Hydrogen production and carbon dioxide enrichment using a catalytic membrane reactor with Ni metal catalyst and Pd-based membrane

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Hydrogen production and carbon dioxide enrichment using a catalytic membrane reactor with Ni metal catalyst and Pd-based membrane Kyung-Ran Hwang, Chun-Boo Lee, Shin-Kun Ryi, Jong-Soo Park* Korea Institute of Energy Research, 71-2 Jang-dong Yuseong-gu, Daejeon 305-343, Republic of Korea

article info

abstract

Article history:

We prepared a catalytic membrane reactor (CMR) by adopting a high-performance metal

Received 21 November 2011

catalyst and PdeAu membrane to investigate the possibility of hydrogen production

Received in revised form

concurrently with carbon dioxide enrichment (up to >80%) in a single-stage reactor from

9 January 2012

a simulated syngas of a coal gasification, via simultaneous WGS reaction and hydrogen

Accepted 11 January 2012

separation process. The CO conversion was above 99% and the H2 recovery was above 94%

Available online 14 February 2012

at del-P ¼ 30 bar in a CMR. The best result for the concentration of the enriched CO2 in the

Keywords:

ratio of 2.0. These results show promise for a feasible simplified process able to achieve CO

Hydrogen

removal from a high-concentration CO mixture gas coming out of coal gasification via

Water-gas shift reaction

a water-gas shift reaction (WGS), to separate hydrogen and also to enrich CO2 for pre-

Catalytic membrane reactor

combustion capture and storage of CO2 (CCS) in substitution for the conventional WGS

retentate side was 85.3% under the conditions of 350  C, del-P ¼ 30 bar and steam to carbon

Carbon capture and storage

and CO2 separation stages in integrated gasification and combined cycle process integrated

Nickel catalyst

with CCS.

Pd-based membrane

Copyright ª 2012, Hydrogen Energy Publications, LLC. Published by Elsevier Ltd. All rights reserved.

1.

Introduction

Pre-combustion capture and storage of CO2 (CCS) from the syngas of integrated gasification and combined cycle (IGCC) power generation has recently received worldwide attention due to concerns about climate change and international cooperation to reduce CO2 emissions in the situation that coal is the dominant fossil fuel for electricity-generation [1]. An IGCC process integrated with CCS mainly consists of the following stages: coal gasification, syngas cleaning, water-gas shift (WGS) reaction, CO2 separation, and a combined cycle gas turbine. In the coal gasification, coal is converted into syngas, where syngas contains about 50e65% of CO [2]. The WGS reaction (CO þ H2O ¼ CO2 þ H2, DH298 ¼ 41.1 kJ/mol) is

an essential step to reduce CO and to maximize H2 production from the syngas, resulting in a reduction of the cost of CO2 capture. The CO2 is separated in a two-stage physical absorption reactor and is compressed for sequestration. The most important barrier for large-scale application of CCS is the loss in electric efficiency, which originates from the large steam requirement of the WGS reaction and the large energy consumption of the CO2 separation and compression. Here, decreasing the steam requirement of the WGS reaction is a significant option for reducing the loss in efficiency [3]. In practice, a large amount of steam is required to reduce CO via the WGS reaction, since coal gasification leads to a relatively high CO concentration in the syngas. Moreover, multi-stage WGS reactors and heat exchangers are needed to achieve

* Corresponding author. Tel.: þ82 42 860 3504; fax: þ82 42 860 3495. E-mail address: [email protected] (J.-S. Park). 0360-3199/$ e see front matter Copyright ª 2012, Hydrogen Energy Publications, LLC. Published by Elsevier Ltd. All rights reserved. doi:10.1016/j.ijhydene.2012.01.048

