Chemical Reaction Engineering and Catalysis Issues in Distributed

Ind. Eng. Chem. Res. 2011, 50, 523–530

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Chemical Reaction Engineering and Catalysis Issues in Distributed Power Generation Systems Paraskevi Panagiotopoulou, Dimitris I. Kondarides, and X. E. Verykios* Department of Chemical Engineering, UniVersity of Patras, GR-26504 Patras, Greece

Three issues are most critical in fuel processors which are integrated with polymer electrolyte membrane (PEM) fuel cells for power generation: the steam reformer configuration, which must ensure very rapid heat transfer to the reformation zone, the development of highly active catalysts for the water-gas shift (WGS) reaction and the development of highly selective catalysts for the methanation reaction. The heat-integrated wall reactor (HIWAR), in either the tubular or the plate form, which offers very rapid heat exchange between a combustion and a reformation zone, is proposed. This configuration results in high efficiency and compact design. Platinum catalysts supported on a “reducible” metal oxide such as TiO2 exhibit high WGS activity, which can be further improved with the addition of alkali or alkaline earth promoters. Titania-supported ruthenium catalysts are capable of completely and selectively methanate CO in the presence of excess CO2, provided that catalyst characteristics are optimized and operating conditions are properly selected. Integration of the three developments can result in efficient power generation systems. 1. Introduction The concept of the “hydrogen economy” seems to be moving into the world of political and strategic planning as well as into business and enterprise. This is primarily due to the fact that it has been widely recognized and accepted that major improvements in energy efficiency of electric power generation must be achieved in order to reduce emissions of pollutants and, particularly, CO2. Currently, power is produced in a centralized manner and distributed over long distances. This mode has the advantage of economy of scale but also major disadvantages such as loss of power due to long distance transmission, infrastructure cost, health and safety issues, and environmental problems. Many analysts predict that centralized power production will be gradually replaced by distributed power generation. This concept implies small-scale production on site and on demand. This mode offers significant advantages such as utilization of thermal energy for district heating, which implies enhanced overall efficiencies, lower maintenance costs, and reduced capital requirements. However, the most critical advantage lies in the possibility for early adaptation of the socalled hydrogen economy, i.e. power production with fuel cells via hydrogen, for reduced pollution and greenhouse gas (GHG) emissions. While hydrogen can be produced from excess renewable electricity (photovoltaic, wind) via electrolysis, with zero GHG emissions, this can not be adapted in distributed power generation due to severe difficulties in hydrogen storage and transportation. To overcome these difficulties, the idea of hydrogen production on site and on demand has been gaining ground. This implies that an appropriate hydrogen carrier, which could be liquid or gas, can be used to extract hydrogen.1-8 Biomass derived fuels, such as bioethanol and biogas, offer very low GHG emissions but are hampered with high capital cost requirements and technically challenging integration issues. On the other hand, fossil fuels, such as natural gas or liquified petroleum gas (LPG) can be used to obtain hydrogen from and produce power with significant reduction of GHG emissions. * To whom correspondence should be addressed. Fax No.: +30 2610 991527. E-mail: [email protected].

Regardless of the type of fuel utilized, the need for reformation and fuel processing is obvious. Fuel processors for distributed power generation systems must possess certain desirable characteristics, such as high efficiency, compactness, no production of atmospheric pollutants, ease of integration with fuel cells, and short startup time. Above all, the fuel processor must produce a hydrogen rich stream suitable to be fed into fuel cells. For low-temperature PEM fuel cells, which are the most advanced such devises, the most critical requirement is that the CO content of the feed stream is less than 30-50 ppm.2,3 A typical block diagram of a fuel processor operating in steam reforming mode is shown in Figure 1A. The process comprises the reformer, the water-gas shift (WGS) reactors (low and high temperature), and the selective methanation reactors which ensure that the CO content is maintained at the desired low levels. Selective methanation is used here instead of selective oxidation because it offers significant advantages. The process flow diagram shown in Figure 1B clearly indicates the complexity of the system, which, apart from the reactors, comprises a number of other devices such as a steam generator, heat exchangers, condensers, burners, metering and control systems, etc. A challenging issue is to utilize the cathode off-gas, which may contain up to 20% of the produced hydrogen, in the reformer so as to enhance the efficiency of the process, as shown in Figure 1B. There are three major issues related with the above process: (1) appropriate configuration of the reformer so as to achieve very rapid heat transport to the reformation zone; (2) development of very active water-gas shift catalysts so as to approach equilibrium at low temperatures; (3) development of selective catalysts for CO methanation in the presence of CO2. These three important issues are discussed in the present manuscript and recent developments in our laboratory are presented. 2. Advanced, Heat-Exchanger Type Microreformer in Tubular or Plate Form The microreformer which has been developed with the goal to obtain very high rates of heat transport is based on the concept of the heat-integrated wall reactor (HIWAR),9-11 which is schematically shown in Figure 2. The reactor comprises a

