depending on the catalyst type being loaded in it. The purpose of the present study is to compare certain aspects of FTS as carried out in the same large-scale SBCR with an iron- or cobalt-based catalyst. The multicomponent, core-annulus multicompartment pseudo-2D two-bubble class model for FTS in a SBCR developed by Iliuta et al.4 was used to perform comparative simulations between the two catalysts. Two mechanistic kinetic models8–10 considering olefin readsorption were employed to describe the formation of paraffins and olefins in the presence of cobalt- and ironbased catalysts as well as the formation of oxygenates for the latter catalyst. In the following, the iron versus cobalt catalyst behaviors will be simulated and discussed in the case of a commercial-size SBCR (7 m wide and 30 m high): (a) influence of the temperature on CO and hydrogen conversions with and without inclusion of the WGS reaction and on CO2 selectivity, (b) influence of the catalyst loading and H2/CO feed ratio on CO and H2 conversions, (c) olefin and paraffin concentrations and the O/P ratio at SBCR exit conditions, (d) liquid and gas alcohol and acid concentrations, and (e) axial temperature profiles in nonisothermal simulations. Discussions relevant to the choice of reactor parameters and the influence of hydrodynamics on the reactor performances will not be discussed because these were the subject of a previous work.4 For the sake of brevity, details of the complete model are also skipped here and may be consulted from ref 4. 2. Model Structure The multicomponent, nonisothermal, core-annulus multicompartment pseudo-2D two-bubble model developed by Iliuta et al.4 was used to simulate the FTS in a SBCR in the presence of iron- and cobalt-based catalysts. The model consisted of a detailed hydrodynamic platform whereupon were tied the FTS and WGS catalytic reactions, the descriptions of thermodynamics (via Peng-Robinson/Marano-Holder thermodynamic model) and thermal effects, the variable gas flow rate due to chemical/ physical contraction, and the gas and slurry (re)circulation and per-phase per-compartment backmixing. The hydrodynamic platform consists of a pseudo-2D axisymmetric hydrodynamic core-annulus model that describes the catalyst sedimentation/ advection/dispersion/(core-annulus) lateral exchange coupled to gas and slurry (re)circulation. The pseudo-2D feature comes from the fact that the model was split into two blocks to handle the axial dependence of the reactor hydrodynamics due to gasphase contraction, besides a radial dependence imposed by the core-annulus flow structure. In the first block, a two-fluid

Table 1. Conditions of Isothermal Simulations D H Cc0 Fp εs

7m 30 m 235-335 kg/m3 950 kg/m3 0.25-0.35

Ug Fg P T H2/CO

0.3 m/s 6.3-6.5 kg/m3 25 bar 493-560 K 1-2.15

turbulent flow model was used for computing the radial profiles of slurry and gas velocities, gas holdup, and bubble size. In the second, a coupled axial multicompartment model accounting for sedimentation, dispersion, advection, and lateral exchanges (liquid-liquid, gas-gas, and solid-solid) was solved for obtaining the longitudinal distributions of the hydrodynamic variables. 2.1. Kinetic Models. The FTS is essentially a polymerization reaction between the primary reactants CO and H2. Ideally, a hydrocarbon chain attached at one end to the catalyst grows by the addition of a single carbon segment at a time. The hydrocarbon products from the FTS are primarily n-paraffins and R-olefins. Cobalt-based catalysts favor the production of n-paraffins, whereas iron-based catalysts, in addition to nparaffins, produce unsaturated compounds, predominantly R-olefins. The iron-based catalysts have relatively high activity for oxygenate formation and a high WGS activity. Because R-olefins produced in FTS reactions can be readsorbed on the surface of the catalyst and further take part in FTS polymerization, detailed mechanistic kinetic models with consideration of olefin readsorption8–10 were used in this work to describe the paraffin and olefin formation in the presence respectively of Co and Fe-Cu-K catalysts. The total amount of oxygenates in FTS products is relatively large over iron-based catalysts compared with cobalt-based catalysts. The kinetic model of oxygenate formation over an industrial Fe-Mn catalyst was described with the model developed by Teng et al.11 Note that our choice of combining kinetic models stemming from two iron catalysts bearing different doping elements was made so as to achieve meaningful simulations in terms of the evolution of the olefin/paraffin ratio with an increase in the carbon number. To the authors’ best knowledge, there is only one kinetic model in the literature describing the formation of paraffins, olefins, and oxygenates.12 However, the carbondependent olefin/paraffin ratio, especially for olefin selectivity with higher molecular weight, still cannot be completely described by the Teng et al.12 kinetic model. Because the production of olefins is very important in the case of an ironbased catalyst, we decided to use the Chang et al.10 kinetic model to describe the paraffin and olefin formation and the kinetic model of Teng et al.11 to describe the oxygenate formation. 2.1.1. Kinetic Model for a Cobalt-Based Catalyst. The mechanistic kinetic model of Anfray et al.8,9 was used to describe paraffin and olefin formation. Two mechanistic pathways were proposed: a dissociative CO adsorption mechanism13 and a nondissociative CO adsorption mechanism.14 In this work, the consumption and formation rates are based on a nondissociative CO adsorption mechanism.8,9 The main expressions for the kinetic rates are summarized in Appendix A (see the Supporting Information), and the numerical values of the different parameters are given in the work of Anfray et al.8,9 An activation energy of 100 kJ/mol was assumed to take into account the variation with temperature of the Fischer-Tropsch intrinsic kinetics. The Fischer-Tropsch consumption and formation rates were estimated with respect to their values at 493 K: Rj,T ) Rj,T)493K exp[100000 ⁄ R(1 ⁄ 493 - 1 ⁄ T)]

