Modeling of a catalytic membrane reactor for CO removal from hydrogen streams – A theoretical study

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Modeling of a catalytic membrane reactor for CO removal from hydrogen streams e A theoretical study M. Teixeira a, L.M. Madeira a, J.M. Sousa a,b,*, A. Mendes a a b

LEPAE e Department of Chemical Engineering, Faculty of Engineering at the University of Porto, 4200-465 Porto, Portugal Department of Chemistry, University of Tra´s-os-Montes e Alto Douro, 5001-911 Vila-Real, Portugal

article info

abstract

Article history:

Typical industrial hydrogen streams arising from reforming processes contain about 1% of

Received 17 December 2009

carbon monoxide (CO). For fuel cell applications hydrogen should contain less than 10 ppm

Received in revised form

of CO, since it poisons the platinum catalysts in the electrodes. Traditionally, this is carried

10 April 2010

out through a selective oxidation reactor e PROX reactor. However, the parallel oxidation of

Accepted 17 April 2010

hydrogen to water should be avoided. This work proposes the use of a catalytic membrane

Available online 21 May 2010

reactor (MR) whose design is based on a CO permselective membrane containing the selective catalyst loaded in the permeate side. It is considered plug-flow pattern and

Keywords:

segregated feed of CO and oxygen. This strategy should improve the selective oxidation, as

Membrane reactor

the permselective membrane enhances the CO/H2 ratio at the catalyst surface.

Hydrogen

The combined process is analyzed using a mathematical model. The performances of the

CO removal

MR and of the PROX reactor at 393 K are compared in terms of selectivity towards CO

Fuel cell

oxidation and catalytic requirements to achieve the desired output CO concentration. The proposed MR is a promising alternative in performing the required purification. ª 2010 Professor T. Nejat Veziroglu. Published by Elsevier Ltd. All rights reserved.

1.

Introduction

Conventional hydrogen production technologies such as hydrocarbons reforming and partial oxidation of methane yield large amounts of CO as by-product [1,2]. When using hydrogen as source for fuel cells, the removal of CO is essential, since it acts as poison to platinum catalysts. Typical CO amounts after hydrocarbons reforming are 1 mol.% and in the case of low temperature Polymer Electrolyte Membrane Fuel Cells (PEMFC) hydrogen streams should have a CO concentration inferior to 10 ppm [3e5]. The removal of CO to ppm levels requires water gas-shift reactors (WGSR) and preferential oxidizers (PROX). In a PROX reactor the CO is catalytically oxidised to CO2 while the oxidation of H2 to water should be minimised [6].

Due to the recent growth in research on fuel cells and fuel processing, a large number of studies on CO preferential oxidation have been published [7e9]. An alternative approach for the removal of residual CO is the use of membranes. When functionalized with transition metals, these membranes can become highly permselective towards carbon monoxide. It has been reported that a strong bond between CO and Cuþ can be established, which can be inferred from the isotherms on several p-complexation sorbents [10e14]. Xie et al. [11] demonstrated also that hydrogen is virtually unadsorbed in these functionalized adsorbents, e.g. Cu(I)/NaY. This suggests that zeolite membranes from Cu(I)/NaY could be promising for CO removal from a hydrogen stream. The objective of this work is to assess the performance of a membrane reactor, which combines membrane

* Corresponding author. LEPAE e Department of Chemical Engineering, Faculty of Engineering at the University of Porto, 4200-465 Porto, Portugal. Tel.: þ351 22 508 1695; fax: þ351 22 508 1449. E-mail address: [email protected] (J.M. Sousa). 0360-3199/$ e see front matter ª 2010 Professor T. Nejat Veziroglu. Published by Elsevier Ltd. All rights reserved. doi:10.1016/j.ijhydene.2010.04.101

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permselectivity and catalytic oxidation selectivity in a single process unit. Indeed, by placing the catalyst on the membrane surface at the permeate side and feeding oxygen directly to the reaction side, at pressures lower than those of the retentate side, it is expectable that one may increase the driving force for CO permeation and the catalyst selectivity towards its oxidation. One undesired reaction which takes place alongside CO oxidation, and that should be minimized, is the hydrogen oxidation to water. In a conventional PROX reactor, for example, a catalytic oxidation selectivity of 50% equates to a loss of approximately 2.7% in fuel efficiency for a 2% CO reformate stream [15]. A large consumption of hydrogen is undesired since it decreases the power generation of the fuel cell [16].