10.1021/ie100132g  2011 American Chemical Society Published on Web 05/17/2010

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Ind. Eng. Chem. Res., Vol. 50, No. 2, 2011

Figure 1. Typical (A) block and (B) process flow diagram of a fuel processor operating in the steam reforming mode.

Figure 2. Schematic diagram of a typical heat-integrated wall reactor (HIWAR).

metallic tube or plate, on one surface of which a combustion catalyst is deposited in the form of a thin film. On the other surface, a reforming catalyst is deposited in the same form. Fuel and air are introduced on the combustion side and combustion of the fuel is taking place catalytically on the catalyst film producing heat. Steam and fuel are introduced on the reformation side and are reacted on the reforming catalyst film. Heat is transported very rapidly through the wall of the metallic tube or plate from the combustion to the reformation zones. The configuration described above offers two important advantages: (a) controlled burning of the fuel since it is taking place catalytically, at reduced temperature and more uniformly

over the entire surface; (b) rapid transport of heat into the reformation zone and achievement of relatively uniform temperatures throughout the surface. This is illustrated in Figure 3 which pertains to ethanol reforming in a tubular HIWAR type microreformer. The reactor was constructed from two concentric tubes (SS 310) of outer diameter of 25 and 12 mm, respectively, with catalytic zones 15 cm in length. The catalyst deposited in the reformation zone was 20% Ni/La2O3/Al2O3, and the quantity deposited was 4 g. The catalyst deposited in the combustion zone was 0.5% Pt/Al2O3, and the quantity deposited was also 4 g. In Figure 3A, the catalyst films were placed symmetrically on either side of the tube wall. The feed in the combustion side consists of 5.5% ethanol, 7.5% steam, balance air, at a total flow rate of 2300 cm3/min. The feed in the reforming side consists of a steam-ethanol mixture (3:1, molar) at a flow rate of 440-1240 cm3/min (curves B-E). Curve A pertains to flow of 300 cm3/min He in the reformation zone, which is used as a reference case. Countercurrent flow of the two streams is employed. The temperature profiles in the reformation side clearly demonstrate that temperature is controlled in a satisfactory manner while it is relatively uniform, as compared to the reference case (use of He). Catalytic results, shown in Figure 3B indicate that complete reformation of ethanol and very high hydrogen selectivities are achieved at flow rates as high as 700 cm3/min, and they decline at higher flow rates. From the temperature profiles, it is also obvious that maximum temperature (reforming zone) is achieved outside the catalytic zone.


Figure 3. Effect of the reforming flow on (A) the temperature profile and (B) the conversion of ethanol and product selectivities in the reaction of ethanol reforming in a tubular HIWAR type microreformer: (experimental conditions) combustion catalyst 0.5% Pt/Al2O3; reformation catalyst 20% Ni/La2O3/Al2O3; combustion feed 5.5% EtOH, 7.5% H2O, 87% air, FT ) 2300 cm3/min; reforming feed (A) 300 cm3/min He, (B-E) EtOH/H2O ) 1/3.

This results in a loss of efficiency. To resolve this, the combustion and reformation zones can be placed at somewhat different positions on the tube surfaces, as illustrated in Figure 4A. In this case, better temperature profiles are achieved. The catalytic results, shown in Figure 4B, show that complete conversion of ethanol and very high hydrogen selectivities are achieved at flow rates as high as 1000 cm3/min. The results discussed above clearly demonstrate the efficiency of application of the HIWAR concept in reformation reactions which have a need for rapid supply of heat in order to overcome the heat of reaction requirements. Application of this concept results in highly compact and highly efficient and safe microreformer configurations. 3. Active Catalysts for the Water-Gas Shift Reaction The WGS reaction is moderately exothermic and equilibriumlimited, and therefore, the desired CO levels can only be achieved at low temperatures. CO + H2O T CO2 + H2,

