WGS reaction in a membrane reactor using a porous stainless steel supported silica membrane

Chemical Engineering and Processing 46 (2007) 119–126

WGS reaction in a membrane reactor using a porous stainless steel supported silica membrane A. Brunetti a,b , G. Barbieri a,∗ , E. Drioli a,b , K.-H. Lee c , B. Sea c , D.-W. Lee c a

Institute for Membrane Technology ITM-CNR, Via Pietro BUCCI, c/o The University of Calabria, cubo 17/C, 87030 Rende CS, Italy Department of Chemical Engineering and Materials, The University of Calabria, Via Pietro BUCCI, cubo 44/A, 87030 Rende CS, Italy c Membrane and Separation Research Center, Korea Research Institute of Chemical Technology (KRICT), 305-600 Daejon, South Korea

b

Received 1 March 2006; accepted 12 May 2006 Available online 22 May 2006

Abstract Water gas shift reaction for hydrogen production was studied in a catalytic membrane reactor using a supported silica membrane at 220–290 ◦ C temperature and 2–6 bar pressure ranges. A CO conversion higher than the thermodynamic equilibrium of a traditional reactor was obtained. The best result, 95% CO conversion, was achieved at 4 bar and 280 ◦ C. The membrane was also characterized in terms of permeance and selectivity by means of permeation tests carried out before and after reaction. In addition, permeance and separation factor were also measured during the reaction. Permeance of all species (H2 : 9.7–29; CO: 0.3–1.1; CO2 : 0.4–1.5 nmol/m2 s Pa), selectivity (H2 /CO, H2 /CO2 and H2 /N2 ) ranging from 15 to 40 and separation factors (H2 /CO = 20–45), showed no dependence on the related permeation driving force. Differences between selectivity and separation factor were registered. Furthermore, no inhibition effects of other gases on the hydrogen flux were observed. The membrane was prepared by the soaking roller procedure depositing a silica layer on a stainless steel support with an intermediate ␥-alumina layer. The membrane reactor allowing selective hydrogen permeation presents a good performance exceeding also the equilibrium conversion of a traditional reactor. © 2006 Elsevier B.V. All rights reserved. Keywords: Hydrogen production; Membrane reactor; Water gas shift; Silica membrane; Fuel cell

1. Introduction Hydrogen as a carrier to be employed in a clean energy process has attracted a great deal of attention in the last few decades and new technologies such as polymer electrolyte fuel cells, requiring CO-free hydrogen, have promoted improvements in the H2 production cycle. Light hydrocarbons represent a more realistic way to produce hydrogen and many studies on the use of catalytic membrane reactors (MRs) in integrated plants for hydrogen production with low CO content have been carried out [5,10]. MRs, in which reaction and separation occur at the same

Abbreviations: MR, membrane reactor; MREC, MR equilibrium conversion; RF, recovery factor (%); TR, traditional reactor; TREC, TR equilibrium conversion; WGS, water gas shift ∗ Corresponding author. Tel.: +39 0984 492029; fax: +39 0984 402103. E-mail address: [email protected] (G. Barbieri). 0255-2701/$ – see front matter © 2006 Elsevier B.V. All rights reserved. doi:10.1016/j.cep.2006.05.005

time and in the same unit, move also in the logic of the Process Intensification Strategy [1], offering an innovative approach to process design and development. An MR with a hydrogen selective membrane increasing the hydrocarbon conversion, reducing the H2 purification section and the energy consumption, allowing better raw material exploitation, follows this new design plant philosophy. Currently the hydrogen produced by reforming and/or partial oxidation of light hydrocarbons, such as natural gas, contains carbon monoxide, carbon dioxide, water and a small amount of CH4 . For this reason one fundamental step of the integrated membrane plant for hydrogen production is the upgrading of the streams coming from the reformer, generally consisting in water gas shift (WGS) reaction, necessary to reduce the CO content and in the meantime producing more hydrogen: CO + H2 O = CO2 + H2 ,

