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Autothermal Reforming of Methane in a Proton-Conducting Ceramic Membrane Reactor Jay Kniep,† Matthew Anderson,† and Y. S. Lin*,† †
School for Engineering of Matter, Transport and Energy, Arizona State University, Tempe, Arizona 85287, United States ABSTRACT: Endothermic steam reforming of methane for hydrogen production requires heat input with selective oxidation of methane. Dense SrCe0.75Zr0.20Tm0.05O3‑δ perovskite membranes were combined with a reforming catalyst to demonstrate the feasibility of a heat-exchange membrane reactor for steam reforming of methane coupled with selective oxidation of permeated hydrogen. The reforming catalyst used was a prereduced nickel based catalyst supported on γ-Al2O3. Hydrogen produced via the steam reforming of methane or water gas shift reaction was able to diffuse through the catalyst bed and transport through the membrane. The permeated hydrogen reacted with oxygen (from air) to produce heat for the steam reforming of methane on the other side of the membrane. The membrane reactor avoids the use of an expensive air separation unit to produce pure oxygen. The influence of experimental conditions, such as temperature, gas hourly space velocity, and the steam to carbon (S/C) ratio, on the membrane reactor was investigated. SrCe0.75Zr0.20Tm0.05O3‑δ showed good chemical stability in steam reforming conditions as X-ray diffraction analysis of the membrane surface exposed to steam-reforming conditions for 425 h showed only minor CeO2 formation. The experimental data demonstrate the feasibility of using a proton conducting ceramic membrane in the heat-exchange membrane reactor for steam reforming of methane coupled with selective oxidation.
1. INTRODUCTION Syngas is a mixture of hydrogen and carbon monoxide and is an important feedstock for synthesis of a large number of chemicals. Syngas is the intermediate product for hydrogen production by reforming fossil fuels. Syngas or hydrogen can be produced by steam reforming (SR), partial oxidative reforming (POX), and autothermal reforming (ATR) of methane or other fossil fuels such as coal. ATR of methane actually combines the endothermic SR and exothermic POX. For hydrogen production, ATR of methane can be expressed as CH4 þ H2 O þ 1=2O2 ¼ > 3H2 þ CO2
ðAÞ
ATR of methane for producing syngas or hydrogen have received much attention in recent years because the process has a higher energy efficiency compared to SR or POX.1,2 The role of oxygen is to react with methane to provide heat for the reforming reaction. One of the key issues in ATR or POX in the conventional reactors is that these processes require an expensive air separation unit to produce pure oxygen for the reactions. Membrane reactors using oxygen semipermeable mixed-conducting ceramic membranes have been studied extensively in the past two decade for ATR of methane to produce syngas or hydrogen.35 In these oxygen permeable membrane reactors, air is fed to one side of the membrane and the methane/steam to the other side of membrane packed with a reforming catalyst. Oxygen from the air permeates through membrane and reacts with methane on the other side of the membrane. Deckman et al.6 recently proposed a heat-exchange membrane reactor using a hydrogen permeable membrane for electrical power generation. In this membrane reactor, steam reforming of methane takes place at high pressure on one side of the membrane. The produced hydrogen permeates through the r 2011 American Chemical Society
membrane and combusts on the other side of the membrane with oxygen (in air) to generate hot steam to drive a turbine. The product in the methane/steam side of the membrane contains mainly carbon dioxide at high pressure ready for sequestration. However, no experimental results demonstrating the performance of such heat-exchange membrane reactor for electrical power generation were reported by these researchers.6 The heat-exchange membrane reactor concept for electrical power generation can be extended for ATR of methane for producing syngas or hydrogen, as shown in Figure 1. In this reactor, the catalytic steam reforming of methane and water gas shift reaction take place on one side of the membrane CH4 þ H2 O ¼ > 3H2 þ CO
ðBÞ
CO þ H2 O ¼ > H2 þ CO2
ðCÞ
A part of the hydrogen produced is transported through the hydrogen permeable membrane and reacts with oxygen (or air if membrane is only permeable to hydrogen) to produce water as H2 þ 1=2O2 ¼ > H2 O
ðDÞ
The heat of reaction is 206 kJ/mol (endothermic) for reaction B, 41 kJ/mol (exothermic) for reaction C, and 241 kJ/mol (exothermic) for reaction D. Thus, heat of combustion from reaction D can be exchanged thermally through the membrane and be used to provide heat for the reforming reaction B. If enough H2 is combusted, the system could theoretically be self-sustained in heat supply. The membrane now serves as a Received: May 16, 2011 Accepted: October 5, 2011 Revised: September 21, 2011 Published: October 05, 2011 12426
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Figure 1. Schematic showing heat-exchange membrane reactor for steam-reforming of methane coupled with selective oxidation using a proton-conducting ceramic membrane.