∆H ) -41.1 kJ/mol (1)

As a result, WGS catalysts for fuel processor applications should be sufficiently active in the temperature range of 200-280 °C, in order to be able to reduce CO concentration to the desired level (0.5-1.0%). In addition, they should be thermally stable, and resistant to poisoning under reforming conditions.2-4 Conventional low-temperature (Cu/ZnO/Al2O3) and high-temperature (Fe3O4/Cr2O3) WGS catalysts, which have been applied

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Figure 4. Effect of catalyst film position on (A) the temperature profile and (B) the conversion of ethanol and products selectivities for the reaction of ethanol reforming in a tubular HIWAR type microreformer. Experimental conditions are the same as those listed in Figure 3.

for several decades in large scale steady-state operations, cannot be used in fuel cell applications due to volume restrictions and also due to the requirements for reduced startup times, durability under steady state and transient conditions and stability to condensation and poisons. This has led to intense interest in the development of novel catalyst formulations for the WGS reaction, which are based on noble metals supported on metal oxide carriers.12-27 A detailed investigation has been carried out in this laboratory in an attempt to identify the key parameters that determine the WGS activity of supported noble metal catalysts and to develop active, selective, and stable materials suitable for practical applications. Results show that catalytic performance depends on several parameters, including the nature and loading of the dispersed metallic phase,12,13 the nature and physicochemical characteristics of the metal oxide support,13,14 and the presence of alkali/alkaline earth promoters.15,16 3.1. Effect of the Nature of the Dispersed Metallic Phase. Results of catalytic performance tests obtained with the use of Pt, Rh, Ru, or Pd catalysts of the same loading (0.5 wt %) supported on TiO2,12 CeO2,13 or Al2O313 showed that Pt exhibits the highest WGS activity. In particular, it was found that the turnover frequency (TOF) of CO conversion at 250 °C varies in the order of Pt > Rh ≈ Ru > Pd, with Pt being 20-50 times more active than Pd, depending on the metal oxide support, whereas Rh and Ru exhibited an intermediate performance.12,13 Similarly, Radhakrishnan et al.26 reported that the order of activity for the WGS reaction is Pt > Rh >Ru > Pd over catalysts supported on ceria-zirconia oxides. Different ranking of noble metals have been reported by Grenoble et al.,25

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Figure 6. Effect of the nature of the support on the turnover frequency of CO conversion of Pt catalysts (0.5 wt %) supported on the indicated metal oxide powders, for the WGS reaction. Experimental conditions are the same as those in Figure 5.

Figure 5. (A) Effect of metal loading on the catalytic performance of Pt catalysts supported on TiO2 for the WGS reaction. (B) Arrhenius plots of reaction rates obtained over Pt/TiO2 catalysts of variable Pt crystallite size. (experimental conditions) Mass of catalyst 100 mg; particle diameter 0.18 < dp < 0.25 mm; feed composition 3% CO, 10% H2O (balance He); total flow rate 200 cm3/min (Reproduced with permission from ref 12. Copyright 2004 Elsevier).

Figure 7. Effect of alkali/alkaline earth-promotion on the catalytic performance of Pt/X-TiO2 (X ) Na, Cs, Ca, Ba) catalysts, for the WGS reaction. Experimental conditions are the same as those in Figure 5.

who found that activity varies in the order of Ru > Pt > Pd ≈ Rh over Al2O3-supported catalysts. These differences may be attributed to the different nature of the oxide carriers used, which, as will be discussed below, affects significantly catalytic activity. 3.2. Effect of Metal Loading and Crystallite Size on Catalytic Performance. Representative results obtained over Pt/TiO2 catalysts of variable Pt content are shown in Figure 5A, where conversion of CO (XCO) is shown as a function of temperature. It is observed that an increase of Pt loading from 0.1 to 5 wt % results in a progressive and significant shift of the CO-conversion curve toward lower temperatures. However, this is solely due to the increased amount of available surface Pt sites: Results of kinetic measurements obtained under differential reaction conditions (Figure 5B) demonstrate that the specific reaction rate (TOF) does not depend on metal loading (0.1-5.0 wt %) or Pt crystallite size (1.2-16.2 nm). Qualitatively similar results were obtained over Pt/CeO2,13 Pt/Al2O3,13 and Ru/TiO212 catalysts. Thus, it can be concluded that the WGS reaction over supported noble metal catalysts is structureinsensitive, as far as the metallic phase is concerned, in agreement with previous studies.24,25,27,29 3.3. Influence of the Nature of the Support on Catalytic Activity. Results of our detailed investigation of a variety of metal-support combinations showed that the key parameters that determine the WGS activity of dispersed noble metal catalysts are related mainly to the nature and the physicochemical characteristics of the metal oxide support. This is clearly shown in the Arrhenius diagram of Figure 6, which summarizes results of catalytic activity measurements obtained over Pt catalysts supported on various oxides. It is observed that specific reaction rate is about 2 orders of magnitude higher when Pt is