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the required conversion and also to remove a large amount of the heat of the reaction. Expanding the number of shift reactors and heat exchangers, however, increases the capital cost linearly. More recently, Marano et al. [4] reported that the integration of H2/CO2 separation membranes into the IGCC process could be considered as a future technology for CCS, since membranes are compact and modular, and could be placed at more than one location to separate H2 from CO2. Furthermore, they reported that enhanced CO conversion and reduced capital cost can be obtained by integrating the membrane separation with the WGS reaction in the same unit. This simultaneous WGS reaction and hydrogen separation process and the related key technologies such as membranes and module configuration have been intensively investigated [5e9] and it has been reported that the CO conversion exceeded the equilibrium conversion of a traditional reactor [7e9]. However, a relatively high CO concentration, about 50e65%, in the syngas from coal gasification was not used as the experimental conditions and also commercial WGS catalyst was applied for the simultaneous reaction and separation process, resulting in still presence of the heat transfer problem [10]. Therefore, a new high-performance WGS catalyst is needed to treat the high-concentration CO mixture gas from coal gasification in the minimum number of reactors and using the minimum steam quantity, while solving the heat transfer problem. We studied the feasibility of separating hydrogen from the 40% CO2 mixture gas for CO2 capture and hydrogen purification using our Pd-based membranes [11]. 90% of the hydrogen could be recovered from the gas mixture at a feed flow rate of 2 L/min, 400  C and 16 bar. The CO2 could be enriched by up to >80% for cost effective CCS by using our membrane and module design. In the present work, we designed and prepared a catalytic membrane reactor (CMR) adopting a highperformance metal catalyst and Pd-based membrane to treat a simulated syngas, 60% CO mixture gas, coming out of a coal gasification in a single step via a simultaneous WGS reaction and hydrogen separation process. Here, we focus on the possibility of hydrogen production concurrently with carbon dioxide enrichment (up to >80%) by using a CMR in a single stage.

2.

Experimental

2.1. Preparation and catalytic reaction test of nickel metal catalyst The plate-type nickel metal catalyst was prepared by potassium modification of the raw nickel material, uniaxial compression and sintering procedures. The potassium modified nickel powder [12] was carefully ground for 30 min and compressed without a binder in a cylindrical metal mould using a homemade press under high pressure (16 MPa). The compressed thin metal plate was further sintered at 700  C under H2 for 2 h to enhance its physical properties. The surface image of the prepared metal catalyst was obtained by scanning electron microscopy (FE-SEM, HITACHIS-4700). The plate-type catalyst was mounted on the standard specimen

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stubs with the help of double adhesive carbon tape and examined by FE-SEM without any pre-coating process. The porosity was measured by mercury intrusion porosimetry (Micromeretics, Autopore IV 9500). The catalyst was also characterized by X-ray powder diffraction using monoA) radiation (XRD, D/MAX IIIC). The chromic Cu Ka (l ¼ 1.54056  2 theta scans were carried out over the range 20 to 80 with a step size of 5 /min. The identification of the XRD pattern was identified by comparing with that included in the Joint Committee on Powder Diffraction Standards files (JCPDS). The WGS reaction test of the prepared metal catalyst was performed using a simulated syngas coming out of a coal gasification, namely, 60.0 vol.% CO, 36 vol.% H2 and 4 vol.% Ar [2]. The steam to carbon ratio (S/C) was 3.0. The prepared catalyst was placed in the unit-cell module reactor consisting of flanges and metal O-rings for the WGS reaction. These parts were tightened together with bolts and nuts through the holes on the edges of the flanges. Two metal O-rings were used to seal the reactor. After being assembled, it was installed in an electric furnace. A K-type thermocouple was fixed just on the upper surface of the catalyst in the reactor to adjust the temperature of the system. The effective area of the catalyst was 16.6 cm2. Reactants were passed vertically through the porous metal catalyst and the produced gases were collected through the channels placed below the catalyst. Argon was used as an internal standard. The feed and converted mixture gases were analyzed by an on-line gas chromatograph (GC6890, Agilent) equipped with a fused silica capillary column (Supelco-carboxen 1010) and a thermal conductivity detector. Before analysis, the produced gases were passed via a cold trap to remove liquid water. The CO conversion and CH4 selectivity were calculated from the following equations: CO conversionð%Þ ¼ ðmoles of CO consumed=moles of CO fedÞ  100 CH4 selectivityð%Þ ¼ ðmoles of CH4 produced = moles of CO consumedÞ  100

ð1Þ

ð2Þ

Here, methane was produced by methanation, which is an undesirable side-reaction that occurs during the WGS reaction at relatively high-temperature operations [13].

