HZSM-5 at

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Methane Nonoxidative Aromatization over Ru-Mo/HZSM-5 at Temperatures up to 973 K in a Palladium-Silver/Stainless Steel Membrane Reactor Maria C. Iliuta, Bernard P. A. Grandjean,* and Faı1ç al Larachi Department of Chemical Engineering and CERPIC, Laval University, St. Foy, Que´ bec, Canada G1K 7P4

Pd-based composite membranes supported on porous stainless steel offer high permeability and good mechanical stability. They represent an appropriate alternative in catalytic membrane reactors. In this study, Pd-Ag/porous stainless steel membranes prepared by electroless plating were used in the oxygen-free methane aromatization over Ru-Mo/HZSM-5 in a catalytic membrane reactor at temperatures up to 973 K. The 0.5% Ru-3% Mo/HZSM-5 catalyst, prepared by incipient-wetness co-impregnation, was highly selective toward benzene production. Despite attained methane conversions well beyond the thermodynamic conversion, the continuous withdrawal of coproduced H2 promoted the formation and deposition of low-H/C carbonaceous species especially at high temperatures. At a methane space velocity of 270 mL(STP)‚h-1‚g-1 and a temperature of 873 K, a maximum conversion of methane to benzene equal to ca. 8% was obtained. When the temperature was raised to 973 K, the maximum conversion culminated at ca. 17% but was accompanied with time by a faster falloff in methane conversion compared to that observed at 873 K. 1. Introduction From the increasingly stringent environmental regulations as well as the quest for efficient multifunctional process intensification means, the use of membrane separation processes is foreseen to become widespread in industrial practice soon. The application of membrane reactors for (oxy)dehydrogenation and decomposition reactions has been receiving greater attention over the past few years. A membrane reactor is a device that takes advantage of the separation function of a membrane to imbalance selectively the species concentration distribution (whether reactant or product) at specific locations within equilibrium-controlled reaction media. Though the principle of such dual processes is encountered in many naturally occurring biological functions, e.g., photosynthesis, digestion, and respiration, it seduced industry in the catalytic (synthetic) membrane reactor technology to find applications in selective recovery and separation of hydrogen that surpass the equilibrium conversion bottleneck.1-5 Continuous withdrawal of hydrogen from the reaction shifts, through the so-called drain-off effect, the equilibrium limitation in the dehydrogenation reactions. Because of their thermal stability, chemical resistance, and mechanical strength, greater attention has been focused on the inorganic membrane design for membrane reactors. Such membranes may be either porous materials or dense impervious films. One of the major concerns in both cases is the film thickness.6 Thin membranes are more desirable because the hydrogen permeation flux is inversely proportional to the film thickness. However, very thin membranes may have the drawback of low mechanical strength and poor hydrogen permselectivity at high temperatures.4 One option to overcome this difficulty is to make composite mem* Corresponding author. Tel.: 418-656-2859. Fax: 418-6565993. E-mail: [email protected].

branes consisting of a thin impervious film deposited on a porous substrate such as glass,7,8,11 polymide,9 ceramics,10 stainless steel (SS),3,6,11 or porous silver disks.12 In view of the reactor design, porous SS materials seem to be a good choice because of their ease in fabrication and processing, lack of corrosiveness, resistance to cracking, and low cost. Preparative techniques of composite membranes include physical and chemical vapor deposition, electroplating, and electroless plating. Electroless deposition has been widely used for the preparation of Pd-based membrane materials.1,6,10,11,13-19 Pure Pd membranes are prone to hydrogen embrittlement that occurs at temperatures below 523 K.18 Hence, silver is added to prevent embrittlement and to improve the hydrogen permeation flux.6,10,14,18 The optimal alloy was found to have around 77% Pd and 23% Ag composition. Because of their high permeability and their mechanical strength, such membranes have found wide application in hydrogen recovery or purification and in the broader area of membrane reactor technology. Methane conversion into more valuable products such as aromatics has received special attention in the past decade. Most studies have focused on methane dehydrogenation and aromatization, especially in an oxygenfree atmosphere, by using Mo-based HZSM-5-supported catalysts.20 Several factors inherent to the zeolite-based catalytic material have been recognized to affect the performance of methane conversion into aromatics.21,22 These include the Brønsted acid sites’ density, the zeolite channel structure, and the oxidation state and location of the molybdenum species. Based on several studies on the interaction between the Mo species and the zeolite, various modes of methane activation have been proposed.20,21,23-26 At low temperatures, methane conversion is severely limited because of its endothermic and equilibriumcontrolled features. Therefore, temperatures in excess of 973 K are required to boost to meaningful levels