1) Ideal gas behavior and isothermal conditions. 2) Plug-flow pattern (negligible axial and radial dispersion) and constant total pressures on both chambers. 3) The transport through the permselective membrane layer is considered to be described by an “overall” permeability coefficient, constant for all the operating conditions. The porous support does not impose any mass transfer resistance. 4) Homogeneous distribution of the catalytic particles over the permselective layer surface, which is considered to be part of the permeate chamber. 5) The reaction rate equations are considered to be valid for all the operating conditions. 6) Negligible film mass transport resistance in the membrane.

3. 2.

Mass balances

Model development The steady-state mass balances are presented in the following sections.

The catalytic membrane reactor considered in this study has the general features depicted in Fig. 1. It consists of retentate and permeate tubular chambers with constant pressure, separated by a hypothetical composite cylindrical membrane permselective towards CO and supporting a catalytic layer on the tube (permeate) side. A typical reformate mixture (50% H2, 16% CO2, 16% H2O, 1% CO and balance He) [17] is fed to the shell side (retentate chamber, high pressure side), while a pure oxygen stream is fed to the permeate side (low pressure side) in stoichiometric excess relative to CO. This study considers the catalytic oxidation of CO, as well as the deeper undesired oxidation of hydrogen e Eqs. (1) and (2).

CO þ 1/2O2 / CO2

(1)

H2 þ 1/2O2 / H2O

(2)

3.1. side




 2prs Li pRi  pPi 1 d Q R pRi þ ¼0 dz d
PR dQ R þ
Separation Layer

O2

   P
2prs Li pRi  pPi i

d

¼0

(3)

(4)

The respective boundary conditions are: and QR ¼ QF,h z ¼ 0, pRi ¼ pF,h i where Q is the volumetric flow rate, L is the permeability coefficient, P and p are total and partial pressures, respectively, rs is the shell radius, z is the axial coordinate, d is the thickness of the permselective layer, is the gas constant and T is the absolute temperature. The superscripts R, P and F refer to the retentate, permeate and feed stream conditions, respectively. The superscript h refers to the high pressure chamber (retentate) and the subscript i refers to the ith component.

The mathematical model proposed comprises the steadystate mass balance equations for the retentate and permeate sides, as well as the respective boundary conditions. The main assumptions are:

CO / H 2

Partial and total mass balances for the retentate

Retentate

PR

Support

Permeate

Catalytic Layer PP

Z Fig. 1 e Schematic diagram of the catalytic membrane reactor (MR).

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3.2. Partial and total mass balances for the permeate side

  2prt Li pRi  pPi 1 d Q P pPi Wcat

  2prt M ni;1 ri;1 d A

 PP d Q P i 
d

 2prt

PR

(10)

The respective boundary conditions are: F;h and Q R ¼ Q F;h z ¼ 0; pR i ¼ pi

ð5Þ

 Wcat X

ni;1 ri;1 AM i ð6Þ

The respective boundary conditions are: P F,l where the superscript l refers to z ¼ 0, pPi ¼ pF,l i and Q ¼ Q the low pressure (permeate) chamber. In Eqs. (5) and (6) (and in all the subsequent ones), vi,j is the stoichiometric coefficient of species i in reaction j ( j ¼ 1 for CO oxidation and j ¼ 2 for H2 oxidation), taken negative for reactants, positive for reaction products and null for the components that do not take part in the reaction. Wcat is the catalyst mass, AM is the permeation surface area of the membrane and rt is the membrane tube radius. It is considered that the thickness of the separation and catalytic layers are negligible; thus the tube radius is assumed to be equal to the shell radius (rt z rs). It was performed a literature review [17e24] on rate equations for both reactions (Eqs. (1) and (2)) and it was concluded that power-law type equations describe adequately these kinetics. So, it was decided to use the kinetic expressions determined by Lee and Kim [17] e Eqs. (7) and (8) below e because they were derived considering the CO and H2 oxidations simultaneously. This information refers to a CuO-CeO2 catalyst operating between 390e500 K and at a total pressure of 1 bar.




