Effect of Pressure on High-Temperature Water Gas ... - ACS Publications

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Effect of Pressure on High-Temperature Water Gas Shift Reaction in Microporous Zeolite Membrane Reactor Seok-Jhin Kim, Zhi Xu, Gunugunuri K. Reddy, Peter Smirniotis, and Junhang Dong* Department of Chemical and Materials Engineering, University of Cincinnati, Cincinnati, Ohio 45220, United States ABSTRACT: A water gas shift (WGS) membrane reactor (MR) has been constructed using a MFI-type zeolite disk membrane packed with a cerium-doped ferrite catalyst. The WGS reaction was performed at high temperatures of 400−550 °C, and the effect of reaction pressure on the MR performance was investigated in a range from 2 to 6 atm with the permeate side swept by nitrogen at atmospheric pressure. Increasing temperature and pressure enhance both the reaction rate and the rate of H2 membrane permeation that in turn significantly enhances the CO conversion. The equilibrium limit of CO conversion can be surpassed in the MR at high pressure and/or high temperature. It has been demonstrated in this study that membranes with moderate H2 selectivity can be effective for enhancing CO conversion at high operation temperature and pressure with the cost of low H2 concentration in the permeate stream. The timely removal of H2 from the catalyst bed dramatically reduced the undesirable methane production because H2 is a reactant for methanation reactions in the WGS system. Both the zeolite membrane and the Fe/Ce catalyst also exhibited good resistances to high concentration of H2S in WGS reactions.

1. INTRODUCTION The water gas shift reaction (WGS) is a key step in hydrogen (H2) production from fossil fuels and biomass through the gasification route. The WGS reaction is moderately exothermic as shown in CO + H2O ↔ CO2 + H2,

Θ ΔH298.15K = − 41.2 kJ/mol

H2 on the permeate side but the membranes have the longstanding issues of sulfur poisoning, hydrothermal embrittlement, and prohibitive costs.5,13−20 The amorphous silica-based membranes also have good H2 selectivity and high permeance but the well-known instability of these membrane materials in hydrothermal conditions still remains a challenge.21−26 The MFI-type zeolite membranes with high silica content, including the silicalite and ZSM-5 zeolites, are hydrothermally stable and intrinsically resistant to sulfur compounds due to their crystalline structure that make them potentially useful for WGS MR.27−32 The hydrophobic surface of siliceous zeolites also makes the membrane H2-selective over water vapor. However, the relatively large pore size of the MFI-type zeolites (effective diameter of 0.56 nm) provides only limited H2 selectivity over CO2 and CO because of the gaseous diffusion mechanism for these small molecules at high temperature. The H2/CO2 selectivity on MFI-type zeolite membranes can be enhanced to over 100 by depositing monosilica inside the zeolitic channels to substantially reduce the pore size which enables the highly selective activated diffusion mechanism.33,34 In our recent studies, alumina-supported MFI-type zeolite membranes were modified by an on-stream catalytic cracking deposition (CCD) process using methyldiethoxysilane (MDES) as precursor.35,36 The modified membrane achieved a high H2/CO2 separation factor and was tested for WGS membrane reaction at near atmospheric pressure and high temperatures of 400−550 °C.37 The modified membrane exhibited good stability in the WGS reaction conditions, and the χCO in this zeolite MR exceeded the equilibrium limit χCO,e at >500 °C. High-temperature (>400 °C) WGS membrane reaction is practically attractive because of the potential to

(1)

Thus, the reaction is kinetically favored at high temperature where the equilibrium CO conversion (χCO,e) is thermodynamically disadvantageous. Current industrial WGS processes typically use two adiabatic packed-bed reactors with the first one operating at high temperature (300−450 °C) for high reaction rate and the second one operating at low temperature (600 °C) and high pressure (>200 psia). Although the χCO,e of the WGS reaction, which has equal volumes in reactant and product, is theoretically independent of pressure in TR, increasing the reaction pressure is expected to enhance the χCO in the MR. At high temperature, where reaction kinetics is not a limiting factor for χCO, higher reaction pressure results in greater H2 partial pressure that creates a larger driving force for the H2 membrane permeation and consequently more effectively enhances the χCO in MR.8,37,38 However, reports on hightemperature and high-pressure WGS MR have been so far very limited. In this work, we use a modified MFI-type zeolite disk membrane to study the effect of reaction pressure on the WGS membrane reaction at high temperatures between 400 and 550 °C and evaluate the MR stability in the presence of highconcentration H2S.

The membrane permeance for gas component i is defined as

Pm,i =

Qi A m t ΔPi

i = H2 , CO2 , ...

