Catalytic Performance of Rhodium-Based Catalysts for CO Preferential

Jul 2, 2008 - reactions. Tests at progressively lower O-to-CO feed ratio and the same WSV value were carried out for the sake of reducing H2 consumpti...
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Ind. Eng. Chem. Res. 2008, 47, 5304–5312

Catalytic Performance of Rhodium-Based Catalysts for CO Preferential Oxidation in H2-Rich Gases Camilla Galletti, Stefania Specchia,* Guido Saracco, and Vito Specchia Dipartimento di Scienza dei Materiali ed Ingegneria Chimica, Politecnico di Torino, Corso Duca degli Abruzzi 24, 10129 Torino, Italy

CO preferential oxidation (CO-PROX) can lead to a reduction of the CO content in the hydrogen-rich gas derived from hydrocarbon re-forming down to at least 10 ppmv or below, so as to enable its direct feed to standard polymer electrolyte membrane fuel cells (PEM FCs). Rh-based catalysts supported on A zeolites (3A, 4A, and 5A), alumina, titania, and ceria were prepared and tested for potential application in CO-PROX operating over a temperature range compatible with PEM FCs (80-100 °C). Among the prepared catalysts, the 1% Rh-zeolite 3A catalyst, tested with a weight space velocity (WSV) of 0.66 N · L · min-1 · gcat.-1, was found to be the most suitable one for the CO-PROX at low temperature: it reduced the inlet CO concentration below 10 ppmv within a temperature range of at least 80-120 °C without the appearance of undesirable side reactions. Tests at progressively lower O-to-CO feed ratio and the same WSV value were carried out for the sake of reducing H2 consumption and improving CO-PROX selectivity. For 1% Rh-3A zeolite the minimum λ value, ensuring a sufficiently wide temperature range of a nearly complete CO conversion at temperatures compatible with PEM FCs operation, was found to be equal to 3. Finally, to decrease the catalyst cost, the Rh load on the catalyst was tentatively reduced from 1 to 0.5%. A better distribution of the active element crystallites over the support surface was even obtained for this last catalyst. When operating at λ ) 3 and at WSV ) 0.66 N · L · min-1 · gcat.-1, the 0.5% Rh-3A catalyst could effectively reduce the inlet CO concentration below 10 ppmv within a temperature range of 100-140 °C, without the appearance of undesired side reactions. 1. Introduction Polymer electrolyte membrane fuel cells (PEM FCs) fed with pure hydrogen can provide power to electrically operated vehicles with virtually no pollutants. At present, significant introduction of PEM FCs is limited by their cost, the unavailability of an H2 distribution infrastructure, and the rather low amounts of H2 gas that can be safely stored on board vehicles.1 A number of research projects involving the authors,2–5 though, are aimed at coupling an internal combustion engine (ICE) for a powertrain with a FC-based auxiliary power unit (APU) for any other on board power requirement. The APU systems are based on a reforming process (either steam or autothermal6–12) of hydrocarbon feedstock fuels, integrated with a PEM FC on board vehicles. As extensive infrastructures already exist for gasoline and diesel oil, these are the preferred sources of hydrocarbon feedstocks for on board reforming.13 The H2 fuel gas for the PEM FC is required to be “nearly CO free” as FC platinum-based anodes get poisoned by traces of CO. The H2rich gas produced by catalytic reforming of fossil fuels, such as gasoline or diesel oil, followed by water gas shift (WGS) reaction, may go through catalytic selective CO oxidation to completely remove its undesired CO content.9,14,15 A simplified flow diagram of an APU system based on hydrocarbon reforming and PEM FC with approximate operating temperatures for each step of the process is shown in Figure 1. The CO concentration of the synthesis gas from the reformer is at first reduced by means of the WGS reaction, which also enriches the H2 concentration in the gas stream: CO + H2O T CO2 + H2 ∆H298 ) -41.2 kJ/mol

(1)

The WGS reaction is generally operated at two different temperature steps (high temperature, HT-WGS, and low tem* To whom correspondence should be addressed. Tel.: +39-0110904608. Fax: +39-011-0904699. E-mail: [email protected].