  0:37  0:62  0:91  94; 400
ri;1 ¼ 3:4  1010 exp pCO pCO2 pH2 O



(7)    0:48  0:69  142; 000 pH2 O ri;2 ¼ 6:1  1013 exp pH2 pCO2
It was also investigated the possible effect of the reverse water gas-shift (RWGS) reaction, but it was concluded that the catalyst considered was inactive towards the RWGS up to 500 K. This is also valid for other PROX catalysts [19,20,22,25].

3.3.2. Dimensionless partial and total mass balances for the permeate side




 d Q P pP i P  GLi pR  Da ni;1 f1 pP þ ni;2 Rr f2 pP ¼0 i  pi i i dz (11)

PP

X
X

 dQ P P G Li pR  Da ni;1 f1 pP i  pi i dz i i
P  þ ni;2 Rr f2 pi ¼0

ð12Þ

The respective boundary conditions are: F;l and Q P ¼ Q F;l z ¼ 0; pP i ¼ pi

where,

P 0:91  P 0:37  P 0:62 pCO2 pH2 O ¼ pCO f1 pP i

(13)


  P  P 0:48  P 0:69 pH2 O f2 pP ¼ pH2 pCO2 i

(14)

Da ¼

k1 Wcat
G¼

AM
Rr ¼

k2 0:09 P : k1 ref

The superscript * means dimensionless conditions and the subscript ref refers to reference conditions. k1 and k2 are the reaction rate constants in reactions (1) and (2), respectively, based on the reaction temperature (393 K) and f1 ðpP i Þ and f2 ðpP i Þ are the dimensionless rate equations. Da is the Damko¨hler number (based on the CO oxidation reaction), G is the dimensionless contact time (ratio between the maximum flux across the membrane for the reference component, that is, permeation of pure species against null permeate pressure, and its molar feed flow rate [26]), Rr is the ratio of the reaction rate constants and z* ¼ z/Z is the membrane dimensionless axial position (Z is the reactor axial length).

4. 3.3.

X
 dQ R P þG Li pR ¼0 i  pi dz i

Model parameters

Dimensionless mass balances

Eqs. (3)e(6) were made dimensionless with respect to the feed conditions (QF,h, PF,h) to component H2 ðLH2 Þ and to the axial dimension of the membrane (z):

3.3.1. Dimensionless partial and total mass balances for the retentate side



 d Q R pR i P þ GLi pR ¼0 i  pi dz

(9)

A systematic study on the effect of the most important operational variables and parameters was made. Table 1 summarizes the values/ranges of the parameters used in the simulations. In this study, it was considered a typical reformate stream [17] that is fed to the retentate side of the membrane reactor. Oxygen was fed to the permeate side at a lower pressure (Table 1), in order to maximize the driving force for CO permeation across the membrane and minimize the undesired reverse permeation of oxygen to the retentate. Moreover, it was considered a ratio between the molar flow rates of oxygen and CO, l, higher than the corresponding stoichiometric one, in order to avoid the depletion of oxygen

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104

0.002e0.05 1e1.5 10 e 50 0.01e0.8

G

PR Da Rr T

100 kPa 0.0001e0.2 5.6  104 393 K

in the reaction medium due to the undesired oxidation of hydrogen: l¼

2Q F;l pF;l O2 Q F;h pF;h CO

(15)