,

(2)

where Qi (mol) is the amount of gas permeated over a time period of t (s); Am (m2) is the active membrane area, which is 2.54 cm2 excluding the area sealed by the graphite gasket; and ΔPi (Pa) is the transmembrane pressure, ΔPi = (Pi)f − (Pi)p, where (Pi)f and (Pi)p are the partial pressures of i in the feed and permeate sides, respectively. The H 2/CO2 perm-selectivity (aH° 2/CO2 ) is defined as the ratio of pure gas permeance:

α°H2 /CO2 =

Pm,H2 Pm,CO2

(3)

The H2/CO2 separation factor (αH2/CO2) for the binary mixture is given by

α H2 /CO2 =

2. EXPERIMENTAL SECTION 2.1. Preparation of Membrane and Catalyst. The MFI zeolite membrane was synthesized on a porous α-alumina disk by the in situ crystallization method. The synthesis procedure was essentially the same as described in our previous publications.35 The porous alumina disk was 2 mm thick and 27 mm in diameter which was made from α-alumina powders (SG-16, Almatis) by dry pressing and sintering processes. The substrate disk had an average pore size of ∼0.1 μm and a porosity of ∼30%. One side of the disk was polished by no. 600 SiC sandpaper for growing zeolite membrane. An aluminumfree clear precursor solution was obtained by mixing 6 g of SiO2, 0.42 g of NaOH, 3 mL of H2O, and 30 mL of 1 M tetrapropylammonium hydroxide ((TPA)OH) at ∼80 °C. The disk substrate was immersed in the precursor solution at the bottom of a Teflon-lined autoclave with the polished side facing upward. The membrane synthesis was conducted at 180 °C with duration of 4 h. After the hydrothermal reaction, the disk membrane was washed thoroughly with DI water and then dried and calcined in air at 550 °C for 6 h to remove the template. The MFI zeolite membrane was then modified by the onstream CCD process using MDES as precursor to deposit monosilica species inside the zeolitic pores. The detailed membrane modification process has been described in our previous publications.35,36 The zeolite membrane was mounted in a stainless steel cell with the membrane surface facing the feed stream. During membrane modification, an equimolar H2/CO2 gas mixture (20 cm3/min) was used as carrier gas which was bubbled through a MDES saturator at room temperature and then fed into the membrane cell. The on-stream CCD modifications was performed at 450 °C under atmospheric pressure with the permeate side swept by helium at a flow rate of 20 cm3/min. The permeate stream was continuously analyzed by the online GC to monitor the membrane modification effect in terms of changes in gas permeance and H2/CO2 separation factor. The membrane modification was accomplished by two steps of CCD operation, and each step lasted for 5 h. Between the two steps of CCD, the membrane was kept at 450 °C for overnight (∼12 h) under the H2/CO2 flow without MDES vapor. After modification, H2 and CO2 single-gas permeation and H2/CO2 binary-gas separation were measured at 400−550 °C.

(yH /yCO )permeate 2 2 (yH /yCO )feed 2

2

(4)

where yH2 and yCO2 are mole fractions of H2 and CO2, respectively. The catalyst used in this work was a cerium-doped ferrite (Fe1.82Ce0.18O3), denoted as Fe/Ce catalyst, which had a Fe:Ce atomic ratio of 9:1. The Fe/Ce catalyst was prepared by the coprecipitation method. The details of preparation and characterization of the Fe/Ce catalyst have been reported elsewhere and briefly explained as follows.39,40 The calculated amounts of iron nitrates and cerium nitrate were first dissolved separately in deionized water and then mixed together. Dilute aqueous ammonia was added dropwise to the mixed solutions under vigorous stirring until precipitation was complete at a pH level of 8.5. The precipitate gel was aged overnight and recovered by filtration. The resulting solid was dried at 80 °C for 12 h and calcined at 500 °C for 3 h in an inert environment using heating and cooling rates of 0.3 °C/min. The fresh catalyst in the hematite form (α-Fe2O3) was transformed into magnetite form (Fe3O4) through controlled reduction of α-Fe2O3 into Fe3O4 in a flow of process gas mixture containing 33.3% CO, 25.0% CO2, 25.0% H2, and 16.7% steam. The process gas was allowed to flow through the catalyst bed for 4 h at 400 °C. Care was taken to properly control the temperature because significant amount of heat is released by the exothermic reactions involved, which could damage the catalyst. It is also important to avoid overreduction and formation of FeO, Fe2C, and metallic iron phases, which catalyze undesired side reactions such as methanation and CO disproportionation. For this particular WGS catalyst, the optimal reduction factor (R) in the process gas mixture was found to be R = 1.4, where R = [CO + H2]/[CO2 + H2O].39 This Fe/Ce has been demonstrated to have good sulfur tolerance in high-temperature WGS reaction.41 2.2. WGS Reaction. The WGS membrane reactor system is schematically shown in Figure 1 which is similar to that used in our previous study37 except for the membrane module. The disk membrane was mounted in a stainless steel cell sealed by soft graphite gaskets (Mercer Gasket & Shim). A total amount of 200 mg of Fe/Ce catalyst powders was spread evenly over the zeolite membrane surface forming a thin catalyst bed. The catalyst bed was covered with a carbon cloth sheet and then a quartz wool pad on top. When mounted in the cell, this quartz1365

dx.doi.org/10.1021/ie201452y | Ind. Eng.Chem. Res. 2012, 51, 1364−1375

Industrial & Engineering Chemistry Research

Article

Figure 1. Schematic showing the membrane reaction system and the MR configuration.

