Zeolite Membrane Reactor for High-Temperature Water-Gas Shift

*Telephone: +1-513-556-3992. ... (χCO) were studied experimentally and by calculations using a simple one-dimensional plug-flow reactor (PFR) model...
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Zeolite Membrane Reactor for High-Temperature Water-Gas Shift Reaction: Effects of Membrane Properties and Operating Conditions Seok-Jhin Kim, Shaowei Yang, Gunugunuri K. Reddy, Peter Smirniotis, and Junhang Dong* Department of Chemical Engineering, University of Cincinnati, Cincinnati, Ohio 45220, United States ABSTRACT: Modified MFI-type zeolite membranes were investigated as high-temperature water-gas shift (WGS) membrane reactors (MRs) in combination with a nanocrystalline Fe/Ce WGS catalyst. The effects of the MR operating conditions and the membrane separation performance on the CO conversion (χCO) were studied experimentally and by calculations using a simple one-dimensional plug-flow reactor (PFR) model. The experimental results showed that, at high temperatures (e.g., >500 °C), the zeolite MR with moderate H2 selectivity (e.g., αH2/CO2 ∼ 31, and αH2/CO ∼ 25) and permeance (Pm,H2 ∼ 0.9 × 10−7 mol s−1 m−2 Pa−1) was capable of overcoming the limit of equilibrium CO conversion and χCO of the MR could be further enhanced by increasing the reaction pressure while keeping the permeate pressure unchanged. At high temperatures and high reaction pressures, CO is rapidly consumed by a fast reaction that minimizes the membrane permeation of unreacted CO; meanwhile, the efficiency of H2 removal is improved as a result of the increased H2 partial pressure difference across the membrane. The model calculations have indicated that the current membrane has the potential to achieve high CO conversion of χCO > 99% under practically meaningful operating conditions.

1. INTRODUCTION The water-gas shift reaction (WGS) is a critical step for the pre-combustion CO2 capture and H2 production from fossil and biomass stocks via the gasification and steam-reforming processes. Hydrogen-permselective microporous ceramic membranes are potentially useful for constructing WGS membrane reactors (MRs). In comparison to the Pd-alloy dense membranes, the microporous ceramic membranes have the advantages of low cost and being resistant to H2S poisoning and hydrothermal embrittlement. However, porous membranes generally have limited H2 selectivity over other gases involved in the WGS reaction system, such as CO2, H2O, and CO, and therefore, H2 produced from such porous MR cannot be of very high purity. The amorphous molecular sieve silica membranes prepared by sol−gel or chemical vapor deposition (CVD) techniques can obtain H2/CO2 selectivity of greater than 100, which is among the best reported for microporous membranes. However, the use of amorphous silica membranes for WGS MR has been impeded by the problem of water condensation reactions in hydrothermal environments that damage the membrane microporous structures and reduces the hydrogen permeability and selectivity over time.1−4 Because of the crystalline nature, siliceous zeolite membranes, particularly the MFI-type zeolite membranes, have excellent hydrothermal stability and chemical resistance to withstand the conditions of a high-temperature syngas WGS reaction.5,6 However, at high temperatures, transport of the small molecules involved in the WGS reaction in the 0.56 nm diameter pores of the MFI-type zeolites is governed by the gaseous diffusion mechanism that leads to poor H2 selectivity over other components.7−9 The pore sizes of the MFI-type zeolite membranes can be effectively reduced to ∼0.36 nm by catalytic cracking of a methyl diethoxysilane (MDES) precursor to deposit mono silica inside the zeolitic channels.10−12 Such modified MFI zeolite membranes have achieved a dramatically increased H2/CO2 © 2013 American Chemical Society