perature, LT-WGS) to minimize the overall amount of catalyst required, which is typically a noble metal supported on ceria. However, the outlet gas from the LT-WGS reactor, still contains approximately 0.5-1% by volume of CO. Tolerable levels of CO concentration for standard-grade PEM FCs are 10 ppmv, with peaks of 50 ppmv for a few minutes, so a further removal of CO from the gas stream has to be performed.16 Preferential oxidation of CO to CO2 (CO-PROX) is a widely studied option to achieve these residual CO levels. CO-PROX is an exothermic process which involves the CO oxidation in the reformate stream to CO2 over a suitable catalyst using molecular O2 fed on purpose. However, during the process, the H2 in the reformate stream (which has a typical concentration of 30-60% by volume after the WGS reactions depending on the reforming process adopted) should not be consumed at all to ensure reasonable fuel efficiency. A decrease in H2 concentration results indeed in a decrease in power generation. Therefore, both active and selective catalysts are required to convert CO to CO2 while minimizing the H2 oxidation to H2O. Various catalysts have been proposed and tested in the literature for selective CO-PROX in H2-rich streams, most notably the noble metals Pt, Pd, Ru, and Rh carried on Al2O3 or zeolites.17–23 Supported Pt catalysts have probably been the most investigated so far.24–30 Oh and Sinkevitch24 in particular carried out CO-PROX tests with synthetic reformate gases (0.85% H2, 900 ppmv CO, and 800 ppmv O2) with Al2O3-supported noble metals (Rh, Ru, Pt, and Pd). The tests for catalyst activity and selectivity were carried out from room temperature to 300 °C. They found that Rh/Al2O3 and Ru/Al2O3 catalysts had a superior activity for CO oxidation and reached a complete CO conversion at temperatures lower than those for Pd/Al2O3 and the most widely used Pt/ Al2O3 catalysts. In addition, Rh/Al2O3 catalyst showed the highest selectivity among all the catalysts tested. Many other

10.1021/ie0713588 CCC: $40.75  2008 American Chemical Society Published on Web 07/02/2008

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Figure 1. Simplified scheme of a liquid hydrocarbon processor for hydrogen production. The average operating temperatures of each process step are indicated.

Figure 2. Microreactor bench test scheme.

Rh-based catalysts have been studied extensively for COPROX31–33 for their good performance at lower temperature compared with other noble metals based catalysts. As a consequence, in the present work, Rh-supported catalysts were prepared using different supports (γ-Al2O3, TiO2, CeO2, and various zeolites) and tested for potential application in COPROX units operating at temperatures compatible with the PEM FCs.18 2. Experimental Section 2.1. Catalyst Preparation and Characterization. The metal oxide materials (γ-Al2O3, TiO2, and CeO2) to be used as Rh carriers were prepared through the solution combustion synthesis (SCS) method.34 Conversely, zeolites with different pore sizes were purchased from Fluka: 3A-type zeolite (K12[(AlO2)12(SiO4)12] · H2O), 4A-type zeolite (Na12[(AlO2)12(SiO4)12] · xH2O), and 5A-type zeolite (CanNa12-2n[(AlO2)12(SiO4)12] · H2O), with pores of about 3, 4, and 5 Å, respectively. The noble metal Rh was added by incipient wetness impregnation (IWI) on the carriers using Rh(NO3)3 as precursor; Rh(NO3)3 was dissolved in distilled water, and the solution was added dropwise over the various carriers meanwhile thoroughly mixing the whole mass. The mixtures were then placed in an oven at 200 °C to remove H2O. The catalyst powders were pressed at 125 MPa into tablets, which were then crushed in an

agate mortar and sieved to separate granules of 0.25-0.42 mm in size. The Rh concentration in all catalysts was equal to 1% by weight, with the only exception of a successively prepared Rh-3A zeolite catalyst, whose Rh concentration was lowered to 0.5% to check the effect of the noble metal load on the catalytic performance. The so-obtained catalysts particles were calcined in pure O2 for 2 h at 350 °C to remove the nitrate ions and, in turn, to form a Rh oxide, which was then reduced in pure H2 for 2 h at 350 °C. X-ray diffraction (XRD) analysis (Philips PW1710 apparatus equipped with a monochromator for the Cu KR radiation) was performed on the prepared catalysts prior to carrying out the catalytic activity tests in order to verify their effective composition and, if visible, to derive a qualitative indication of the presence of comparatively large Rh metal crystallite. Then, the obtained Rh-based catalysts were analyzed by highresolution transmission electron microscopy (HRTEM, Jeol JEM 2010 apparatus) to investigate the metal dispersion on the supports. Temperature-programmed desorption (TPD) of CO2 was performed on the zeolite 3A, 4A, and 5A supports by a TPD/ R/O apparatus (Thermoquest TPD/R/O 1100 analyzer equipped with a Baltzer Quadstar 422 quadrupole mass spectrometer). CO2 saturation was preliminarily performed by flowing 50 N · cm3 · min-1 of CO2 for 1 h at 40 °C; then a He flow rate of