carbon cloth pad presses to fix the catalyst bed. The carbon cloth/quartz wool blanket also served as a gas diffusion layer for the feed stream. The external gas tightness of the membrane cell and the tubing connections were verified by soap bubble test when the system was charged with 2 psig helium gas. The internal membrane sealing effect was examined by measuring the H2/CO2 gas separation at 400−550 °C. Before the WGS reaction, the MR was maintained at 250 °C under N2 gas flows (10 cm3/min) on both the catalyst and permeate sides for overnight. The process gas mixture was then introduced to activate the catalyst under the above-described conditions. In the WGS reaction experiments, the pressure on the reaction side was controlled by a back-pressure regulator connected to the exit of the retentate side of MR. The permeate side was swept by a N2 flow at atmospheric pressure. The flow rates of CO and N2 were controlled by mass flow controllers (GFC, Aalborg), and the flow rate of water was controlled by a precision syringe pump (KDS 410, KD Scientific). The liquid water was vaporized at ∼150 °C by a heating jacket prior to entering the preheating coiling located in the furnace. Preheating coils were employed for both the feed and sweeping gases to ensure that the gases reach the set temperature before entering the reactor. The outlet streams from retentate and permeate sides passed through an ice bath to remove most of the water vapor. The flow rates of the streams exiting the MR were frequently checked by a soap film flow meter after the ice traps. The retentate and permeate gases were analyzed by an online GC (Agilent, 6890N) equipped with an Alltech Hayesep DB packed column and a thermal conductivity detector (TCD). A heating or cooling rate of 0.3 °C/min was used when switching the reaction temperature. Before taking data at each of the set conditions, the membrane reaction was

allowed to stabilize for 1 h which was found to be sufficient at high temperatures. The online sampling and GC analysis of the retentate and permeate streams were typically performed twice for each point of operation with 30 min between the two sampling actions. The data reported in this paper were averaged from the values of the two samples. Normally, results of the two samples had relative deviations within ±2% in GC peak areas. In rare cases, the relative deviation between the first two samples was greater than 2% and then a third sample was taken. Byproducts of side reactions such as methane and carbon deposits were not appreciable for this Fe/Ce catalyst in the WGS reaction at >400 °C and atmospheric pressure.39,40 However, methanation reactions are associated with volume reduction which are thermodynamically favored at high pressure. In this work, the reaction products were analyzed by the mass spectrometer (MS) to more closely observe the influence of operation conditions on reaction selectivity. The methane selectivity was found to be far less than 1% in most of the tested conditions in this study. Therefore, the CO conversion (χCO) was calculated on the basis of the CO feed in out out flow rate (FCO ) and the rates of CO2 (FCO2 ) and CH4 (FCH ) 4 exiting the reactor and confirmed by the value calculated out using the amounts of CO entering and exiting (FCO ) the reactor:

χCO =

out out FCO + FCH 2 4 in FCO

F out ≈ 1 − CO in FCO

(5)

out The approximation in eq 5, i.e., neglecting FCH ,4 was assumed in most cases except for in TR mode at >500 °C and >2 atm where CH4 selectivity exceeded 0.1%. The H2 recovery (RH2) is defined by

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R H2 =

(amount of H2 in permeate) (total amount of H2 generated by reaction)

Article

where JH2 , JCO2, JCH4, and JCO are respectively the H2, CO2, CH4, and CO fluxes permeating through the membrane. The reaction selectivity of methanation (SCH4) is calculated by

(6)

The permeate side H2 concentration (yH2,P) is defined on a dry gas basis by

yH ,P = 2

SCH4 =

JH

ζC =

(7)

(FCO + FCO2 + FCH4)retentate,out + (FCO + FCO2 + FCH4)permeate,out −1 (FCO + FCO2)feed,in

In general, the carbon balance was quite good with ζC values typically within the range from −3 to +2%. The deviations may be attributed mainly to the errors of flow rate measurement 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.39,40 When the membrane-mounted cell was used in TR operation mode, 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. The main operating conditions of the WGS reactions are given in Table 1.