(1)

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Figure 1. Influence of the temperature on the CO and H2 conversions: (a) cobalt-based catalyst; (b) iron-based catalyst.

The FTS and WGS (CO + H2O f CO2 + H2) reactions occur on different active sites with no interactions between the respective surface reactions.15 The presence of the WGS reaction with a cobalt catalyst was accounted for by using a kinetic model developed by Keyser et al.15 (Appendix B, see the Supporting Information). RWGS in eq B.1 was added to eq A.1 to calculate the net CO consumption. Similarly, it was rested from eq A.2 to account for the hydrogen produced. In addition, the CO2 selectivity was calculated with the following expression: CO2 selectivity ) (VSBIεSBICCO2,SBI|z)H + VSBIIεSBIICCO2,SBII|z)H +

⁄[

in VLBεLBCCO2,LB|z)H) UgCCO - (VSBIεSBICCO,SBI|z)H +

VSBIIεSBIICCO,SBII|z)H + VLBεLBCCO,LB|z)H)

]

(2)

2.1.2. Kinetic Model for an Iron-Based Catalyst. Paraffins and r-Olefins Kinetic Model.10 Chang et al.10 developed a comprehensive model based on the alkylidene mechanism for FTS and the formate intermediate mechanism for the WGS reaction. Three kinds of reactions are considered in the model: (1) formation of primary hydrocarbons, (2) secondary reactions of primary 1-olefins, and (3) WGS reactions. FTS and WGS reactions occur on different active sites, named as site θ (primary

Figure 2. Influence of the temperature on the CO2 selectivity: (a) cobaltbased catalyst; (b) iron-based catalyst.

paraffin and olefin formation and re-enter propagation), site σ (secondary hydrogenation), and site ψ (WGS). No interaction occurs between the respective surface reactions. The primary olefins are assumed to be able to readsorb on the catalyst surface and then re-enter the reaction chains on site θ or directly hydrogenate to the corresponding paraffins on the site σ. The kinetic model is summarized in Appendix C (see the Supporting Information), and the numerical values of the different parameters can be taken from the work of Chang et al.10 FTS CO and H2 consumption rates were evaluated using equations similar to eqs A.1 and A.2. The presence of the WGS reaction under a Fe-Cu-K catalyst was accounted for by using the kinetic model developed by Chang et al.,10 which assumes that the slowest step in the WGS reaction is the desorption of gaseous carbon dioxide via formate intermediate species (eq C.11). Oxygenates Kinetic Model.11 The model assumes that oxygenate and hydrocarbon formation reactions occur on different active sites with no interaction between the respective surface reactions. A CO insertion mechanism was used to describe the alcohol and acid formation: alcohols are formed by successive hydrogenation of acyl intermediates, and acids

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Figure 3. CO and H2 conversions vs volumetric catalyst fraction: (a) cobaltbased catalyst; (b) iron-based catalyst.

are formed by the reaction between acyl and hydroxyl species. The kinetic model is summarized in Appendix D (see the Supporting Information), and the numerical values of the different parameters can be taken from the work of Teng et al.11 For the purpose of applying the kinetic model for an ironbased catalyst, the partial pressure Pj of component j is calculated with its corresponding equilibrium concentration in the liquid phase. The compositions at the liquid and vapor sides of the interfaces are evaluated using vapor-liquid equilibrium (VLE) flash calculations at the interfacial temperatures corresponding to the contiguous vapor and liquid compartments (SBI-LI, LB-LI, and SBII-LII).4 The VLE were evaluated using a Peng-Robinson/Marano-Holder thermodynamic model.16 For alcohols and acids, the VLE were evaluated assuming a constant enthalpy of the solution.17 3. Results and Discussion 3.1. Isothermal Simulations. This section presents the simulation results for the isothermal case. The operating conditions and geometry of the reactor are listed in Table 1. The simulations were conducted bearing in mind two features