heat-exchanger. If the membrane is only permeable to hydrogen, this allows for the use of air instead of pure oxygen, thus avoiding an air separation unit. The net effect of such a membrane reactor is steam-reforming of methane to produce H2 and CO (and CO2) with air, methane, and steam as the feeds, without the need of a prereactor air separation and removal of water vapor from the product stream. For use in the proposed heat-exchange membrane reactor for ATR of methane, the membrane should be semipermeable to hydrogen and be able to operate with one surface exposed to a reducing gas and another to air. Dense, proton-conducting perovskite structured ceramic membranes, based on either SrCeO3 or BaCeO3, offer high hydrogen selectivity, thermal stability, and mechanical strength.7,8 Protons can absorb into the membrane in either wet or dry hydrogen atmospheres and then migrate through the membrane by hopping between adjacent lattice oxygen with the driving force for migration being a hydrogen partial pressure difference on either side of the membrane. Proton-conducting ceramic membranes normally work under the conditions with one surface exposed to a hydrogen containing gas and another surface to oxygen or air, which is required in order to maintain high electronic conductivity on that surface.9,10 Furthermore, operating temperatures for proton-conducting ceramic membranes are in the range required for steam-reforming of methane. These properties make the proton-conducting ceramic membranes good candidates for use in the heat-exchange membrane reactor described above. The proton-conducting SrCeO3 or BaCeO3 based membranes generally have lower hydrogen permeability than metal membranes, mainly due to low electronic conductivity. In order to improve the electronic conductivity of these materials, various trivalent cations have been partial substituted for cerium which causes oxygen vacancies to form within the membrane to maintain electroneutrality. Examples of this doping strategy include SrCe0.95Yb0.05O3‑δ and SrCe0.95Tm0.05O3‑δ.11,12 In addition to improving the electronic conductivity of proton conducting perovskites, partial substitution of zirconium for cerium has been shown to increase the chemical stability of the membranes.13,14 Exposure of the membrane surface at high temperatures to carbon dioxide can cause the surface to decompose into carbonates and various metal oxides.15 This has a harmful effect on the membrane’s mechanical strength and hydrogen flux through the membrane as the carbonates and metal oxides inhibit the surface reactions. Recently, our group studied the effect of zirconium doping on the hydrogen permeation through SrCe0.95‑xZrxTm0.05O3‑δ (0 e x e 0.40) dense
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membranes. SrCe0.75Zr0.20Tm0.05O3‑δ was shown to have improved chemical stability and a higher hydrogen flux than SrCe0.95Tm0.05O3‑δ in a long-term hydrogen permeation test with a carbon dioxide containing environment.16 This paper reports an experimental study on steam-reforming of methane on a membrane reactor using the improved protonconducting ceramic membrane SrCe0.75Zr0.20Tm0.05O3‑δ. The performance of the membranes and membrane reactor under various operating conditions as well as the characterization of the membranes and catalysts are discussed in detail. The main objective of the work is to demonstrate experimentally the feasibility of the heat exchange membrane reactor using a proton-conducting ceramic membrane for ATR of methane for production of syngas or hydrogen.