2.2. Preparation and permeation test of Pd-based membrane The plate-type PdeAu membrane was prepared by sputtering and sintering procedures. The Pd (4 mm) and Au (50 nm) were deposited in consecutive order over the pre-polished porous nickel metal support. After the PdeAu deposition, it was sintered at 700  C for 2 h. For more detailed information, refer to the previous work [11]. A surface image of the prepared membrane was obtained by FE-SEM (HITACHIS-4700). Hydrogen permeance tests of the prepared membrane were carried out at 350  C, 370  C and 390  C with no sweep gas in a membrane module [14]. The permeated hydrogen was metered by a digital flow meter (ADM 2000, Agilent Technologies) and the hydrogen flux (mol/m2/s) was calculated by dividing the permeated flow into an effective membrane area (16.6 cm2). The effect of the inhibitor, CO and steam, on the

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hydrogen permeance through the prepared membrane was analyzed by means of permeation tests at different CO (or steam) concentration (0e40 (v/v)%) in the mixture gas and pressures.

2.3.

WGS reaction test in a catalytic membrane reactor

2.3.1.

Preparation of a catalytic membrane reactor (CMR)

An outline (a) and photograph (b) of an assembled CMR are shown in Fig. 1. The CMR consisted of the plate-type nickel catalysts, Pd-based membrane, double o-rings, two flanges and bolts and nuts. The reactant that entered was passed vertically through the first catalyst and then the products were gathered in a collector placed just below the first catalyst. The collected products went into the space between the second catalyst and the membrane, where the second catalyst was the ancillary one for the WGS reaction. While the products were passed through the space, the WGS reaction was further carried out on the second catalyst and the hydrogen was separated through the membrane. The CO2-rich retentate gas was finally vented to the retentate port. Here, the distance between the second catalyst and the membrane was 0.6 mm and this narrower distance resulted in a faster linear flow velocity of the gas mixture, resulting in a reduction of the concentration polarization [11]. Several o-rings were used for the tight sealing of the reactor. The CMR was tightened with bolts and nuts on the edges of the flanges. The volume of the assembled CMR was about 480 mL. The assembled CMR was inserted in the high-pressure chamber, which takes two roles; (a) safer sealing of the CMR by the offsetting of the pressure difference between the outer and inner of the CMR and (b) preheating of reactant gases before they reach the catalysts and membrane.

2.3.2.

WGS reaction test in a CMR

The experimental schematic diagram for the WGS reaction in the CMR is presented in Fig. 2. The assembled CMR was installed in the high-pressure chamber and then set in the furnace. The reactant mixture was a simulated feed of the gas coming out of a coal gasification, namely, 60.0 vol.% CO, 36 vol.% H2 and 4 vol.% Ar. The steam to carbon ratio (S/C) was 3.0. The feed pressure (P2) was controlled using the backpressure regulator in the vent line of the retentate side and the pressure of the permeate (P1) side was always atmospheric pressure. Before the gas entered into a GC system, the liquid water was removed by a stainless steel cold-trap (Fig. 2(7)), which was connected to a chiller. The CO conversion, gas composition in the retentate side, and the hydrogen flow rate in the permeate side were analyzed using GC (Agilent 6890) and a digital flow meter (ADM 2000, Agilent Technologies), respectively. The hydrogen recovery was calculated from the following equation;

H2 recoveryð%Þ ¼

3.

Results and discussion

3.1.