◦ H298 = −41 kJ/mol

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This reaction is exothermic and characterized by no variation in the number of moles. Thus, the CO equilibrium conversion is favoured by a low temperature and is independent of the reaction pressure in a traditional reactor (TR). A catalyst is used because of the very low reaction kinetics. A promising approach is the use of MRs, combining the reaction and H2 separation through a selective membrane. In this way the equilibrium can be shifted, achieving a higher CO conversion [2–6]. Furthermore, as reported in previous work [2,7] the MR reduces the catalyst amount necessary for a given conversion; thus reaction (or plant) volume reduces consequently. In the last two decades, H2 -separation membranes have been developed using various materials, such as palladium and its alloys, silica and alumina, etc. Palladium membranes formed on porous alumina supports by electroless plating and chemical vapour deposition methods were reported to show high H2 permselectivity at 250–300 ◦ C [8,9], but they have disadvantages, such as high costs, degradation of H2 -separation performance, and embrittlement. Ceramic materials, on the other hands, can be hopefully expected to be membrane materials of high stability at high temperatures. Considerable attention was attracted by silica membranes for the production of hydrogen streams free of CO for feeding PEMFCs. Silica membranes exhibit lower module costs, high permeating fluxes, no inhibition effects by chemical species, and high thermal stability, on the contrary to Pd-based membranes generally used for H2 -separation. Therefore, they appear attractive for hydrogen purification [10]. However, the improvement of H2 selectivity of this membrane type is an important issue. Since permeance is normally enhanced through the reduction of selectivity, a balance between permeance and selectivity is essential to achieve optimum performance. The presence of an intermediate layer [13,14] between the active layer and the support tube is expected to improve both permeance and selectivity. One of the attractive candidates for the high perm-selective membrane is a silica layer formed on mesoporous ␥-alumina film, supported on porous stainless steel. In this work the WGS reaction in an MR with a porous stainless steel supported silica membrane was analyzed. Furthermore, the membranes were characterized by means of permeation tests before and after the reaction. In addition, permeances and separation factors were measured also during the reaction.

2. Materials and methods 2.1. Experimental apparatus An MR prototype using a flat membrane was sealed in an SS shell by using silicon gaskets. The experimental apparatus used in permeation and reaction tests is showed in Fig. 1. The stainless steel module containing the membrane was placed in a temperature controlled electric furnace (with PID control). An HPLC pump was utilized to feed the liquid water, while a coil in the furnace allows water vaporization before its mixing with CO. Mass flow controllers (Brooks Instrument 5850S) were used for feeding all the inlet gaseous streams. The flow rates of the outlet streams were measured by means of bubble soap flow-meters. No automated instruments (e.g., mass flow controller) can be utilized because the outlet composition is unknown a priori. The chemical analyses on the retentate and permeate streams were performed by means of a gas chromatograph (Agilent 6890N) with two parallel analytical lines. Each line is equipped with two columns: an HP-Plot-5A (for separating permanent gases such as H2 , N2 and CO) and an HP-Poraplot-Q (for other species) and a TCD. The permeation driving force for each species is the transmembrane pressure difference (PiTM ) between reaction and permeate sides. Therefore, any pressure variation generated on outlet streams affects the permeation through the membrane. A GC with one analytical line requires a switching valve; the switching changes the pressures. The GC configuration adopted in this plant allows (a) the analysis of both outlet (retentate and permeate) streams at the same time and (b) avoidance of any pressure variation on the reaction and permeation sides, with no modification of the steady-state. The reaction was performed packing a commercial CuO/CeO2 based catalyst (Engelhard). For reaction tests in the TR the same catalyst was used and the results were also compared with those obtained with another commercial catalyst (Cu/Zn oxide, Hal dor Topsoe, LK821-2). 2.2. Silica membrane A porous stainless steel (SUS 316) disk (Mott Co., US) was used as the support to give the membrane mechanical strength. The macropores of the support disk were modified by packing

Fig. 1. Experimental laboratory-scale plant. MFC, mass flow controller; TC, thermocouple; GC, gas chromatograph.