2. EXPERIMENTAL METHODS 2.1. Preparation and Characterization of Membranes. SrCe0.75Zr0.20Tm0.05O3‑δ membranes were prepared using the liquid citrate method. In this method, stoichiometric amounts of the corresponding metal nitrates Sr(NO3)2 (Alpha Aesar, 99.0%), Ce(NO3)3•6H2O (Alpha Aesar, 99.5%), ZrO(NO3)2•xH2O (Sigma-Aldrich, 99.0%), and Tm(NO3)3•5H2O (Alpha Aesar, 99.9%) were mixed with citric acid in distilled water. The amount of citric acid was three times the total molar amount of metal ions present. The transparent liquid was heated to 95100 °C and remained there under reflux and stirring for four hours for the polymerization reaction to occur. The lid was then removed, and excess water was evaporated off at 100 °C leaving the solution a viscous gel. The gel was dried for 24 h at 110 °C, and the resulting brittle, porous material was heated to 400 °C for self-ignition to burn out the organics that were present. The material was then ground with a mortar and pestle for 15 min followed by calcination at 850 °C for 8 h (ramp rate = 5 °C/min). Samples of the calcined powders were put into a die with a diameter of 2.30 cm and pressed with a hydraulic press (Carver, Model#3853) to 180 MPa. The resulting green disks were sintered in air at 1525 °C for 24 h in a furnace (Thermolyne, 46100) with a ramp rate of 2 °C/min. The gas tightness of each membrane was verified using a room temperature unsteady state permeation system with helium. The membrane was considered gastight if the He permeance through the membrane was less than 1010 mol/m2 3 Pa 3 s. X-ray diffraction (XRD) (Bruker, CuKα) was used to characterize the phase structure of each membrane. Characterization of the membranes was evaluated in the 2θ range of 20° to 70° with a step size of 0.02°/s. 2.2. Preparation and Characterization of Reforming Catalyst. Appropriate amounts of Ni(NO3)2 3 6H2O (Alfa Aesar, 98% Purity) and γ-Al2O3 (Alfa Aesar) were weighted out. The γAl2O3 support was ground with a mortar and pestle and sieved with a #80 sieve so the support particles were 180 μm or less. The Ni(NO3)2 3 6H2O and γ-Al2O3 were mixed with deionized water at room temperature. The solution was then heated to 80 °C to allow the excess water to evaporate off. The resulting slurry was then dried overnight at 120 °C. The dried sample was ground with a mortar and pestle and then calcined for 4 h at 700 °C (ramp rate = 5 °C/min). Once the catalyst has been synthesized, samples were reduced in a 10% H2/He gas mixture for 6 h at 600 °C before use in a steam reforming of methane membrane reactor. Calcined and reduced catalyst samples were characterized with XRD in the 2θ range of 30° to 80° with a step size of 12427
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Figure 2. Experimental high temperature membrane reactor set up.
0.02°/s. The BET surface area, pore volume, and average pore size were characterized using nitrogen porosimetry (Micromeritics ASAP2020). Nitrogen adsorption/desorption isotherms were measured at the temperature of liquid nitrogen (77 K). The catalyst sample was degassed at 350 °C for 2 h before adsorption to remove any physically absorbed components. 2.3. Hydrogen Permeation and Reforming of Methane in Membrane Reactor. Steam reforming of methane coupled with selective oxidation on a dense SrCe0.75Zr0.20Tm0.05O3‑δ membrane reactor was conducted in the temperature range of 750900 °C using a high temperature gas permeation system. A schematic of the system is shown in Figure 2. The gas permeation module (Probostat, Norwegian Eletro Ceramics AS) utilizes spring force and metal seals to ensure sealing for the H2 flux measurements. The feed gas mixture contained CH4 and steam (with the balance He), while the sweep gas as the oxidizing agent was a 20% O2/Ar gas mixture to simulate the air. A gastight membrane and silver seal (Alfa Aesar, 99.9%) were mounted on the inner alumina tube and held in place by spring pressure applied by the alumina spacer on top of the membrane. Once the membrane and seal were mounted, 300 mg of prereduced catalyst was placed on top of the membrane. The sealing procedure consisted initially of heating the setup from ambient conditions to about 950 °C (ramp rate = 1 °C/min) to soften the silver ring. Next, He and Ar were introduced on the feed and sweep side, respectively. The flow rate of the inert gases on either side of the membrane was 30 mL/min, which was regulated by mass flow controllers (MKS, Model 1179) and a 4 channel readout (MKS, Type 247). The amount of He in the Ar stream (and therefore, the leakage rate through the seal) was determined by running gas samples through a gas chromatographer (Agilent, 6890N) with a packed column (2836PC, Alltech) and a TCD detector. Once the He content in the sweep stream was minimized, the system was ramped down (1 °C/min) to experimental conditions. For H2 permeation flux measurements with a 20% O2/Ar sweep, the relative humidity in the effluent of the sweep gas was measured by a thermohygrometer (Cole Palmer, 37950) and used to calculate the flux of H2 through the membrane. The use of Ar instead of N2 in the O2/Ar feed would allow one to check if the permeation/reaction system leaked by detection of N2 in the system. The retentate stream was analyzed using the gas chromatographer. For membrane reactor measurements, the CO selectivity was calculated as the amount of CO formed with relation to all carbon species (CO or CO2) formed from the
Figure 3. XRD pattern of calcined, reduced, and used 10% Ni/γ-Al2O3 catalyst.
