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Catalytic Pd0.77Ag0.23 alloy membrane reactor for high temperature water-gas shift reaction: Methane suppression
Chemical Engineering Journal 362 (2019) 116–125
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Catalytic Pd0.77Ag0.23 alloy membrane reactor for high temperature watergas shift reaction: Methane suppression Subhasis Pati, Ashok Jangam, Wang Zhigang, Nikita Dewangan, Wai Ming Hui, Sibudjing Kawi
T
⁎
Department of Chemical and Bio-Molecular Engineering, Engineering Drive-4, National University of Singapore, Singapore 117585, Singapore
H I GH L IG H T S
G R A P H I C A L A B S T R A C T
reactor prepared using Pd• Membrane Ag membrane packed with Ni-SiO 2
• • •
catalyst. Ultra thin (1 µm) Pd-Ag membrane showed high selectivity and adherence properties. CO conversion was enhanced and methanation was suppressed in CMR compared to FBR. The membrane showed excellent stability for 100 h of WGS reaction at 600 °C.
A R T I C LE I N FO
A B S T R A C T
Keywords: Pd-Ag membrane Catalytic membrane reactor High temperature water-gas shift reaction Methane suppression Hydrogen recovery
A catalytic membrane reactor was constructed using inner coated Pd0.77Ag0.23 (mass fraction) alloy membrane on Al2O3 hollow fiber substrate, prepared by electroless plating method and packed with Ni-Phyllosilicate catalyst for high temperature water gas shift (WGS) reaction. The performance of the ultra-thin ∼1 µm membrane was evaluated in a temperature range of 300–500 °C and hydrogen partial pressure of 50–200 kPa. The membrane showed high hydrogen permeance of 5.11 × 10−4 mol.m−2 s−1 Pa−0.65 at 100 kPa pressure and 500 °C. WGS reaction was carried out using the membrane reactor at temperature range of 400–500 °C, the effect of pressure, steam to carbon ratio and GHSV on CO conversion and methane formation were evaluated. Results showed increase in H2 permeation with increase in temperature and pressure, which eventually helps surpass the thermodynamic equilibrium CO conversion at high temperatures. H2 recovery and CO conversion decreased with increase in GHSV from 2896 h−1 to 11008 h−1. In this study, the shifting of equilibrium conversion and suppressed methane formation is also demonstrated at a temperature range of 400–600 °C. WGS reaction was performed in a catalytic membrane reactor at 600 °C for 100 h and it was found to be stable in terms of CO conversion and H2 recovery for the entire range of investigation. Finally, the membrane was tested for its performance using a reformate gas mixture similar to the gasification plant without any diluents and the conversion was found to be enhanced by 28% in terms of CO conversion at 500 °C using a membrane reactor.
1. Introduction Increase concern about the adverse effect of greenhouse gases on environmental conditions and climate change imparts emerging of
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alternative energy sources. Hydrogen being environmentally benign and a potential candidate of future energy carrier, attracted many attentions. Besides energy carrier hydrogen is a vital element for several chemical industries, preparation of semiconductors, desulphurization of
Corresponding author. E-mail address: [email protected] (S. Kawi).
https://doi.org/10.1016/j.cej.2018.12.112 Received 8 October 2018; Received in revised form 27 November 2018; Accepted 19 December 2018 Available online 20 December 2018 1385-8947/ © 2018 Published by Elsevier B.V.
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using a 29 µm thick Pd membrane and commercial Fe-Cr catalyst. They reported a shift in CO conversion from thermodynamic equilibrium value of 35% to 85.0% meanwhile achieved 82% of hydrogen recovery at 6 bar pressure. Tosti et al. [27] studied WGS reaction using 50 µm thick Pd-Ag foil at a temperature range of 325–330 °C and 100 kPa pressure and obtained CO conversion of 100% which is 20% higher than the equilibrium conversion. Mostly, the studied membrane reactor is focused on shifting the equilibrium within a temperature range of 300–450 °C and seldom reported information about the methane formation. Another crucial parameter which is not frequently reported is the stability of the membrane reactor under high temperature conditions. For combination of high temperature WGS reaction in a CMR with IGCC, its performance should be evaluated at high temperatures for a prolonged period of time. Further, the adverse effect of CO on the membrane surface can be suppressed by operating membrane at higher temperatures (> 350 °C) [28–30]. Both the conditions indicate the importance of high temperature WGS reaction in a membrane reactor and its stability over a longer duration. Pd0.77Ag0.23 alloy membrane is widely studied for hydrogen purification applications. Addition of Ag to Pd reduces the noble metal cost and hydrogen embrittlement problems. However, for this particular alloy composition Pd0.77Ag0.23 the hydrogen flux is higher than that of pure Pd membranes. Hence, in this study the alloy composition was chosen to be Pd0.77Ag0.23 for membrane reactor application. Most of the membranes studied in literature are either self-supported or composite membranes using tubular alumina/stainless steel substrate with the outer surface coated and used commercial Fe-Cr catalyst was used. In contrast, porous alumina hollow fiber substrates have advantages over other in terms of uniform pore distribution, higher packing density and high surface area. Use of the outer coated membranes in CMR causes diffusion of catalyst particles on the membrane surface and delamination of membrane from the support surface at high temperatures, causing a loss in performance of both catalyst activity and membrane permeance. In contrast, the inner coating helps to prevent the direct interaction of membrane material and catalyst, thereby preventing the diffusion of metal particles [31]. Maneerung et al. [32] also reported that the use of inner coated membrane results in better adherence properties compared to outer coated membranes. Herein, we report the performance of CMR for WGS reaction using Ni-Phyllosilicate catalyst at a temperature range of 400–600 °C. Ultrathin (1 µm) Pd0.77Ag0.23 alloy membrane coated in the inner side of Al2O3 hollow-fiber support used as a H2 permeable membrane. The performance of the membrane was evaluated at different temperature and trans-membrane pressure. Activity of catalytic membrane reactor was studied and compared with the conventional fixed bed reactor (FBR). To assess the applicability of the membrane reactor for high temperature WGS reaction, the stability of the catalytic membrane reactor was measured at 600 °C for 100 h.
oil, reduction of metal and food industries [1–3]. Presently about 48% of the hydrogen is produced from coal/fossil fuel gasification by steam reforming followed by water-gas shift (WGS) reaction [4]. The WGS reaction is one of the process for producing hydrogen from syn-gas by reaction with steam [5].
CO + H2 O↔ CO2 + H2 (Δr H 298 = −41.1 kJ/mol)
(1)
The reaction is moderately exothermic in nature, hence the equilibrium constant (Kp) decreases with temperature and becomes thermodynamically limited at high temperatures. However, the reaction kinetics is favoured at high temperature resulting in high throughput of hydrogen production. In the industries, the reaction is carried out in two steps; high temperature WGS reaction at 350–550 °C using Fe-Cr based catalyst to swiftly convert CO and a low-temperature shift reaction at 200–350 °C over Cu/ZnO/Al2O3 catalyst to get the equilibrium conversion [6–8]. Ni based catalysts are broadly studied for high-temperature WGS reactions in place of Fe-Cr catalysts which is toxic due to the Cr content and pyrophoric in nature in its reduced state [9–12]. Though, the Ni-based catalysts are highly active for WGS reaction, suppressing methane formation is still a challenge [13]. Methane formation takes place through hydrogenation of carbon species and consumes hydrogen which is a major product of WGS reaction [14].
CO + 3H2 ↔ CH 4 + H2 O (Δr H 298 = −206.1 kJ/mol)
(2)
CO2 + 4H2 ↔ CH 4 + 2H2 O (Δr H 298 = −165 kJ/mol)
(3)
2CO + 2H2 ↔ CH 4 + CO2 (Δr H 298 = −247 kJ/mol)
(4)
All the aforementioned reactions, illustrated different ways of methane formation. Yet, the methane formation should be avoided to enhance the overall performance of the system and improve the hydrogen recovery. This can be achieved by removing produced hydrogen during the reaction thereby decreasing the availability of H2 for further hydrogenation reaction. Several studies have been carried out in the past to suppress the methane formation reaction using the Ni-based catalyst. A study by Ashok et al. [15] reported suppressed methane formation using Ni-phyllosilicate catalyst. The performance of catalyst was found to be superior to the Ni-supported on SiO2 catalyst. The results were ascribed to the presence of the surface hydroxyl (–OH) group in Ni-phyllosilicate structure, which eventually decreases the methane formation rate. The combination of membrane with catalytic reactor often termed as catalytic membrane reactor (CMR) is an alternative for multi-stage WGS reaction and CO2 scrubbing [16–18]. Cooling of hot gas in a twostage WGS reaction adds capital cost to the process and decreases the power generation in integrated gasification combined cycle (IGCC). The combination of IGCC and CMR is beneficial in terms of performance and economic value. High temperature WGS reaction in a membrane reactor allows CO conversion to surmount the thermodynamic equilibrium limitations by selectively removing H2 from the product stream during the reaction. Catalytic membrane reactor operated at high temperatures benefits in following different ways (1) shifts the thermodynamic equilibrium, (2) producing high purity hydrogen at a faster rate, and (3) suppress the methane formation [16]. However, a highly permeable membrane with excellent selectivity is imperative for production of high purity hydrogen with enhanced throughput. Due to the unparallel potential of Pd-based membranes to selectively separate hydrogen from the stream containing impurities, it has been extensively studied for hydrogen purification applications and in catalytic membrane reactors for WGS reactions [19–24]. Bi et al. [25] studied WGS in a catalytic membrane reactor using a 1.4 µm thick Pd membrane supported on a porous ceramic support and Pt/Ce0.6Zro.4O2 catalyst at 325–400 °C and 12 bar pressure. They observed CO conversion higher than that of thermodynamic equilibrium value up to a high gas hourly space velocity (GHSV) of 14500 h−1 and a reduction in H2 recovery with a rise in GHSV. Pinacci et al. [26] studied WGS reaction at 410 °C
2. Experimental 2.1. Preparation of Pd-Ag membrane and Ni-phyllosilicate catalyst Pd-Ag alloy membrane was deposited in the inner side of the alumina-hollow fiber substrate by electroless plating method. The Al2O3 hollow fiber was prepared by phase inversion method previously reported by Sun et al. [33]. Presence of minor organic impurities or any passive layer on the substrate surface results in poor adherence and low membrane stability. Surface treatment in all plating process is a preliminary measure in order to achieve a homogenous, defect-free coating. This step involves cleaning and activation of the surface. Pre-treatment of calcined hollow fiber substrate was done as per the following procedure: a. Soaked in 1 M HNO3 for 6 h followed by ultrasonication in DI H2O b. Cleaned in an alkali solution containing 0.5 M NaOH, 0.5 M Na2CO3, 117
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Table 1 Electroless plating bath composition. Chemical
Pd-bath
Ag-Bath
PdCl2 AgNO3 Na2EDTA NH4OH (25%) N2H4(1 M) pH Temperature
2 g/l – 38 g/l 198 ml/l 5 ml/l 11 60 °C
– 0.38 g/l 38 g/l 198 ml/l 5 ml/l 11 60 °C
and 10 ml/L of soap solution, c. Finally, the hollow fiber substrate was thoroughly washed with DI water and dried overnight in the oven at 120 °C. Further, the substrate surface was activated by alternative dipping in SnCl2 and PdCl2 solution followed by intermediate rinsing in DI water. The process promotes growth of nano-layer of Pd over the substrate surface as per Eq. (5), the deposited Pd grains further helps in nucleation and growth of Pd during electroless plating.