400−550

reaction pressure at exit, atm

2−6

permeate pressure, atm steam-to-CO ratio, RH2O/CO, mol/mol

1 1.0−3.5

weight hourly space velocity (WHSVa), h−1 N2 sweeping flow rate, FN2, cm3(STP)/min Fe/Ce catalyst load (mcat.), g catalyst packing density (ρcat.), g/cm3

(9)

3. RESULTS AND DISCUSSION 3.1. Zeolite Membrane Properties. The MFI-type zeolite membrane had a thickness of 2−3 μm according to the SEM observations. Before template removal, a perfect membrane should be impermeable to any gases due to the complete block of zeolite channels by the TPA molecules. The as-synthesized membrane of this work had low helium permeances of 400 °C, including ∼1100 h under WGS reaction conditions and ∼700 h under dry gas permeation. 3.2. WGS Reaction. 3.2.1. Low-Pressure WGS Reaction. The disk zeolite membrane MR was first examined for WGS reactions at 400−550 °C with feed side pressure of 2 atm, the N2 sweeping flow rate of 20 cm3(STP)/min, WHSV = 7500 h−1, and RH2O/CO = 3.5. The results of WGS membrane reaction are presented in Figure 2 in comparison with data of TR operation mode. The previously reported data of WGS membrane reactions obtained on a tube MR under similar operation conditions are also included in Figure 2.37 The tubular membrane in ref 37 was modified by the same CCD process, which had slightly larger H2 permeance of (1.0−1.4) × 10−7 mol/(m2·s·Pa) but significantly higher H2/CO2 separation factor (αH2/CO2 = 38−45) compared to the present disk membrane. The tube MR was packed with the same Fe/Ce catalyst at the same load of 79 mg of catalyst/(cm2 of membane). Although the membrane geometry of this study (disk) was different from that of the reference (tube),37 the two MRs had similarly small reactor sizes with ∼3 mm space above the membrane in the disk MR and 3.5 mm of inner radius for the tube MR. Thus,

Table 1. WGS Membrane Reaction Conditions reaction temperature, °C

(8)

The carbon balance (ζC) was checked for each reaction condition based on the actually measured gas flow rates and compositions:

2

JH + JCH + JCO + JCO 2 4 2

moles of CH4 in exiting streams of both sides moles of CO reacted

7 500−60 000 0−40

0.2 2.0

a

H2 CO2 H2O CO H2 CO2 WHSV = (vfeed CO + vfeed + vfeed + vfeed )/(mcat./ρcat.); vfeed , vfeed , vfeed , H2O are volumetric flow rates of components in the feed streams and vfeed tot. CO H2 CO2 = vfeed + vfeed + vfeed + at STP. The total volumetric flow rates (vfeed H2O ) of the entering feed stream were 12.5, 25, 50, and 100 vfeed cm3(STP)/min at WHSV of 7500, 15 000, 30 000, and 60 000 h−1, respectively; the feed rate of CO is defined by the RH2O/CO and dry gas CO tot. composition, vfeed = vfeed yCO/(1 + RH2O/CO), where yCO is the CO mole fraction in dry gas.

A series of WGS experiments were carried out in both the MR and TR operation modes. The effect of reaction pressure on CO conversion has been studied for various temperatures, steam-to-CO ratios, weight hourly space velocity (WHSV), and sweep flow rates. The effects of reaction pressure on the undesirable methanation were also investigated under the MR and TR operations. The WGS reaction experiments were further performed for a simulated syngas containing H2, CO, and CO2 as well as for feed streams containing H2S which is a major contaminant in coal-derived syngas. 1367



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Table 2. Gas Permeance (Pm,i, 10−7 mol/(m2·s·Pa)) and H2 Selectivity for the Modified Membrane after operating for 1800 h at >400 °C

fresh after modification T (°C) 400

property

H2

CO2

H2

CO2

H2O

N2

CO

Pm,ia

1.20

0.08 14.2

0.83

0.09 9.23

0.31 2.64

0.06 12.7

0.07 12.6

Pm,ib αH2/i

0.85

0.09 9.76

0.62

0.08 7.47

Pm,ia α H2/i°°

1.25

0.08 15.6

0.90

0.09 10.3

0.33 2.77

0.06 14.2

0.07 12.5

Pm,ib αH2/i

0.93 -

0.09 10.9

0.64

0.08 7.79

Pm,ia α°H2/i

1.33

0.07 18.1

1.03

0.09 11.4

0.34 2.99

0.06 16.1

0.07 14.3

Pm,ib αH2/i

0.99

0.09 11.1

0.65

0.08 8.28

Pm,ia αH2/i

1.45

0.07 18.8

1.03

0.08 12.2

0.35 2.95

0.07 15.9

0.07 14.5

Pm,ib αH2/i

1.01

0.09 11.3

0.66

0.08 8.6

αH° 2/i °

450

500

550

a

Single-gas permeance. bPermeance for H2/CO2 equimolar mixture.