Figure 4. CO and H2 conversions vs H2/CO molar ratio: (a) cobalt-based catalyst; (b) iron-based catalyst.

relevant to the industrial context of FTS: (i) the gas superficial velocities correspond always to the churn turbulent flow regime; (ii) the vessel sizes (height and diameter) are compatible with commercial-scale SBCR units. Figures 1 and 2 respectively show the influence of the temperature on the CO and H2 conversions and on the CO2 selectivity under constant and variable gas superficial velocity conditions for cobalt (a) and iron (b) catalysts. The constant Ug case corresponds to neglect of the chemical contraction caused by reactant (and product) chemical consumption (and production) and prevailing partition due to thermodynamic VLE. The variable Ug case accounts for the gas residence time increasing as a result of the same phenomena. Higher temperatures in the reactor mean higher CO and H2 consumption rates and consequently higher CO and H2 conversions. The influence of the temperature on CO and H2 conversions is more evident in the case of a cobalt-based catalyst. In the presence of an iron-based catalyst, the temperature effect is more pronounced in the case of CO conversion because of the WGS reaction. The presence of the WGS reaction amplifies significantly the CO consumption rate, and the result is a higher CO conversion in the reactor. The selectivity of CO2 is not strongly dependent on the temperature. An iron-based catalyst

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Figure 5. Influence of the temperature on CO and H2 conversions with/ without the WGS reaction: (a) cobalt-based catalyst; (b) iron-based catalyst.

produces a high quantity of CO2. The selectivity of CO2 varies between 44 and 48%, which is consistent with the experimental results obtained by Chang et al.10 In the presence of a cobaltbased catalyst, the temperature influence on H2 conversion is more important than that in the case of an iron-based catalyst. H2 conversion is higher than CO conversion because the molar H2/CO ratio is low and because cobalt-based catalysts have relatively low activity for the WGS reaction. A cobalt-based catalyst produces a low quantity of CO2. CO2 accumulation diminishes significantly when the H2/CO molar ratio approaches the value 2. In addition, the order of magnitude regarding CO2 selectivity is consistent with the experimental results reported by Keyser et al.15 The superficial gas velocity directly affects the reactor performance via the mean residence time. CO and H2 conversions are higher under variable gas velocity because the residence time of the reactants increases. Therefore, reliable calculation of a change in the superficial gas velocity is important for proper predictions of the reactor performance. The change in the gas flow rate was calculated from the overall mass balance equation that includes both small and large bubbles.4 Figure 1 shows that CO and H2 conversions are very sensitive to gas-phase contraction effects. When gas conversion is high,

Figure 6. Paraffin (a) and olefin (b) distributions in the liquid phase at the SBCR exit for a cobalt catalyst.

the contraction in the gas phase is significant so that CO and H2 conversions become appreciably affected by changes in Ug. When gas conversion is low, gas contraction is not considerable and CO and H2 conversions are barely influenced by changes in the local superficial gas velocity. Parts a and b of Figure 3 show the effect of the catalyst concentration on CO and H2 conversions under constant and variable gas superficial velocity conditions for cobalt and iron catalysts, respectively. For both catalysts, the increase of the catalyst concentration amplifies the consumption rates of CO and H2, and the result is an increased gas conversion. With a cobalt-based catalyst, increasing the catalyst inventory by 40% (from εs ) 0.25 to 0.35) translates in ca. 45% improvement in CO conversion and ca. 50% in H2 conversion. The increase is moderate for an iron-based catalyst: ca. 30% improvement in CO conversion (especially due to WGS) and only 10% in H2 conversion. Note that the influence of catalyst holdup is not only a positive factor in terms of reaction but also a negative factor on the total gas holdup, which decreases as the solids holdup is increased.18 Figure 4 shows the CO and H2 conversions as a function of the H2/CO molar ratio for cobalt (a) and iron (b) catalysts. The conversions of CO and H2 depend on the H2/CO feed mole ratio

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Figure 7. Paraffins (a) and olefin (b) distributions in the liquid phase at the SBCR exit for an iron catalyst. Figure 9. Alcohol concentration in the liquid (a) and gas (b) phases in the exit reactor section at different reactor temperatures for an iron-based catalyst.