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methane conversion to aromatics within the traditional fixed-bed catalytic reactors (CRs). Most of the experimental studies reported by the research community in the open literature were performed at 973 K or above. Nonetheless, high-temperature operation leads to serious deactivation of Mo-based catalysts by coke deposition and/or by Mo loss through volatilization.27 Coke deposition is one important drawback to methane aromatization that is responsible for the decrease in conversion into useful products. Much of the current work has been targeting improved conversion into aromatics and reduction in coke formation.20,27-35 In our previous work,36 the oxygen-free methane aromatization over Ru-Mo/HZSM-5 was studied in a commercial membrane reactor (REB Research & Consulting, Ferndale, MI) hosting four 178 mm long and 2.38 mm external diameter Pd-coated tantalum-niobium alloy tubular membranes exhibiting high hydrogen permeability. Methane aromatization was evaluated under two sets of conditions: without hydrogen permeation in a conventional fixed-bed CR and with hydrogen permeation in a inert membrane CR (IMCR). The reaction temperature was limited to 873 K because of the membrane poor stability at higher temperatures. However, the methane conversion into benzene steeply increases with increasing reaction temperature; the corresponding thermodynamic values are 5.2% at 873 K and 11.5% at 973 K. Moreover, our previous work on the IMCR configuration at 873 K36 revealed the possibility of boosting conversions well beyond the thermodynamic value at the same temperature. It is expected therefore that temperatures higher than 873 K in the IMCR configuration would yield remarkable levels of methane conversion into benzene. This investigation reports on oxygen-free methane aromatization in membrane CR at temperatures up to 973 K using a Ru-Mo/HZSM-5 catalyst. A dense permeable and hydrogen permselective Pd-Ag membrane (70 mm in length and 9.5 mm external diameter) was prepared by electroless plating and used in the IMCR. The membrane was characterized by scanning electron microscopy (SEM) and hydrogen permeation tests. The IMCR performances are reported and discussed. 2. Experimental Section 2.1. Membrane Preparation. Specimens of porous 316L SS tubes having a 9.5 mm o.d. and a 0.2 µm particle retention size were purchased from Mott Metallurgical. The tubes were cut to a length of 70 mm and welded on each side to a dense SS tube of the same diameter. The effective membrane area was ca. 21 × 10-4 m2. The Pd-Ag membranes were grown on the outer shell of the porous SS tube by electroless plating. Because of the quite large length of the porous SS tube and the rather wide pore size distribution in the substrate, the main experimental difficulty during deposition of Pd and Ag was the stabilization and maintenance of a uniform flow of the solution around the tube in order to remove the incipient N2 formed during the reaction on the porous surface. The preparation of the membrane includes the pretreatment, activation, deposition, and thermal treatment steps as discussed below. 2.1.1. Pretreatment. Good cleaning was found to be very important in the membrane preparation. The tube was first immersed in an ultrasonic bath of carbon tetrachloride for 1 h. It was followed by an advanced cleaning under vacuum using the same solvent. Both

Table 1. Composition of the Activating Solutions Sn-based solution (kg/m3)

precursor SnCl2‚2H2O Pd(NH3)4(NO3)2 HCl (37%)

Pd-based solution (kg/m3)

0.5 0.168 0.001

0.001

Table 2. Electroless Plating Baths Palladium Plating component

(A) dilute solution

(B) concentrated solution

Pd(NH3)4(NO3)2 Na2EDTA‚2H2O NH4OH (28%) N2H4 (1 M)

0.45 kg/m3 50.0 kg/m3 0.200 m3/m3 0.004-0.01 m3/m3

4.5 kg/m3 50.0 kg/m3 0.200 m3/m3 0.004-0.01 m3/m3

Silver Plating component

(C) Ag bath

AgNO3 Na2EDTA‚2H2O NH4OH (28%) N2H4 (1 M)