As can be seen in Figure 2b, the H2 molar concentration in the permeate stream (yH2,p) of the disk MR was notably lower than that in the tube MR because the disk MR had a much lower H2/CO2 selectivity. The H2 recovery in the disk MR was only slightly less than that in the tube MR because of the small difference in H2 permeance between the two membranes. The disk MR had lower χCO than the tubular MR, but the difference of χCO between the two MRs diminished when the reaction temperature was raised to above 500 °C. Because the two membranes had similar H2 permeances, the main cause of the lower χCO in the disk MR is likely to be its low H2 selectivity which allows for larger permeation of CO. At relatively low temperature, the slow reaction keeps the CO partial pressure high in the feed side that causes more CO permeation under the same WHSV. Thus, the H2 selectivity has a greater impact on the χCO of the MR at relatively low temperature. In contrast, at high temperature, the fast reaction rate dramatically reduces the amount of unreacted CO in the feed side that lowers the driving force for CO permeation and thus reduces the effect of H2 selectivity on χCO of the MR. The fast reaction at high temperature also increases the CO2 permeation through the membrane that reduces the yH2,p; however, the removal of CO2 from the catalyst bed also enhances the χCO just the same as does the removal of H2. These findings suggest that membranes with moderate H2 selectivity can still be effective for χCO enhancement at high operation temperature but will compromise the effectiveness of simultaneous H2 purification in the MR. However, the reduction of χCO caused by the permeation of CO and H2O could become a serious issue for low RH2O/CO feeds if the H2 selectivity of the membrane is too low. The present disk membrane had H2/CO and H2/H2O permselectivities of about 14 and ∼3.0, respectively, as measured after 1800 h of gas permeation and WGS reaction at >400 °C. 3.2.2. Effect of Pressures. The WGS reaction experiments were performed with feed side pressures varying from 2 to 6 atm. The permeate side was swept by a N2 flow at atmospheric pressure, and the WHSV and RH2O/CO were fixed at 7500 h−1 and

Figure 2. Results of WGS reaction in the disk MR compared with data of TR and tube MR in ref 37: (a) χCO and (b) RH2 and yH2,p.

under high space velocity, the variations caused by the differences in reactor geometry and flow configuration are expected to be insignificant that warrants reasonable comparability between the two MRs when other operation conditions are kept identical. 1368



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Figure 3. Effect of reaction pressure and temperature on χCO, RH2, and yH2,p in the disk MR: (a) χCO vs pfeed at 400 °C, (b) χCO vs pfeed at 450 °C, (c) χCO vs pfeed at 500 °C, (d) χCO vs pfeed at 550 °C, (e) RH2 and yH2,p as functions of pfeed, and (f) χCO as a function of temperature.

3.5, respectively. Figure 3 compares the χCO between MR and TR operation modes. The RH2 and yH2,P in MR are also presented in the Figure 3e as a function of pressure at temperatures of 400, 450, 500, and 550 °C, respectively. The χCO of the MR is much higher than that of the TR and increases with reaction pressure (feed side, pfeed) at all temperatures. The χCO in the TR exhibited a slight increase with reaction pressure which is probably caused by more effective utilization of internal surface in the nanocrystalline catalyst bed under high pressure. In the MR, a higher reaction

pressure leads to larger H2 partial pressure in the feed side that creates greater driving force (ΔPH2 = (PH2)f − (PH2)p) for H2 membrane transport when the permeate side pressure is kept relatively constant by the sweeping gas. Thus, H2 is removed more efficiently at large pfeed to shift the WGS reaction toward the right side and meanwhile achieve in higher H2 recovery in the permeate stream as shown in Figure 3e. However, the enhanced H2 removal leads to a smaller value of (yH2/yCO2)feed in the reaction side because of H2 depletion. Also the αH2/CO2 is rather insensitive to pressure at high temperature where the 1369



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lectivity of ∼3.0 and thus steam loss may become an issue under high pressure and low RH2O/CO. The WGS reaction in the MR was studied at 550 °C for RH2O/CO = 1.0, 1.5, 2.0, and 3.5, respectively. Figure 4 presents the experimental χCO as a function of RH2O/CO at WHSV = 7500 h−1 and feed pressure of 2 and 6 atm, respectively. The results show that χCO increases with RH2O/CO at both feed pressures because the excessive steam favors χCO,e and high steam pressure enhances the reaction rate.37 It is also seen in