separation factor (αH2/CO2) through the realization of certain size-exclusion effects for the relatively larger molecules, such as CO2, CH4, and CO, at the reduced pore opening as well as the molecular size-sensitive activated diffusion mechanism inside the narrowed zeolitic channels.13 The molecular silica deposited by the catalytic cracking deposition (CCD) method is chemically bonded to the zeolite internal surface that makes the modified membrane chemically stable in hydrothermal conditions.13,14 In our recent work, WGS MRs were constructed using the CCD-modified MFI zeolite membranes packed with a ceriumdoped ferrite (Fe/Ce) catalyst.5,15 The zeolite membrane reactors were successfully tested for WGS reactions in a high-temperature range of 400−550 °C. The zeolite membrane showed good stability after operating for more than 2000 h in WGS feed streams with and without the presence of H2S. The zeolite membranes with αH2/CO2 between 10 and 45 and H2 permeance (Pm,H2) of 400 °C and atmospheric pressure.5,16 However, methanation reactions are associated with volume reduction, which are favored at high pressures. For the Fe/Ce catalyst and operation conditions used in this study, our previous work found that the methane selectivity was 2 bar and 550 °C. The deviations of carbon balance may be attributed mainly to the errors of flow rate measurements and GC analysis as well as minor carbon formation on the oxidized stainless-steel tubing and reactor surface. Carbonization on the catalyst surface was not observed in our previous studies.16,17 When the membrane-mounted cell was operated as a packed-bed TR, the entering sweeping gas was removed and the exit of the reaction side was connected to the original sweeping inlet. The gas stream from the reaction side thus passed through the permeate chamber to exit from the permeate side. A series of WGS experiments were carried out in both the MR and TR operation modes to study the effects of the reaction pressure, temperatures, steam/CO ratios, WHSV, and sweep flow rates. The main conditions of the experiments and model calculations are summarized in Table 1.

⎛ Ea, i ⎞ o Pm, i = Pm, ⎟ i exp⎜ − ⎝ RT ⎠

experimental

calculation

400−550 2−6 1 1.0−3.5

400−550 2−50 1 1.0−10.0

WHSVa (h−1) CO feed flow rate (kg m−2 of membrane h−1) sweep gas flow rate, FN2 [cm3 (STP)/min] catalyst load, m (mg of catalyst/cm2 of membrane)

7500−60000 0.83−6.56

7500−60000 0.83−6.56

0−40

0−100

79

39−790

(9)

2.2.3. Reactor Model. Several general assumptions are made for reactor modeling, including (i) isothermal steady-state operation, (ii) ideal gas behavior and pressure-independent permeance, (iii) negligible side reactions, and (iv) negligible mass-transfer resistances in the thin catalyst layer (∼60 μm thick) and the macroporous substrate. The model is first validated by comparing to experimental data and then used for simulation of the MR performance beyond the experimental conditions. Because of the very small size of the MR, two reactor models were evaluated, including the 1D PFR and the continuous stirred-tank reactor (CSTR) models. In the PFR model, both the reaction (feed) side and permeate side are considered under plug-flow conditions, and for the CSTR model, both sides of the membrane assume a perfect mixing effect. Figure 2 shows schematically the membrane reactor structure and

Table 1. WGS Membrane Reaction Conditions temperature range (°C) reaction side pressure at exit (atm) permeate pressure (atm) RH2O/CO (mol/mol)

(i = H 2 , CO2 , CO, and H 2O)

H2 CO2 H2 O a CO WHSV = (νCO feed + νfeed + νfeed + νfeed )/(mcat/ρcat), where νfeed, H2 CO H2O , νfeed2, and νfeed are volumetric rates of components in feed νfeed stream at standard temperature and pressure (STP).

2.2. Reactor Modeling. 2.2.1. Reaction Rates. The WGS reaction is moderately exothermic.

CO + H 2O ↔ CO2 + H 2 ,

Θ ΔH298.15 K = − 41.2 kJ/mol

(6)

The empirical power-law equation is employed here for the WGS reaction rates because of its simplicity and proven effectiveness in correlating the experimental data especially for WGS reactions over doped iron catalysts.18,19 The general power-law WGS reaction rate equation is given by18

⎛ pa p b ⎞⎛ ⎞ 1 pCO2 pH2 ⎟ CO H O rA = k ⎜ c 2d ⎟⎜⎜1 − ⎜p p ⎟ Ke pH O pCO ⎟⎠ ⎝ CO2 H2 ⎠⎝ 2

⎛ −E ⎞ and k = k 0 exp⎜ a ⎟ ⎝ RT ⎠

Figure 2. Schematic showing the gas flow arrangement and mass balance in the MR.