0.1-0.5 kg/m3 50.0 kg/m3 0.200 m3/m3 0.003-0.007 m3/m3

procedures were found to ensure very good cleaning of the substrate, not only at the surface but also deep in the pores. 2.1.2. Substrate Activation. For reducing the long induction period at the beginning of Pd deposition, the substrate was preseeded with Pd nuclei. Uemiya et al.15 suggested a two-step activation procedure with Sn2+ and Pd2+ solutions, used alternatively 10 times for each solution. Keuler et al.37 reported that small quantities of tin might remain on the substrate surface after activation, thus influencing negatively the membrane properties due to the lower tin melting point. The presence of tin was also confirmed in our study by means of energy-dispersive X-ray analysis (EDX). As a consequence, we modified the activation procedure by treating the substrate eight times (each lasting for 5 min) with a dilute Sn2+ solution and 10 times with a Pd2+ solution (for 5 min the first eight dips and for 10 min the last two dips). Activation started using the Sn2+ solution first, and subsequent dips took place in pristine solutions. The tube was thoroughly cleaned with deionized water after each dip and finally dried at 120 °C. The solution compositions are given in Table 1. This procedure leads to the formation of palladium nuclei mainly on the surface of the substrate. To allow the formation of palladium nuclei also into the pores, an additional activation step under vacuum using the same solutions as those described in Table 1 was also performed. 2.1.3. Electroless Plating. The deposition of Pd and Ag was carried out in a Pyrex bath vessel whose dimensions were fitted with those of the porous SS tube to prevent any unwanted deposition on the dense SS extension parts. The height of the Pyrex enclosure was 0.09 m, and its i.d. was 33 mm. The apparatus was surrounded with a water jacket for temperature control. The compositions of different electroless plating baths used are given in Table 2. The presence of silver in the Pd-based plating bath was found to inhibit the palladium deposition, with the silver being preferentially deposited on the substrate.6,14 A faster deposition rate of palladium with respect to silver through use of a higher hydrazine concentration in the plating bath14 could, however, lead to better simultaneous Pd-Ag codeposition, but the palladium-silver membrane composition control was found to be rather complex.

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Figure 1. Experimental apparatus.

Pd deposition was realized from dilute (Table 2A) and concentrated solutions (Table 2B), for comparison. For preparation of the Pd membrane, the plating procedure starts with a silver-free solution using excess hydrazine. To control Pd deposition at the beginning and to avoid a too fast surface deposition that would impede penetration into the pores (especially for the concentrated Pdcontaining solutions), only 25% of hydrazine was added in the first 1 h of deposition. Silver was then directly deposited on the Pd surface membrane, from a Pd-free bath (Table 2C). The amount of AgNO3 in the bath was calculated by taking into account the amount of Pd deposited. 2.1.4. Thermal Treatment. The Pd-Ag membranes were thermally treated above the Tamman temperatures of the components (550 °C) for 300 min in the presence of hydrogen. Before introduction of hydrogen, the increase in temperature was realized in an argon atmosphere. After thermal treatment, a cooling was also performed in an argon atmosphere. 2.1.5. Membrane SEM. The surface morphology and microstructure of the Pd films were observed using SEM. A JEOL JSM-840A apparatus equipped with an energy-dispersive X-ray analyzer for analyzing the film composition was used. 2.1.6. Gas Permeation Measurements. Gas permeation tests were performed in the tubular membrane reactor using the volumetric method. Argon was used as the permeating gas at room temperature to verify whether the membrane was impervious. The Ar stream was introduced in the reaction side under constant pressure. The permeated argon flow rate, if any, was monitored using a digital bubble flowmeter (model 520, Fisher Scientific) on the permeability side of the reactor maintained at atmospheric pressure. Prior to the H2 permeation tests, the reactor was first heated at 573 K in an argon atmosphere to check for membrane leaks. Then, hydrogen was introduced into the reactor side at temperatures over 573 K and controlled pressures while the permeation side was maintained at atmospheric pressure. The hydrogen flow rate in the permeation side was measured using a digital bubble flowmeter. An electric oven was used to heat/regulate the membrane cell temperature, which was measured using a thermocouple placed in the middle of the tube. Hydrogen permeability was investigated experimentally to verify the selectivity of the membrane toward hydrogen at 773, 873, and 973 K.