Catalytic activity of the porous nickel metal catalyst

A photograph and SEM image (a) and an XRD pattern (b) of the prepared porous metal disc catalyst are shown in Fig. 3. The plate-type nickel metal catalyst with a depth of 1.65 mm and a diameter of about 51 mm had wormlike pores and its porosity as measured by mercury porosimetry was 29.8%. The XRD peaks of the catalyst at 2q ¼ 44.5, 51.8 and 76.4 were identified as the typical diffraction peaks of Ni metal and specific peaks for impregnated potassium (0.3 wt.%) were not observed. The CO conversion and CH4 selectivity for the WGS reaction under the conditions of gas-hourly space velocity (GHSV) ¼ 20 000 h1 and S/C ¼ 3.0 are shown in Fig. 4. The CO conversion was slightly reduced with an increase in temperature and the methane selectivity was below 1%. The CO conversion decreases with increasing temperature in the reactor since the WGS reaction is a moderately exothermic reversible reaction and at the used conditions conversion is closed to the equilibrium one. Commercially the WGS reaction is carried out in two adiabatic reactors, namely a hightemperature shift reactor followed by a low-temperature shift reactor, to take advantage of kinetics and thermodynamics. In the case of the high-temperature shift reactor, the inlet temperature is normally maintained at 350  C and this inlet temperature gives a maximum temperature of approximately 550  C at the exit, while the 10% CO concentration reduces to 3% [15,16]. In order to convert 60% of the CO concentration to CO2 and H2, a multi-stage WGS reactor and several heat-exchangers should be required, since a large amount of heat was produced by the reaction. However, in the present results, the CO conversion was 95.4% over the porous nickel metal catalyst in a single unit module at 380  C and 20 000 h1 and it followed the equilibrium conversion. This might be attributed to the fact that a metal catalyst having extremely high heat conductivity was used for the WGS reaction, resulting in a rapid release of the heat of reaction. The methane selectivity could be maintained below 1% in the relatively high-temperature regions due to the potassium modification of the nickel catalyst as reported in previous work [12]. The GHSV effects on the CO conversion and hydrogen production rate for the WGS reaction under the conditions of 400  C and S/C ¼ 3.0 are shown in Fig. 5. The CO conversions are slightly reduced from 94.7% to 94.3% and the hydrogen production rate increased linearly with an increase of GHSV. The conversions still followed the equilibrium conversion (94.2%) within the margin of error. The H2 and CO2 concentrations in the product gases were about 58.0 (v/v)% and 36.7 (v/v)%, respectively, and about 17.2 L/h of hydrogen was produced at the condition of 20000 h1. The carbon balance,

flowrate of H2 in permeate  100 flowrate of H2 in permeate þ flowrate of H2 in Retentate

(3)

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Fig. 1 e Outline (a) and photograph (b) of an assembled catalytic membrane reactor.

calculated from the number of carbon molecules of the reactant (CO) and product (CO2, CH4 and CO) for all experiments, was 1.0  0.02, indicating the absence of any carbon formation on the nickel catalyst or in the reactor. Consequently, these results show the possibility of the treatment of a high-concentration CO stream via a WGS reaction with the prepared porous metal catalyst in a singlestage reactor. The simultaneous reaction and hydrogen separation process adopting this high-performance metal catalyst will further produce the high-purity hydrogen in the permeate side and enrich the carbon dioxide in the retentate side, simultaneously.

3.2.

Hydrogen permeability of the PdeAu membrane

Although Pd membrane is highly hydrogen-permeable membrane capable of theoretically infinite selectivity, commercial application has been limited by the high cost of palladium, deactivation by carbon and sulfur compounds and hydrogen embrittlement. Therefore, several Pd alloy composite membranes have been intensively investigated. Among the Pd-based membranes, PdeAu membrane shows high tolerance to both sulfur [17,18] and carbon compounds [17,19] with high hydrogen permeability and selectivity. A photograph and SEM image of the prepared PdeAu membrane,

Fig. 2 e Experimental schematic diagram for the WGS reaction in the catalytic membrane reactor.

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Fig. 3 e Photograph and SEM image (a) and XRD pattern (b) of the prepared porous metal catalyst.

Fig. 4 e CO conversion (gray bar) and CH4 selectivity (dark gray bar) for the WGS reaction under the conditions of GHSV [ 20 000 hL1 and S/C [ 3.0 (equilibrium conversion (black dotted line) was estimated by Aspen Plus).