gases are essentially nonadsorbing in the membrane and exhibit ideal gas behavior.36 Thus, on the basis of eq 4, reducing (y H 2 / y CO 2 ) feed decreases the value of (y H 2 / y CO 2 ) permeate (=αH2/CO2(yH2/yCO2)feed), which explains the trend of yH2,p decreasing with increasing RH2 and χCO under higher feed pressures. At a given WHSV, overcoming the equilibrium limit χCO,e by H2 removal in the MR may be realized by increasing the H2 partial pressure in the reaction side through increasing the temperature for faster H2 generation (thus approaching the equilibrium state faster) and/or increasing reaction pressure. Results in Figure 3a−d demonstrate that the necessary reaction temperature for surpassing the χCO,e can be lowered by increasing the reaction pressure. Figure 3f, on the other hand, shows that the reaction pressure necessary for overcoming the χCO,e can be reduced when the temperature is raised. It is also observed that χCO tends to level off after certain points in both cases of increasing pressure at a fixed temperature and increasing temperature under a fixed pressure. The χCO in a MR depends on multiple factors including membrane performance, catalyst properties, and operation conditions. The level-off behavior indicates that adjusting or improving one individual factor alone is insufficient for achieving nearly complete CO conversion. For example, at 550 °C and 6 atm, a high χCO of 98% was achieved at a RH2 of 64%, where CO concentration became very low and is no longer able to rapidly generate H2 and create high H2 partial pressure even under high pressure and temperature. In this case, unless the permeate side H2 is completely removed and the overall H2 transport resistance is near zero, further increasing the χCO is difficult. Moreover, for complete CO conversion, the H2 selectivity of the membrane must be high enough to prevent permeation of unreacted CO which directly reduces the χCO. The current disk membrane is not expected to achieve complete CO conversion due to its moderate H2 selectivity. Although the transmembrane CO partial pressure feed permeate − PCO ) became rather small when difference (ΔPCO = PCO approaching the exit of MR, the ΔPCO in the membrane area near the MR entrance could be large enough to allow for significant CO permeation, especially under high reaction pressure. At 500 °C, the amount of CO in the permeate side was only 29% of the total unreacted CO exiting the MR at 2 atm; this percentage increased to 45% when pressure rose to 6 atm because raising the operation pressure increases membrane fluxes for all components. 3.2.3. Steam-to-CO Ratio. The ability to achieve high χCO with low RH2O/CO is an important advantage of WGS MR because lowering RH2O/CO reduces the reactor size and energy consumption for steam recycling. A low RH2O/CO is also desired for minimizing the adverse effect of H2 dilution by excessive steam in the reaction zone and thus maintaining high efficiency of H2 transport through the membrane.37 A perfect H2-selective membrane such as the dense Pd-alloy membrane, which is impermeable to H2O and CO, can theoretically achieve nearly complete CO conversion even for a feed with stoichiometric RH2O/CO = 1. However, permeation of steam and CO is inevitable in porous membranes because their molecular sizes are very close to H2. The permeance of H2O is much greater than CO because the molecular size of H2O is smaller than that of CO. The loss of steam from a feed with RH2O/CO = 1.0 would be directly translated into the same amount of decrease in χCO. The current disk membrane had a rather modest H2/H2O permse-

Figure 4. Effect of feed RH2O/CO on χCO and RH2 at 550 °C and WHSV = 7500 h−1 under different pressures.

Figure 4 that the enhancement of χCO is more pronounced for a RH2O/CO increase from 1 to 1.5 than for an increase from 2 to 3.5. Changing the RH2O/CO from 1 to 1.5 in the MR operation not only thermodynamically and kinetically favors the χCO but also avoids the stoichiometric steam deficit (RH2O/CO < 1) caused by the H2O permeation at RH2O/CO = 1. Although further increasing the RH2O/CO leads to greater χCO,e, it causes more H2 dilution in the reaction side to hinder the H2 transfer from the catalyst bed to the permeate side. The countering effects of high RH2O/CO on H2 transfer is evidenced by the nearly unchanged RH2 with increasing RH2O/CO as shown in Figure 4. 3.2.4. Effect of WHSV. Figure 5 presents the results of WGS reaction in the disk MR and TR at 550 °C under reaction pressures of 2 and 6 atm, respectively. The WHSV value was varied from 7500 to 60 000 h−1 at a fixed RH2O/CO = 3.5. At both operation pressures, χCO in the MR increased with decreasing WHSV because of longer residence time for reaction and H2 permeation at smaller WHSV. The χCO can be enhanced by simultaneously increasing feed pressure and lowering the WHSV. Also, when WHSV decreased from 60 000 to 7500 h−1 at 6 atm, the RH2 was found to increase from 41 to 64% while H2 purity had a rather moderate decrease over the range of WHSV. However, increasing pressure compromises yH2,p as shown in Figure 5b due to relatively larger CO2 flux when H2 in the reaction side depleted under high RH2. 3.2.5. Effect of Sweep Flow. The permeate side sweep gas flow rate affects the H2 membrane flux (JH2) by changing the driving force (ΔpH2): 1370



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Figure 5. Effect of WHSV on (a) the χCO and (b) RH2 and yH2,P.

factor of the overall mass-transfer resistance and obtaining a higher χCO and RH2 will rely on simultaneously improving other steps of the transport process. 3.2.6. Methanation. Methane is one of the main undesirable byproducts of WGS reaction.38,43−47 Methanation reactions in the WGS reaction system are exothermic and involve an overall volume decrease which are therefore thermodynamically favored at high pressure and kinetically favored at high temperature. The selectivity of methanation was examined under both MR and TR operation modes at pressures of 2−6 atm and temperatures of 500 and 550 °C, respectively. The WHSV and RH2O/CO were fixed at 7500 h−1 and 3.5, respectively. The amounts of methane in permeate and retentate streams were analyzed by GC/MS due to the small quantity. Figure 7 presents the comparison of methane selectivity (SCH4) between TR and MR operations. As expected, in all cases, SCH4 increased with operation temperature and pressure. The SCH4 was much larger at 550 °C than at 500 °C. The SCH4 increased rapidly from 0.02 to 0.92% under the TR operation mode as the feed pressure increased from 2 to 6 atm. The methane selectivity at 550 °C in the MR