2.2. Reaction Testing. 2.2.1. Synthesis and Pretreatment of a Ru-Mo/ HZSM-5 Catalyst. The method is similar to that presented previously.36 A 0.5% Ru-3% Mo/HZSM-5 catalyst was prepared by incipient-wetness co-impregnation of the ammonium form of the zeolite [NH4ZSM5] (Si:Al ) 15, supplied by Zeolyst) using the required amount of aqueous ammonium heptamolybdate [(NH4)6Mo7O24‚4H2O] and ruthenium chloride both supplied by Aldrich. The catalyst was air-dried at ambient temperature for 12 h and then for 2 h at 393 K, and it was finally air-calcined for 4 h at 873 K. The solid samples were pressed, crushed, and sieved to separate catalyst granules in the size range of 20-35 mesh for subsequent use in aromatization reactions. Before exposure to the reactant, the catalyst underwent a gradual heating under argon in temperatureramped mode up to 873 K and was maintained at this temperature for 4 h. Then, the catalyst was treated in flowing air for 30 min at 973 K and in pure hydrogen at 573 K. 2.2.2. Reaction. Figure 1 is a sketch of the experimental setup which consists of a 17 mm i.d. membrane reactor hosting the 70 mm long and 9.5 mm external diameter Pd-Ag tubular membrane. The reactor loaded with 3 g of catalyst was connected to a gas feeding unit and analytical equipment. The reaction temperature, varied between 873 and 973 K, was controlled by a sheathed thermocouple. After catalyst pretreatment, a 90:10 (v/v %) CH4-Ar mixture was fed to the reactor at a mass-flow-controlled rate of 270 mL(STP)‚h-1‚g-1 methane hourly space velocity. All methane aromatization tests were run at atmospheric pressure. Hydrogen permeation was ensured using a vacuum pump connected to the permeation zone providing a sufficient driving force for withdrawal of the hydrogen produced from the aromatization out of the reaction zone. The gaseous reaction products were analyzed by means of a Perkin-Elmer gas chromatograph equipped with a flame ionization detector connected to a GS-Q capillary megabore column (30 m length and 530 µm i.d.) supplied by J&W Scientific and a thermal conductivity detector connected to a Carboxen 1010 capillary column (30 m length and 530 µm i.d.) supplied by Supelco. All detected gaseous products in the reactor exit stream were analyzed: benzene, toluene, ethylene, and ethane. Benzene was the major hydrocarbon prod-

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Fast deposition from concentrated Pd solutions leads to rapid formation of palladium films topping the support surface. This is not propitious to the formation of dense impervious membranes. Consecutive depositions from dilute Pd solutions allow more efficient in-depth incursions of Pd into the pores. However, this process brings about the formation of large Pd crystals and nonimpervious membranes.38 When following the conventional pretreatment and activation procedures,6,15 EDX showed a large number of seeds on the surface of the substrate that contain palladium and a small amount of tin. Consequently, the substrate was activated with Sn2+ and Pd2+ solutions using a modified activation procedure as described in section 2.1.3. The presence of divalent Sn was speculated to allow complete replacement of Pd2+ by Pd0 via the following redox reaction:13

Sn2+ + Pd2+ f Sn4+ + Pd0

Figure 2. SEM micrograph of the activated substrate: (a) ×1000; (b) ×3500.

uct, whereas naphthalene was detected in trace amounts mainly during the IMCR operation. Naphthalene was hence ignored in the carbon balance calculations. Also discounted in the carbon balance was the carbon contributed by the coke deposits, which were found to be negligible. Taking into account that benzene was the major reaction product, methane conversion into benzene was calculated based on the benzene concentration. After about 4000 min of reaction in CR and IMCR modes, the increase in the catalyst weight represented less than 0.1% of carbon contained in the total amount of methane used at 873 K and about 0.2% at 973 K. It is important to mention that blank experimental tests have shown that no CH4 conversion occurs on the wall of the membrane in the absence of catalyst. 3. Results and Discussion 3.1. Membrane Preparation, Pd and Ag Deposition, and Membrane Characterization. Commercially available sintered porous SS substrates in the form of either sheets or tubes exhibit very large pores because of limitations of powder metallurgical technology. SEM photographs of the porous SS tubes indicate that the pore size distribution is rather wide with a particle retention size of 0.2 µm (Figure 2a). To fabricate impervious composite membranes, a relatively thick Pd film should be deposited on the substrate surface by electroless plating and the deposition process must be carefully controlled to ensure complete pore coverage.