Pd2 + + Sn2 + → Pd0 + Sn4 +
Fig. 1. Schematic of catalytic membrane reactor.
passing pure H2 and N2. Nitrogen permeation flux was also measured over a temperature range of 25–500 °C and 100 kPa pressure and maximum nitrogen flux of 1.91 × 10−4 mol m−2 s−1 was observed at 500 °C. Hydrogen permeation experiments were performed at a temperature range of 300–500 °C and trans-membrane pressure of 50–200 kPa using hydrogen and nitrogen mixture. The trans membrane pressure could not be increased from 200 kPa because of the limitations of the sealant used to join hollow-fiber with solid tube. Prior to membrane permeation, the membrane was activated by heating it at 400 °C under hydrogen atmosphere for 120 min. Helium was used as a sweep gas at a rate of 20 ml/min to create a partial pressure difference across the membrane. Concentration of permeated gas was analyzed online using gas chromatography (GC, HP 6890) using a thermal conductivity detector (TCD) and the flow rate was measured using an Agilent ADM digital flow meter.
(5)
Binary Pd-Ag alloy membrane was synthesized by sequential electroless plating method using bath composition and experimental condition given in Table 1. The inner side of the hollow fiber was coated using a peristaltic pump (Masterflex® L/S® Digital drive) equipped with a pump head for high performance tubes. The detailed procedure and the bath composition is reported elsewhere [32]. The Ag layer is sandwiched in between Pd layers. The deposited membrane was annealed at 500 °C under 1:1 H2 and He mixture for 8 h to ensure the formation of Pd-Ag solid solution. The formation of Pd-Ag solid solution was characterized by X-ray diffraction. The XRD patterns of the membrane were taken by grinding the membrane and separating the metal particles from the alumina powder. Surface morphology of the membrane was characterized by scanning electron microscopy (SEM) using (SEM, JEOL2872) at 15 keV electron beam energy and composition of the membrane was analysed by energy dispersive spectroscopy (EDS) analysis at different morphological positions.
2.3. WGS reaction in fixed bed and catalytic membrane reactor WGS reaction was carried out using the membrane reactor shown in Fig. 1. The reactor used is 8 mm ID and 150 mm length quartz tube. The membrane was sealed from both sides using high-temperature sealants with stainless steel tubes and the annular space around the membrane was filled with Ni-PS catalyst and packed with quartz wool from both the ends. HPLC pump (Shimadzu LC-20AT HPLC Pump) was used to feed water and mass flow controllers (Brook) were used to supply the feed gas mixture. A preheated assembly kept at 200 °C was used to evaporate liquid H2O to steam and the feed gas was mixed prior to the reaction chamber. WGS reaction was carried out using 50 mg of catalyst and feed gas composed of CO, H2O balance He at a gas hourly space velocity (GHSV) of 5793 h−1. Three different steam to carbon ratio of 1.5, 3 and 5 were used in this study and the effect of GHSV on CO conversion and H2 recovery was analysed. The un-reacted moisture after the reactor was collected using a cold trap maintained at 5 °C before the gas was analysed by HPGC equipped with a Hayesep-D column. The reactor was pressurised using a back pressure regulator connected after the cooling trap. CO conversion, CH4 formation and H2 recovery was calculated to evaluate the performance of membrane reactor. The CO conversion (Xco) was calculated using Eq. (6).
2.1.1. Preparation of Ni-Phyllosilicate catalyst WGS reaction was carried out using Ni-Phyllosilicate catalyst prepared using ammonia evaporation method reported by Ashok et al. [15]. Ni(NO3)2·6H2O and 28% NH3 solution were dissolved in de-ionized water in a molar ratio Ni:NH3 = 8 followed by adding colloidal silica 50% in water and stirred for 2 h. The whole complex was stirred at 60 °C for 7 h to evaporate all ammonia (pH = 7) and then precipitate was oven dried and calcined at 973 K for 5 h. The formation of phyllosilicate structure was analysed by X-ray diffraction (XRD) analysis using Shimadzu XRD-6000 diffractometer having Cu-Kα radiation operated at 40 kV voltage and 30 mA current. The samples were scanned at a 2 ϴ range of 20–60°at a rate of 2°/min. Calcined and reduced catalysts were also characterized by using high-resolution TEM using JEOL JEM-2100F instrument. 2.2. Membrane testing
X CO = [CO2 ]out − [CO2 ]in /{[CO]out + [CO2 ]out + [CH 4 ]out }
Synthesised membrane was tested using a gas permeation setup. The membrane used in this study is coated on hollow fiber having ID = 1.6 mm, length 100 mm and the surface area is 5.02 cm2. The membrane was joined with stainless steel tube using high temperature silicon sealant. The outer side of the membrane was filled with nitrogen at 200 kPa pressure and kept at 400 °C for 4 h and pressure drop was measured. A pressure drop of ≤5 mbar was observed for 4 h, which ensures the membrane to be defect free. Ideal hydrogen to nitrogen (H2/N2) selectivity was tested at 500 °C and 100 kPa pressure by
(6)
where, [CO2] out, [CO] out and [CH4] out are the molar flow rates of CO2, CO and CH4 in the outlet side respectively and [CO2] in, is the molar flow rate of CO2 in the inlet side. Similarly, the fraction of carbon converted to CH4 is calculated using Eq. (7).
X CH4 = [CH 4 ]out /{[CO]out + [CO2 ]out }
(7)
where, [CH4]out is the molar flow rate of CH4 in the outlet side of reactor. Hydrogen recovery in membrane reactor was calculated using the 118
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Fig. A3. The XRD pattern of calcined catalyst shows diffractions due to NiO and Ni-phyllosilicate phases. After reducing in hydrogen atmosphere for about 2 h, the NiO phase is converted to metallic Ni phase and the peak due to the phyllosilicate phase are absent. This may be attributed to the collapse of the PS structure during reduction and forms Ni metal embedded in silica matrix and this observations are similar to that reported by Ashok et al. [15]. The morphology of formed phyllosilicate structures was further confirmed by TEM. The TEM image of calcined catalyst is shown in Fig. A4. The needle-like characteristic phyllosilicate structures was observed from the TEM image. 3.3. Hydrogen permeation Nitrogen permeation flux was also measured over a temperature range of 25–500 °C and 100 kPa pressure to determine the selectivity of membrane. However, no nitrogen signal was detected at room temperature indicating a defect free membrane. Yet, maximum nitrogen flux of 1.91 × 10−4 mol m−2 s−1 was observed at 500 °C which is negligible compared to hydrogen flux. Hydrogen permeation flux through the membrane was evaluated at a temperature range of 300–500 °C. Flux at different temperatures and trans-membrane partial pressure of hydrogen is shown in Fig. 3. Hydrogen permeation through metallic membrane is a mass transport limited process and is dependent on temperature and pressure. Hydrogen flux through the membrane increased with increase in temperature and trans-membrane pressure of hydrogen as is evident from Fig. 3. However, the flux through the membrane follows a linear relation with ΔpH20.65. Here, the deviation of pressure exponent from 0.5 for very thin membrane (≤1µm) was obtained by fitting the value of flux as a function of feed pressure in a least squire regression method. The deviation of pressure exponent from n = 0.5 can be ascribed to the grain size, surface morphology or microstructure of the membrane. The hydrogen permeance obtained in this study along with the literature reported data is summarized in Table 3 which shows good agreement. For a better comparison of the literature reported data the unit of permeance is converted to mol m−2 s−1 Pa−1. Permeability was calculated using the formula given in Eq. (9).