Figure 8. Olefin-to-paraffin ratios vs carbon number for an iron-based catalyst.

for both catalysts, but in a different way. With a cobalt catalyst, low H2/CO feed ratios slow down CO and H2 consumption and corresponding overall conversions. Coherent with the findings

of Anfray,8 the simulations revealed that high H2/CO ratios favor the production of a large quantity of light n-paraffins because of favored H2 adsorption and dissociation and thereafter hydrogenation of hydrocarbon chains. The CO2 formation rate decreases significantly with an increase of the H2/CO feed molar ratio for a cobalt catalyst, and the result is a decrease in CO2 selectivity (between 0.2 and 0.1). In the case of an iron catalyst, the formation rates of the higher hydrocarbons decrease with increasing H2/CO feed molar ratio because the chain growth factor decreases. The result is a decrease of the CO and H2 consumption rates in the FTS process. However, the CO overall conversion increases in all likelihood because of consumption of CO in the WGS reaction. Unlike the cobalt catalyst, hydrogen consumption by the FTS reactions is counterbalanced through hydrogen production due to WGS. However, this latter reaction was unable to reverse the decreasing H2 conversion trend as the H2/CO ratio is increased. Apart from the fact that an iron catalyst produces ca. 2 times more CO2 (CO2 selectivity ) 45-47%) than a cobalt catalyst, the influence of the H2/CO molar ratio on CO2 selectivity is less marked with an iron catalyst.

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Figure 10. Acid concentration in the liquid (a) and gas (b) phases in the exit reactor section at different reactor temperatures for an iron-based catalyst. Table 2. Conditions of Nonisothermal Simulations D H Cc0 Fp CCo

7m 30 m 285 kg/m3 950 kg/m3 21% w/w

Ug Fg P Tg,in εs

0.2-0.3 m/s 6.5 kg m3 25 bar 240-252 °C 0.3

Parts a and b of Figure 5 show the influence of the temperature on CO and H2 conversions, respectively, for cobalt and iron catalysts, assuming variable Ug. The simulated trends are shown when the WGS reaction is disabled and enabled. A large conversion of CO is found by enabling the WGS reaction for the iron catalyst (Figure 5b). The influence of WGS for a cobalt catalyst is moderate at lower temperatures but may become significant at higher temperatures. Hydrogen consumption by the FTS reactions is counterbalanced by hydrogen production due to the WGS reaction. However, H2 conversion appears to be less influenced by the WGS reaction. The slight increase in hydrogen conversion when the WGS is enabled could be explained by the fact that the FTS reaction rates are boosted by hydrogen local overconcentration, which would contribute for further CO conversion.

Figure 11. Axial variations of the gas and liquid temperatures within the SBCR core (a) [small and large bubble classes SBI and LB and liquid LI] and annulus (b) [small bubble class SBII and liquid LII] compartments for a cobalt-based catalyst. Tg,in ) Tw ) 513 K.

The cobalt catalyst produces mainly paraffins. Figure 6a shows the paraffin distribution in the liquid phase at the reactor exit (T ) 513 K). The simulated average carbon number of Fischer-Tropsch wax is 26.2 (H2/CO ) 2.15; Ug,in ) 0.3 m/s) and is in good agreement with the experimental results reported by Anfray.8 For the simulated conditions with the cobalt catalyst, the liquid concentrations of olefins are typically an order of magnitude less than their paraffin counterparts up to C15 (see Figure 6b). The decline in the concentrations in olefins higher than C15 is even more pronounced and so continues below the 0.001 mol/L level shown in Figure 6b. An iron catalyst produces paraffins and olefins in comparable proportions. Parts a and b of Figure 7a,b show, respectively, the distribution of paraffins and olefins in the liquid phase at the column exit (T ) 530 K). According to the simulation results, the average carbon number of Fischer-Tropsch wax is 27.7 (H2/CO ) 1.4) in good agreement with the experimental results of Wang et al.19 Figure 8 shows the dependence of the olefin-to-paraffin ratio in the liquid phase as a function of the carbon number. The kinetics model predicts qualitatively the correct decreasing trend at high carbon number of the olefinto-paraffin ratio.11,20