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10 N · cm3 · min-1 was fed to the reactor meanwhile increasing the temperature at 5 °C · min-1 rate up to 450 °C. The same apparatus was used for a test of TPD of H2 on 1% Rh-3A and 0.5% Rh-3A catalysts, pretreated by flowing 50 N · cm3 · min-1 of O2 and then 50 N · cm3 · min-1 of H2 at 500 °C for 1 h. After cooling to ambient temperature, H2 desorption was performed in 20 N · cm3 · min-1 Ar flow by heating each catalyst up to 600 °C at a rate of 10 °C · min-1 and then maintaining this temperature for 20 min. The concentration of the desorbed H2 was monitored with a thermal conductivity detector (TCD). A deeper study on metal dispersion and distribution was carried out only on 1 and 0.5% Rh on 3A zeolite with a chemisorption analysis (Micromeritics 2750 Pulse Chemisorb). H2 saturation was first performed by flowing 20 N · cm3 · min-1 of H2 for 2 h at 350 °C; then a He flow rate of 20 N · cm3 · min-1 for 1.5 h was fed to the reactor increasing the temperature to 370 °C. Then, at room temperature, a mixture of 10% CO in He was injected in pulses of 500 N · µL each, until the fulfillment of constant outlet peaks. The amount of adsorbed gas was determined as the difference between the total injected volume and the residual escaped one. Then, the metal dispersion on the carrier surface was determined through a suitable mathematical expression by considering the CO adsorbed molecules, the catalyst mass, and the noble metal load. 2.2. Reactor System and Analytical Methods. A fixed bed tubular microreactor (a Pyrex tube of 4 mm i.d.), heated up by a PID-regulated oven and containing 150 mg of catalyst particles held in place by flocks of quartz wool, was used for the selective CO oxidation reaction. A K-type thermocouple was inserted into the reactor to measure the temperature of the catalytic bed. A schematic representation of the microreactor bench test system is shown in Figure 2. The total inlet gas was fed at a flow rate of 100 N · mL · min-1 and with the following bv standard composition: 37% H2, 5% H2O, 18% CO2, 1% CO, 2% O2, and He as balance. This composition (He apart) is representative of a typical LT-WGS reactor outlet following a gasoline autothermal reformer.2 Mass flow controllers (Brooks) were used to regulate the feed flow rate of each component; a pressure transducer (VIKA) was conversely used to monitor the pressure at the bed inlet to check its possible clogging. The outlet gas stream was analyzed through a gas chromatograph (Varian CP-3800) equipped with a thermal conductivity detector (TCD), a “Poraplot Q” column (0.53 mm diameter, 30 m length) to separate CO2 and H2O, and a “Molsieve 5A” column (0.53 mm diameter, 25 m length) to separate CO, H2, and O2. The two columns, connected in series by a six-way valve, were kept at 70 °C; the sample injection was accomplished using He as carrier gas at a flow rate of about 2.8 N · mL · min-1. The CO detection limit was 10 ppmv. The experimental tests were carried out in the temperature range of 50-250 °C, at different O-to-CO feed ratio and with a weight space velocity WSV ) 0.66 N · L · min-1 · gcat.-1. The conversion of CO (ζCO) and O2 (ζO2), as well as the O2 selectivity to CO oxidation (σO2, assuming neither CO formation by reverse water gas shiftsRWGSsnor CO consumption by methanation) were calculated as follows: ζCO ) 1 -

[CO]out [CO]in

(2)

ζO2 ) 1 -

[O2]out [O2]in

(3)

σO2 )

1 [CO]in - [CO]out 2 [O2]in - [O2]out

(4)