Fig. 2. SEM image of Pd-Ag coated hollow fiber membrane (a) before annealing, (b) after annealing at 500 °C, and (c) cross-sectional image.
formula
%H2 (R) = 100 × H2(perm) /{H2(perm) + H2(feed)}
(8)
where, H2 (perm) and H2 (feed) are the molar flow rates of hydrogen in the permeate and feed side respectively. 3. Results and discussion 3.1. Characterization of the membrane The SEM images of plated Pd-Ag membrane before and after annealing are shown in Fig. 2. The figure indicates formation of homogeneous and dense Pd-Ag membrane on the alumina substrate. Some globular shapes and axial growth of particles on the surface can be attributed to the faster and dendritic growth of Ag on Pd surface. The difference in reduction potential of Pd and Ag causes a faster plating of Ag compared to Pd [34]. After annealing, no pinhole formation was noticed and increase in grain size was observed. The cross-sectional image of the membrane after annealing is shown in Fig. 2(c), it is apparent that the membrane is adherent to the substrate surface and the thickness of the membrane is ∼1 µm. The supported ultra-thin membrane is efficient in terms of performance and economic value for purification of hydrogen. EDS analysis at different morphological positions confirms the composition to be Pd0.77Ag0.23. The EDS image of the membrane is shown in Fig. A1. The composition of membrane was also calculated from the weight gain after each plating experiment; gravimetric analysis result is also in agreement with the EDS analysis value and is summarized in Table 2. The formation of Pd-Ag alloy phase was confirmed by XRD analysis of the membrane before and after annealing at 500 °C for 8 h. The XRD patterns are given in Fig. A2, which clearly indicates the individual metal peaks after plating Pd and Ag. However, after annealing at 500 °C for 8 h the individual peaks of Pd and Ag merged to form Pd-Ag solid solution phase.
Pe = F × L/ p10.65 − p20.65
(9) −1 −1
−0.65
), L is the thickness of where, Pe is permeability (mol m s Pa membrane (m), F is flux (mol m−2 s−1), p1 and p2 (Pa) are the partial pressure of hydrogen in the feed side and permeate side respectively.
3.2. Characterization of catalyst Formation of phyllosilicate structure and phase purity of the catalyst was analysed by TEM and XRD. The XRD spectrum of the prepared Ni-PS catalyst after calcination and reduction in hydrogen is shown in Table 2 Thickness and composition by different methods. Method
Composition
Thickness
Gravimetric SEM/EDS
Pd0.78Ag0.22 Pd0.77Ag0.23
1.3 µm ≤1 µm
Fig. 3. Variation of hydrogen flux with pressure at different temperatures. 119
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Table 3 Hydrogen permeation results and comparison with literature reported data. Membrane
Support material
Thickness (µm)
T/°C
ΔP /kPa
Selectivity (H2/N2)
Permeance/10−7 Mol m−2 s−1 Pa−1
Refs.
Pd Pd Pd Pd-Ni Pd0.59Cu0.41 Pd-Au Pd0.90Ag0.10 Pd0.77Ag0.23 Pd0.88Ag0.12 Pd-Ag Pd0.77Ag0.23
Al2O3 Al2O3 Al2O3 Al2O3 Self SS-ZrO2 Al2O3 YSZ- α-Al2O3 Al2O3 Hast-x Al2O3
2 7 0.5–1 1 16.7 10 2.2 1.2 11 4–5 1
500 400 450–500 500 400 400 600 500 550 400 500
100 100 100 68 120 100 100 100 413 100 100
2020 7500 100–1000 317 105 10,000 > 1000 1583 2000 200000 2380
22.1 23 5–10 21 10 14 40 57 12 10 45
[35] [36] [37] [38] [39] [40] [41] [32] [42] [43] This work
3.4. WGS activity in membrane reactor 3.4.1. Effect of temperature, steam to carbon ratio and GHSV on CO conversion and H2 recovery Pd0.77Ag0.23 membrane was tested for WGS reaction using different steam to carbon ratios (S/C) using Ni-phyllosilicate (Ni-PS) catalyst in a catalytic membrane reactor as shown in Fig. 1. Activity of the catalyst for WGS reaction was carried out using three different steam to carbon ratios of 1.5, 3 and 5 between 400 and 500 °C. The CO conversion as a function of temperature together with calculated equilibrium conversion values is shown in Fig. 5. According to figure the CO conversion is highest at 400 °C at all S/C conditions and decreased with increase in temperature. Highest CO conversion of 93.95% is achieved for S/C ratio = 5. The CO conversion for S/C ratio of 5 and 3, experimental value is less than the equilibrium conversion value at lower temperatures but approaches to equilibrium conversion at 500 °C. However, the CO conversion for S/C equals 1.5, approaches equilibrium at 400 °C and at higher temperature the value is higher than the equilibrium conversion. This can be accredited to the percentage of H2 recovered in the permeate side for different S/C ratios. Similarly, the CH4 formation for three different S/C ratios as a function of temperature is shown in Fig. 6. This figure shows that with decrease in steam to carbon ratio the formation of methane increased. CH4 formation in a CMR is below 0.26% at all reaction conditions, and the yield decreases to zero at 500 °C. Both the phenomena are related to the percentage of hydrogen recovery under different conditions. Hydrogen recovery (%H2 (R) in the permeate side as a function of temperature is shown in Fig. 7. The %H2 (R) recovered is highest in case of S/C = 1.5 as is evident from Fig. 7.
Fig. 4. Plot of ln(Pe) vs 1000 K/T at different trans-membrane pressure.
Permeability was found to increase with increase in temperature and follows Arrhenius type behaviour as given in Eq. (10).
Pe = P0 e−Ea / RT
(10)
Arrhenius plot ln(Pe) vs. 1/T was constructed to evaluate the activation energy of hydrogen permeation as shown in Fig. 4 and it was found to be 13.4 ± 0.1 kJ/mol. Hydrogen permeation through the membrane takes place by several steps like (a) physisorption of H2 on metal surface, (b) dissociative chemisorption of H2 to H atoms, (c) diffusion of hydrogen atom through the bulk metal and (d) association of hydrogen atoms on the permeate side and diffusion of H2 from the membrane surface [44–46]. Each process is kinetically controlled and contributes to the hydrogen permeation. Thus, the activation energy of hydrogen permeation depends on several factors such as thickness, surface morphology, microstructure and grain size etc. Ward and Dao [45] described a clear mechanical aspects of hydrogen permeation in the Pd membranes. Their calculations shows that diffusion of hydrogen to be the limiting step for clean Pd membranes of thickness up to 1 µm without presence of any mass transfer resistance and temperature above 300 °C. In the present study, the activation energy obtained also suggests that hydrogen permeation through the membrane is controlled by bulk diffusion of hydrogen [32]. The reproducibility of hydrogen permeation was tested using two more membranes. The membranes are denoted as M2 and M3, the hydrogen permeance results are shown in Supplementary Information in Fig. A5. The hydrogen flux of the membranes M2 and M3 at 100 kPa and 500 °C are 0.43 and 0.39 mol m−2 s−1 respectively. The characteristics of all three membranes are given in Table-A1 in the Supplementary Information.
Fig. 5. % CO conversion for different S/C ratios at different temperatures. 120
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Fig. 8. Effect of GHSV on CO conversion and H2 recovery in a CMR at 500 °C and 100 kPa pressure.
Fig.6. %CH4 formation as a function of temperature.
carbonaceous species as the ratio decreases. However, at higher steam condition the presence of excess steam dilutes concentration of available H2 and the CH4 formation decreases. With temperature CH4 formation decreased, which is due to the higher permeation of hydrogen through the Pd-Ag membrane at higher temperature and low hydrogenation of carbonaceous species. The effect of GHSV on the performance of membrane reactor was studied at 500 °C and 1 bar feed pressure at S/C = 1.5 and is shown in Fig. 8. The CO conversion was found to decrease from 88% to 83.6% with increasing the space velocity from 2896 h−1 to 11008 h−1 this may be due to decreased residence time of reactant species with the catalyst. However, at all conditions the CO conversion was above the equilibrium conversion of 81.7%. The H2 recovery also follows a similar decline with increase in GHSV and the value decreased from 40% to 14.2% with change in GHSV from 2896 h−1 to 11008 h−1. The H2 recovery can be increased by operating the CMR at high pressure conditions. Most importantly it was also observed that at very low GHSV the formation of methane also increased which was not seen at high flow rates. The increase in residence time also facilitates the unwanted methanation reaction which consumes the produced hydrogen from the reaction stream. Thus an optimised GHSV of 5793 h−1 was used in this study. It is noteworthy to mention here that, concentration polarization effect (CPE) also affects the hydrogen permeation and recovery, this phenomena happens due to the presence of non-permeating gases in the gas mixture. Boundary layer formation and the interaction of contaminating gases on the membrane surface are mainly responsible for CPE. When, the boundary layer diffusion of non-permeating gas to hydrogen is higher, the boundary layer effect is minimum, and the effect is minimized by operating at higher flow rates and high temperature [47,48]. In the present study, the inert diluent gas used is helium and He/H2 boundary layer diffusion coefficient is much higher, and the total GHSV used high enough for restricting the boundary layer formation. The polarization effect is also prominent when the temperature of operation is low and the interaction parameter for physorption of the contaminating species is a dominating factor. However, the interaction parameters of CO and CO2 are almost negligible above 400 °C [28–30,47]. Thus in this experiment the conditions such as high flow and high temperature of reaction is supposed to minimize the concentration polarization effect. WGS reaction is limited by thermal equilibrium and the CO conversion is penalized by increase in temperature but, increase of pressure has minimal impact on equilibrium conversion in a packed bed reactor. However, in a catalytic membrane reactor, the increase in pressure
Fig. 7. H2 recovery at different temperature for S/C = 5, 3 and 1.5.