supported on TiO2, as compared to SiO2. Generally, it has been shown that noble metals exhibit significantly higher activities when supported on “reducible” (e.g., TiO2, CeO2, La2O3) compared to “irreducible” oxides (e.g., Al2O3, MgO, SiO2).13,17 This is in accordance with the results of Sandoval et al.30 which pertains to gold supported catalysts. It has also been shown that the activity of Pt/TiO2 catalysts depends strongly on the primary crystallite size of the support.14 Generally, activity can be improved when the reducible metal oxide (MOx) is dispersed on a high surface area support, such as Al2O3.17 This has been attributed to the higher reducibility of smaller MOx crystallites which may affect the WGS activity either directly (redox properties) or indirectly (population and reactivity of surface hydroxyl groups).17 3.4. Effects of Alkali/Alkaline Earth-Promotion on Catalytic Performance. The influence of selected alkalis and alkaline earths added on TiO2 on catalytic activity was also investigated and representative results are shown in Figure 7, where XCO is shown as a function of reaction temperature for Pt/TiO2 samples promoted with Na, Cs, Ba, or Ca. The catalytic performance of the unpromoted catalyst is also shown for comparison. It is observed that, in all cases, addition of alkalis and alkaline earths results in a significant shift of the conversion curve toward lower reaction temperatures. This is more pronounced for the Na- and Ca-promoted samples, compared to samples containing Cs or Ba. The Na- and Ca-promoted catalysts are able to reach equilibrium CO conversion at temperatures around 325 °C, which is about 75 °C lower than that obtained for the unpromoted catalyst. Qualitatively similar results were obtained for Ru and Pd catalysts.15,16 Results of mechanistic and kinetic studies15,16 indicated that added alkalis interact strongly with the TiO2 surface and result in the creation


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Figure 9. Catalytic performance of Rh (0.5 wt %) supported on the indicated commercial oxide carriers for the selective methanation of CO: (solid symbols) CO conversion, (open symbols) CO2 conversion; (experimental conditions) mass of catalyst 150 mg; particle diameter 0.18 < dp < 0.25 mm; feed composition 1% CO, 15% CO2, 50% H2 (balance He); total flow rate 200 cm3/min. Figure 8. Effect of space velocity on the catalytic performance of 0.5% Pt/1% CaO-TiO2 catalyst under realistic feed composition of (A) HT WGS: 9.7% CO, 38.7% H2O, 44.8% H2, 6.8% CO2 and (B) LT WGS: 1.6% CO, 29.9% H2O, 52.2% H2, and 16.3% CO2.