A hypothetical CO permselective membrane, such as a functionalized zeolite membrane, was considered. While such a membrane is yet to be assembled, copper-functionalized adsorbents with high adsorption selectivity towards CO (up to 393 K) have been reported [11]. A defect-free membrane synthesized from such a zeolite should be highly permselective to CO, since the high coverage of adsorbed species blocks the transport of the others [27]. The membrane permselectivity values towards CO, defined as the ratio between the permeability coefficients of CO and of the reference component, were considered in the range from 10 to 50 (cf. Table 1), based on the above mentioned adsorption selectivity data. The temperature value considered in this study for the MR was 393 K, at which CO adsorbs significantly in the membrane, thus blocking the permeation of hydrogen. As the temperature increases, the effect of the adsorption in the membrane decreases, thus becoming unreasonable to consider an adsorption controlled transport phenomena in the membrane. In order to compare the performances of the proposed membrane reactor and of the conventional PROX reactor, the extent of both oxidations (conversion) and the selectivity towards the desired reaction (CO oxidation) were calculated. The conversion Xi of component i can be expressed by: Xi ¼ 1 

R R Q P pP i þ Q pi F;l F;l F;h Q pi þ Q pF;h i

(16)

5.

(17)

Solution of the model equations

The general strategy used for solving the equations is the same as used previously [29]. A time derivative term was added to the right-hand side of Eqs. (9) and (11) transforming this problem into a pseudo-transient one. The resulting partial differential equations, combined with Eqs. (10) and (12), were spatially discretised using orthogonal collocation [30]. A suitable variable transformation of the collocation points in the axial direction as a function of the Damko¨hler number was applied, in order to obtain a solution with high accuracy and

λ=1-COmolar fraction λ=1.15-COmolar fraction

102

0.1

λ=1-O2 conversion λ=1.15-O2 conversion

0.01 0 .1 Da

Fig. 2 e CO molar fraction and oxygen conversion as a function of Da in a PROX reactor at 393 K, for two different l ratios. P [ 100 kPa.

low computational effort [31]. It was considered 11 internal collocation points in all simulations. The time integration routine LSODA [32] was then used to integrate the resulting set of equations until a steady-state solution was reached.

6.

Results and discussion

6.1.

PROX reactor

For simulation purposes, it was considered that the performance of a conventional plug-flow PROX reactor could be represented by the results of a MR equipped with a nonpermeable membrane, fed with the same stream and with the catalyst placed on the membrane retentate side. This assumption was assessed comparing the catalytic performance of the plug-flow reactor reported by Lee and Kim [17] and the MR for the experimental conditions they have employed. The experimental values of the CO conversion

1.0

CO Conversion/Selectivity

sCO

103

101 0.01

The selectivity for CO oxidation is defined as the ratio of net moles of CO reacted (or CO2 formed) per mole of oxygen reacted [28], adjusted for stoichiometry, and is given by:   R R F;h F;h 0:5 Q P pP pCO2 CO2 þ Q pCO2  Q   ¼  P P R R Q F;l pF;l O2  Q pO2 þ Q pO2

1

Oxygen Conversion

PP l LCO*

CO Concentration (ppm)

Table 1 e Parameter ranges used in the simulations.

0.8

0.6

0.4

Conversion Selectivity

0.2

0.0 0 .0 1

Da

0.1

Fig. 3 e CO conversion and selectivity in a PROX reactor as a function of Da. T [ 393 K; l [ 1.15; Feed: 100 mL/min, 1% CO, 1% O2, 50% H2, 16% CO2, 16% H2O, balance He; P [ 100 kPa.