Then, maintaining the same CO inlet concentration, the influence of O2 over CO ratio in the microreactor feed gas stream on the catalytic performance of 1% Rh-3A catalyst was analyzed by reducing the O2 inlet concentration from 2 to 0.5 vol % (the minimum stoichiometric O2 concentration) at 0.5% progressive downward steps. 3. Results and Discussion. A preliminary characterization of the as-prepared Rh-based catalysts was carried out with XRD analysis (Figures 3 and 4) in order to evaluate the crystalline structure and to check the presence of Rh clusters detectable with the XRD analysis. From the XRD spectra in Figure 3, Rh peaks appeared only over CeO2 and TiO2 carriers, but they were very small and not well-defined. The carrier CeO2 appeared a little bit more crystallized than the TiO2 one, with higher and thinner peaks. Al2O3 has shown a very amorphous spectrum as the carrier was γ-Al2O3, and, then, the Rh peaks, if present, could be hidden and not detectable. On the contrary, the zeolite structures (Figure 4) appeared very crystallized and Rh clusters were not visible, indicating a good metal dispersion on those carriers. A preliminary screening of the catalytic performance of Rhbased catalysts on different carriers, in terms of ζCO and ζO2, as well as σO2, was carried out. A comparison of ζCO, for 1% Rh catalysts supported on TiO2, Al2O3, CeO2, and 3A zeolite (as a representative of A zeolite materials) is at first shown in Figure 5. The activity tests were carried out at two different λ ratios (λ(2O2/CO) ) 2 and 4), in order to work with two different maximum selectivity, 50 and 25%, respectively, values reachable when both ζCO and ζO2 are simultaneously equal to 1. The main goal of the experiments was, in fact, to reach the complete conversion of CO, with at the same time a complete conversion of O2, in a sufficient wide temperature range (suitable for the control of the PROX reactor), feeding the lowest O2 excess rate to maintain as low as possible the H2 loss due to oxidation; anyway, the lowest O2 consumption is not a mandatory issue, if compared with the necessity of avoiding the FC electrocatalyst poisoning by CO. At λ ) 4, for the 1% Rh-Al2O3 catalyst, complete ζCO associated with ζO2 ) 1 was observed only at 80 °C; at this temperature also the maximum σO2, 25%, was obtained. On the contrary, O2 having completely reacted, a complete ζCO was not reached with the CeO2- and TiO2-supported catalysts: the maximum conversion with 1% Rh-CeO2 was observed at 70 °C with a residual CO concentration equal to 700 ppmv, and at 100 °C with an outlet CO concentration of about 687 ppmv, for 1% Rh-TiO2. The best performance was obtained with 1% Rh-3A: complete ζCO with a residual CO concentration below 10 ppmv was observed from 80 to 120 °C, a temperature range in agreement with the operating PEM FCs range (80-100 °C).6–8 When λ was decreased to 2, complete ζCO with ζO2 ) 1 was still reached with the catalyst supported on zeolite, but this happened at higher temperature (110 °C), whereas 1% Rh-Al2O3 catalyst did not more completely convert CO, notwithstanding all the fed O2 reacted. Moreover, the ceria-supported catalyst performance decreased and the residual CO concentration was 1350 ppmv at 80 °C. Finally, with 1% Rh-TiO2 catalyst ζCO practically remained at the same value, but the maximum conversion (92-96%) shifted toward too high temperature (from 120 to 200 °C). Then, the most promising catalyst, 1% Rh-3A, showed the best performance at λ ) 4, although with a penalty of the lower selectivity. In addition, no undesirable side reactions (RWGS;

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Figure 3. XRD spectra for 1% Rh catalysts on TiO2, CeO2, and Al2O3.

methanation) occurred to an appreciable extent for such a catalyst, as shown in Figure 7. Therefore, two other A-type zeolites, with similar structure but larger pore sizes, were tested; they exhibited some differences in catalytic activity, as shown in Figures 6 and 7. A complete CO oxidation was observed in the range of 80-120 °C for 1%Rh-3A catalyst, of 80-140 °C for 1%Rh-4A, and of 100-200 °C for 1%Rh-5A. In addition, no methanation reaction was detected with 1% Rh-3A catalyst for all the temperature test conditions, while a significant methanation activity was observed with the other two,

with the difference that 1% Rh-4A catalyst started at a higher temperature (CH4 formation at T > 160 °C, with 216 ppmv at 180 °C and 219 ppmv at 200 °C) than 1% Rh-5A catalyst (CH4 production at T > 120 °C, with 218 ppmv at 140 °C and 235 ppmv at 200 °C). The data drawn in Figure 7 show the methanation capability of the three zeolite-based catalysts that might also be due to the reaction between H2 and CO2: 4H2+CO2 f CH4+2H2O