conversion of methane on the reactor side of the membrane. The system was allowed to equilibrate for 5 h after a gas mixture change or 1 h after a temperature change for all experiments. The error in determining the H2 permeation flux using this procedure is approximately (8%. The standard temperature and pressure used for determining the hydrogen permeation are 0 °C and 1 atm. All of the thermodynamic equilibrium data calculations were performed using the HSC Chemistry for Windows 5.0 software (Outokumpu Research Oy).
3. RESULTS AND DISCUSSION 3.1. Membrane Characteristics and Performance. The gastightness of each SrCe0.75Zr0.20Tm0.05O3‑δ membrane made from the citrate method was confirmed using a room temperature helium permeation test, with a helium permeance of 1010 mol/m2 3 Pa 3 s or less. The room temperature XRD patterns of the membranes show a single perovskite phase and are consistent with the previously published pattern for SrCe0.75Zr0.20Tm0.05O3‑δ.16 The reforming catalyst prepared in this work for the membrane reactor was prereduced 10% Ni/γ-Al2O3. The XRD patterns of calcined, reduced, and used catalysts are shown in Figure 3. The peaks indicative of NiO in the calcined sample or of Ni in the prereduced sample are not that visible due to peak overlap of the γ-Al2O3 support. The more pronounced Ni peaks in the used catalyst indicate that the prereduced catalyst was not completely reduced prior to use in the membrane reactor. Based on nitrogen adsorption/desorption isotherms, the BET surface area was 153.8 m2/g, the pore volume was 0.46 cm3/g, and the average pore diameter was 8.85 nm. To determine the feasibility of SrCe0.75Zr0.20Tm0.05O3‑δ as a material for use as a membrane in a reforming membrane reactor, 1 mm thick membranes with and without a reforming catalyst were subjected to practical working steam reforming conditions at 850 °C. The inlet methane concentration was 5% with a steam to carbon ratio (S/C) of 3.0 and the balance helium while the sweep side of the membrane consisted of 20% O2/Ar to ensure a low sweep side hydrogen partial pressure. An S/C ratio of 3.0 was chosen to ensure the SRM reaction took place with a higher than stoichiometric S/C to avoid carbon deposition on the catalyst or membrane surface. The comparison of the two tests can be found in Figure 4. 12428
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Figure 4. Comparison of SrCe0.75Zr0.20Tm0.05O3‑δ SRM membrane reactors without and with Ni/γ-Al2O3 catalyst (T = 850 °C, S/C = 3.0).
Figure 5. XRD pattern of SrCe0.75Zr0.20Tm0.05O3‑δ membrane exposed to SRM conditions (feed side) and 20% O2/Ar (sweep side) at 850 °C for 30 h.
As shown in Figure 4, the experiment conducted without a reforming catalyst had a methane conversion of roughly 20%, which is considerably lower than the thermodynamic equilibrium conversion of ∼96% at 850 °C. Due to the low methane conversion, the hydrogen concentration in the retentate stream was extremely small and did not provide a sufficient hydrogen partial pressure differential on either side of the membrane. Therefore, the hydrogen flux through the membrane under those conditions was below the detection limit of the high temperature permeation system. Figure 5, the XRD pattern of surface of the membrane exposed to SRM conditions (without catalyst), does not contain additional peaks and is identical to the surface exposed to 20% O2/Ar, indicating the membrane is chemically stable under steam reforming conditions. The 300 mg of 10% Ni/γ-Al2O3 used in the SrCe0.75Zr0.20Tm0.05O3‑δ steam reforming membrane reactor in Figure 2 corresponds to a bed height of 4 mm. The methane conversion was
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Figure 6. Influence of reactor temperature on the methane conversion, CO selectivity, and H2 flux (reactor conditions: methane inlet concentration = 5%, S/C = 3.0, GHSV = 4500 h1, 20% O2/Ar sweep rate = 100 mL/min).