Use of higher percentage of moisture dilutes the hydrogen in the feed side and decreases the partial pressure of hydrogen which ultimately decreases the hydrogen permeation. In this case for S/C = 1.5, removal of higher hydrogen from the product, facilitates the shifting of the thermodynamic equilibrium towards higher conversion which is in consistence with Le Chatelier's principle. Further, the CO conversion is supposed to increase with S/C, however removal of H2 from the product stream plays an important role in CO conversion for CMR. As mentioned above, that the H2 recovery is low for higher S/C ratios. This indicates that for S/C = 3 the removal of higher hydrogen helps in improving the CO conversion more than that for S/C = 5. This makes the CO conversion for S/C = 3 and S/C = 5 nominally close to each other. Decrease in H2 recovery with increase in S/C ratio was also reported by Augustine et al. [47]. They observed lower H2 recovery by using excess amount of H2O, and reported that the use of an optimum amount of steam to carbon ratio of 1.6 gives a high CO conversion and H2 recovery. Similarly, negative effect of excess moisture content on hydrogen permeation was also reported by Tosti et al. [27]. They explained the adsorption of H2O on the surface of membrane causes a reduction in hydrogen permeation sites and decreases the permeability. Further, the decrease in CH4 formation with increase in S/C ratio is due to the increased probability of interaction of hydrogen with
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Fig.10. % CO conversion in FBR and CMR for S/C = 1.5 as a function of temperature.
Fig. 9(a). Hydrogen recovery as a function of temperatures at different feed pressures.
conversion decreases with increase in temperature, however in the membrane reactor, percentage of hydrogen permeated through the membrane and the kinetics of reaction contributes towards shifting the equilibrium. The shift is significant at higher temperature and the CO conversion reaches 85.5% which is 12% higher than the thermodynamic equilibrium value of 73%. This shift is attributed to the higher rate of hydrogen permeation through the Pd0.77Ag0.23 membrane as a function of temperature. Increase in H2 recovery with temperature could also be due to reduced detrimental effect of CO on Pd-Ag membrane at higher temperatures (> 400 °C). Dissociative chemisorption of H2 on Pd-Ag alloy membrane surface is one of the central steps of hydrogen transport through the membrane, at lower temperature (< 350 °C) CO blocking on the surface penalize the hydrogen permeance [28–30]. However, the decremented effect of CO on hydrogen permeation reduces at higher temperature and also the rate of hydrogen diffusion increases with temperature [50]. Both the effect contributes to a higher rate of recovery of hydrogen from the reaction site and enhances the CO conversion. The concentration of permeated hydrogen was also measured at different temperatures during WGS reaction. The purity of the permeated hydrogen changed from ≥99.9% at 400 °C to ∼99.5% at 600 °C. Stability of the catalyst in FBR and CMR for 24 hrs was evaluated at 400 °C, % CO conversion and the methane formation was compared in both the cases shown in Fig. 11. The activity of catalyst in FBR and CMR
Fig.9(b). CO conversion at different temperatures and feed pressure.
increases the driving force for hydrogen permeation which shifts the equilibrium [49]. Recovery of hydrogen as a function of temperature at different pressures in a CMR is shown in Fig. 9(a) and CO conversion for respective conditions is shown in Fig. 9(b) and compared with the thermodynamic equilibrium calculated for 100 kPa pressure. Fig. 9(a) depicts higher H2 recovery with pressure at a particular temperature and with rise in temperature, the recovery is increased. The recovery of higher percentage of hydrogen results in increasing the CO conversion as evident from Fig. 9(b). The value of CO equilibrium conversion decreases linearly with increase in temperature, however, in case of membrane reactor small decrease in CO conversion is observed as the temperature increases. 3.4.2. Comparison of CO conversion and methane formation in a FBR and CMR High temperature WGS reaction was carried out at a temperature range of 400–600 °C in fixed bed reactor (FBR) and catalytic membrane reactor (CMR) using a steam to carbon ratio 1.5 and the results are compared in Fig. 10. It was observed that the CO conversion in the CMR is higher than the FBR and also the thermodynamic equilibrium conversion value at all experimental conditions. The thermodynamic CO
Fig. 11. Stability of catalyst in FBR and CMR for 24 h at 400 °C. 122
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Fig. 12. Stability of membrane reactor at 600 °C for 100 h. Fig. 13(a). % CO conversion as a function of temperature using reformate gas mixture (30% CO, 15% CO2, 25% H2 and 30% H2O).
was found to be stable for 24 h and the CO conversion is higher in CMR. However, methane formation was found to be suppressed in CMR compared to FBR. According to Eqs. 2–4, the methane formation depends upon the availability of H2 in the product and each mol of CH4 consumes 2–4 mol of hydrogen, which eventually decreases the hydrogen concentration in the product. In case of membrane reactor the removal of H2 from the product stream suppresses the methane formation and enhances the throughput of the reaction. The hydrogen permeance was found to be 1.1 × 10−4 mol m−2 s−1 Pa0.65 under the reaction conditions at 400 °C and was stable for 24 h of operation. The stability of CMR was tested for 100 h at high temperature of 600 °C using S/C ratio of 1.5 and is shown in Fig. 12. The performance of the membrane reactor was found to be stable over the entire period of investigation, however, a small decline in CO conversion is observed after 90 h. This may be due to some carbon deposition on the surface of the catalyst which blocks a few active surface sites for reaction. The H2 recovery was also monitored over the period of 100 h, the hydrogen permeability was found to be 2.5 × 10−4 mol m−2 s−1 Pa0.65, which is stable at around 43%. During the stability test at 600 °C, in the permeate side small CO peak was observed in the GC signal after ∼96 h of operation. This may be due to the formation of leakage through sealing at high temperature or formation of small defects on the thin membrane. Either of the causes are responsible for possible transport of CO through the membrane. Due to the possible leakage the purity of permeated hydrogen decreased from ≥99.5% to 98.6%.
Fig. 13(b). CH4 formation and H2 recovery as a function of temperature using reformate gas mixture (30% CO, 15% CO2, 25% H2 and 30% H2O).
from Fig. 13(b) that the formation of methane is drastically reduced from 7% in FBR to 2% in membrane reactor. 3.4.4. Post reaction membrane characterization The Pd-Ag membrane was characterized by SEM to analyse the change in morphology and stability after prolonged WGS reaction and heating-cooling cycles. The surface of the membrane was found to be stable and well adherent to the Al2O3 substrate even after prolonged heating at 600 °C for 100 h. The cross-sectional SEM image of the membrane after 100 h heating at 600 °C is shown in Fig. 14, which indicates better adherent property of inner coated membrane. It is a well-known fact that the difference in thermal expansion coefficient (α) between Pd-Ag (α = 10.24 × 10−6 K−1) and Al2O3 (1 × 10−6 K−1) [23] results in delamination of membrane from the substrate surface after few thermal cycles. However, this effect can be controlled by coating Pd-Ag membrane in the inner side of tube. The inner coated membrane expands along the direction the substrate at high temperatures and exerts a positive force on the substrate surface. Thus it remains attached with the substrate surface and resists to the thermal expansion [32]. Carbon deposition on the membrane surface was analysed by using thermo-gravimetric analysis. The membrane sample was
3.4.3. Effect of reformate gas on CO conversion in a FBR and CMR WGS reaction was performed using a reformate gas mixture of 30% CO, 15% CO2, 25% H2 and 30% H2O which is nearly similar with the outlet syn-gas composition of a gasifier plants [51,52]. The reaction was carried out at temperature range of 400–500 °C and 1 bar pressure, the CO conversion, H2 recovery and CH4 formation were evaluated. The CO conversion in CMR exceeds thermodynamic equilibrium and CO conversion in FBR, which is evident from Fig. 13(a). However, the shift in CO conversion is much higher compared to the previous cases, which, can be ascribed to the amount of higher concentration of hydrogen in the reaction mixture. Higher concentration of hydrogen in the reaction mixture increases the partial pressure of hydrogen in the feed side, which eventually increases the chemical potential of hydrogen. The increase in chemical potential in the feed side ultimately increases the driving force for hydrogen transport through the membrane and results in high hydrogen permeation [38]. Since higher quantity amount of hydrogen is permeated through the membrane to the permeate side, the probability of methane formation decreases in a CMR. This is evident 123
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Kawi, Ultra-thin (< 1 μm) internally-coated Pd–Ag alloy hollow fiber membrane with superior thermal stability and durability for high
Fig. 14. Cross sectional view SEM of Pd-Ag membrane after WGS reaction.
heated in a TGA instrument under air up to 800 °C with a ramping rate of 10 °C/min. The TGA graph of the membrane is shown in Fig. A6. The mass loss in TGA was analysed and result showed negligible carbon deposition on membrane. 4. Conclusions In the present study ultra-thin (∼1µm) Pd0.77Ag0.23 alloy membrane was synthesized by electroless plating method in the inner side of Al2O3 hollow fiber support. Hydrogen permeation characteristic of the membrane was evaluated at different temperatures and trans-membrane hydrogen pressure. Hydrogen permeance at 500 °C and 100 kPa was found to be 5.11 × 10−4 mol−1 m−2 Pa−0.65 and ideal H2/N2 selectivity of the membrane was found to be 2380 at similar conditions. WGS reaction was carried out in CMR at a temperature range of 400–600 °C and CO conversion, H2 recovery and CH4 formation were evaluated. CMR was found to be efficient in shifting the thermodynamic equilibrium for a steam to carbon ratio of 1.5. Use of excess steam was found to dilute the H2 concentration and reduced H2 recovery in the permeate side. At 600 °C, a stable CO conversion of 85.5% was achieved which is ∼12% higher than the equilibrium value of 73%. CO conversion was found to be enhanced and methane formation was suppressed in a CMR compared to a FBR. The CMR was found to be stable for 100 h at 600 °C. About 60% of hydrogen was recovered and methane formation was suppressed from 6.5% to 0.5% at 500 °C using a reformate gas mixture of 30% CO, 15% CO2, 25% H2 and 30% H2O. In terms of methane suppression, hydrogen production and CO conversion the CMR was found to have superior properties compared to FBR. Acknowledgements The authors generously thank financial support from Singapore Agency for Science, Technology and Research (A*STAR) AME IRG grant (No. A1783c0016) and National Environment Agency (NEA) of Singapore (WTE-CRP 1501-103). Appendix A. Supplementary data Supplementary data to this article can be found online at https:// doi.org/10.1016/j.cej.2018.12.112. References [1] K. Mazloomi, C. Gomes, Hydrogen as an energy carrier: prospects and challenges,
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Catalytic Pd0.77Ag0.23 alloy membrane reactor for high temperature watergas shift reaction: Methane suppression Subhasis Pati, Ashok Jangam, Wang Zhigang, Nikita Dewangan, Wai Ming Hui, Sibudjing Kawi
T
⁎
Department of Chemical and Bio-Molecular Engineering, Engineering Drive-4, National University of Singapore, Singapore 117585, Singapore
H I GH L IG H T S
G R A P H I C A L A B S T R A C T
reactor prepared using Pd• Membrane Ag membrane packed with Ni-SiO 2
• • •
catalyst. Ultra thin (1 µm) Pd-Ag membrane showed high selectivity and adherence properties. CO conversion was enhanced and methanation was suppressed in CMR compared to FBR. The membrane showed excellent stability for 100 h of WGS reaction at 600 °C.