of *s-Ti3+ defects (where *s denotes an oxygen vacancy) which, in turn, affect the chemisorptive properties of noble metal atoms located at the metal-support interface. The so formed NM-*s-Ti3+ sites were proposed to be the catalytically active sites for the WGS reaction, since they are capable of adsorption and activation of both CO and H2O.15,16,21 3.5. Catalytic Performance under Realistic Reaction Conditions. The performance of optimized WGS catalysts was investigated under realistic reaction conditions, i.e., with the use of catalysts in the form of 1/16-in. pellets and feed compositions relevant to the high and low temperature shift (HTS), as applied in hydrogen generation. Typical results obtained with the use of 0.5% Pt/(1% CaO-TiO2) catalyst are summarized in Figure 8, where conversion of CO is shown as a function of temperature for space velocities ranging between 4000 and 9800 1/h. It is observed that, under HTS conditions (Figure 8A), conversion of CO reaches equilibrium at temperatures lower than ca. 350 °C even at the highest space velocity investigated. Further reduction of CO content to the desired levels (∼0.5%) can be achieved under LTS conditions at reaction temperatures around 250 °C (Figure 8B). 4. Selective Catalysts for the Methanation of CO In fuel processors, the CO methanation reaction (eq 2) can be used as the final purification step of reformate gas to reduce concentrations of CO to the extremely low levels ( YSZ > Al2O3 > CeO2 > SiO2 with TOF being ca. 2 orders of magnitude higher when Rh is supported on TiO2, compared to CeO2. Qualitatively similar results were obtained for Ru catalysts.34 4.2. Effect of the Nature of the Metallic Phase on Catalytic Activity. Representative results obtained over M/Al2O3 catalysts (0.5 wt %) are summarized in the Arrhenius diagram of Figure 10. It is observed that catalytic activity for CO hydrogenation depends appreciably on the nature of the dispersed metallic phase and follows the order of Rh > Ru . Pt > Pd, with Rh and Ru catalysts being 1-2 orders of magnitude more active than Pt and Pd, at 250 °C. A different ranking of noble metals is observed with respect to their activity for CO2 hydrogenation, where the reaction rate at 350 °C decreases by about 1 order of magnitude in the order Pt > Ru > Rh ∼ Pd (Figure 10). However, in our previous work it has been found that the RWGS reaction is operable already at 230 °C for Pt and Pd catalysts,28 which makes the determination of the specific rate for methane formation difficult. The activity order of noble metals may be similar to that observed in the methanation of CO2 (in the absence of CO), i.e. Ru > Rh > Pt > Pd,28 taking into account the selectivity toward methane. Results of Figure 10 are in agreement with those reported by Utaka et al.,35 who found that supported Ru catalysts consumed hydrogen, mainly by methanation of CO and CO2, while Pt/Al2O3 demonstrated high CO conversion, without methanation, due to the RWGS reaction. 4.3. Effect of Noble Metal Crystallite Size on Catalytic Activity. An important parameter, which may affect reaction rate is metal crystallize size. Results obtained over Ru/TiO2 catalysts of variable metal loading (0.5-5 wt %), shown in Figure 11, demonstrate that the CO/CO2 hydrogenation reactions are structure sensitive, i.e., the reaction rate per surface metal atom (turnover frequency, TOF) depends on crystallite size of the dispersed metallic phase. In particular, TOFs of both CO at 215 °C (Figure 11A) and CO2 at 330 °C (Figure 11B) over Ru/TiO2 catalysts increase by more than 1 order of magnitude with increasing Ru crystallite size from 2.1 to 4.5 nm. The effect is smaller for Ru supported on Al2O3, where TOFs of CO and CO2 conversion increase by a factor of 3.5 and 17, respectively, with increasing metal crystallite size from 1.3 to 13.6 nm. In contrast to what is observed for Ru catalysts, specific activity of Rh/Al2O3 for CO and CO2 conversion decreases by a factor of 3 and 2.3, respectively, with increasing Rh crystallite size from 1.4 to 5.1 nm (Figure 11). The structure sensitivity of dispersed noble metals for the methanation of CO, CO2, and their mixtures has been investi-

Figure 11. Effect of metal crystallite size (dM) on the turnover frequencies of (A) CO conversion at 215 °C and (B) CO2 conversion at 330 °C obtained from Ru/Al2O3, Rh/Al2O3, and Ru/TiO2 catalysts, for the selective methanation of CO. Experimental conditions are the same as those in Figure 9.