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Table 1 Geometric membrane characteristics Selective layer Type Silica thickness (␮m) Average pore size (nm) Alumina thickness (␮m)

SiO2 1 <1 5

Support

Fig. 2. Nitrogen adsorption–desorption isotherm for a ␥-alumina film prepared from boehmite sol.

silica xerogel (500 nm) with a pressure under 10 MPa and by coating the intermediate layer of ␥-alumina. A mesoporous ␥-alumina layer was coated on the SUS support disk from a boehmite sol (␥-AlOOH) by the soaking–rolling procedure [14]. After this coating process, repeated three times, the thickness of the ␥-alumina layer was about 5 ␮m, and the top surface was smooth and defect free. After the film was heat-treated at 700 ◦ C for 1 h in air, the typical adsorption–desorption isotherm (Fig. 2) for the ␥-alumina film was of type IV corresponding to the mesoporous solid (2–50 nm pore diameters). The average pore size of the ␥-alumina coating layer was determined with a BET unit (Micromeritics, ASAP 2200) and was about 5–7 nm in diameter. To improve the gas permselectivity, amorphous silica was formed on the ␥-alumina layer by the thermal decomposition of tetraethoxysilane (Aldrich) at 600 ◦ C for 1 h. Details of these procedures for SUS support modification and membrane preparation were reported in previous studies [11,13,14]. Fig. 3 shows a SEM photo of a typical membrane cross-section. A thin silica layer effective for selective H2 permeation, can be seen on a ␥-alumina intermediate layer, covering completely the SUS support. This intermediate layer reduces the pinholes on the separative layer, allowing a better permselectivity. Table 1

Type Disk diameter (mm) Thickness (mm) Surface area (cm2 ) Average pore size (nm) Porosity (%)

316L porous stainless steel (SUS) 30 1.5 2.8 500 30–40

summarizes the membrane characteristics. Before permeation and reaction tests the membrane was heated at 1 ◦ C/min in hydrogen flux up to 290 ◦ C. 3. Results and discussion 3.1. Permeation tests Permeation tests were carried out before and after the reaction in order to estimate the eventual effect due to reaction. 3.1.1. Before reaction Permeation tests before the reaction were carried out at 250 and 290 ◦ C with typical WGS reaction gasses (H2 , CO, CO2 ) and with N2 . Pure gas and mixtures were tested, as reported in Table 2. The pressure drop method was used for the single gas experiments and thus a trans-membrane pressure difference acted as the driving force for the permeation during the measurements. No sweep gas was used. The permeation tests measure the permeating flux through the membrane and then the gas permeance can be evaluated as:   mol permeating fluxi permeancei = (1) m2 s Pa PiTM where PiTM is the trans-membrane pressure difference which acts as the driving force for permeation. Table 2 Permeation tests operating conditions Temperature range (◦ C) Feed pressure range (bar) Permeate pressure (bar) Pure gases Gas mixtures

Fig. 3. Fractured section of silica membrane formed on alumina coated SUS support.

Feed flow rate (cm (STP)/min)

250–290 2–6 1 H2 , N2 , CO H2 :N2 = 50:50 H2 :CO = 50:50 H2 :CO2 = 50:50 5–40

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The ratio of two pure gas permeances measured at the same temperature is the selectivity: permeancei selectivityi,j = (2) permeancej While, the ratio of the permeances measured in a gas mixture is the separation factor: permeancei separation factori,j = (3) permeancej The permeation tests show a linear dependence of the flux with the corresponding driving forces for all the species analyzed (Fig. 4); therefore a constant permeance value can be assumed. In all the cases, the flux increases with the temperature because the permeation is favoured by a high temperature. However, comparing the figures, the H2 flux is at least one order of magnitude higher than that of the other species. Species flux measured both in a single gas and in a gas mixture permeation tests is fitted by the same line (Fig. 4), at a fixed temperature. Both N2 and CO do not inhibit the hydrogen flux through the membrane, showing an inert gas behaviour. Permeance and selectivity values for the gas tested are reported in Table 3. The permeance, of hydrogen in particular, is not very high; however a comparison with literature data [11–14] of other SiO2 membranes shows a good agreement (Fig. 5) even if some difference with the data of Lee et al. [14] are present due to a different membrane preparation. H2 selectivities with respect to CO, CO2 and N2 (Table 3) are very interesting; specifically, the H2 /CO selectivity is the highest. This improves the separation of products (H2 and CO2 ) from reagents (e.g., CO) during the WGS reaction, allowing an H2 permeate rich stream with low CO content. A separation factor higher than the selectivities is present at all the temperature investigated (Table 3). This difference can be attributed to the different gas behaviour in the two permeation tests. In pure gas tests, in fact, the whole membrane surface is available for the permeation of the only species present in the retentate volume. In gas mixture tests the permeation of a gas is conditioned by the presence of the other gas, e.g., the CO permeation is hindered by the H2 (the faster