A R T I C LE I N FO
A B S T R A C T
Keywords: Pd-Ag membrane Catalytic membrane reactor High temperature water-gas shift reaction Methane suppression Hydrogen recovery
A catalytic membrane reactor was constructed using inner coated Pd0.77Ag0.23 (mass fraction) alloy membrane on Al2O3 hollow fiber substrate, prepared by electroless plating method and packed with Ni-Phyllosilicate catalyst for high temperature water gas shift (WGS) reaction. The performance of the ultra-thin ∼1 µm membrane was evaluated in a temperature range of 300–500 °C and hydrogen partial pressure of 50–200 kPa. The membrane showed high hydrogen permeance of 5.11 × 10−4 mol.m−2 s−1 Pa−0.65 at 100 kPa pressure and 500 °C. WGS reaction was carried out using the membrane reactor at temperature range of 400–500 °C, the effect of pressure, steam to carbon ratio and GHSV on CO conversion and methane formation were evaluated. Results showed increase in H2 permeation with increase in temperature and pressure, which eventually helps surpass the thermodynamic equilibrium CO conversion at high temperatures. H2 recovery and CO conversion decreased with increase in GHSV from 2896 h−1 to 11008 h−1. In this study, the shifting of equilibrium conversion and suppressed methane formation is also demonstrated at a temperature range of 400–600 °C. WGS reaction was performed in a catalytic membrane reactor at 600 °C for 100 h and it was found to be stable in terms of CO conversion and H2 recovery for the entire range of investigation. Finally, the membrane was tested for its performance using a reformate gas mixture similar to the gasification plant without any diluents and the conversion was found to be enhanced by 28% in terms of CO conversion at 500 °C using a membrane reactor.
1. Introduction Increase concern about the adverse effect of greenhouse gases on environmental conditions and climate change imparts emerging of
⁎
alternative energy sources. Hydrogen being environmentally benign and a potential candidate of future energy carrier, attracted many attentions. Besides energy carrier hydrogen is a vital element for several chemical industries, preparation of semiconductors, desulphurization of
Corresponding author. E-mail address: [email protected] (S. Kawi).
https://doi.org/10.1016/j.cej.2018.12.112 Received 8 October 2018; Received in revised form 27 November 2018; Accepted 19 December 2018 Available online 20 December 2018 1385-8947/ © 2018 Published by Elsevier B.V.
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using a 29 µm thick Pd membrane and commercial Fe-Cr catalyst. They reported a shift in CO conversion from thermodynamic equilibrium value of 35% to 85.0% meanwhile achieved 82% of hydrogen recovery at 6 bar pressure. Tosti et al. [27] studied WGS reaction using 50 µm thick Pd-Ag foil at a temperature range of 325–330 °C and 100 kPa pressure and obtained CO conversion of 100% which is 20% higher than the equilibrium conversion. Mostly, the studied membrane reactor is focused on shifting the equilibrium within a temperature range of 300–450 °C and seldom reported information about the methane formation. Another crucial parameter which is not frequently reported is the stability of the membrane reactor under high temperature conditions. For combination of high temperature WGS reaction in a CMR with IGCC, its performance should be evaluated at high temperatures for a prolonged period of time. Further, the adverse effect of CO on the membrane surface can be suppressed by operating membrane at higher temperatures (> 350 °C) [28–30]. Both the conditions indicate the importance of high temperature WGS reaction in a membrane reactor and its stability over a longer duration. Pd0.77Ag0.23 alloy membrane is widely studied for hydrogen purification applications. Addition of Ag to Pd reduces the noble metal cost and hydrogen embrittlement problems. However, for this particular alloy composition Pd0.77Ag0.23 the hydrogen flux is higher than that of pure Pd membranes. Hence, in this study the alloy composition was chosen to be Pd0.77Ag0.23 for membrane reactor application. Most of the membranes studied in literature are either self-supported or composite membranes using tubular alumina/stainless steel substrate with the outer surface coated and used commercial Fe-Cr catalyst was used. In contrast, porous alumina hollow fiber substrates have advantages over other in terms of uniform pore distribution, higher packing density and high surface area. Use of the outer coated membranes in CMR causes diffusion of catalyst particles on the membrane surface and delamination of membrane from the support surface at high temperatures, causing a loss in performance of both catalyst activity and membrane permeance. In contrast, the inner coating helps to prevent the direct interaction of membrane material and catalyst, thereby preventing the diffusion of metal particles [31]. Maneerung et al. [32] also reported that the use of inner coated membrane results in better adherence properties compared to outer coated membranes. Herein, we report the performance of CMR for WGS reaction using Ni-Phyllosilicate catalyst at a temperature range of 400–600 °C. Ultrathin (1 µm) Pd0.77Ag0.23 alloy membrane coated in the inner side of Al2O3 hollow-fiber support used as a H2 permeable membrane. The performance of the membrane was evaluated at different temperature and trans-membrane pressure. Activity of catalytic membrane reactor was studied and compared with the conventional fixed bed reactor (FBR). To assess the applicability of the membrane reactor for high temperature WGS reaction, the stability of the catalytic membrane reactor was measured at 600 °C for 100 h.
oil, reduction of metal and food industries [1–3]. Presently about 48% of the hydrogen is produced from coal/fossil fuel gasification by steam reforming followed by water-gas shift (WGS) reaction [4]. The WGS reaction is one of the process for producing hydrogen from syn-gas by reaction with steam [5].
CO + H2 O↔ CO2 + H2 (Δr H 298 = −41.1 kJ/mol)
(1)
The reaction is moderately exothermic in nature, hence the equilibrium constant (Kp) decreases with temperature and becomes thermodynamically limited at high temperatures. However, the reaction kinetics is favoured at high temperature resulting in high throughput of hydrogen production. In the industries, the reaction is carried out in two steps; high temperature WGS reaction at 350–550 °C using Fe-Cr based catalyst to swiftly convert CO and a low-temperature shift reaction at 200–350 °C over Cu/ZnO/Al2O3 catalyst to get the equilibrium conversion [6–8]. Ni based catalysts are broadly studied for high-temperature WGS reactions in place of Fe-Cr catalysts which is toxic due to the Cr content and pyrophoric in nature in its reduced state [9–12]. Though, the Ni-based catalysts are highly active for WGS reaction, suppressing methane formation is still a challenge [13]. Methane formation takes place through hydrogenation of carbon species and consumes hydrogen which is a major product of WGS reaction [14].
CO + 3H2 ↔ CH 4 + H2 O (Δr H 298 = −206.1 kJ/mol)
(2)
CO2 + 4H2 ↔ CH 4 + 2H2 O (Δr H 298 = −165 kJ/mol)
(3)
2CO + 2H2 ↔ CH 4 + CO2 (Δr H 298 = −247 kJ/mol)
(4)
All the aforementioned reactions, illustrated different ways of methane formation. Yet, the methane formation should be avoided to enhance the overall performance of the system and improve the hydrogen recovery. This can be achieved by removing produced hydrogen during the reaction thereby decreasing the availability of H2 for further hydrogenation reaction. Several studies have been carried out in the past to suppress the methane formation reaction using the Ni-based catalyst. A study by Ashok et al. [15] reported suppressed methane formation using Ni-phyllosilicate catalyst. The performance of catalyst was found to be superior to the Ni-supported on SiO2 catalyst. The results were ascribed to the presence of the surface hydroxyl (–OH) group in Ni-phyllosilicate structure, which eventually decreases the methane formation rate. The combination of membrane with catalytic reactor often termed as catalytic membrane reactor (CMR) is an alternative for multi-stage WGS reaction and CO2 scrubbing [16–18]. Cooling of hot gas in a twostage WGS reaction adds capital cost to the process and decreases the power generation in integrated gasification combined cycle (IGCC). The combination of IGCC and CMR is beneficial in terms of performance and economic value. High temperature WGS reaction in a membrane reactor allows CO conversion to surmount the thermodynamic equilibrium limitations by selectively removing H2 from the product stream during the reaction. Catalytic membrane reactor operated at high temperatures benefits in following different ways (1) shifts the thermodynamic equilibrium, (2) producing high purity hydrogen at a faster rate, and (3) suppress the methane formation [16]. However, a highly permeable membrane with excellent selectivity is imperative for production of high purity hydrogen with enhanced throughput. Due to the unparallel potential of Pd-based membranes to selectively separate hydrogen from the stream containing impurities, it has been extensively studied for hydrogen purification applications and in catalytic membrane reactors for WGS reactions [19–24]. Bi et al. [25] studied WGS in a catalytic membrane reactor using a 1.4 µm thick Pd membrane supported on a porous ceramic support and Pt/Ce0.6Zro.4O2 catalyst at 325–400 °C and 12 bar pressure. They observed CO conversion higher than that of thermodynamic equilibrium value up to a high gas hourly space velocity (GHSV) of 14500 h−1 and a reduction in H2 recovery with a rise in GHSV. Pinacci et al. [26] studied WGS reaction at 410 °C
2. Experimental 2.1. Preparation of Pd-Ag membrane and Ni-phyllosilicate catalyst Pd-Ag alloy membrane was deposited in the inner side of the alumina-hollow fiber substrate by electroless plating method. The Al2O3 hollow fiber was prepared by phase inversion method previously reported by Sun et al. [33]. Presence of minor organic impurities or any passive layer on the substrate surface results in poor adherence and low membrane stability. Surface treatment in all plating process is a preliminary measure in order to achieve a homogenous, defect-free coating. This step involves cleaning and activation of the surface. Pre-treatment of calcined hollow fiber substrate was done as per the following procedure: a. Soaked in 1 M HNO3 for 6 h followed by ultrasonication in DI H2O b. Cleaned in an alkali solution containing 0.5 M NaOH, 0.5 M Na2CO3, 117
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Table 1 Electroless plating bath composition. Chemical
Pd-bath
Ag-Bath
PdCl2 AgNO3 Na2EDTA NH4OH (25%) N2H4(1 M) pH Temperature
2 g/l – 38 g/l 198 ml/l 5 ml/l 11 60 °C
– 0.38 g/l 38 g/l 198 ml/l 5 ml/l 11 60 °C
and 10 ml/L of soap solution, c. Finally, the hollow fiber substrate was thoroughly washed with DI water and dried overnight in the oven at 120 °C. Further, the substrate surface was activated by alternative dipping in SnCl2 and PdCl2 solution followed by intermediate rinsing in DI water. The process promotes growth of nano-layer of Pd over the substrate surface as per Eq. (5), the deposited Pd grains further helps in nucleation and growth of Pd during electroless plating.