∼88% (Figure 4), which is still lower than the thermodynamic equilibrium conversion of ∼96%. The reason for the lower methane conversion is due to the high reaction gas hourly space velocity (GHSV) of the system. With an inlet gas flow rate of 50 mL/min, the GHSV is ∼4500 h1 which means a low contact time for the catalyst and therefore a lower conversion. However, with a steady state conversion of 88%, the concentration of H2 in the retentate stream reached ∼13% which imposed a H2 partial pressure differential on either side of the membrane and resulted in a H2 permeation flux through the membrane of 0.015 mL (STP)/cm2 3 min. Although this system needs to be optimized, this experiment proves that H2 produced via the steam reforming reaction can diffuse through the catalyst bed to the membrane surface and then be transported through the membrane. 3.2. Membrane Reactor Performance. Figure 6 shows the influence of the temperature of the membrane reactor on the methane conversion, CO selectivity, and H2 flux through the membrane. The other parameters of Figure 6 include a methane inlet concentration of 5%, a S/C ratio of 3.0, a GHSV of 4500 h1, and a 20% O2/Ar sweep rate of 100 mL/min. The experimental methane conversion ranged from 81 to ∼89% over the temperature range tested, while the thermodynamic equilibrium conversions range from 88 to 97%. Again, the reason for the lower experimental methane conversions is the GHSV is high enough to limit the contact time for the catalyst resulting in lower conversions. The CO selectivity is fairly independent of temperature and is roughly 90% over the temperature range tested. The high selectivity of CO can be explained by two points. First, at higher temperatures, the steam reforming reaction kinetics (reaction B) are more favorable than the WGS reaction kinetics (reaction C), and the amount of H2 removed from the system through the membrane is not enough to significantly shift the WGS reaction (reaction B) to the right side. The H2 permeation flux through the membrane increases over the temperature range tested due to a combination of factors. As the temperature increases, the methane conversion also increased which yields a higher H2 concentration in the retentate and a larger H2 partial pressure difference on either side of the membrane. Also, the protonic and electronic conductivity of SrCe0.75Zr0.20Tm0.05O3‑δ increase with increasing temperature, which leads to more favorable transport properties.16 12429
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Figure 7. Influence of reactor gas hourly space velocity on methane conversion (reactor conditions: temperature = 900 °C, methane inlet concentration = 1%, S/C = 3.0, amount of catalyst = 150 mg, 20% O2/Ar sweep rate = 100 mL/min).
Figure 8. Influence of the S/C ratio on the methane conversion and CO selectivity (reactor conditions: temperature = 850 °C, methane inlet concentration = 5%, GHSV = 4500 h1, 20% O2/Ar sweep rate = 100 mL/min).
The GHSV of the reactor system was varied by changing the flow rate of the inlet methane, steam, and helium mixture (the ratios of the mixture remained the same) while maintaining the same amount of catalyst used, and the effect of the GHSV on the system is shown in Figure 7. The other parameters of Figure 7 are a temperature of 900 °C, methane inlet concentration of 1%, S/C ratio of 3.0, 150 mg of 10% Ni/γ-Al2O3, and a 20% O2/Ar sweep flow rate of 100 mL/min. As the GHSV increased from 1800 to 14,500 h1, the methane conversion linearly decreased as expected. At the lowest GHSV of 1800 h1, the methane conversion was over 96%, which is in good agreement with the SRM thermodynamic equilibrium of ∼97%. As mentioned earlier, a high GHSV limits the contact time for the catalyst resulting in lower conversions. In addition, as the GHSV increases, the residence time of the reactor is reduced, and this shortens the
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Figure 9. XRD pattern of SrCe0.75Zr0.20Tm0.05O3‑δ membrane exposed to SRM conditions for 425 h.