(7) where rA is the reaction rate (mol g−1 of catalyst s−1), Pi is the partial pressure of component i in the reaction zone (Pa), Ke is the WGS reaction equilibrium constant, k0 is the pre-exponential factor, Ea is the activation energy (kJ/mol), R is the gas constant (kJ mol−1 K−1), T is the temperature (K), and a, b, c, and d are the apparent reaction orders of CO, H2O, CO2, and H2, respectively. The equilibrium constant (Ke) of the reaction is given by

⎛p p ⎞ CO2 H2 ⎟ Ke ≅ ⎜⎜ ⎟ p ⎝ H2O pCO ⎠e

⎛E ⎞ and Ke = Ko exp⎜ r ⎟ ⎝ RT ⎠

concurrent cross-flow arrangements used in both experiments and model calculations. The CSTR model is described by the following equation of the mass balance around the catalyst bed: Fi ,out − Fi ,in = ni − Q i

(10)

where Fi is the molar flow rate on the feed side (mol/s). For gas i, the flow rate through the membrane (Qi, mol/s) and the rate of generation by the reaction (ni, mol/s) are given by Q i = Pm, iAΔPi

(8)

(11)

ni = νirAρB V

where the subscript “e” denotes the equilibrium state. The energy constant Er of the exothermic WGS reaction is a positive value; i.e., Er = 38.06 kJ/mol, with Ko = 1.3168 × 10−2.20

(12) 3

where V (=δA) is the volume of the catalyst bed (m ) [δ is the thickness of the catalyst bed (m) and A is the membrane area (m2)], ρB is the 4473

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density of the packed catalyst (g/m3), and vi is the stoichiometric coefficient of species i. For the 1D PFR model, the mass balance equation for a differential section of the reactor (shown in Figure 2) is obtained as follows:

dFi = Fi|A + dA − Fi|A = dni − dQ i

(13)

dni = νirAρB δ dA

(14)

dQ i = Pm, i(ΔPi)A dA

(15)

Membrane MA had higher H2 selectivity and higher H2 permeance than membrane MB. The performance difference between membranes MA and MB is likely caused by the fresh membrane quality variation, which is common to zeolite membranes prepared by in situ crystallization, even under nominally same conditions. Unmodified MFI zeolite membranes of reasonable quality are CO2-selective against H2 (i.e., αH2/CO2 < 1.0) at room temperature because preferentially adsorbed CO2 hinders the entry of non-adsorbing H2.12 A CCD-modified MFI zeolite membrane often exhibits higher αH2/CO2 when the fresh membrane has lower αH2/CO2 at room temperature, which indicates less defects that may remain after modification.21 For the homemade small-size supports of this work, some minor differences in film formation conditions, such as localized uneven surface and pinholes, could cause significant quality variation in the in situ synthesized zeolite membranes. Although the two membranes were made under similar conditions, the unmodified membrane MA had a room temperature αH2/CO2 of 0.38, which was much smaller that of the unmodified membrane MB (αH2/CO2 = 0.78), and hence, the former had better separation performance after modification. 3.1.2. Reaction Rates. The WGS reaction rate equation was experimentally determined by varying the individual gas partial pressure under TR operation mode at 400−550 °C. The values of a, b, c, d, and k in eq 7 were first determined at each temperature by a procedure shown in Figure 3a for 500 °C as an example. The constants ko and Ea in eq 7 were then obtained from the Arrhenius plot of k in Figure 3b. The values of a, b, c, and d were found to be reasonably independent of the temperature, and therefore, the average values of different temperatures are used as constants for eq 7. The final power-law reaction rate equation for the WGS reaction over the current Fe/Ce catalyst is given below by eq 16. It should be noted that eq 16 is a lumped reaction rate equation only applicable to the specific reactor structure and flow conditions of this study.