Pd2 + + Sn2 + → Pd0 + Sn4 +
Fig. 1. Schematic of catalytic membrane reactor.
passing pure H2 and N2. Nitrogen permeation flux was also measured over a temperature range of 25–500 °C and 100 kPa pressure and maximum nitrogen flux of 1.91 × 10−4 mol m−2 s−1 was observed at 500 °C. Hydrogen permeation experiments were performed at a temperature range of 300–500 °C and trans-membrane pressure of 50–200 kPa using hydrogen and nitrogen mixture. The trans membrane pressure could not be increased from 200 kPa because of the limitations of the sealant used to join hollow-fiber with solid tube. Prior to membrane permeation, the membrane was activated by heating it at 400 °C under hydrogen atmosphere for 120 min. Helium was used as a sweep gas at a rate of 20 ml/min to create a partial pressure difference across the membrane. Concentration of permeated gas was analyzed online using gas chromatography (GC, HP 6890) using a thermal conductivity detector (TCD) and the flow rate was measured using an Agilent ADM digital flow meter.
(5)
Binary Pd-Ag alloy membrane was synthesized by sequential electroless plating method using bath composition and experimental condition given in Table 1. The inner side of the hollow fiber was coated using a peristaltic pump (Masterflex® L/S® Digital drive) equipped with a pump head for high performance tubes. The detailed procedure and the bath composition is reported elsewhere [32]. The Ag layer is sandwiched in between Pd layers. The deposited membrane was annealed at 500 °C under 1:1 H2 and He mixture for 8 h to ensure the formation of Pd-Ag solid solution. The formation of Pd-Ag solid solution was characterized by X-ray diffraction. The XRD patterns of the membrane were taken by grinding the membrane and separating the metal particles from the alumina powder. Surface morphology of the membrane was characterized by scanning electron microscopy (SEM) using (SEM, JEOL2872) at 15 keV electron beam energy and composition of the membrane was analysed by energy dispersive spectroscopy (EDS) analysis at different morphological positions.
2.3. WGS reaction in fixed bed and catalytic membrane reactor WGS reaction was carried out using the membrane reactor shown in Fig. 1. The reactor used is 8 mm ID and 150 mm length quartz tube. The membrane was sealed from both sides using high-temperature sealants with stainless steel tubes and the annular space around the membrane was filled with Ni-PS catalyst and packed with quartz wool from both the ends. HPLC pump (Shimadzu LC-20AT HPLC Pump) was used to feed water and mass flow controllers (Brook) were used to supply the feed gas mixture. A preheated assembly kept at 200 °C was used to evaporate liquid H2O to steam and the feed gas was mixed prior to the reaction chamber. WGS reaction was carried out using 50 mg of catalyst and feed gas composed of CO, H2O balance He at a gas hourly space velocity (GHSV) of 5793 h−1. Three different steam to carbon ratio of 1.5, 3 and 5 were used in this study and the effect of GHSV on CO conversion and H2 recovery was analysed. The un-reacted moisture after the reactor was collected using a cold trap maintained at 5 °C before the gas was analysed by HPGC equipped with a Hayesep-D column. The reactor was pressurised using a back pressure regulator connected after the cooling trap. CO conversion, CH4 formation and H2 recovery was calculated to evaluate the performance of membrane reactor. The CO conversion (Xco) was calculated using Eq. (6).
2.1.1. Preparation of Ni-Phyllosilicate catalyst WGS reaction was carried out using Ni-Phyllosilicate catalyst prepared using ammonia evaporation method reported by Ashok et al. [15]. Ni(NO3)2·6H2O and 28% NH3 solution were dissolved in de-ionized water in a molar ratio Ni:NH3 = 8 followed by adding colloidal silica 50% in water and stirred for 2 h. The whole complex was stirred at 60 °C for 7 h to evaporate all ammonia (pH = 7) and then precipitate was oven dried and calcined at 973 K for 5 h. The formation of phyllosilicate structure was analysed by X-ray diffraction (XRD) analysis using Shimadzu XRD-6000 diffractometer having Cu-Kα radiation operated at 40 kV voltage and 30 mA current. The samples were scanned at a 2 ϴ range of 20–60°at a rate of 2°/min. Calcined and reduced catalysts were also characterized by using high-resolution TEM using JEOL JEM-2100F instrument. 2.2. Membrane testing
X CO = [CO2 ]out − [CO2 ]in /{[CO]out + [CO2 ]out + [CH 4 ]out }
Synthesised membrane was tested using a gas permeation setup. The membrane used in this study is coated on hollow fiber having ID = 1.6 mm, length 100 mm and the surface area is 5.02 cm2. The membrane was joined with stainless steel tube using high temperature silicon sealant. The outer side of the membrane was filled with nitrogen at 200 kPa pressure and kept at 400 °C for 4 h and pressure drop was measured. A pressure drop of ≤5 mbar was observed for 4 h, which ensures the membrane to be defect free. Ideal hydrogen to nitrogen (H2/N2) selectivity was tested at 500 °C and 100 kPa pressure by
(6)
where, [CO2] out, [CO] out and [CH4] out are the molar flow rates of CO2, CO and CH4 in the outlet side respectively and [CO2] in, is the molar flow rate of CO2 in the inlet side. Similarly, the fraction of carbon converted to CH4 is calculated using Eq. (7).
X CH4 = [CH 4 ]out /{[CO]out + [CO2 ]out }
(7)
where, [CH4]out is the molar flow rate of CH4 in the outlet side of reactor. Hydrogen recovery in membrane reactor was calculated using the 118
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Fig. A3. The XRD pattern of calcined catalyst shows diffractions due to NiO and Ni-phyllosilicate phases. After reducing in hydrogen atmosphere for about 2 h, the NiO phase is converted to metallic Ni phase and the peak due to the phyllosilicate phase are absent. This may be attributed to the collapse of the PS structure during reduction and forms Ni metal embedded in silica matrix and this observations are similar to that reported by Ashok et al. [15]. The morphology of formed phyllosilicate structures was further confirmed by TEM. The TEM image of calcined catalyst is shown in Fig. A4. The needle-like characteristic phyllosilicate structures was observed from the TEM image. 3.3. Hydrogen permeation Nitrogen permeation flux was also measured over a temperature range of 25–500 °C and 100 kPa pressure to determine the selectivity of membrane. However, no nitrogen signal was detected at room temperature indicating a defect free membrane. Yet, maximum nitrogen flux of 1.91 × 10−4 mol m−2 s−1 was observed at 500 °C which is negligible compared to hydrogen flux. Hydrogen permeation flux through the membrane was evaluated at a temperature range of 300–500 °C. Flux at different temperatures and trans-membrane partial pressure of hydrogen is shown in Fig. 3. Hydrogen permeation through metallic membrane is a mass transport limited process and is dependent on temperature and pressure. Hydrogen flux through the membrane increased with increase in temperature and trans-membrane pressure of hydrogen as is evident from Fig. 3. However, the flux through the membrane follows a linear relation with ΔpH20.65. Here, the deviation of pressure exponent from 0.5 for very thin membrane (≤1µm) was obtained by fitting the value of flux as a function of feed pressure in a least squire regression method. The deviation of pressure exponent from n = 0.5 can be ascribed to the grain size, surface morphology or microstructure of the membrane. The hydrogen permeance obtained in this study along with the literature reported data is summarized in Table 3 which shows good agreement. For a better comparison of the literature reported data the unit of permeance is converted to mol m−2 s−1 Pa−1. Permeability was calculated using the formula given in Eq. (9).