(1)

Figure 2 shows the SEM photographs of the activated substrate. As seen in Figure 2b, combination of the proposed membrane preparation method and activation under vacuum as described in section 2.1.3 was shown to improve Pd deposition because of better penetration deep into the pores. Ten consecutive depositions using at each dip pristine but dilute Pd solutions (Table 2A) led, because of the very low Pd concentration in the solution, to a membrane characterized by beautiful large crystals, as shown in Figure 3 at two different magnification scales.38-40 It seems that the agitation conditions might influence the crystal formation, but at the same time the agitation rate is also correlated with the removal rate of the incipient N2 formed in the plating reaction. However, as was also reported by Shu et al.,38 the membrane was found to be too permeable to Ar even at room temperature. Repeating several times the deposition experiments according to this protocol and modifying the Pd deposition rate led to the same conclusion. Dense impervious Pd membranes exclusively permselective to hydrogen were obtained by deposition from a concentrated Pd solution (Table 2B). The crystal dimensions are much smaller (Figure 4) than those in the case of deposition from the dilute Pd solutions. Also, the overwhelming amount of small particles yielded membranes completely impervious to Ar. The estimated membrane thickness assuming a Pd density equal to 11 900 kg/m3 was about 15 µm (i.e., 17 mg/cm2) over the 70 mm active length. The deposition efficiency calculated by gain weight before and after 8 h of Pd deposition was ca. 90%. Silver deposition occured after palladium deposition. Ag was deposited quite easily, with the deposition system being stabilized by an excess of the complexating agent. A SEM micrograph obtained after Ag deposition on the Pd membrane surface is shown in Figure 5. Silver deposition was found to be more efficient when using dilute AgNO3 solutions (Table 2C). The deposit composition analysis determined by EDX at various points of the membrane revealed a PdAg ratio ranging between 2.8 and 3.3. After deposition, thorough washing with water, and drying at 120 °C, hydrogen permeation tests were performed. Figure 6 illustrates the hydrogen permeation rate measurements between 573 and 773 K for a PdAg membrane obtained by a two-step procedure for palladium (concentrated solution) and silver depositions. The permeation fluxes per unit membrane surface area,


Figure 3. SEM micrographs of Pd deposited from a dilute Pd solution.

Figure 4. SEM micrographs of Pd deposited from a concentrated Pd solution.

for a pressure difference of 202 kPa, ranged between 2.31 × 10-3 m3‚m-2‚s-1 (at 773 K) and 3.21 × 10-3 m3‚m-2‚s-1 (at 973 K). The hydrogen permeation flux is proportional to the difference of the square roots of hydrogen partial pressure in the reaction and permeation sides:

P (r) 0.5 (r) 0.5 (p) 0.5 (p) 0.5 J ) B[(PH ) - (PH ) ] ) [(PH ) - (PH ) ]) 2 2 2 2 L DS D°S° (r) 0.5 (p) 0.5 ) - (PH ) ]) ‚ [(PH exp[-(∆HS + 2 2 L L (r) 0.5 (p) 0.5 ED)/RT][(PH ) - (PH ) ] (2) 2 2 (r) (p) where P, L, D, S, ∆HS, ED, PH , and PH are respec2 2 tively the permeability, the membrane thickness, the diffusion coefficient, the solubility coefficient, the molar enthalpy for hydrogen adsorption, the diffusivity activation energy, and the hydrogen partial pressures on the reaction and permeation sides. The permeabilitytemperature dependence recast in terms of Arrhenius plots and taking into account the ∆HS value for hydrogen adsorption on Pd41 yields a value for the diffusivity activation energy, ED ) 16.8 kJ‚mol-1. The membrane was then used in the reaction tests. 3.2. Reaction. As previously observed for a temperature of 873 K,36 the impact of hydrogen permeation on

Figure 5. Ag deposition on a Pd membrane.

methane conversion is quite modest at high methane space velocity. Above 600 mL(STP)‚h-1‚g-1, methane conversions for both CR and IMCR modes are almost identical. The IMCR mode becomes increasingly advantageous when the methane space velocity decreases, particularly below ca. 400 mL(STP)‚h-1‚g-1. Therefore, the reaction tests reported here were carried out at a methane space velocity of 270 mL(STP)‚h-1‚g-1.