(5)

The superior performance of 1% Rh-3A catalyst, with regard to the methanation inhibition, might be due to its structure. In

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Figure 5. CO conversion vs temperature for 1% Rh catalysts on different supports at λ ) 2 and 4, with standard feed gas composition (WSV ) 0.66 Nl · min-1 · gcat.-1). Figure 4. XRD spectra for 1% Rh catalysts on different A type zeolite carriers.

contrast with the 4A and 5A zeolite pore sizes (4 and 5 Å, respectively), the pore size of 3A zeolite (3 Å), close to the diameter of the CO2 molecule (2.5 Å), could make difficult the CO2 admittance into the pores themselves. Also the different basicity of zeolites 3A, 4A, and 5A, related in turn to the different nature of their counterions (K+, Na+, and Ca2+, respectively) may have an influence: this could explain the different interaction of supports with CO2 and a consequent different promotion of the methanation reaction.35,36 CO2 TPD results for 3A, 4A, and 5A zeolites (drawn in Figure 8), carried out to evaluate the interaction between CO2 and these materials, illustrate indeed a different interaction with CO2: zeolite 3A shows a small peak at about 80 °C, whereas zeolite 4A desorbs a much higher CO2 amount at about 100 °C. Zeolite 5A, instead, shows two separated CO2 desorption peaks at about 90 and 230 °C, respectively. The limited interaction of zeolite 3A with CO2, whatever the origin (i.e., chemical interaction or geometrical hindrance) could account for the higher selectivity and the total CO conversion (up to 10 ppmv residual) in a wide temperature range (80-120 °C) of 1% Rh-3A catalyst. The low CO2 concentration in the neighborhood of the catalytic active sites, due to its low affinity with the support, could indeed be responsible for (i) enhancement of the CO-PROX reaction and (ii) inhibition of the methanation reaction (eq 5).

Figure 6. CO conversion vs temperature for 1% Rh catalysts on different A type zeolite carriers with standard feed gas composition (λ ) 4, WSV ) 0.66 Nl · min-1 · gcat.-1).

Because the main purpose of the present study was to develop a suitable catalyst for CO-PROX operating in a temperature range compatible with that of PEM FCs, and, in addition, without undesirable side reactions, the 1% Rh-3A catalyst was considered as the best catalytic material and employed in the further tests.

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Figure 7. CH4 formation vs temperature for 1% Rh catalysts on different A type zeolite carriers with standard feed gas composition (λ ) 4, WSV ) 0.66 Nl · min-1 · gcat.-1).

Figure 8. Temperature-programmed CO2 desorption for 1% Rh catalysts on the various A zeolite carriers.

Figure 9. CO conversion vs temperature for 1% Rh-3A catalyst at constant WSV ) 0.66 Nl · min-1 · gcat.-1 and different λ values (standard gas feed composition with decreasing O2 concentration).

As above-mentioned, to improve the O2 selectivity to CO oxidation with the lowest possible H2 parasitic oxidation, the O2 concentration in the feed gas was decreased. The obtained results are shown in Figures 9 and 10. The maximum ζCO observed at stoichiometric O2 condition (0.5%, λ ) 1) was 78.8% at 140 °C (Figure 9). By doubling the O2 concentration to 1% (λ ) 2), a complete ζCO was observed in the temperature range of 120-140 °C. After a further O2 concentration increase to 1.5% (λ ) 3), the conversion/temperature curve remained practically equal to 1

Figure 10. (A) Oxygen selectivity vs temperature for 1% Rh-3A at different λ values; (B) temperature range of maximum selectivity vs λ for 1% Rh3A catalyst at WSV ) 0.66 Nl · min-1 · gcat.-1 (hatching ) highest selectivity zone; standard gas feed composition with decreasing O2 concentration).