time for the H2 produced to diffuse through the catalyst to the membrane surface.17,18 In a conventional steam reforming fixed bed reactor, the S/C ratio is usually in the range of 3.0 to 4.0.18 By maintaining an S/C ratio well above the stoichiometric ratio of S/C ensures that coke will not deposit on the catalyst, which would lead to lower methane conversion and reactor performance. In addition, a high steam partial pressure in the membrane reactor could also drive the SRM reaction (reaction A) or WGS reaction (reaction B) to the right.17 The influence of the S/C ratio on the methane conversion and CO selectivity of a SrCe0.75Zr0.20Tm0.05O3‑δ SRM membrane reactor is shown in Figure 8. The other parameters of the membrane reactor in Figure 8 are a temperature of 850 °C, an inlet methane concentration of 5%, a GHSV of 4500 h1, and a 20% O2/Ar sweep flow rate of 100 mL/min. The methane conversion increases as the S/C ratio increases from 2.5 to 3.5 but slightly decreases at a S/C ratio of 5.0. Chen et al.17,18 also report a slight increase in the methane conversion in the S/C ratio range of 2.5 to 3.5, but higher or lower S/C ratios are not reported. There are conflicting reports in the literature concerning the effect of the S/C ratio on the methane conversion in a conventional fixed bed reactor.18,19 The CO selectivity is fairly independent of the S/C ratio in Figure 8, while the CO selectivity in a Pd based membrane reactor decreased with increasing S/ C.18,19 Therefore, it is sufficient to say that the effect of the S/C ratio on methane conversion and CO selectivity is not independent of the other parameters of the system, specifically temperature and GHSV. For the other experiments reported in this work, a S/C ratio of 3.0 was used, and spent catalyst characterized by XRD showed no coke formation. The XRD patterns of the surfaces of a SrCe0.75Zr0.20Tm0.05O3‑δ membrane exposed to various steam reforming conditions for 425 h is shown in Figure 9. When compared to the SrCe0.75Zr0.20Tm0.05O3‑δ XRD patterns in Figure 5, the feed side surface (the side exposed to the methane, steam, and helium mixture) pattern has new peaks which indicate the formation of CeO2 on the membrane surface. The sweep side surface of the membrane also contains minor new peaks indicating some CeO2 formation on the surface. However, there are no new peaks that would be a sign of carbonate formation on the membrane surfaces. While CeO2 on the membrane 12430
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Figure 10. Comparison of temperature vs time for either type of cooling down experiment.
surface will hurt the performance of the membrane, SrCe0.75Zr0.20Tm0.05O3‑δ has been shown to have an appreciable H2 flux capabilities even with CeO2 present on the surface of the membrane.16 3.3. Verifying Performance of Heat-Exchange Membrane Reactor. To determine the feasibility of SrCe0.75Zr0.20Tm0.05O3‑δ as a candidate for use in a heat-exchange membrane reactor, two types of experiments were conducted. In the first set of experiments, a SrCe0.75Zr0.20Tm0.05O3‑δ membrane was heated up to 900 °C with the inert gases helium and argon flowing on either side of the membrane, respectively. After the system had dwelled at 900 °C for a period of time, the power was cut to the tubular furnace and the temperature of the membrane was recorded over time. The thermocouple in the high temperature permeation system is located directly next to the membrane so an accurate membrane temperature is ensured. In the second set of experiments, a SrCe0.75Zr0.20Tm0.05O3‑δ steam reforming membrane reactor under the same conditions as listed in Figure 6 was allowed to reach steady state. Once steady state conditions had been verified, the power was cut to the tubular furnace, and the temperature of the membrane reactor was recorded over time. A comparison of the temperature vs time profile for both types of experiments is shown in Figure 10. Two runs were conducted for either type of experiment, and the repeatability of the experiments is excellent as shown in Figure 10. The SrCe0.75Zr0.20Tm0.05O3‑δ steam reforming membrane reactors cooled down at a slower rate than the membrane in inert conditions due to the combustion of H2 on the sweep side of the membrane. However, the SrCe0.75Zr0.20Tm0.05O3‑δ steam reforming membrane reactors still cools down over time as the amount of heat produced by combustion and thermally transported through the membrane is less than the energy needed to drive the endothermic steam reforming reaction. An energy balance was then conducted on the steady state conditions of Figure 6 at 900 °C in order to determine the H2 flux and combustion needed to balance the energy of the system. At 900 °C, the steady state methane conversion was 88.4%. Assuming the methane is converted only by the steam reforming reaction (reaction B), the amount of energy needed for the reaction is 0.31 J/s. With a steady state H2 flux through the membrane of 0.023 mL (STP)/cm2 3 min, the amount of heat produced by the