JH = Pm,H2ΔpH = Pm,H2(pH ,f − pH ,p ) 2 2 2 2 = Pm,H2pH ,f (1 − pH ,p /pH ,f ) 2 2 2

(10)

Figure 6 shows χCO and RH2 as a function of N2 sweeping flow rate (FN2) at feed pressures of 2 and 6 atm, respectively. The FN2 was varied from 5 to 40 cm3/min at 500 °C under a fixed WHSV = 7500 h−1 and RH2O/CO = 3.5. Increasing the sweeping flow rate reduces the permeate side H2 partial pressure pH2,p that increases JH2 (i.e., rate of H2 removal) and

Figure 6. Effect of N2 sweeping flow rate on χCO and RH2 for a fixed WHSV = 7500 h−1 at 500 °C.

consequently enhances the χCO. On the other hand, the pH2,f increases with reaction pressure; therefore, the χCO is less sensitive to the pH2,p variation or change of sweeping flow rate at high reaction pressure. This explains why χCO in the MR is more sensitive to FN2 at an operation pressure of 2 atm than at 6 atm. This observation implies the possibility to operate the MR without a sweeping flow at sufficiently high reaction pressure. In Figure 6, both χCO and RH2 level off as FN2 reaches certain levels, more noticeably for reaction pressure of 2 atm. The permeate side H2 partial pressure is only one of the many factors influencing the overall rate of H2 removal from catalyst surface to the permeate stream, including the transport resistances in the catalyst layer, between the catalyst and membrane surface, through the membrane pores, through the substrate, and from membrane to permeate. The plateau of χCO and RH2 in Figure 6 indicate that pH2,p is no longer a significant

Figure 7. Methane selectivity in the MR and TR under a fixed WHSV = 7500 h−1 and RH2O/CO = 3.5.

was dramatically reduced to 0.01% at 2 atm and 0.03% at 6 atm, respectively. The inhibition of methanation in the MR is attributed to the removal of H2, a reactant for methane formation, from the catalyst surface. The ability of the MR to inhibit methanation is an important advantage for WGS MR because methanation consumes the useful H2 and creates undesirable CH4 impurity in the product. 3.2.7. CO-H2−CO2 Mixture Feed. The disk MR was tested for WGS reaction for a mixture feed containing 40% CO, 30% 1371



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H2, and 30% CO2. The experiments were performed with a fixed WHSV = 7500 h−1 and a RH2O/CO = 3.5 in temperature and pressure ranges of 400−550 °C and 2−6 atm, respectively. The results are shown in Figure 8. In general, the temperature and pressure dependences of χCO, RH2, and yH2,p for the ternary mixture feed were similar to those for the pure CO feed. The χCO in the MR exceeded the χCO,e at >450 °C, and the enhancement of χCO in the MR became more significant as reaction temperature and pressure increased. The RH2 and yH2,p

RH2 and yH2,p have been explained in section 3.2.2. Increasing the feed side pressure creates greater ΔpH2 that leads to a higher rate of H2 removal and H2 recovery. On the other hand, the CO2 mole fraction in the reaction side becomes larger as more H2 is removed that results in an increase of CO2 in the membrane permeation to lower the H2 concentration in the permeate stream. The methanation selectivity in WGS reaction of the mixture feed was also examined, and the results are shown in Figure 9. The methane selectivity SCH4 for the mixture feed was higher than that for the pure CO feed because of the large amounts of preexisting CO2 and H2 which are reactants of methanation reactions. The SCH4 was dramatically reduced in the MR operation as compared to the TR operation that is consistent with the result observed in pure CO feed. At reaction pressure of 6 atm and

Figure 9. Methane selectivity for the CO−H2−CO2 mixture feed in the MR and TR under a fixed WHSV = 7500 h−1 and RH2O/CO = 3.5.