For CSTR model, the equations are solved by an iterative process. For the PFR model, the differential equations are solved numerically with the reactor divided into 150 sections of equal membrane area (i.e., equal amounts of catalyst). By setting Qi = 0, eqs 10 and 13 describe CSTR and PFR models, respectively, under TR operation modes for the same reactor.

3. RESULTS AND DISCUSSION 3.1. Membrane Properties and Reaction Rate Equation. 3.1.1. Membrane Properties. Table 2 lists the values of Pom,i and Table 2. Pom,i and Ea,i Values for Equation 9 and Membrane Properties at 500 °C H2

H2O

CO2

CO

MA

(×10 , mol s m Pa ) Ea,i (kJ/mol) Pm,i (×10−8, mol s−1 m−2 Pa−1) αH2/i

45.5 10.4 8.77

3.50 3.16 2.13 4.12

0.82 6.82 0.29 30.7

1.21 7.95 0.35 24.8

MB15

Pom,i (×10−8, mol s−1 m−2 Pa−1) Ea,i, (kJ/mol) Pm,i (×10−8, mol s−1 m−2 Pa−1) αH2/i

15.7 5.58 6.52

4.44 2.57 2.99 2.18

1.35 3.39 0.79 8.25

1.12 3.78 0.63 10.4

Pom,i

−8

−1

−2

−1

Ea,i in eq 9 for gases involved in the WGS reaction together with the gas permeance and H2 selectivity data at 500 °C. These Pom,i and Ea,i values were obtained through regressions of the permeation data of H2/CO2 and H2/CO binary mixtures and pure water vapor, which were measured in the temperature range of 400−550 °C under a feed-side pressure of 1.5 atm and a permeate-side pressure of 1 atm. The Pom,i and Ea,i values for membrane MB were obtained from ref 15. The gas permeation data for both membranes MA and MB were measured in the catalyst-packed MR after performing the WGS reaction and gas permeation experiments at >400 °C for more than 1000 h.

⎛ −121 ± 1.00 ⎞ 1.05 ± 0.05 0.07 ± 0.03 ⎟P rA = 104.455 ± 0.05 exp⎜ PH2O ⎝ 10−3RT ⎠ CO ⎛ 1 PCO2PH2 ⎞⎟ −0.06 ± 0.02 −0.04 ± 0.02⎜ 1 PCO P − H ⎜ 2 2 Ke PCOPH2O ⎟⎠ ⎝

(16)

3.1.3. Model Selection. After obtaining the gas permeance equation (eq 9) and reaction rate equation (eq 16), calculations

Figure 3. WGS reaction rates in TR mode: (a) relationships between CO, CO2, H2, and H2O partial pressures and reaction rates at 500 °C and (b) Arrhenius plot of k. 4474

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were first carried out for WGS reactions in MR using membrane MB and TR operation mode in the same reactor. Figure 4

two models make little difference for such a small reactor size when fast reaction at such a high temperature can bring χCO close to χCO,e, regardless of the gas mixing effect. However, in the MR operation, while the PFR model gave good prediction of χCO as a function of the reaction pressure, the CSTR model predicted χCO significantly lower than the experimental values. The largely underestimated χCO by the CSTR model for the MR mode may be attributed mainly to the overestimated mixing effect on both sides of the membrane that could lead to a significantly underestimated driving force for H2 permeation. Thus, the simple 1D PFR model is selected for MR simulations in the present work. 3.2. Comparison between Membranes MA and MB. Figure 5 shows the results of the WGS membrane reaction in membrane MA at 400 and 550 °C in comparison to the results on membrane MB from ref 15. In comparison to the TR operation mode, χCO values in the MRs were significantly increased because of the H2 removal during the reaction. The enhancement of χCO in the MRs became more significant when the reaction pressure (feed side, pfeed) increased. The reason is that, when pH2,p remains very small under sweeping flow at 1 atm, a higher pfeed results in a greater pH2,f and, hence, a larger driving force, ΔpH2 (=pH2,f − pH2,p), is created. Thus, H2 can be removed to a greater extent that leads to more effective χCO enhancement. Meanwhile, the increasing H2 flux results in higher RH2 (panels b and d of Figure 5). However, the more complete removal of H2 reduces the value of (yH2/yCO2)feed that, when αH2/CO2 is nearly constant, leads to a smaller (yH2/yCO2)permeate [=αH2/CO2(yH2/yCO2)feed] and, thus, decreases the H2 purity in the permeate side (yH2,p), as observed in panels b and d of Figure 5. The microporous zeolite membranes, because of their