gated by several authors over various metal-support combinations.31,36-39 For example, Kowalczyk et al.38 reported that TOFs for both CO and CO2 methanation reactions increase with decreasing Ru dispersion over Ru/Al2O3, Ru/MgO, and Ru/ MgAl2O4 catalysts. Ojeda et al.39 found that, although CO conversion over Rh/Al2O3 catalysts decreases on larger rhodium particles, TOF increases by a factor of 4 with increasing metal particle size from 5 to 30 nm. Regarding combined methanation of CO/CO2 mixtures, Dagle et al.31 reported that the performance of Ru/Al2O3 catalysts is markedly affected by Ru crystallite size, in the range of 7.5-34.2 nm. Results presented in Figure 11A and B may be explained by considering that the surface structure of metal particles changes by varying particle size, i.e., the fraction of coordinatively unsaturated metal atoms at edges and corners relative to the total metal atoms at the surface decreases with increasing crystallite particle size. For ruthenium catalysts, the correlation between particle size and TOF (Figure 11) indicates that hydrogenation of CO/CO2 preferably occurs on flat Ru surfaces. It has been proposed that both reactions proceed via dissociation of carbon monoxide to C and O atoms, followed by their hydrogenation into CH4 and H2O.37,39,40 Since both CO and CO2 hydrogenation require a CO dissociation step, this probably indicates that larger Ru particles facilitate the cleavage of the C-O bond. In contrast, the key step in CO/CO2 hydrogenation over Rh catalysts seems to be favored at the edge and corner sites of the metal crystallites. However, a clear conclusion regarding the different effects of crystallite size observed for Ru and Rh catalysts cannot be obtained from the present results. 4.4. Catalytic Performance under Realistic Reaction Conditions. The reformate gas obtained under realistic conditions contains considerable amounts of steam, which affects CO removal. Results obtained over 5% Ru/TiO2 catalysts (Figure 12) show that addition of up to 30% water vapor in the feed does not practically affect conversion of CO but retards CO2 methanation, thereby expanding the temperature window of


Figure 12. Effect of addition of water vapor in the feed (0-30%) on the catalytic performance of 5% Ru/TiO2 catalyst for the selective methanation of CO: (solid symbols) CO conversion; (open symbols) CO2 conversion (Reproduced with permission from ref 34. Copyright 2009 Elsevier).

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“reducible” rather than on “irreducible” oxides. The specific reaction rate (TOF) does not depend on metal loading, dispersion, or crystallite size for all metal-support combinations investigated. Titania-supported Pt catalysts exhibit sufficiently high activity, which can be further improved by addition of small amounts of alkali or alkaline earth promoters. Catalytic activity for CO/CO2 hydrogenation is much higher for Ru and Rh catalysts, compared to Pd or Pt, and is significantly improved when supported on TiO2, compared to Al2O3, CeO2, YSZ, or SiO2. Both hydrogenation reactions are structure sensitive, and the specific activity (TOF) increases substantially with increasing Ru crystallite size or with decreasing Rh crystallite size. Among the various noble metal-support combinations investigated, optimal results were obtained over the 5% Ru/TiO2 catalyst, which is very active, selective, and stable under realistic reaction conditions and, therefore, is a promising candidate for use in the selective methanation of CO for fuel cell applications. Literature Cited

Figure 13. Catalytic performance of 5% Ru/TiO2 catalyst for the selective methanation of CO, under realistic reaction conditions: (experimental conditions) feed composition 0.5% CO, 14% CO2, 55.5% H2, and 30% H2O; SV ) 5000 1/h.

operation (shadowed area) for the title reaction. It has also been found that among the various noble metal-support combinations investigated, optimal results are obtained over the 5% Ru/TiO2 catalyst, which exhibits excellent activity, selectivity, and stability for extended time on stream.34 Thus, the catalytic performance of 5% Ru/TiO2 catalyst in the form of 1/16-in. pellets has been investigated under realistic reaction conditions using a feed stream consisting of 0.5% CO, 14% CO2, 55.5% H2, and 30% H2O and a space velocity of 5000 h-1. Results obtained are summarized in Figure 13, where the catalytic performance of 5% Ru/TiO2 catalyst in powder form obtained with a space velocity of 48 800 1/h is also shown for comparison. It is observed that both conversion curves of CO and CO2 are shifted toward lower temperatures with decreasing space velocity from 48 800 to 5000 1/h. It is important to note that under realistic reaction conditions the 5% Ru/TiO2 catalyst is able to completely and selectively methanate CO at temperatures around 155 °C and, therefore, is a promising candidate for use in selective methanation of CO for fuel cell applications. 5. Summary and Conclusions Fuel processing for hydrogen production for fuel cell applications possesses three significant requirements: rapid heat transport to the catalytic sites, active WGS catalysts, and selective CO methanation catalysts. The HIWAR reactor in either the tubular or the plate form offers very rapid heat exchange which results in high efficiency and compact design of the reformer. Catalytic activity for the WGS reaction of supported noble metal catalysts depends strongly on the nature of both the metallic phase and the metal oxide support employed. Platinum catalysts are generally more active than Ru, Rh, and Pd and exhibit significantly higher activity when supported on

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ReceiVed for reView January 20, 2010 ReVised manuscript receiVed April 26, 2010 Accepted April 30, 2010 IE100132G

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