from 80 to 120 °C, showing a wider, and at a lower level, temperature range as compared to that for λ ) 2. The 2% O2 concentration (λ ) 4) did not further increase the temperature range of complete ζCO, which remained equal to that obtained when using λ ) 3. These results are mostly in agreement with Oh and Sinkevitch’s findings on Pt, Ru, and Rh aluminasupported catalysts:24 the increase of O2 concentration above a minimum limit value (1.5%, in our tests) did not enlarge the temperature range where a complete CO removal conversion occurs. The lines in Figure 10A, drawn for test conditions of complete conversion of both CO and O2, show, as expected, that as the O2 concentration increases above that necessary for CO stoichiometric oxidation, the selectivity of the catalyst decreased, as a result of more O2 molecules free to react with the H2 ones, given that the catalyst surface was no longer mostly covered by adsorbed CO. The maximum selectivity was σO2 ) 0.50 for λ ) 2, reduced to σO2 ) 0.333 at λ ) 3 and then to σO2 ) 0.25 at λ ) 4. As a consequence, when increasing the O2 concentration from 1.5 to 2% (from λ ) 3 to λ ) 4) only a further decrease of σO2 without beneficial enlargement of the temperature range of complete ζCO was obtained. Figure 10B shows the trend of the maximum selectivity temperature range, obtained at complete ζCO and ζO2, by varying the λ value: for λ ) 3 and λ ) 4 the range width was the same (40 °C, from 80 to 120 °C), but for λ ) 2 the temperature window was tighter (about

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Figure 12. HRTEM micrographs of Rh-catalysts: (A) 1%Rh-3A; (B) 0.5%Rh-3A.

Figure 11. CO conversion vs temperature for 0.5 and 1% Rh on 3A catalysts with standard feed gas composition (λ ) 3, WSV ) 0.66 Nl · min-1 · gcat.-1).

20 °C) and moved up to higher temperature levels (maximum and minimum temperature equal to 120 and 140 °C, respectively). Taking into account the possibility to achieve complete ζCO at a temperature range compatible with PEM FCs and as large as possible for a satisfactory PROX reactor controllability with, at the same time, acceptable H2 consumption, the feed O2 concentration of 1.5 vol % (λ ) 3) was considered the best one. As visible in Figures 9 and 10A, at temperature higher than 160-180 °C, both ζCO and σO2 of 1% Rh-3A catalyst decreased, excluding the methanation reaction between H2 and CO: CO + 3H2 f CH4+H2O

(6)

(inhibited for 1% Rh-3A catalyst, as shown in Figure 7); that consuming, to some extent, CO could help in maintaining high CO conversion values, the above decreases seem to be related to the R-WGS reaction: CO2+H2 f CO + H2O

(7) 13

limiting the CO conversion at high temperatures. Another important point for the catalyst developed concerns the considerable cost of Rh precursor (the Rh market price is considerably higher than that of Pt: 6,215.77 vs 1,292.93 U.S. dollars/oz (USD/oz) in 200737); then, a decrease of the noble metal load required for complete CO oxidation in a satisfactory temperature range can potentially reduce the investment cost of the CO-PROX stage19. Therefore, the effect of the Rh load on catalyst activity and selectivity was investigated at λ ) 3; catalysts with 1 and 0.5% Rh on zeolite 3A are compared in Figure 11 in terms of ζCO. Again, the 0.5% Rh-supported catalyst was prepared with the IWI impregnation method. At 80 °C, while the 1% Rh-3A catalyst completely converted the fed CO, the 0.5% Rh-3A catalyst exhibited a very poor ζCO (23.8%, point not reported in Figure 11, drawn with the aim to enlarge the phenomena in the ζCO range larger than 95%); the latter started to completely convert CO at 100 °C and continued this performance until 140 °C. For the 1% Rh-3A and the 0.5% Rh-3A catalysts, the temperature range of complete ζCO was, therefore, 80-120 and 100-140 °C, respectively. Then, both catalysts showed complete ζCO, without residual O2, for the same temperature range of 40 °C, but the 0.5% Rh-3A catalyst shifted this range to a little bit higher level compared with the 1% Rh-3A one.

Figure 13. TPD results with H2 for 0.5 and 1% Rh-3A catalysts.