Figure 4. Comparisons between the experimental and model-calculated χCO data for TR and MR operations using membrane MB15 at 550 °C (WHSV = 7500 h−1; RH2O/CO = 3.5; pperm = 1 atm; and FN2 = 20 cm3/min).

presents the calculated χCO as a function of the reaction (feed side) pressure (pfeed) at 550 °C using the 1D PFR and CSTR models. Other reaction conditions used in the calculations include WHSV of 7500 h−1, RH2O/CO of 3.5, permeate pressure (pperm) of 1 atm, and FN2 of 20 cm3/min. The experimental data of the WGS membrane reaction in membrane MB are obtained from our previous publication.15 As seen in Figure 4, under TR operation mode, χCO values calculated by both PFR and CSTR models agree well with the experimental data. This is because the

Figure 5. χCO, RH2, and yH2,p as a function of the WGS reaction pressure in membranes MA and MB15 at (a and b) 400 °C and (c and d) 550 °C. 4475

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finite selectivity and permeance, are incapable of achieving complete CO conversion (i.e., χCO = 100%) as a result of the inevitable permeation of CO and diminishing ΔpH2 along the reactor length. At 400 °C, the reaction rate is relatively low. Therefore, a large ΔpCO (=pCO,feed − pCO,perm) exists in a big portion of the MR, leading to significant permeation of unreacted CO to limit the χCO in the MR (Figure 5a). Figure 6 shows the H2 and CO fluxes

pressure along the MR length in the reaction side. The rapidness of the increase in H2 flux and decrease of CO permeation in the beginning part of the reactor strongly depends upon the reaction temperature and pressure. High temperatures and pressures favor both the CO consumption (and H2 generation) and H2 permeation that lead to greater enhancement of χCO. At both reaction temperatures, membrane MA outperformed membrane MB at all reaction pressures because membrane MA had higher H2 selectivity and higher H2 permeance for more prompt removal of H2 from the catalyst bed with less permeation of unreacted CO than membrane MB. However, as shown in Figure 5c, when operating at high temperature (550 °C) and high pressure, both membranes MA and MB were able to surpass χCO,e and the difference of χCO between the two MRs becomes less significant because the very rapid consumption of CO by fast reaction drastically reduced CO permeation even in the beginning sections of the reactors. The RH2 and yH2,p, however, are much higher in membrane MA than in membrane MB because of the greater H2 selectivity and permeance in membrane MA. At 550 °C, 6 atm and WSHV of 7500, χCO, yH2,p, and RH2 in membrane MA reached 98.5, 92.4, and 73.2%, respectively. 3.3. Model Validation. The WGS membrane reaction was performed by membrane MA under various temperatures (400− 550 °C), reaction pressures (2−6 atm), WHSVs (7500−60000 h−1), and sweep gas flow rates (0−40 m3/min). The experimental results are compared to the 1D PFR model calculations. Figure 7 shows the comparisons between the experimental and calculated χCO as a function of pfeed for MR and TR operations. The calculated results agree well with the experimental data for both MR and TR operations, especially at temperatures above 450 °C. Figure 8 presents the calculated and experimental values of χCO as a function of RH2O/CO in MR and TR operations under pfeed of

Figure 6. H2 and CO fluxes along the MR length (pfeed = 2 atm; WHSV = 7500 h−1; RH2O/CO = 3.5; and FN2 = 20 cm3/min).