Fig. 4. Hydrogen, carbon monoxide, carbon dioxide, nitrogen permeating flux as a function of the respective driving force PiTM at different temperatures in pure gas (full symbols) and mixtures (open symbols).

species), which occupying the membrane pores reduces the CO permeation. 3.2. Reaction tests Reaction tests were carried out on TR and MR at the operating conditions reported in Table 4. An equimolecular H2 O/CO stream was used in the reaction experiments, on the basis of previous studies [15]. No sweep gas was employed during the tests. The CO conversion of a TR was

Table 3 Permeation tests before the reaction—selectivities, separation factors and measured permeance Type

Permeance (nmol/m2 s Pa) H2

N2

CO

CO2

1.898

0.119

0.05

0.054

H2 /N2

H2 /CO

H2 /CO2

Selectivity: 16

Selectivity: 42

Selectivity: 35

225 ◦ C Pure gases Mixture H2 :CO = 50:50 252 ◦ C Pure gases Mixture H2 :CO = 50:50 291 ◦ C Pure gases Mixtures H2 :N2 = 50:50 H2 :CO = 50:50

1.77 2.20

0.04 0.15

2.57

0.06

Separation factor: 45 0.06

Selectivity: 15

0.06

4.32

0.20

3.94 4.59

0.20

0.14

Selectivity: 37 Separation factor: 45

0.189

Selectivity: 22

Selectivity: 31

Separation factor: 19 0.13

Selectivity: 35

Separation factor: 35

Selectivity: 23

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Fig. 5. H2 (full symbols) and N2 (open symbols) permeance as function of temperature: comparison between experimental and literature data.

calculated as following: CO conversion = 1 −

OUT FCO Feed FCO

=

OUT FCO 2 Feed FCO

(4)

whereas, it was calculated by means of the following equation for an MR: CO conversion = 1 − =

Retentate + F Permeate FCO CO Feed FCO

Retentate + F Permeate FCO CO2 2

(5)

Feed FCO

The most valid kinetics for Cu-based catalyst is that proposed by Moe (Amadeo–Laborde) [16] (Eq. (6)): rCO = (0.86 + 0.14P

Reaction

)1.85

 ×10−5 e12.88−1855.5/T PCO PH2 O 1 −

PH2 PCO2 PH2 O PCO Keq

 (6)

It considers a slight dependence on the reaction pressure. Therefore, a slightly positive effect is expected also on the WGS reaction, characterized by no variation of mole number and, therefore, no effect is foreseen by thermodynamics.

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Fig. 6. TR CO conversion as a function of temperature at different PReaction values. Experimental data: Haldor Thopsoe catalyst (filled symbols); Engelhard catalyst (empty symbols). TREC values (dashed line).

CuO/CeO2 ) and for the latter the reaction was carried out at different feed pressures, in order to evaluate the influence of this parameter on the reaction conversion. The positive feed pressure effect on the kinetic, and as consequence, on the CO conversion is more visible at low temperatures, where the kinetic is very low and, thus, it is the rate determining step, whereas it is quite absent at temperatures higher than 280 ◦ C, where the reaction is faster but close to the thermodynamic equilibrium. This fact induces a decrement in the TR conversion that follows the TREC in an asymptotic way. In fact, a maximum in the CO conversion as a function of temperature was observed between 260 and 280 ◦ C (Fig. 6) for both the catalysts analyzed. 3.2.2. Membrane reactor The (Engelhard) catalyst (3.4 g) was packed on the opposite separating layer, in order to preserve the membrane top layer. Fig. 7 shows the MR CO conversion as a function of temperature at different feed pressures. The MR was successfully used in the WGS reaction; measured CO conversion is an increasing function of the temperature

3.2.1. Traditional reactor Both commercial catalysts were tested for the WGS reaction in the TR: (a) Haldor Thopsoe (low medium temperature CuO/ZnO) and (b) Engelhard (low medium temperature Table 4 Reaction tests operating conditions Temperature range (◦ C) Feed pressure range (absolute bar) Permeate pressure (bar) H2O/CO molar feed ratio (–) Space velocity (h−1 ) CO feed flow rate (cm3 (STP)/min)

220–300 2–6 1 1 2000 60

Fig. 7. MR CO conversion as a function of temperature at different PReaction . Experimental data (symbols). TREC values (dashed line). PPermeate = 1 bar; m = 1; SV = 2000 h−1 .