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Fig. 4 e CO concentration in the retentate (A and B) and permeate (C and D), as a function of G, operating either in the very low Da region (A and C) or the high Da region (B and D), for LCO* ˛ [10e50]. PR [ 100 kPa, PP [0:005, l [ 1.15, T [ 393 K.

obtained by Lee and Kim [17] were compared with the simulated results for ambient pressure, 393 K and for different catalyst loads; it was verified that the model reproduces quite well the experimental data for the entire range of catalyst loads (data not shown). Fig. 2 shows the CO molar fraction and oxygen conversion as a function of Da in a PROX reactor, for two different l ratios. For l ¼ 1, oxygen is depleted in the reactor before the desired CO level of 10 ppm is achieved. So, higher values of l are required to meet the target CO concentration. After conducting a series of simulations, it was concluded that the optimum value of l e minimum value that guarantees a CO outlet maximum concentration of 10 ppm e is 1.15. The corresponding catalyst oxidation selectivity for these conditions is approximately 88% e Fig. 3.

6.2.

Membrane reactor (MR)

6.2.1.

Effect of permselectivity

Fig. 4 shows the CO concentration at the retentate and permeate sides as a function of the contact time; it illustrates the effect of the CO membrane permselectivity. Fig. 4A and C shows the retentate and permeate molar fractions of CO when operating in the very low Da region, while Fig. 4B and D illustrates the composition patterns of CO in both streams in the high Da region. When the reaction extent is negligible (very low Da), the compositions of the retentate and permeate exiting streams are determined mostly by the composition and flux of the permeating stream (membrane separation effect). High G values lead to high stage cut values (ratio between the permeation rate and the feed flow rate) and because the membrane is permselective towards CO such an increase

leads to retentate streams poorer in CO e Fig. 4A. Moreover, when more permselective membranes are used, the permeant flux of CO is also higher, thus decreasing its retentate fraction value and increasing its permeate concentration e Fig. 4A and C. Furthermore, Fig. 4C shows a maximum on the permeate composition of CO as a function of G. At low G values the permeating stream is highly enriched in CO, yet it becomes progressively impoverished for higher G values. Moreover, for high G values the effect of the membrane permselectivity on the permeate composition of CO decreases due to the depletion of CO on the retentate side. Nonetheless, the use of a simple membrane process is not enough to bring the CO

Fig. 5 e CO Concentration in the mixed stream (in ppm), as a function of G and for LCO* ˛ [20e50]. PR [ 100 kPa, PP [0:005, l [ 1.15, Da [ 0.08, T [ 393 K.

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A

B

103

105

104

102

103

101 100 10-1 10-2

0.0001

102 PP*=0.002 PP*=0.005 PP*=0.05

101

CO Permeate Concentration (ppm)

CO Retentate concentration (ppm)

104

100 0.01

Da

0.1

0.0001

0.01

0.1

Da

Fig. 6 e CO Concentration in the retentate (A) and permeate (B) as a function of Da, for PP ˛½0:002L0:05. G [ 0.5, LCO* [ 20, PR [ 100 kPa, l [ 1.15, T [ 393 K.

6.2.2.

Permeate pressure effect

This section aims to analyze the effect of the pressure of the permeate chamber on the performance of the MR. Fig. 6 shows the effect of Da on the composition of CO in the retentate (Fig. 6A) and permeate (Fig. 6B) streams, for different values of the dimensionless permeate pressure ðPP ˛½0:002e0:05Þ. As mentioned above, when the reaction extent is negligible (very low Da), the composition of CO in both streams depends solely on the composition and flux of the permeating stream. The use of lower pressure values on the reaction chamber translates into lower partial pressure values for all species, thus maximizing the permeation driving force for CO and minimizing the undesired reverse permeation of oxygen. Consequently, a smaller fraction of CO leaves the MR through the retentate e Fig. 6A. However, the permeating fluxes of H2, CO2 and H2O also increase. Thus, the composition of the permeating stream depends on the relative enhancement of the flux for each component. Because CO is the component fed in lower concentration it is subject to the largest permeation flux enhancement and therefore its concentration in the