calculated by the 1D PFR model along the reactor length for membrane MA. The maxima of H2 flux observed along the membrane length resulted from the competition between the H2 generation and H2 permeation. For CO, the flux decreases monotonically because of the continuously decreasing CO partial

Figure 7. Experimental and simulated χCO in MR (membrane MA) and TR (WHSV = 7500 h−1; RH2O/CO = 3.5; FN2 = 20 cm3/min; and pperm = 1 atm) at (a) 400 °C, (b) 450 °C, (c) 500 °C, and (d) 550 °C. 4476

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Figure 8. Experimental and calculated χCO as a function of RH2O/CO in MA and TR at pfeed of (a) 2 atm and (b) 6 atm, respectively (T = 550 °C; WHSV = 7500 h−1; FN2 = 20 cm3/min; and pperm = 1 atm).

Figure 9. Effect of WHSV on χCO in MR and TR at 550 °C, RH2O/CO of 3.5, pperm of 1 atm, and FN2 of 20 cm3/min at pfeed of (a) 2 atm and (b) 6 atm.

much lower feed flow rate in the reactor. The high feed flow rate can reduce the mass-transport resistance in the packed bed, and therefore, the actual H2 permeance may be much greater than that used in the calculations. The sweep flow rate FN2 affects the H2 partial pressure in the permeate side and the driving force for H2 membrane permeation. Figure 10 shows the dependence of χCO upon FN2 obtained from experiments and calculations for a reaction temperature of 500 °C, WHSV of 7500 h−1, and RH2O/CO of 3.5. The χCO in the MR increased with an increasing FN2 because a larger FN2 results in a smaller H2 partial pressure in the permeate side (pH2,p) and a greater driving force (ΔpH2) to facilitate H2 transport through the membrane. The effect of FN2 on χCO was more pronounced for a low reaction pressure (pfeed = 2 atm) than for a high reaction pressure (pfeed = 6 atm). This can be explained by the fact that a small change in pH2,p may cause a significant difference to ΔpH2 (=pH2,f − pH2,p) when pH2,f is also small (at low pfeed), but such a difference of ΔpH2 becomes less significant for a large pH2,f (at high pfeed). The underpredicted χCO at high FN2 for a low reaction pressure is again possibly caused by the underestimation of Pm,H2 using eq 9, which was determined under low FN2. 3.4. Model Calculations. To investigate the possibility for the current membranes to achieve near-completion CO conversion under practically meaningful conditions, the 1D PFR model was used to simulate the MR performance for operations beyond the experimental conditions used in this study. Figure 11 presents the calculated χCO in membranes MA and MB as functions of the reaction temperature (T) and

2 and 6 atm, respectively. The model predicted correctly the trend of χCO increase with an increasing RH2O/CO, and the deviations between experimental and calculated data are quite small for RH2O/CO > 1.5. The model, however, underestimated the χCO in the MR at low RH2O/CO (e.g., RH2O/CO = 1−1.5). One possible reason is that the calculation used χCO = 0 for the entering point of the MR, where χCO could actually be much greater than 0 because the feed stream must diffuse through the catalyst bed to reach the membrane surface; on the other hand, CO and H2O partial pressures at the catalyst/membrane interface are lower than the average values especially in the early section of MR. Thus, for the initial part of the MR, CO and H2O fluxes are overestimated, while the H2 flux is underestimated. The overestimated reactant loss can lead to significantly underestimated χCO when RH2O/CO is close to a stoichiometric ratio of 1. The impact of overestimating reactant loss and underestimating H2 permeation on the χCO calculation becomes smaller when steam is largely excessive (RH2O/CO > 1.5). The influence of errors in gas permeation calculations is also evidenced by the fact that χCO was not underestimated by the 1D PFR model for TR operation, where gas permeation is not involved. The increase of WHSV causes χCO to decrease in both MR and TR operations, as shown in Figure 9, because a larger WHSV means a shorter time for reaction and H2 permeation. The calculated χCO values are consistent with the experimental results for WHSV < 30 000 h−1. However, the χCO predicted by the model is significantly below the experimental value at high WHSV of 60 000 h−1. The large deviation at high WHSV may be caused by underestimation of H2 flux at a high feed flow rate because the permeance used in the calculations was obtained at a 4477

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Figure 10. Effect of FN2 on χCO for the MR at a reaction temperature of 500 °C, WHSV of 7500 h−1, and pperm of 1 atm at pfeed = (a) 2 atm and (b) 6 atm.