Fig. 5 e GHSV effects on the CO conversion (gray bar) and hydrogen production rate (black solid line) for the WGS reaction under the conditions of 400  C and S/C [ 3.0 (equilibrium conversion (black dotted line) was estimated by Aspen Plus).

which has a diameter of about 51 mm, are shown in Fig. 6. The thickness of the defect-free Pd-based membrane film supported by porous nickel was about 4 mm. The permeation test was conducted with a single gas of H2 as a function of transmembrane pressure difference at 350  C, 370  C and 390  C as displayed in Fig. 7. As expected, the hydrogen flux increased proportionally with the pressure difference across the membrane and temperature. For thick Pd membranes (Pd film thickness  10 mm), the rate-determining step in the permeation process is the bulk metal diffusion of atomic hydrogen, for which the n-value should be 0.5 in the equation of Sieverts’ Law [20]. With decreasing Pd membrane thickness, the n-value increases from 0.5 to 1 and the hydrogen dissociation adsorption on the surface of the Pd membrane, rather than bulk diffusion in the Pd layer, dominates the permeation rate of hydrogen [21]. As indicated by the graph, the linear regression of the data gave an n-value of 0.65, meaning that both the surface process and bulk diffusion are responsible for determining H2 permeation rate. H2 permeance, which is the slope  P0.65 ), was 1.01  104 (mol H2/m2/ in plots of H2 flux vs. (P0.65 2 1 0.65 2  s/Pa ) for 390 C and the R value for the best fit lines through

Fig. 6 e Photograph and SEM image of the prepared PdeAu membrane.

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Fig. 7 e Results of permeation test with H2 gas as a function of trans-membrane pressure difference at 350  C (white dot), 370  C (gray dot) and 390  C (black dot).

SigmaPlot for the flux vs. (P0.65  P0.65 ) data was 0.9999. The 2 1 temperature dependence of the H2 permeance (k) can be obtained by an Arrhenius-type relation between k0 and (1/T) [22] and an Arrhenius expression for the permeance of Pd-based membranes is presented below, where temperature is in Kelvin and R is the universal gas constant, 8.314 J/mol/K:   ð14:62kJ=molÞ 0 k ¼ 1:437  103 exp RT

(4)

H2/He selectivity was evaluated using hydrogen permeation and helium leakage tests and was 1500 at 390  C and 1 bar Table 1 summarized the H2 permeance, selectivity and activation energy of various Pd and Pd-based membranes. The permeance of the prepared PdeAu membrane was slightly lower, but the selectivity was higher than those of Pd and PdeAg membranes. The apparent activation energy, 14.62 kJ/

mol, obtained in our work is in good agreement with the reported values. Note that if bulk diffusion or both bulk diffusion and surface process is the rate-limiting step, the activation energy lies between 8 and 15 kJ/mol [21]. The CO and steam affect the hydrogen permeation of the membrane as inhibitors [32e34]. The effect of steam on the hydrogen permeating flux through the prepared Pd-based membrane at 400  C is shown in Fig. 8(a). The ideal value is the theoretical hydrogen flux, corresponding to the partial pressure of hydrogen in the mixture of steam and hydrogen. The deviation from the ideal value increased with increasing steam composition in the mixture. It may be caused by the competitive adsorption of steam and hydrogen on the active sites of the membrane surface, leading to a decrease in the actual permeation area of the Pd membrane. The inhibitor effect of steam was higher than that of CO at the high concentration in the mixture gas at low del-P. In our module design, which results in a faster linear flow velocity of the gas mixture due to the narrower distance between the catalyst and the membrane [11], the experimental data reached the theoretical hydrogen flux at feeding pressures higher than about 7.5 bar. Note that the experimental data (in Fig. 8) was within the margin of error. The effect of CO on the hydrogen permeating flux through the prepared Pd-based membrane at 400  C is shown in Fig. 8(b). The hydrogen flux reduced with the CO content at relatively lower feed pressures and the deviation from the ideal value for all CO mixtures was about 10% at higher feed pressure ranges (about 5 bare12 bar). Barbieri et al. [33] explained this phenomenon by a model equation, which is a complication of Sieverts and Langmuir’s adsorption laws and is dependent on the reduction factor and surface coverage of the inhibitor gas. For further increases of feed pressure (more than 12 bar), the ratio of the experimental value to the theoretical hydrogen flux seemed to be close to 1. Consequently, although CO and steam affected the hydrogen permeation of the membrane at relatively low system