Moreover, between 140 and 200 °C the 0.5% Rh-3A catalyst seemed to exhibit a slightly higher ζCO than the 1% Rh-3A one. Parts A and B of Figure 12 show the HRTEM micrographs of 1 and 0.5% Rh-3A catalysts, respectively. The Rh dispersion on zeolite appears quite good for both the catalysts; however, the particle dispersion seems better, even if the cluster sizes are probably a little bit larger, for the 0.5% Rh-3A one. Both the different Rh dispersion (see Figure 12) and Rh load decrease on the catalysts could have caused these results. Moreover, the catalyst with the lower Rh load inhibited much more parasite reactions. This could explain the larger ζCO decrease at high temperature (>140 °C) for the 1% Rh-3A catalyst and the shift of the complete ζCO range to higher temperatures for the 0.5% Rh-3A one. Furthermore, to evaluate the Rh dispersion, TPD analyses with H2 were carried out on 0.5% Rh-3A and 1% Rh-3A catalysts. The obtained results are shown in Figure 13. The 0.5% Rh-3A catalyst presented a “bulging” TPD curve that should be due to a large number of different nature active sites and a good dispersion of Rh over the support, whereas the 1% Rh3A curve showed a peak indicating the presence of H2-active adsorption sites of a similar nature. These results are in agreement with the morphology represented in HRTEM micrographs showing a better Rh dispersion for the 0.5% Rh-3A catalyst as opposed to its more heavily loaded counterpart. Moreover, on the grounds of calibration data, the amount of exposed reactive Rh on the zeolite support was calculated by built-in TPD interpretation software and resulted to be higher for 0.5% Rh-3A catalyst (approximately 74%) than for 1% Rh3A one (approximately 50%). The previous results were also validated by CO chemisorption analyses carried out on both catalysts with different Rh load. The 0.5% Rh-3A catalyst showed a percentage distribution of noble metal equal to 55%, against the 46% obtained for the 1% Rh-3A one. These results confirmed those obtained with

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Figure 14. CO conversion vs temperature for 0.5% Rh-3A at two different λ values and WSV ) 0.66 Nl · min-1 · gcat.-1 (standard gas feed composition with decreasing O2 concentration).

HRTEM and H2-TPD analyses, enlightening again the better metal dispersion for 0.5% Rh-3A catalyst. As a result of the above investigations the following conclusions may be drawn: (1) Both catalysts are active for selective CO oxidation; however, the 0.5% Rh-3A catalyst exploits its activity in a temperature range slightly higher than that of its 1% Rh-3A counterpart. (2) As a consequence, despite the lower fraction of exposed Rh, 1% Rh-3A catalyst presents a globally higher metal active sites, so it ignites CO oxidation at temperatures lower than 0.5% Rh-3A catalyst (see Figure 11). (3) It is likely, however, that 1% Rh-3A catalyst is also more active toward parasite or competing reactions, which limits its application at high temperatures. (4) Hence, depending on the operating temperatures, either 1% Rh-3A or 0.5% Rh-3A catalyst give a good performance (Figure 11). An activity test on 0.5% Rh-3A was then carried out again with an O2 concentration of 1 vol % (λ ) 2); the obtained results are compared in Figure 14 with those at λ ) 3. When decreasing the O2 concentration to 1 vol % (λ ) 2), the complete CO conversion with 0.5% Rh-3A catalyst was reached only at 120 °C; however, at higher temperatures the conversion was slightly decreased to about 97-98%. Therefore, the performance of 0.5% Rh-3A catalyst was similar to that of the 1% Rh-3A one, in particular at λ ) 3; therefore, the former could still be considered for application in the CO-PROX reactor operating at temperatures compatible with the PEM FCs, in this way reducing the noble metal load and consequently the catalyst and reactor cost. The obtained results can be considered very promising, taking into account that the catalyst performance was maintained unchanged for more than 50 h, under the described operating conditions. Anyway, to complete the picture of the catalyst performance, its stability and operating life need to be verified for a longer time. 4. Conclusions CO-PROX tests carried out under realistic conditions with Rh catalysts supported on A zeolites (3A, 4A, and 5A), Al2O3, TiO2, and CeO2, indicated as the most suitable catalyst for application in CO-PROX operating at a temperature range compatible with the PEM FC (80-100 °C) 1% Rh-zeolite 3A.