Figure 11. Calculated χCO as a function of the reaction temperature and pressure in (a) membrane MA and (b) membrane MB.

maximum χCO appears at a certain level of Pm,H2, after which further increasing the Pm,H2 causes χCO to decrease. The decrease of χCO in a MR with very high gas permeance but limited αH2/CO is caused by excessive permeation of unreacted CO. This result suggests that, for a membrane with limited αH2/CO, a very high Pm,H2 is not necessarily always beneficial to the MR performance. In principle, the MR performs with best efficiency when the H2 generation rate is matched up by the H2 permeation rate. The reaction rate depends upon the reaction temperature, pressure, catalyst load and activity, and mass-transfer efficiency. In this work, χCO was calculated for membranes MA and MB as a function of the reaction pressure (pfeed) and catalyst load (m) at 400 and 550 °C. The results of the calculations are shown in Figure 13. In the figure, the catalyst load is given by “m/m0”, where “m” is the catalyst load used in the calculation and “m0” is the load used in the experiments (i.e., 78.7 mg of catalyst/cm2 of membrane). It should be noted that the variations of masstransport resistance for different catalyst loads are not considered in the calculations. At 400 °C, χCO increases dramatically as m/m0 increases from 0.5 to ∼2, where the slow reaction rate is the limiting factor to the χCO enhancement. The χCO starts to level off after m/m0 of ∼3 in membrane MB and m/m0 of ∼2 in membrane MA, suggesting that Pm,H2 becomes the limiting factor

pressure (pfeed). Other operating conditions used in the calculations include catalyst load (m) of 78.7 mg of catalyst/cm2 of membrane, WHSV of 7500 h−1, and RH2O/CO of 3.5, which were the same as those used in the experiments. The results show that an increasing temperature and reaction pressure both enhance the χCO in the MRs. However, χCO tends to plateau above a certain temperature and pressure. The highest χCO values of 99.2% in membrane MA and 98% in membrane MB are obtained at T > 500 °C and pfeed > 30 atm, which are practically possible conditions. The difference in maximum χCO obtained by membranes MA and MB shown in Figure 11 demonstrates the importance of membrane properties (i.e., H2 selectivity and H2 permeance) to the WGS MR performance. The effect of the membrane separation performance on the WGS membrane reaction were further studied through simulations of χCO as a function of H2/CO selectivity (αH2/CO) and H2 permeance (Pm,H2). The simulation was carried out for operation conditions, including m of 78.7 mg of catalyst/cm2 of membrane, WHSV of 7500 h−1, RH2O/CO of 3.5, reaction temperature of 550 °C, pfeed of 2 atm, and pperm of 1 atm. As shown in Figure 12, the results indicate that improving αH2/CO for a fixed Pm,H2 always helps to enhance χCO because of the reduced permeation of unreacted CO in membranes with higher αH2/CO. It is interesting that, for MR with a specific αH2/CO, a 4478

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catalyst load of m/m0 < 1.0. This indicates that the membrane separation performance is the limiting factor for χCO enhancement in the MR because of the fast reaction at such a high temperature. It is important to notice that, at 550 °C, increasing the m/m0 from 0.5 to ∼1.5 causes an increase in χCO for membrane MB but has almost no effect on membrane MA. This again demonstrates that increasing the reaction rates by a higher catalyst load or activity or a higher temperature is more important to MRs with relatively poor H2 separation performance. The above simulation results indicate that χCO in the zeolite MR with moderate H2 selectivity and permeance can be improved significantly by selecting proper operation conditions, including the catalyst load. More simulations were carried out to further investigate the feasibility of the current membrane MA for achieving χCO > 99.5%, which is the final conversion level of the multiple reactor systems used in the industry. The results have shown that χCO > 99.5% could be obtained in membrane MA but will require operation conditions beyond those used in the current industrial processes. An example of the simulation results is given in Figure 14, which presents the χCO as a function of the catalyst load at pfeed of 2, 25, and 50 atm and other conditions, including WHSV of 7500 h−1, RH2O/CO of 5.0, and pperm of 1 atm. It can be seen that, for the current