Table 1 e Permeance, selectivity and activation energy of various Pd and Pd-based membranes. Membrane Pd/a-Al2O3 Pd/PSSb Pd46Cu54/ZrO2e PSS PdeCu/a-Al2O3 Pd/CeO2ePSS Pd/g-Al2O3eaAl2O3 PdeAg/PSS PdeAg/(pd)-gAl2O3/PSS Pd/Al2O3/PSS PdeAu/PNSe a b c d e

Preparation

Thickness (mm)

T ( C) 

Electroless plating þ osmosis techniques Electroplating Electroless plating

10.3 20 10

350 593e753

Electroless plating Electroless plating Electroless plating

11 13 15

450 550 400

Electroless plating Electroless plating

20e26 16

Electroless plating Sputtering

4.4 4

n: Pressure exponent. PSS: Porous stainless steel. Selectivity at 500  C. Selectivity at 8 bar. PNS: Porous nickel support.

na ()

Permeance (mol/m2/s/Pan)

0.65

3.7  106

0.5 1.0

Selectivity () at 1 bar

Ea (kJ/mol)

Ref.

1000 (H2/N2)

e

[23]

2.61  104 e

e e

e 15.4

[24] [25]

1.0 0.5 0.61

2.32  106 1.4  107 8.7  107

1400 (H2/N2) Infinity (H2/He) 1000 (H2/N2)

e 16.2 10

[26] [27] [28]

450 400

0.5 0.5

3.1  104 5.2  104

954 (H2/He) 300 (H2/N2)c

8e15 12.3

[29] [30]

500 390

0.5 0.65

2.94  103 1.01  104

1124 (H2/He)d 1500 (H2/He)

e 14.62

[31] This work

577

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pressure by acting as diluents, this effect could be overcome at higher feed pressures of more than about 15 bar in our module configuration.

3.3.

WGS reaction in a CMR

In general, the simultaneous WGS reaction and hydrogen separation process could obtain the enhanced CO conversion and recover the high-purity hydrogen. In the present work, carbon dioxide could also be simultaneously enriched in the retentate side from the high-concentration CO mixture gas in one step by using a high-performance metal catalyst and membrane. The CO conversion, CO2 concentration in the retentate side, and H2 recovery in the permeate side of a CMR under the conditions of S/C ¼ 3.0, GHSV ¼ 7000 h1 and 340  C as a function of del-P are shown in Fig. 9(a). The CO conversion increased from 98.6% to 99.1% with increasing system pressure difference across the membrane (0e25 bar), both of which values exceeded the equilibrium conversion, 96.9%. The CO2 concentration and the H2 recovery also increased with the pressure difference across the membrane. The initial CO2 concentration of about 40% was enriched to 76% and about 85.4% of hydrogen was recovered through the Pd-based

Fig. 8 e Steam (a) and CO (b) effect on the hydrogen permeating flux through the Pd-based membrane at 400  C (*: volume percentage in a mixture gas).

membrane in a CMR at del-P ¼ 25 bar. The GHSV effect on the CO conversion, CO2 concentration in the retentate side and H2 recovery in the permeate side in a CMR under the conditions of S/C ¼ 3.0, del-P ¼ 25 bar and 340  C are presented in Fig. 9(b). The CO conversion was affected by the residence time of the feed gas. The CO conversion decreased according to the reduction of hydrogen recovery. The performance of the simultaneous reaction and removal process was affected by at least the following interrelated factors: (a) reaction rate over the WGS catalyst, (b) hydrogen separation rate through the membrane and (c) hydrogen diffusion from the reaction zone to the surface of the membrane. In the present work, the last factor could be ignored because the CMR was designed to minimize the space between the metal catalyst and the membrane, resulting in a short H2 diffusion distance to the surface of the membrane. The first factor could be also discarded as a reason for the decrease of CO conversion at GHSV ¼ 20000 h1, since the porous metal catalyst could achieve equilibrium conversion at higher GHSV, as shown in