Fig. 2. SEM image of Pd-Ag coated hollow fiber membrane (a) before annealing, (b) after annealing at 500 °C, and (c) cross-sectional image.
formula
%H2 (R) = 100 × H2(perm) /{H2(perm) + H2(feed)}
(8)
where, H2 (perm) and H2 (feed) are the molar flow rates of hydrogen in the permeate and feed side respectively. 3. Results and discussion 3.1. Characterization of the membrane The SEM images of plated Pd-Ag membrane before and after annealing are shown in Fig. 2. The figure indicates formation of homogeneous and dense Pd-Ag membrane on the alumina substrate. Some globular shapes and axial growth of particles on the surface can be attributed to the faster and dendritic growth of Ag on Pd surface. The difference in reduction potential of Pd and Ag causes a faster plating of Ag compared to Pd [34]. After annealing, no pinhole formation was noticed and increase in grain size was observed. The cross-sectional image of the membrane after annealing is shown in Fig. 2(c), it is apparent that the membrane is adherent to the substrate surface and the thickness of the membrane is ∼1 µm. The supported ultra-thin membrane is efficient in terms of performance and economic value for purification of hydrogen. EDS analysis at different morphological positions confirms the composition to be Pd0.77Ag0.23. The EDS image of the membrane is shown in Fig. A1. The composition of membrane was also calculated from the weight gain after each plating experiment; gravimetric analysis result is also in agreement with the EDS analysis value and is summarized in Table 2. The formation of Pd-Ag alloy phase was confirmed by XRD analysis of the membrane before and after annealing at 500 °C for 8 h. The XRD patterns are given in Fig. A2, which clearly indicates the individual metal peaks after plating Pd and Ag. However, after annealing at 500 °C for 8 h the individual peaks of Pd and Ag merged to form Pd-Ag solid solution phase.
Pe = F × L/ p10.65 − p20.65
(9) −1 −1
−0.65
), L is the thickness of where, Pe is permeability (mol m s Pa membrane (m), F is flux (mol m−2 s−1), p1 and p2 (Pa) are the partial pressure of hydrogen in the feed side and permeate side respectively.
3.2. Characterization of catalyst Formation of phyllosilicate structure and phase purity of the catalyst was analysed by TEM and XRD. The XRD spectrum of the prepared Ni-PS catalyst after calcination and reduction in hydrogen is shown in Table 2 Thickness and composition by different methods. Method
Composition
Thickness
Gravimetric SEM/EDS
Pd0.78Ag0.22 Pd0.77Ag0.23
1.3 µm ≤1 µm
Fig. 3. Variation of hydrogen flux with pressure at different temperatures. 119
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Table 3 Hydrogen permeation results and comparison with literature reported data. Membrane
Support material
Thickness (µm)
T/°C
ΔP /kPa
Selectivity (H2/N2)
Permeance/10−7 Mol m−2 s−1 Pa−1
Refs.
Pd Pd Pd Pd-Ni Pd0.59Cu0.41 Pd-Au Pd0.90Ag0.10 Pd0.77Ag0.23 Pd0.88Ag0.12 Pd-Ag Pd0.77Ag0.23
Al2O3 Al2O3 Al2O3 Al2O3 Self SS-ZrO2 Al2O3 YSZ- α-Al2O3 Al2O3 Hast-x Al2O3
2 7 0.5–1 1 16.7 10 2.2 1.2 11 4–5 1
500 400 450–500 500 400 400 600 500 550 400 500
100 100 100 68 120 100 100 100 413 100 100
2020 7500 100–1000 317 105 10,000 > 1000 1583 2000 200000 2380
22.1 23 5–10 21 10 14 40 57 12 10 45
[35] [36] [37] [38] [39] [40] [41] [32] [42] [43] This work
3.4. WGS activity in membrane reactor 3.4.1. Effect of temperature, steam to carbon ratio and GHSV on CO conversion and H2 recovery Pd0.77Ag0.23 membrane was tested for WGS reaction using different steam to carbon ratios (S/C) using Ni-phyllosilicate (Ni-PS) catalyst in a catalytic membrane reactor as shown in Fig. 1. Activity of the catalyst for WGS reaction was carried out using three different steam to carbon ratios of 1.5, 3 and 5 between 400 and 500 °C. The CO conversion as a function of temperature together with calculated equilibrium conversion values is shown in Fig. 5. According to figure the CO conversion is highest at 400 °C at all S/C conditions and decreased with increase in temperature. Highest CO conversion of 93.95% is achieved for S/C ratio = 5. The CO conversion for S/C ratio of 5 and 3, experimental value is less than the equilibrium conversion value at lower temperatures but approaches to equilibrium conversion at 500 °C. However, the CO conversion for S/C equals 1.5, approaches equilibrium at 400 °C and at higher temperature the value is higher than the equilibrium conversion. This can be accredited to the percentage of H2 recovered in the permeate side for different S/C ratios. Similarly, the CH4 formation for three different S/C ratios as a function of temperature is shown in Fig. 6. This figure shows that with decrease in steam to carbon ratio the formation of methane increased. CH4 formation in a CMR is below 0.26% at all reaction conditions, and the yield decreases to zero at 500 °C. Both the phenomena are related to the percentage of hydrogen recovery under different conditions. Hydrogen recovery (%H2 (R) in the permeate side as a function of temperature is shown in Fig. 7. The %H2 (R) recovered is highest in case of S/C = 1.5 as is evident from Fig. 7.
Fig. 4. Plot of ln(Pe) vs 1000 K/T at different trans-membrane pressure.
Permeability was found to increase with increase in temperature and follows Arrhenius type behaviour as given in Eq. (10).
Pe = P0 e−Ea / RT
(10)
Arrhenius plot ln(Pe) vs. 1/T was constructed to evaluate the activation energy of hydrogen permeation as shown in Fig. 4 and it was found to be 13.4 ± 0.1 kJ/mol. Hydrogen permeation through the membrane takes place by several steps like (a) physisorption of H2 on metal surface, (b) dissociative chemisorption of H2 to H atoms, (c) diffusion of hydrogen atom through the bulk metal and (d) association of hydrogen atoms on the permeate side and diffusion of H2 from the membrane surface [44–46]. Each process is kinetically controlled and contributes to the hydrogen permeation. Thus, the activation energy of hydrogen permeation depends on several factors such as thickness, surface morphology, microstructure and grain size etc. Ward and Dao [45] described a clear mechanical aspects of hydrogen permeation in the Pd membranes. Their calculations shows that diffusion of hydrogen to be the limiting step for clean Pd membranes of thickness up to 1 µm without presence of any mass transfer resistance and temperature above 300 °C. In the present study, the activation energy obtained also suggests that hydrogen permeation through the membrane is controlled by bulk diffusion of hydrogen [32]. The reproducibility of hydrogen permeation was tested using two more membranes. The membranes are denoted as M2 and M3, the hydrogen permeance results are shown in Supplementary Information in Fig. A5. The hydrogen flux of the membranes M2 and M3 at 100 kPa and 500 °C are 0.43 and 0.39 mol m−2 s−1 respectively. The characteristics of all three membranes are given in Table-A1 in the Supplementary Information.
Fig. 5. % CO conversion for different S/C ratios at different temperatures. 120
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Fig. 8. Effect of GHSV on CO conversion and H2 recovery in a CMR at 500 °C and 100 kPa pressure.
Fig.6. %CH4 formation as a function of temperature.
carbonaceous species as the ratio decreases. However, at higher steam condition the presence of excess steam dilutes concentration of available H2 and the CH4 formation decreases. With temperature CH4 formation decreased, which is due to the higher permeation of hydrogen through the Pd-Ag membrane at higher temperature and low hydrogenation of carbonaceous species. The effect of GHSV on the performance of membrane reactor was studied at 500 °C and 1 bar feed pressure at S/C = 1.5 and is shown in Fig. 8. The CO conversion was found to decrease from 88% to 83.6% with increasing the space velocity from 2896 h−1 to 11008 h−1 this may be due to decreased residence time of reactant species with the catalyst. However, at all conditions the CO conversion was above the equilibrium conversion of 81.7%. The H2 recovery also follows a similar decline with increase in GHSV and the value decreased from 40% to 14.2% with change in GHSV from 2896 h−1 to 11008 h−1. The H2 recovery can be increased by operating the CMR at high pressure conditions. Most importantly it was also observed that at very low GHSV the formation of methane also increased which was not seen at high flow rates. The increase in residence time also facilitates the unwanted methanation reaction which consumes the produced hydrogen from the reaction stream. Thus an optimised GHSV of 5793 h−1 was used in this study. It is noteworthy to mention here that, concentration polarization effect (CPE) also affects the hydrogen permeation and recovery, this phenomena happens due to the presence of non-permeating gases in the gas mixture. Boundary layer formation and the interaction of contaminating gases on the membrane surface are mainly responsible for CPE. When, the boundary layer diffusion of non-permeating gas to hydrogen is higher, the boundary layer effect is minimum, and the effect is minimized by operating at higher flow rates and high temperature [47,48]. In the present study, the inert diluent gas used is helium and He/H2 boundary layer diffusion coefficient is much higher, and the total GHSV used high enough for restricting the boundary layer formation. The polarization effect is also prominent when the temperature of operation is low and the interaction parameter for physorption of the contaminating species is a dominating factor. However, the interaction parameters of CO and CO2 are almost negligible above 400 °C [28–30,47]. Thus in this experiment the conditions such as high flow and high temperature of reaction is supposed to minimize the concentration polarization effect. WGS reaction is limited by thermal equilibrium and the CO conversion is penalized by increase in temperature but, increase of pressure has minimal impact on equilibrium conversion in a packed bed reactor. However, in a catalytic membrane reactor, the increase in pressure
Fig. 7. H2 recovery at different temperature for S/C = 5, 3 and 1.5.