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Fig. 8. MR CO conversion as a function of PReaction at different temperature. Experimental data (symbols). TREC values (dashed lines). PPermeate = 1 bar; m = 1; SV = 2000 h−1 .

Fig. 9. Recovery factor of i-specie as a function of Pi at different temperatures. Experimental data: H2 (); CO (䊉); CO2 (). PPermeate = 1 bar; m = 1; SV = 2000 h−1 .

and always exceeds that of a TR and at a temperature higher than 250 ◦ C also the TREC limit. However, the conversion presents a maximum as a function of reaction pressure at a high (280 ◦ C) reaction temperature: at 6 bar it is lower than that at 4 bar. In fact, a pressure increase favours permeation, thus, if on the one hand higher H2 removal takes place from reaction volume, on the other hand reactant removal also occurs with a consequent lower CO conversion. This behaviour is more visible at a high temperature (Fig. 8) since temperature favours reaction and permeation: product permeation is not in competition with reaction, whereas reactant permeation competes with reaction, depleting the CO conversion. With the temperature the competitive permeation mechanism increases and, at 280 ◦ C and 6 bar, prevails over the reaction. The conversion difference between the MR and TR increases with temperature at the different reaction pressures studied. The increase with temperature at 6 bar is lower than that obtained at lower reaction pressure, confirming the previous considerations about the influence of pressure on selective permeation. The best operating condition for the MR was individuated at 280 ◦ C and 4 bar, obtaining a CO conversion increment of 8% with respect to TR and of 5.5% from the TREC limit, although the highest difference was individuated at 250 ◦ C and 6 bar. In order to evaluate the recovery capability of this MR, the recovery factor, defined as:

these species suggests the MR works well allowing H2 selective permeation, at least, 15 times higher than CO and CO2 . The H2 RF reaches 30% giving several advantages in terms of Process Intensification Strategy: it reduces the volume required by a factor of 1/3 and consequently the catalyst amount.

recovery factori = RFi =

FiPermeate Permeate Fi + FiRetentate

3.3. Permeation tests after reaction During the reaction tests the permeances and the separation factors of all the species were also evaluated (Table 5). In addition, other permeation tests with pure gas were carried out after the reaction at the same operating conditions of the permeation tests (before reaction) and with these compared in order to evaluate eventual changes in the membrane behaviour. A significant difference in permeances measured during and after the reaction tests with respect to those measured before the reaction was observed (Fig. 10). This is probably due to an incomplete stabilization of the membrane properties, achieved progressively during the reaction, as confirmed also by the agreement of the results obtained during the reaction with those obtained after the reaction tests, and also to structural changes in the sepaTable 5 Permeances, selectivities and separation factors measured during and after the reaction tests T (◦ C)

(7)

estimated from the reactor outlet streams, represents the species fraction permeated through the membrane. Fig. 9 reports the RF of each species as a function of the related driving force, the partial pressures difference between the retentate and permeate side. Two zones are individuated: the first includes a high RF at a low Pi and the second a low RF at a high driving force. The first area is typical of the highly permeable species like H2 ; whereas the second one is characteristic of those less permeable species such as CO and CO2 that permeate slowly also at a higher partial pressure difference. The distribution of

Permeance (nmol/m2 s Pa)

Separation factor (–)

H2

CO2

H2 /CO

H2 /CO2

0.4 0.6 1.8

28.5 21.6 21.5

26.2 22.2 15.4

CO

During the reaction 225 9.7 0.3 250 12.3 0.6 280 27.1 1.3 T (◦ C)

Permeance (nmol/m2 s Pa)

Selectivity (–)

H2 After the reaction 220 11.5 250 15.7 280 29.1

CO

CO2

H2 /CO

H2 /CO2

0.3 0.6 1.1

0.4 0.5 1.5

37.2 27.5 27.0

29.8 30.0 20.1

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force for the CO2 and H2 permeation is much higher than that of CO, reducing in this way the amount of this species in the permeate stream. Fig. 11 also shows the influence of the reaction pressure on the CO level in the permeate stream; it increases with pressure when the permeation driving force increases. 4. Conclusion