Fig. 9 e CO conversion (gray bar), CO2 concentration in the retentate side (white dotted line) and H2 recovery in the permeate side (black solid line) in a catalytic membrane reactor at S/C [ 3.0, GHSV [ 7000 hL1 and 340  C (a) and at S/C [ 3.0, del-P [ 25 bar and 340  C (b).

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Table 2 e Summary of the results of WGS reaction in a catalytic membrane reactor. GHSV (h1) 7000 7000 5000 5000 5000

del-P (bar)

Temp. ( C)

S/C ()

CO conversion (%)

H2 recovery (%)

CO2 concentration (%)

25 25 30 30 25

340 375 340 350 360

2.0 2.0 2.0 2.0 2.0

99.1 99.4 99.5 99.8 99.7

85.6 88.1 94.8 95.0 87.5

80.5 81.1 82.0 85.3 81.6

Fig. 4. Since the inhibitive effect of gases on the hydrogen permeation of the membrane increases with decreasing temperature [35], sufficient residence time for hydrogen separation through the membrane is required. Therefore, it is supposed that in the case of such a higher GHSV value, the membrane performance was a crucial factor in our CMR design. The CO conversion, H2 recovery and CO2 concentration under several experimental conditions are summarized in Table 2 in order to search for the conditions that can achieve over 80% enrichment of CO2 in the retentate side. The CO conversion values were all above 99% due to the high-pressure operation and the H2 recovery values through the membrane were above 94% at del-P ¼ 30 bar. The concentration of enriched CO2 was above 80% for all experimental conditions and the best result, 85.3%, was obtained at the condition of 350  C, del-P ¼ 30 bar and S/C ¼ 2.0. The carbon balance was 1.0  0.02 for all experiments.

4.

Conclusions

We designed and prepared a catalytic membrane reactor (CMR) adopting a high-performance metal catalyst and Pdbased membrane to investigate the possibility of hydrogen production concurrently with carbon dioxide enrichment (up to >80%) in a single step from a simulated syngas, a 60% CO mixture gas coming out of a coal gasification via simultaneous WGS reaction and hydrogen separation process. The porous nickel metal catalyst, having extremely high heat conductivity, showed a high-performance for WGS reactions under severe conditions, which followed thermodynamic equilibrium conversion. In our module design, the effect of CO and steam on the hydrogen permeation of the membrane could be overcome at relatively high feed pressures of more than about 15 bar and the H2 permeance of the prepared PdeAu membrane was 1.01  104 mol H2/m2/s/Pa0.65 for 390  C. The CO conversion, CO2 concentration in the retentate side and H2 recovery in the permeate side in a CMR as a function of del-P, GHSV and reaction temperature have been investigated. The CO conversion was above 99% and the H2 recovery was above 94% at del-P ¼ 30 bar in a CMR. The best result for the concentration of enriched CO2 was 85.3% at the conditions of 350  C, del-P ¼ 30 bar and S/C ¼ 2.0. These results show the promise of a feasible simplified process that is able to achieve CO removal from a high-concentration CO mixture gas coming out of coal gasification via WGS reaction, to separate pure hydrogen and to enrich CO2 for CCS, where the porous metal catalyst, membrane and reactor design seem to be key factors to substitute the conventional WGS reactor and CO2 separator in IGCC process integrated with CCS.

Acknowledgements The authors would like to acknowledge the financial support of the Korea Institute of Energy Technology Evaluation and Planning (KETEP) under “Energy Efficiency & Resources Programs” (Project No. 2010201010008B) of the Ministry of Knowledge Economy, Republic of Korea.

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