Use of higher percentage of moisture dilutes the hydrogen in the feed side and decreases the partial pressure of hydrogen which ultimately decreases the hydrogen permeation. In this case for S/C = 1.5, removal of higher hydrogen from the product, facilitates the shifting of the thermodynamic equilibrium towards higher conversion which is in consistence with Le Chatelier's principle. Further, the CO conversion is supposed to increase with S/C, however removal of H2 from the product stream plays an important role in CO conversion for CMR. As mentioned above, that the H2 recovery is low for higher S/C ratios. This indicates that for S/C = 3 the removal of higher hydrogen helps in improving the CO conversion more than that for S/C = 5. This makes the CO conversion for S/C = 3 and S/C = 5 nominally close to each other. Decrease in H2 recovery with increase in S/C ratio was also reported by Augustine et al. [47]. They observed lower H2 recovery by using excess amount of H2O, and reported that the use of an optimum amount of steam to carbon ratio of 1.6 gives a high CO conversion and H2 recovery. Similarly, negative effect of excess moisture content on hydrogen permeation was also reported by Tosti et al. [27]. They explained the adsorption of H2O on the surface of membrane causes a reduction in hydrogen permeation sites and decreases the permeability. Further, the decrease in CH4 formation with increase in S/C ratio is due to the increased probability of interaction of hydrogen with
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Fig.10. % CO conversion in FBR and CMR for S/C = 1.5 as a function of temperature.
Fig. 9(a). Hydrogen recovery as a function of temperatures at different feed pressures.
conversion decreases with increase in temperature, however in the membrane reactor, percentage of hydrogen permeated through the membrane and the kinetics of reaction contributes towards shifting the equilibrium. The shift is significant at higher temperature and the CO conversion reaches 85.5% which is 12% higher than the thermodynamic equilibrium value of 73%. This shift is attributed to the higher rate of hydrogen permeation through the Pd0.77Ag0.23 membrane as a function of temperature. Increase in H2 recovery with temperature could also be due to reduced detrimental effect of CO on Pd-Ag membrane at higher temperatures (> 400 °C). Dissociative chemisorption of H2 on Pd-Ag alloy membrane surface is one of the central steps of hydrogen transport through the membrane, at lower temperature (< 350 °C) CO blocking on the surface penalize the hydrogen permeance [28–30]. However, the decremented effect of CO on hydrogen permeation reduces at higher temperature and also the rate of hydrogen diffusion increases with temperature [50]. Both the effect contributes to a higher rate of recovery of hydrogen from the reaction site and enhances the CO conversion. The concentration of permeated hydrogen was also measured at different temperatures during WGS reaction. The purity of the permeated hydrogen changed from ≥99.9% at 400 °C to ∼99.5% at 600 °C. Stability of the catalyst in FBR and CMR for 24 hrs was evaluated at 400 °C, % CO conversion and the methane formation was compared in both the cases shown in Fig. 11. The activity of catalyst in FBR and CMR
Fig.9(b). CO conversion at different temperatures and feed pressure.
increases the driving force for hydrogen permeation which shifts the equilibrium [49]. Recovery of hydrogen as a function of temperature at different pressures in a CMR is shown in Fig. 9(a) and CO conversion for respective conditions is shown in Fig. 9(b) and compared with the thermodynamic equilibrium calculated for 100 kPa pressure. Fig. 9(a) depicts higher H2 recovery with pressure at a particular temperature and with rise in temperature, the recovery is increased. The recovery of higher percentage of hydrogen results in increasing the CO conversion as evident from Fig. 9(b). The value of CO equilibrium conversion decreases linearly with increase in temperature, however, in case of membrane reactor small decrease in CO conversion is observed as the temperature increases. 3.4.2. Comparison of CO conversion and methane formation in a FBR and CMR High temperature WGS reaction was carried out at a temperature range of 400–600 °C in fixed bed reactor (FBR) and catalytic membrane reactor (CMR) using a steam to carbon ratio 1.5 and the results are compared in Fig. 10. It was observed that the CO conversion in the CMR is higher than the FBR and also the thermodynamic equilibrium conversion value at all experimental conditions. The thermodynamic CO
Fig. 11. Stability of catalyst in FBR and CMR for 24 h at 400 °C. 122
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Fig. 12. Stability of membrane reactor at 600 °C for 100 h. Fig. 13(a). % CO conversion as a function of temperature using reformate gas mixture (30% CO, 15% CO2, 25% H2 and 30% H2O).
was found to be stable for 24 h and the CO conversion is higher in CMR. However, methane formation was found to be suppressed in CMR compared to FBR. According to Eqs. 2–4, the methane formation depends upon the availability of H2 in the product and each mol of CH4 consumes 2–4 mol of hydrogen, which eventually decreases the hydrogen concentration in the product. In case of membrane reactor the removal of H2 from the product stream suppresses the methane formation and enhances the throughput of the reaction. The hydrogen permeance was found to be 1.1 × 10−4 mol m−2 s−1 Pa0.65 under the reaction conditions at 400 °C and was stable for 24 h of operation. The stability of CMR was tested for 100 h at high temperature of 600 °C using S/C ratio of 1.5 and is shown in Fig. 12. The performance of the membrane reactor was found to be stable over the entire period of investigation, however, a small decline in CO conversion is observed after 90 h. This may be due to some carbon deposition on the surface of the catalyst which blocks a few active surface sites for reaction. The H2 recovery was also monitored over the period of 100 h, the hydrogen permeability was found to be 2.5 × 10−4 mol m−2 s−1 Pa0.65, which is stable at around 43%. During the stability test at 600 °C, in the permeate side small CO peak was observed in the GC signal after ∼96 h of operation. This may be due to the formation of leakage through sealing at high temperature or formation of small defects on the thin membrane. Either of the causes are responsible for possible transport of CO through the membrane. Due to the possible leakage the purity of permeated hydrogen decreased from ≥99.5% to 98.6%.
Fig. 13(b). CH4 formation and H2 recovery as a function of temperature using reformate gas mixture (30% CO, 15% CO2, 25% H2 and 30% H2O).
from Fig. 13(b) that the formation of methane is drastically reduced from 7% in FBR to 2% in membrane reactor. 3.4.4. Post reaction membrane characterization The Pd-Ag membrane was characterized by SEM to analyse the change in morphology and stability after prolonged WGS reaction and heating-cooling cycles. The surface of the membrane was found to be stable and well adherent to the Al2O3 substrate even after prolonged heating at 600 °C for 100 h. The cross-sectional SEM image of the membrane after 100 h heating at 600 °C is shown in Fig. 14, which indicates better adherent property of inner coated membrane. It is a well-known fact that the difference in thermal expansion coefficient (α) between Pd-Ag (α = 10.24 × 10−6 K−1) and Al2O3 (1 × 10−6 K−1) [23] results in delamination of membrane from the substrate surface after few thermal cycles. However, this effect can be controlled by coating Pd-Ag membrane in the inner side of tube. The inner coated membrane expands along the direction the substrate at high temperatures and exerts a positive force on the substrate surface. Thus it remains attached with the substrate surface and resists to the thermal expansion [32]. Carbon deposition on the membrane surface was analysed by using thermo-gravimetric analysis. The membrane sample was
3.4.3. Effect of reformate gas on CO conversion in a FBR and CMR WGS reaction was performed using a reformate gas mixture of 30% CO, 15% CO2, 25% H2 and 30% H2O which is nearly similar with the outlet syn-gas composition of a gasifier plants [51,52]. The reaction was carried out at temperature range of 400–500 °C and 1 bar pressure, the CO conversion, H2 recovery and CH4 formation were evaluated. The CO conversion in CMR exceeds thermodynamic equilibrium and CO conversion in FBR, which is evident from Fig. 13(a). However, the shift in CO conversion is much higher compared to the previous cases, which, can be ascribed to the amount of higher concentration of hydrogen in the reaction mixture. Higher concentration of hydrogen in the reaction mixture increases the partial pressure of hydrogen in the feed side, which eventually increases the chemical potential of hydrogen. The increase in chemical potential in the feed side ultimately increases the driving force for hydrogen transport through the membrane and results in high hydrogen permeation [38]. Since higher quantity amount of hydrogen is permeated through the membrane to the permeate side, the probability of methane formation decreases in a CMR. This is evident 123
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Fig. 14. Cross sectional view SEM of Pd-Ag membrane after WGS reaction.
heated in a TGA instrument under air up to 800 °C with a ramping rate of 10 °C/min. The TGA graph of the membrane is shown in Fig. A6. The mass loss in TGA was analysed and result showed negligible carbon deposition on membrane. 4. Conclusions In the present study ultra-thin (∼1µm) Pd0.77Ag0.23 alloy membrane was synthesized by electroless plating method in the inner side of Al2O3 hollow fiber support. Hydrogen permeation characteristic of the membrane was evaluated at different temperatures and trans-membrane hydrogen pressure. Hydrogen permeance at 500 °C and 100 kPa was found to be 5.11 × 10−4 mol−1 m−2 Pa−0.65 and ideal H2/N2 selectivity of the membrane was found to be 2380 at similar conditions. WGS reaction was carried out in CMR at a temperature range of 400–600 °C and CO conversion, H2 recovery and CH4 formation were evaluated. CMR was found to be efficient in shifting the thermodynamic equilibrium for a steam to carbon ratio of 1.5. Use of excess steam was found to dilute the H2 concentration and reduced H2 recovery in the permeate side. At 600 °C, a stable CO conversion of 85.5% was achieved which is ∼12% higher than the equilibrium value of 73%. CO conversion was found to be enhanced and methane formation was suppressed in a CMR compared to a FBR. The CMR was found to be stable for 100 h at 600 °C. About 60% of hydrogen was recovered and methane formation was suppressed from 6.5% to 0.5% at 500 °C using a reformate gas mixture of 30% CO, 15% CO2, 25% H2 and 30% H2O. In terms of methane suppression, hydrogen production and CO conversion the CMR was found to have superior properties compared to FBR. Acknowledgements The authors generously thank financial support from Singapore Agency for Science, Technology and Research (A*STAR) AME IRG grant (No. A1783c0016) and National Environment Agency (NEA) of Singapore (WTE-CRP 1501-103). Appendix A. Supplementary data Supplementary data to this article can be found online at https:// doi.org/10.1016/j.cej.2018.12.112. References [1] K. Mazloomi, C. Gomes, Hydrogen as an energy carrier: prospects and challenges,
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