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Figure 6. Hydrogen permeability through a Pd-Ag alloy membrane at different temperatures: before the reaction tests (773973 K) and after the reaction tests (973 K). Figure 8. Comparison between the time evolution of methane conversion to benzene at 873 and 973 K, over Ru-Mo/HZSM-5 catalyst at 270 mL(STP)‚g-1‚h-1 in cycled CR-IMCR modes. One CR step followed by one IMCR step defines a cycle.

Figure 7. Time evolution of methane conversion to benzene over Ru-Mo/HZSM-5 catalyst at different temperatures (873-973 K) and 270 mL(STP)‚g-1‚h-1 in the CR mode.

As illustrated in Figure 7, methane aromatization in an oxygen-free environment proceeds via an induction period36 prior to the formation of hydrocarbon products: ethylene, ethane, benzene, toluene, and naphthalene. The hexavalent molybdenum initially available in the zeolite channels likely as MoO3, or as (Mo2O5)2+ dimers, is first reduced by CH4 into active MoCx carbide species, which are believed to provide the active sites for C-H bond activation into CHx (x < 4) fragments and their condensation into, mainly, ethylene and ethane. Such products are further converted into aromatics on the acidic sites via oligomerization, cracking, and cyclization. Once the induction period is over, the RuMo/HZSM-5 catalyst activity levels off at conversions parametrized by the methane space velocity (or contact time).36 The methane conversion is strongly dependent on the temperature, as was expected from the theoretical conversion trend. With prolonged time exposures in the CR mode, the catalyst was found to exhibit good stability, as was also previously reported.36 No significant activity decline was detected during continuous operation at temperatures ranging between 873 and 973 K and lasting for as long as 1400 min. The membrane influence (hydrogen permeation) on methane conversion is shown in Figure 8, representing the time evolution of methane to benzene conversion

Figure 9. Temperature dependence of methane to benzene conversion in the absence of H2 permeation (CR mode) and with H2 permeation across the membrane (IMCR mode), with respect to the calculated equilibrium curve (dotted line).

during a CR-IMCR cycle spanning a time interval of 1600 min. Note that one CR step followed by one IMCR step defines a cycle. Right after one switched from the CR mode to the IMCR mode, benzene conversion climbed sharply to culminate at about 17%. As long as H2 bled across the membrane, benzene conversion decreased steadily from 17% to 14.6% after the first 600 min of the IMCR step, representing almost a 16% loss of the peak conversion. Conversely, at 873 K, this loss represents only 11% for the equivalent reaction time. To better illustrate the temperature influence on the reaction in the CR and IMCR modes, methane conversion as a function of temperature is presented in Figure 9 in the absence of H2 permeation (CR mode) and with H2 permeation across the membrane (IMCR mode). The IMCR conversion curve lies far remote on top of the calculated equilibrium curve (dotted line). Considering the beneficial effect of temperature on both equilibrium conversion and hydrogen permeability, it can be seen that the increase of temperature leads to a considerable increase of methane conversion in the IMCR mode


position is thermodynamically more favorable with respect to other compounds (hydrocarbons) at temperatures higher than 823 K, as confirmed by previous experimental works.43 Though Pd-Ag membranes are stable at high temperature from a hydrogen separation standpoint, their permeability might depend on their sensitivity to the reaction type. Coke laydown during methane nonoxidative aromatization both on the catalyst and on the membrane surface is inevitable especially at high temperatures.1,21,23,31 Despite the fact that the catalyst was shown to be quite stable at 973 K in the CR mode, we preferred to work at a lower temperature, e.g., 873 K, yielding lower conversion but securing greater membrane permeability for much longer operation time intervals. However, further studies concerning the improvement of the thermal stability of the hydrogen permeable membrane which seem so efficient especially in these very low conversion reactions are planned. Figure 10. Time evolution of methane conversion to benzene over a Ru-Mo/HZSM-5 catalyst at 270 mL(STP)‚g-1‚h-1 and 973 K in cycled CR-IMCR modes.