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combustion reaction (reaction D) is 0.0064 J/s. Therefore, a H2 flux approximately 50 times higher is needed just to balance the energy between the two reactions. As the assumption that all of thermal energy produced by reaction D will transport through the membrane to the reforming reaction side is most likely not valid, an even higher H2 flux will be needed. Hydrogen permeation through proton-conducting ceramic membranes is controlled by the chemical diffusion of proton in the bulk phase of the membrane even for thin membranes due to fast surface reaction kinetics as compared to oxygen permeation through mixed-conducting ceramic membranes.20 Hamakawa et al.21 reported bulk diffusion limited hydrogen permeation through a ytterbium doped SrCeO3 with thickness down to 2 μm. Therefore it is possible that a SrCe0.75Zr0.20Tm0.05O3‑δ membrane with a reduced thickness could be used to balance the energy required for the reforming reaction of methane. Since the membrane used in this work was 1 mm thick, a 50-fold increase in hydrogen permeance requires a 50-fold decrease in the membrane thickness assuming a rate-limiting step of bulk diffusion. This means that a SrCe0.75Zr0.20Tm0.05O3‑δ membrane thinner than 20 μm is required to provide sufficiently high hydrogen flux for energy balance. Such a thin proton-conducting ceramic membrane should be coated on porous ceramic or metallic support to provide adequate mechanical strength. Since the membrane also serves as a heat-exchanger wall, a porous metal alloy support is more desirable due to its higher thermal conductivity. Clearly more study on membrane material development, synthesis of the thin proton-conducting membranes, and design and understanding the performance of the membrane reactor is needed in order to advance the heat-exchange membrane reactor concept to the stage attractive for industrial application.
4. CONCLUSIONS Dense SrCe0.75Zr0.20Tm0.05O3‑δ membranes prepared using the liquid citrate method were combined with a 10% Ni/γ-Al2O3 reforming catalyst for use in a heat-exchange membrane reactor for steam reforming of methane coupled with oxidation of hydrogen that permeated through the membrane. Hydrogen produced by either the steam reforming of methane or water gas shift reaction was able to diffuse through the catalyst bed and transport through the membrane and subsequently reacted with oxygen to produce heat. At 900 °C, with a GHSV of 4500 h1 and a S/C ratio of 3.0, a steady state H2 permeation flux of 0.023 mL (STP)/cm2 3 min was measured. A SrCe0.75Zr0.20Tm0.05O3‑δ membrane exposed to steam reforming of methane conditions for 425 h exhibited good chemical stability. The results demonstrated feasibility of the heatexchange membrane reactor for steam reforming of methane with heat supplied by oxidation of hydrogen permeated through the membrane. However, a substantial increase in the hydrogen permeation flux through reducing the thickness of the SrCe0.75Zr0.20Tm0.05O3‑δ membranes is required in order to provide sufficient heat to balance the energy of the system. ’ AUTHOR INFORMATION Corresponding Author
*E-mail:
[email protected].
’ ACKNOWLEDGMENT This project was supported by the Department of Energy through grant (DE-FG36-07GO17001). We acknowledge 12431
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Industrial & Engineering Chemistry Research Mr. Fred Pe~ na for his assistance in building the high temperature permeation and membrane reactor set up.
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(21) Hamakawa, S.; Li, L.; Li, A.; Iglesia, E. Synthesis and Hydrogen Permeation Properties of Membranes Based on Dense SrCe0.95Yb0.05O3‑α Thin Films. Solid State Ionics 2002, 48, 71–81.
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