Fig. 10. Hydrogen experimental permeance as a function of the temperature.

rative layer for the exposure to water vapour [17,18] and to the modification of the support external area that changes its morphology (superficial metal oxide reduction) owing to a continuous contact with the H2 stream, producing micro fractures in the separating layer. The permeance and selectivities exhibited after the reaction always prove to be higher than those obtained during the reaction itself (Table 5). The presence of several gases in the retentate volume, also in this case, controls the permeation behaviour with hindrance effects, preferring the permeation of the gas with a low kinetic diameter (e.g., H2 ). The CO content in the MR permeate stream has specific importance (Fig. 11), in order to use the obtained hydrogen rich stream for feeding a PEMFC. It ranges from 1 to 10% depending on the operating conditions. Specifically, the CO level increases with temperature, registering a maximum at 250 ◦ C, at any pressure investigated. At low temperature both permeation and kinetics are very slow and they increase with a temperature increase. At higher temperatures both mechanisms are favoured and the CO conversion increases, reducing the CO amount in the reaction side available for the permeation, so increasing the amount of CO2 and H2 . Therefore, the driving force of each species is modified in dependence on its amount on the reaction side. At 280 ◦ C the conversion is the highest and the driving

The water gas shift (WGS) reaction in a catalytic membrane reactor (MR) was studied using a porous stainless steel supported silica membrane. The membrane was prepared with the soaking–rolling procedure, modifying the macropores of the support disk by packing silica xerogel (500 nm) and coating the intermediate layer of ␥-alumina in order to improve the H2 membrane selectivity. Permeance of all species (H2 : 9.7–29; CO: 0.3–1.1; CO2 : 0.4–1.5 nmol/m2 s Pa) and selectivity (H2 /CO, H2 /CO2 and H2 /N2 ) ranging from 15 to 40 and separation factors (H2 /CO = 20–45), showed no dependence on the permeation driving force. Differences between selectivity and separation factor were also registered. Furthermore, no inhibition effect of other gases on the hydrogen flux were observed. The MR was successfully used in WGS reaction, measured CO conversion is an increasing function of the temperature and always exceeds that of a traditional reactor (TR) and at a temperature higher than 250 ◦ C also the TREC (TR equilibrium conversion) limit. However, the conversion presents a maximum as a function of the reaction pressure at a high (280 ◦ C) reaction temperature: at 6 bar it is lower than that at 4 bar. In fact, a pressure increase favours permeation, thus, if on the one hand higher H2 removal takes place from the reaction volume, on the other hand the reactant removal also occurs with a consequent lower CO conversion. This behaviour is more visible at a high temperature since temperature favours the reaction and permeation: product permeation is not in competition with reaction, whereas reactant permeation competes with reaction, depleting the CO conversion. With temperature the competitive permeation mechanism increases and, at 280 ◦ C and 6 bar, prevails over the reaction. The recovery factor of each species and its dependence on the driving force were also studied, showing a species distribution with H2 at least 15 times higher than CO and CO2 . A CO content in the permeate stream ranging from 1 to 10% was measured depending on operating conditions. Silica membranes can be employed for the WGS reaction, obtaining conversions also higher than the TREC at a temperature higher than 250 ◦ C. However, the permeate H2 rich stream obtained with this membrane cannot be directly used as feed of a PEMFC, because the CO content is higher than the limit (10 ppm) required by the PEMFC catalysts, thus a further integrated purification step is necessary for the residual CO reduction. Acknowledgements

Fig. 11. CO content in permeate stream as a function of temperature at different PReaction .

The Italian Ministry for Foreign affairs, Direzione generale per la promozione e la Cooperazione Culturale, Rome, Italy

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is gratefully acknowledged for cofunding this research. Oleg M. Ilinitch (Engelhard, USA) and Poul Erik Hojlund Nielsen (Haldor-Topsoe, Norway) are gratefully acknowledged for the catalysts supplied. Appendix A. Nomenclature

J k Keq m P SF T

permeating flux (nmol/m2 s) kinetic constant (gcatalyst min/mol bar2 ) equilibrium constant H2 O/CO feed ratio pressure (bar) separation factor temperature (◦ C)

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