compared to the CR mode. Hence, hydrogen permeation entrained an increase in benzene conversion of about 150% of the equilibrium value over the whole range of reaction temperature explored. If at 873 K methane conversion was found to keep practically the same value for about 6000 min (in accordance with previous work36), at 973 K the catalyst started losing its activity after ca. 4000 min (Figure 10), suggesting an acceleration in the formation and buildup of carbonaceous species. This loss in activity at higher temperatures is very likely triggered by a larger probability in the formation of carbonaceous species during aromatization. A sharper decrease in conversion at 973 K (around 23% conversion loss) was further observed during the second IMCR step, where conversion tended toward that obtained in the CR mode. The beneficial effect of H2 permeation in shifting equilibrium toward benzene and hydrogen production weakened with time. However, methane conversion attained in the second and third CR steps at 973 K approached practically the same conversion profile as that in a straight-run CR mode (Figure 7). The global CR methane conversion loss is about 8% from its maximum value over a total reaction period of ca. 4000 min. This might suggest that the strongest decline in conversion at 973 K was mainly due to a limited blockage of the catalyst active sites that is more important at higher temperatures. On the other hand, no evidence of a decrease of hydrogen permeation through the membrane was observed at 873 K. On the contrary, at 973 K a slight decrease of hydrogen permeability was detected after the reaction tests (Figure 6, dotted line). This might be a reason for the observed methane conversion decrease in the IMCR mode. However, taking into account that methane conversion is rather low even at high temperatures, the membrane permeation capacity is overdesigned for the studied reaction. Therefore, a small reduction in membrane permeability does not influence practically hydrogen withdrawal from the reaction side. The applicability of the membrane cannot be limiting at least for a reasonable reaction time. The observed decrease of membrane permeability might be due to a small deposition of lowH/C-ratio carbonaceous species on the membrane surface.42 The C-formation reaction from methane decom-

4. Conclusion Pd-Ag/porous SS membranes obtained by an electroless plating technique were used in the oxygen-free methane aromatization over Ru-Mo/HZSM-5 in a catalytic membrane reactor at temperatures up to 973 K. The 0.5% Ru-3% Mo/HZSM-5 catalyst, prepared by incipient-wetness co-impregnation, was highly selective toward benzene production. In the IMCR mode, hydrogen permeation entrained a significant increase in methane conversion especially at higher temperatures (973 K), reaching levels well beyond the thermodynamic conversion but followed, however, by a quite sharp decrease. Continuous withdrawal across the membrane of H2 promoted the formation of carbonaceous species under very low hydrogen pressure, especially at high temperatures, thus negatively affecting mainly the catalyst activity. Acknowledgment Financial support from the Natural Sciences and Engineering Research Council of Canada (NSERC), the Fonds pour la formation de chercheurs et d’aide a` la recherche (Que´bec, Canada), and the Natural Gas Technologies Center (Boucherville, Que´bec, Canada) is gratefully acknowledged. Literature Cited (1) Shu, J.; Grandjean, B. P. A.; Van Neste, A.; Kaliaguine, S. Catalytic Palladium-Based Membrane Reactors: A Review. Can. J. Chem. Eng. 1991, 69, 1036. (2) Jemaa, N.; Grandjean, B. P. A.; Kaliaguine, S. Diffusion Coefficient of Hydrogen in a Pd-Ag Membrane: Effect of Hydrogen Solubility. Can. J. Chem. Eng. 1995, 73, 405. (3) Li, A.; Liang, W.; Hughes, R. Characterisation and permeation of palladium/stainless steel composite membranes. J. Membr. Sci. 1998, 149, 259. (4) Ho¨llein, V.; Thornton, M.; Quicker, P.; Dittmeyer, R. Preparation and characterization of palladium composite membranes for hydrogen removal in hydrocarbon dehydrogenation membrane reactors. Catal. Today 2001, 67, 33. (5) She, Y.; Han, J.; Ma, Y. H. Palladium membrane reactor for the dehydrogenation of ethylbenzene to styrene. Catal. Today 2001, 67, 43. (6) Shu, J.; Grandjean, B. P. A.; Ghali, E.; Kaliaguine, S. Simultaneous deposition of Pd and Ag on porous stainless steel by electroless plating. J. Membr. Sci. 1993, 77, 181.

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Received for review July 1, 2002 Revised manuscript received November 4, 2002 Accepted November 12, 2002 IE020486N

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