temperature of 550 °C, the SCH4 values were ∼1.5% in TR operation and ∼0.07% in the membrane reaction, respectively. 3.2.8. Stability in H2S. H2S is a major impurity in the coalderived syngas, which is poisonous to the WGS catalysts and Pd-based membranes.16,48 The zeolite membrane reactor was tested for WGS reaction with a CO feed containing 400 ppm H2S. The WGS membrane reaction was performed under two general conditions: the first was at a reaction pressure of 2 atm and a fixed RH2O/CO of 3.5 with temperature varying from 400 to 550 °C; and the second was at a reaction pressure of 2 atm and a fixed temperature of 550 °C while varying the RH2O/CO from 1 to 3.5. The WGS reaction with the H2S-containing feed was allowed for 4 h of stabilization at each reaction condition before taking data. The results are presented in Figure 10, which shows essentially no difference in χCO between feeds with and without H2S. The MR was then further tested for WGS reaction with 1000 ppm H2S in the CO feed stream for an extended time of 100 h. The reaction was conducted at 500 °C with a fixed WHSV = 7500 h−1 and a RH2O/CO = 3.5. Prior to introducing the H2S, the membrane WGS reaction was stabilized for 3 h with a sulfur-free feed stream. Figure 11 shows the resultant χCO versus time-on-stream. The χCO in the MR was quite stable with only 1.5% decrease after 100 h of operation. This result indicates that modified MFI-type membrane is chemically resistant to H2S, and the Fe/Ce catalyst is tolerant to high H2S content at high reaction temperature. Also, during the WGS reaction with 1000 ppm H2S, H2S was not appreciable in the permeate stream by GC analysis while a

Figure 8. Results of WGS reaction in the MR for a feed stream containing 40% CO, 30% H2, and 30% CO2.

had opposite trends with changing pressure, namely, the RH2 increased while yH2,p decreased upon raising the reaction pressure. The causes of these distinct pressure dependences of 1372



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Figure 10. WGS membrane reaction with and without H2S in feed at a fixed WHSV = 7500 h−1: (a) χCO as a function of temperature at a RH2O/CO = 3.5 and (b) χCO as a function of RH2O/CO at 550 °C.

water vapor, H2S can physically adsorb on silica surface at relatively low temperatures and desorb at high temperature but does not chemically adsorb or react with silica even at high temperatures.49,50 Therefore, one possible reason for the more rapid increase of H2 permeance is the reduction of H2S adsorption at the membrane surface and pore entrance at high temperature that increased the permeation of the nonadsorbing H2. At 550 °C, the membrane showed H2 permeance of 9.0 × 10−8 mol/(m2·s·Pa), H2S permeance of 1.7 × 10−9 mol/(m2·s·Pa) and a high H2/H2S separation factor of ∼53.

4. CONCLUSIONS A modified MFI-type zeolite disk membrane reactor (MR) packed with a cerium-doped ferrite (Fe/Ce) catalyst was used for high-temperature WGS reaction to study the effect of reaction pressure in a range of 2−6 atm. It was found that increasing reaction pressure in the MR benefits the enhancement of the χCO and dramatically reduces the undesirable side reactions of methanation. Both the enhancement of χCO and the inhibition of methanation in the MR are attributed to the more effective removal of H2 from the catalyst bed at higher pressures, because H2 is a product of WGS reaction and meanwhile a reactant for methanation reactions. The results of this study indicate that membranes with moderate H 2 selectivity can still be effective for χCO enhancement at high operation temperature but will compromise the effectiveness of simultaneous H2 purification in the MR. Althrough the zeolite membrane used in this work had moderate H2/CO2 selectivity and H2 permeance, χCO in the MR exceeded the χCO,e at >450 °C and the χCO enhancement became greater at higher temperature and pressure. However, for porous membrane with imperfect H2 selectivity, the RH2 and yH2,p exhibit opposite trends with pressure, namely, increasing RH2 but decreasing yH2,p with raising reaction pressure, because the relative amount of CO2 in the membrane flux increases when H2 is depleted in the reaction side at high H2 recovery. The zeolite membrane and the Fe/Ce catalyst also showed good stability and tolerance to high concentration H2S in high-temperature WGS reaction. It is clear that the improvement of the WGS MR depends not only on the development of better membrane and catalyst but also on the optimization of the operation conditions.

Figure 11. χCO vs WGS reaction time for a feed containing 1000 ppm H2S.

large H2S peak was found in the GC spectrum of the retentate sample. This suggests that the membrane also has good separation for H2 and H2S. The separation of H2/H2S was then examined for the membrane using gas mixture containing 1% H2S and balance H2 at temperatures from 25 to 550 °C and atmospheric pressure in both the feed and permeate sides. The permeate side was swept by N2 gas at a flow rate of 20 cm3(STP)/min. Figure 12 presents the H2/H2S separation factor and individual gas permeance as a function of temperature. Both H2 and H2S permeance increased with increasing temperature because of the activated diffusion mechanism in the modified zeolite pores. However, the magnitude of H2 permeance increase was greater than that of H2S when temperature rose from 25 to 550 °C. It has been reported in the literature that, with or without the presence of

Figure 12. Results of membrane separation mixture of 1% H2S. 1373



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AUTHOR INFORMATION

Corresponding Author

*Tel.: +1 (513) 556-3992. Fax: +1 (513) 556-3474. E-mail: [email protected].

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ACKNOWLEDGMENTS Research supported by the U.S. DOE/NETL (Grant DE-FG36GO15043), the Ohio Air Quality Development Authority (Grant AY08-09-C21), and the National Science Foundation (Grant CBET-0854203).

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