CO Mixed Stream Concentration (ppm)

concentration to the 10 ppm threshold, regardless of the membrane permselectivity. In the high Da region, the composition of CO in the permeate (and therefore in the retentate) depends on the combined effect of permeation through the membrane and reaction on the permeate side. The enhancement of the CO conversion leads to permeate streams impoverished in CO (Fig. 4D). Thus, the permeation driving force for CO is enhanced, which translates into higher permeation fluxes. Consequently, the concentration of CO in the retentate also decreases (Fig. 4B), and the target concentration value of 10 ppm can be realized. Furthermore, when more permselective membranes are used, the flux of CO towards the permeate side becomes higher. Conversely, the rate of oxidation is favored by higher concentrations of CO (cf. Eq. (7)). The permeate molar fraction of CO is the result of the balance between these two factors; at high Da the more relevant effect is the oxidation of CO (except for very low G values), leading to lower molar fractions of CO in the permeate e Fig. 4D e and lower molar fraction values of CO in the retentate e Fig. 4B. It can be seen in Fig. 4D that for LCO* ¼ 50 the CO concentration is very low in both the retentate and permeate streams, clearly showing the synergetic effect of reaction and permeation for the effective CO removal in the MR. Nonetheless, the CO content in the permeate stream still remains above the desired 10 ppm level. To avoid using an additional separation unit for the purification of the permeate stream, it is proposed to mix the retentate and permeate streams (although this requires compressing the permeate stream up to the retentate pressure); this arrangement allows obtaining the required maximum CO concentration of 10 ppm. Fig. 5 exhibits the effect of the CO membrane permselectivity on the CO concentration in the mixed stream as a function of G, for Da ¼ 0.08. It can be concluded that, for Da ¼ 0.08, an MR achieves the desired mixed stream purity if it features a membrane with a minimum CO permselectivity of 40. It also shows that when more permselective membranes are used, the proposed MR reaches the desired stream purity at lower G values. This trade-off between G, Da and the membrane permselectivity is analyzed in Section 6.2.3.

104

103

102

101

100 0.0001

PP*=0.002 PP*=0.005 PP*=0.05

0.01

Da

0.1

Fig. 7 e CO concentration in the mixed stream as a function of Da, for PP ˛½0:002L0:05. G [ 0.5, LCO* [ 20, l [ 1.15, T [ 393 K.

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B

A

Fig. 8 e Minimum Da to achieve 10 ppm of CO on the mixed stream (A) and CO selectivity (B) as a function of G, for LCO* ˛ [15e50]. PP [0:005, l [ 1.15, T [ 393 K.

permeating stream increases (this increase is very small and can not be distinguished in Fig. 6B). For Da > 0.0002 the reactions conversion becomes noticeable on the CO concentration in the retentate and permeate. The extent of both oxidation reactions is favored at lower permeate pressures e Eqs. (13) and (14), yet the enhancement observed is greater for the oxidation rate of hydrogen. For Da < 0.1 the controlling effect is the enhancement in the oxidation of CO, and the best results are obtained for lower permeate pressures, PP ¼ 0:002. However, for Da > 0.1 the consumption of hydrogen increases rending the decrease in the molar fraction of CO to level off; in the conditions of Fig. 6 the best results are obtained for intermediate permeate pressures, PP ¼ 0:005. Because the permeate concentration of CO is higher, the composition of the mixed stream, shown in Fig. 7, follows the trend of the composition of the permeate stream. The range of interest of Da for the intended application is Da > 0.1, since it is when the mixed stream concentration of CO reaches 10 ppm. Thus, the optimum dimensionless permeate pressure value is PP ¼ 0:005 and it is used in the subsequent simulations.

6.2.3.

Damko¨hler, permselectivity and stage cut trade-off

The use of a MR targets to achieve the desired CO purity with low Da, G and permselectivity values. This trade-off is illustrated in Fig. 8A, where the minimum Da required to reach 10 ppm of CO on the mixed stream is plotted as a function of G, for various permselectivity values between 15 and 50. For comparison purposes, the Da required by the PROX reactor to reach the same performance is also plotted (cf. Fig. 2). Fig. 8B shows the comparison between the selectivity obtained with a PROX reactor and the MR, for the same sets of Da and G values. Fig. 8A shows that the MR exhibits lower catalytic requirements than the PROX reactor. The use of a permselective membrane raises the molar fraction of CO at the catalyst surface, enhancing the CO oxidation rate, thus the Da required to achieve the same reaction extent is smaller than for a PROX reactor. At the same time, in a MR, the hydrogen concentration at the catalyst surface is smaller leading to lower undesired hydrogen oxidation when compared to the PROX reactor e Fig. 8B. Moreover, it is observed in all curves the existence of a minimum G value (Gmin) for effective CO removal, dependent