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Figures 11 and 12 show the axial temperature profiles in the different gas and liquid phase compartments for respectively the cobalt and iron catalysts. SBI and LB represent core small and large bubble classes and LI core liquid. Similarly, SBII and LII designate respectively the small bubble class and the liquid in the annulus region. The WGS reaction is enabled, and the inlet gas temperature is taken to be equal to the wall temperature imposed by the heat exchanger. For these simulation conditions, the temperature in the liquid compartments LI and LII is nearly uniform despite a tiny dip toward the reactor inlet. The nonisothermal simulations reveal also that inasmuch as heat removal is well managed from the heat exchange area, the reactor operation can be considered as quasi-isothermal. 4. Conclusion The influence of the catalyst type on the conversion levels of CO and H2, on the CO2 selectivity, and compositions in the FTS in SBCRs of large scale was analyzed using a detailed multicomponent transport-reaction model. The model takes into account the thermodynamic and thermal effects, the change in the gas flow rate along the reactor as a result of chemical/ physical contraction, the recirculation of gas and slurry phases, and backmixing. Two mechanistic models accounting for olefin readsorption were employed to describe the paraffin and olefin formation with cobalt and iron catalysts. The cobalt catalyst produces predominantly paraffins, and the iron-based catalyst produces paraffins and olefins. The ironbased catalyst has a relatively high activity for oxygenate formation and, unlike the cobalt-based catalyst, produces high CO2 quantity. The influence of the temperature and superficial gas velocity on CO and H2 conversion is more evident in the case of the cobalt-based catalyst. For both catalysts, the spacedependent superficial gas velocity directly affects the mean gas residence time, and so the temperature and CO conversions were influenced by the actual local gas velocity. Reliable estimation of the gas velocity due to chemical contraction is critical when conversion exceeds 50%. For both catalysts, nearly isothermal operation is achieved because heat removal is well managed. Figure 12. Axial variations of the gas and liquid temperatures within the SBCR core (a) [small and large bubble classes SBI and LB and liquid LI] and annulus (b) [small bubble class SBII and liquid LII] compartments for an iron-based catalyst. Tg,in ) Tw ) 525 K.

Unlike the cobalt catalyst, the iron-based catalyst exhibits a relatively high activity for oxygenate formation, for the most part alcohols. The distributions of alcohols and acids in the liquid (Figures 9a and 10a) and gas (Figures 9b and 10b) phases at the reactor exit are illustrated for three different inlet temperatures. The concentrations of alcohols and acids in the liquid phase decrease as expected with an increase in the carbon number. 3.2. Nonisothermal Simulations. This section briefly discusses some simulation results for the nonisothermal case. The operating conditions and geometry of the SBCR are listed in Table 2. The simulations are shown for the case where an intense lateral mass transfer takes place between the core small bubbles SBI and large bubbles LB to allow homogenization of the species compositions within the two bubble classes. A value for the mass-transfer coefficient kSBI-LB equal to 350 s-1 was chosen to force the same compositions in both bubble classes. The parameter kSBI-LB allows modulation of the species compositions in the small and large bubble classes as a result of complete or partial equilibrium between bubble coalescence and breakup rates.

Acknowledgment Drs. S. Savin, J. Bousquet, and F. Luck are greatly acknowledged for fruitful discussions on the Fischer-Tropsch process. Supporting Information Available: Fischer-Tropsch intrinsic kinetics for a cobalt-based catalyst (Appendix A), WGS intrinsic kinetics for a cobalt-based catalyst (Appendix B), Fischer-Tropsch intrinsic kinetics for an iron-based catalyst (Appendix C), and an oxygenates kinetic model for an ironbased catalyst (Appendix D). This material is available free of charge via the Internet at http://pubs.acs.org. Nomenclature aw ) heat transfer area, m2/mreactor3 C ) concentration of species j, kmol/mi3 D ) reactor diameter, m H ) reactor length, m kSBI-LB ) LB-SBI interaction cross-flow coefficient, s-1 Pj ) partial pressure of species j, Pa T ) temperature, K Vi ) interstitial velocity of compartment i, m/s Ug ) superficial gas velocity, m/s z ) axial coordinate of the reactor, m Greek Letters

Ind. Eng. Chem. Res., Vol. 47, No. 11, 2008 3869 εi ) volumetric fraction of compartment i, m /mreactor F ) density, kg/m3 Subscripts g ) gas g,in ) gas inlet conditions LI ) liquid in the core region LII ) liquid in the annulus region LB ) large bubbles p ) particle SBI ) small bubbles in the core region SBII ) small bubbles in the annulus region w ) wall WGS ) water-gas shift 3

3

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ReceiVed for reView December 25, 2007 ReVised manuscript receiVed March 3, 2008 Accepted March 13, 2008 IE701764Y