This was due to its ability to reduce the CO concentration under the test conditions (WSV ) 0.66 N · L · min-1 · gcat.-1) to below 10 ppmv within a temperature range of at least 80-140 °C for λ ) 4, during which no undesirable side reactions were detected. To optimize the O2 selectivity of this catalyst, the λ ratio was reduced. For 1% Rh-3A zeolite catalyst, a λ value of 3 was found to give, at complete CO and O2 conversions, a maximum selectivity σO2 ) 0.333 with a temperature range for complete CO conversion of 80-120 °C. With the aim to reduce the catalyst cost, a test decreasing the Rh load from 1 to 0.5% was carried out. No significant change in the catalytic activity was noticed: the load reduction did not affected the width of complete CO conversion temperature range. There was only a slight shift to higher temperature values than for 1% Rh-3A catalyst, probably due to the lower Rh load that, on the other side, inhibited the parasite reactions shifting them to higher temperatures. Therefore, the 0.5% Rh-3A catalyst operating at λ ) 3 could potentially be used for CO-PROX applications at temperatures compatible with the PEM FCs: it was able to reduce the inlet CO concentration below 10 ppmv within a temperature range of 100-140 °C, with no appearance of undesired side reactions and, in addition, at a possibly more acceptable cost for the catalyzed reactor. Literature Cited (1) Hefner, R. A., III. Toward Sustainable Economic Growth: The Age of Energy Gases. Int. J. Hydrogen Energy 1995, 20, 945. (2) EU project PROFUEL: On-Board Gasoline Processor for Fuel Cell Vehicle Application; ERK6-CT-1999-00023; www.europa.cordis.eu. (3) EU project MINIREF: Miniaturized Fuel Processor for Fuel Cell Vehicle Application; ENK6-CT-2001-00515; www.europa.cordis.eu. (4) EU project BIOFEAT: Biodiesel Fuel Processor for Fuel Cell Auxiliary Power Unit for a Vehicle; ENK5-CT-2002-00612; www.europa.cordis.eu. (5) EU project HYTRAN: Hydrogen and Fuel Cell Technologies for Road Transport; TIP-CT-2003-502577; www.europa.cordis.eu. (6) Heinzel, A.; Vogel, B.; Hubner, P. Reforming of Natural GassHydrogen Generation for Small Scale Stationary Fuel Cell Systems. J. Power Sources 2002, 105, 202. (7) Beckhaus, P.; Heinzel, A.; Mathiak, J.; Roes, J. Dynamics of Hydrogen Production by Steam Reforming. J. Power Sources 2004, 127, 294. (8) Faur Ghenciu, A. Review of Fuel Processing Catalysts for Hydrogen Production in PEM Fuel Cell Systems. Curr. Opinion Solid State Mater. Sci. 2002, 6, 389. (9) Specchia, S.; Tillemans, F. W. A.; van den Oosterkamp, P. F.; Saracco, G. Conceptual Design and Selection of a Biodiesel Fuel Processor for a Vehicle Fuel Cell Auxiliary Power Unit. J. Power Sources 2005, 145, 683. (10) Petrachi, G. A.; Negro, G.; Specchia, S.; Saracco, G.; Maffettone, P. L.; Specchia, V. Combining Catalytic Combustion and Steam Reforming in a Novel Multifunctional Reactor for On-Board Hydrogen Production from Middle Distillates. Ind. Chem. Eng. Res. 2005, 44, 9422. (11) Cutillo, A.; Specchia, S.; Antonini, M.; Saracco, G.; Specchia, V. Diesel Fuel Processor for PEM Fuel Cells: Two Possible Alternatives (ATR versus SR). J. Power Sources 2006, 154, 379. (12) Specchia, S.; Cutillo, A.; Saracco, G.; Specchia, V. Concept Study on ATR and SR Fuel Processors for Liquid Hydrocarbons. Ind. Eng. Chem. Res. 2006, 45, 5298. (13) Korotkikh, O.; Farrauto, R. Selective Catalytic Oxidation of CO in H2: Fuel Cell Applications. Catal. Today 2000, 62, 249. (14) Oetjen, H. F.; Schmidt, V. M.; Stimming, U.; Trila, F. Performance Data of a Proton Exchange Membrane Fuel Cell Using H2/CO as Fuel Gas. J. Electrochem. Soc. 1996, 143, 3838. (15) Dudfield, C. D.; Chen, R.; Adcock, P. L. A Compact CO Selective Oxidation Reactor for Solid Polymer Fuel Cell Powered Vehicle Application. J. Power Sources 2000, 86, 214. (16) Manasilp, A.; Gulari, E. Selective CO Oxidation over Pt/Alumina Catalysts for Fuel Cell Applications. Appl. Catal., B 2002, 37, 17.

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ReceiVed for reView October 9, 2007 ReVised manuscript receiVed May 16, 2008 Accepted May 19, 2008 IE0713588