on the membrane permselectivity. When higher permselectivity values are considered the desired concentration of CO can be attained at lower G. For example, a MR with LCO* ¼ 15 requires a minimum stage cut value of 0.45 while a MR with LCO* ¼ 50 can operate with a stage cut value of 0.15; this can be translated into a MR three times smaller.

7.

Conclusions

In this work it was simulated the CO removal from a reformate stream by selective oxidation, using a conventional PROX reactor and a MR equipped with a CO permselective membrane loaded with an oxidation catalyst on the permeate side. It was concluded that the PROX unit required a minimum ratio between the oxygen and carbon monoxide flows (l) of 1.15. to decrease the concentration of CO below the 10 ppm threshold (at 393 K). The effect of the membrane permselectivity on the performance of a membrane reactor was analyzed. It was concluded that operating with membranes that exhibit higher permselectivities enhances the permeation of CO to the reaction chamber, hence boosting the conversion of CO and lowering its retentate output molar fraction. It was seen that for Da > 0.1 very low concentrations of CO were observed in the output retentate stream. Nevertheless, the CO concentration in the permeate stream was above the desired 10 ppm level. To avoid using an additional separation unit for its purification, it was proposed to mix the retentate and permeate streams. While the membrane reactor was able to reach 10 ppm of CO on the mixed stream with a permselectivity value as low as 15 ðPP ¼ 0:005Þ, higher permselectivity values are required to effectively remove CO at stage cut values smaller than 0.45; a MR with permselectivity 50 performs the purification with a stage cut of 0.15. More permselective membranes also had lower catalytic activity requirements to achieve the desired purity. Nonetheless, it was observed that regardless the permselectivity, the catalytic requirements were always inferior to those of the conventional PROX reactor.

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The proposed MR showed to be a promising alternative to the current PROX technology. This indicates the need for developing CO permselective membranes for operating at temperatures around 400 K coupled with stable low temperature CO oxidation catalysts.

Acknowledgements Miguel Teixeira is grateful to the Portuguese Foundation for Science and Technology (FCT) for his doctoral grant (reference SFRH/BD/24768/2005). The authors also acknowledge financing from FCT through the projects POCTI/EQU/59344/ 2004 and POCTI/EQU/59345/2004.

Appendix. Nomenclature

AM Da k1 k2 L p P Q r Rr < T Wcat X z Z

membrane surface area (m2) Damko¨hler number () CO oxidation kinetic constant 1 kPa0.08) (mol kg1 catalyst s hydrogen oxidation kinetic constant 1 kPa0.17) (mol kg1 catalyst s permeability coefficient (mol m m2 s1 kPa1) partial pressure (kPa) total pressure (kPa) volumetric flow rate (m3 s1) radius (m) ratio of the reaction rate constants () gas constant (J mol1 K1) absolute temperature (K) catalyst mass (kg) conversion () axial coordinate (m) reactor length (m)

Greek symbols d selective layer membrane thickness (m) G dimensionless contact time () v stoichiometric coefficient () selectivity towards CO oxidation () sCO l molar inlet ratio between O2 and CO () Subscripts i species i j reaction j ref relative to the reference conditions (feed) or component (H2) Supersripts F relative to feed stream conditions h relative to high pressure stream conditions l relative to low pressure stream conditions P relative to permeate stream conditions R relative to retentate stream conditions s relative to the shell t relative to the tube * dimensionless variable

references

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i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 5 ( 2 0 1 0 ) 1 1 5 0 5 e1 1 5 1 3

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