Figure 12. Effect of αH2/CO and Pm,H2 on χCO of the MR (WHSV = 7500 h−1; RH2O/CO = 3.5; pfeed = 2 atm; pperm = 1 atm; FN2 = 20 cm3/min; and T = 550 °C).

when sufficient catalyst is provided to achieve a high reaction rate because the H2 permeation rate becomes less than the rate of H2 generation. At 550 °C, χCO was found to plateau at a much lower

Figure 13. Effect of the catalyst load and reaction pressure on χCO in the MRs (WHSV = 7500 h−1; RH2O/CO = 3.5; pfeed = 2 atm; pperm = 1 atm; and FN2 = 20 cm3/min) for (a) membrane MB at 400 °C, (b) membrane MA at 400 °C, (c) membrane MB at 550 °C, and (d) membrane MA at 550 °C. 4479

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ACKNOWLEDGMENTS This research was supported by the United States Department of Energy’s National Energy Technology Laboratory (U.S. DOE/NETL) (Grant DE-FG36-GO15043), the Ohio Air Quality Development Authority (OAQDA) (Grant AY08-09-C21), and the National Science Foundation (NSF) (Grant CBET-0854203).

■

Figure 14. Simulated χCO for MA as a function of the catalyst load at 550 °C (RH2O/CO = 5.0, and WHSV = 7500 h−1).

MA, when the catalyst load is greater than 196.7 mg of catalyst/cm2 of membrane, the χCO can exceed 99.03, 99.30, and 99.50% under reaction pressures of 2, 25, and 50 atm, respectively.

4. CONCLUSION The experimental and simulation studies of this work demonstrated that the porous MFI-type zeolite membranes with moderate H2 selectivity and H2 permeance can be useful as WGS MRs to effectively enhance the χCO and overcome the equilibrium limit χCO,e. To achieve high χCO, the MR must be operated at high temperatures and high reaction pressures. The high operation temperature provides fast reaction to rapidly consume CO that consequently minimizes the membrane permeation of unreacted CO, and the high reaction pressure results in a large driving force for H2 permeation that improves the efficiency of H2 removal. The fast consumption of CO may also be achieved by increasing the catalyst load if mass-transfer resistance is not an issue (or catalyst activity, which is not studied in this work). The H2 selectivity and permeance of the membrane are critical to the MR performance. Increasing the H2 selectivity and permeance improves the χCO in the MR especially when the reaction rate is the limiting factor. For the current small-size MR, the 1D PFR model was found to work well for simulating the WGS membrane reaction especially for operations under high temperatures and high pressures. The simulation results suggested that the current zeolite MR (membrane MA), although only possessing moderate H2 selectivity (αH2/CO2 ∼ 31, and αH2/CO ∼ 25) and permeance (Pm,H2 < 10−7 mol s−1 m−2 Pa−1), could achieve CO conversion of ≥99% under realistic operation conditions (e.g., at >500 °C, ∼30 atm, and RH2O/CO of ∼3.5). A CO conversion of >99.5% may be obtained by membrane MA at 550 °C if a higher pressure (∼50 atm) and RH2O/CO (∼5.0) are employed. Because of its excellent hydrothermal stability and chemical resistance in a high WGS reaction environment, the modified MIF-type zeolite membranes are potentially useful for constructing MR for high-temperature WGS reaction of coal-derived syngas.

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AUTHOR INFORMATION

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*Telephone: +1-513-556-3992. Fax: +1-513-556-3474. E-mail: [email protected]. Notes

The authors declare no competing financial interest. 4480

dx.doi.org/10.1021/ef302014n | Energy Fuels 2013, 27, 4471−4480