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Development of A zeolites-supported noble-metal catalysts for CO preferential oxidation: H2 gas purification for fuel cell
Applied Catalysis B: Environmental 48 (2004) 195–203
Development of A zeolites-supported noble-metal catalysts for CO preferential oxidation: H2 gas purification for fuel cell Ilaria Rosso∗ , Camilla Galletti, Guido Saracco, Edoardo Garrone, Vito Specchia Dipartimento di Scienza dei Materiali e Ingegneria Chimica, Politecnico di Torino, Corso Duca degli Abruzzi, 24, 10129 Turin, Italy Received 27 April 2003; received in revised form 7 August 2003; accepted 20 October 2003
Abstract Even traces of CO in the hydrogen-rich feed gas to proton exchange membrane fuel cells (PEMFC) poison the platinum anode electrode and dramatically decrease the power output. In this work, a variety of catalytic materials consisting of noble metals supported on A zeolites were synthesised, characterised and tested under realistic conditions in the quest of a catalyst for the removal of CO via the CO preferential oxidation (CO-PROX) reaction. Pt, Pd and Ru-based catalysts, prepared by wet impregnation and characterised by XRD and HRTEM, were investigated in a fixed bed reactor, by determining CO conversion and selectivity through the outlet concentrations of CO and O2 . In contrast to supported Pd and Ru catalysts, Pt-catalysts showed complete CO-conversion and a comparatively high selectivity. The 1% Pt-3A catalyst showed the best performance: it kept the complete CO-conversion in a wide temperature range, showing the highest selectivity for CO oxidation with minimal involvement in side reactions, such as H2 oxidation and RWGS reaction. Experimental data proved that the RWGS outcome is directly related to the support structure. The rather high temperature (≈260 ◦ C) at which complete CO conversion is achieved by the 1% Pt-3A catalyst enables to locate the CO-PROX unit immediately after the low temperature water-gas shift unit of the fuel processor converting hydrocarbons into hydrogen-rich gas. © 2003 Elsevier B.V. All rights reserved. Keywords: Pt/Pd/Ru-A zeolite; CO preferential oxidation; Return water gas shift
1. Introduction Because of the perspectives of a significantly higher efficiency and of almost no emission of pollutants, proton exchange membrane fuel cells (PEMFC) have been extensively studied in the last two decades for many applications and especially for low emissions vehicles [1,2]. Pure hydrogen is the ideal fuel for the PEMFC. Currently, however, there is no available technology for the safe and economical storage of enough hydrogen to give to a PEMFC powered vehicle an acceptable driving range. On-demand generation of hydrogen by reforming methanol or other liquid hydrocarbon fuels in a fuel processor is a viable means of providing the needed hydrogen and the driving range. To this purpose, a complex system has to be developed including feeding, production and purification of the gas rate flowing to the fuel cells. The development of a complete on-board fuel processor for the production of clean hydrogen from commercial gaso∗ Corresponding author. Tel.: +39-011-5644710; fax: +39-011-5644699. E-mail address: [email protected] (I. Rosso).
0926-3373/$ – see front matter © 2003 Elsevier B.V. All rights reserved. doi:10.1016/j.apcatb.2003.10.016
line is the objective of an industrial-type EU project, PROFUEL, which involves, besides our group, several partners from the automotive and catalyst manufacturing industries [3]. The fuel processor is schematically described in Fig. 1. It consists of: a dosing system, an autothermal reformer, a high-temperature water-gas-shift (HTWGS) reactor, a desulphuriser, a low-temperature water-gas-shift (LTWGS) reactor and a CO preferential oxidation unit (CO-PROX). All primary fuel processing reactions, which take place in the reformer, produce substantial amount of carbon monoxide, which is, unfortunately, incompatible with direct feeding of PEMFC stacks, as CO strongly chemisorbs on the platinum-based electrodes and deactivates them. In order to reduce drastically the CO concentration and increase the hydrogen concentration, the water-gas-shift reactors (CO + H2 O → CO2 + H2 ) are placed immediately after the reformer. This reaction reduces the CO content of the hydrogen-rich gas stream to about 0.5–1% (5000–10,000 ppm) [4], still far above the 10 ppm tolerance limit of a typical anode catalyst, so that a third process unit is required. Palladium-based membrane purification, catalytic methanation and catalytic CO preferential oxida-
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Fig. 1. Schematic description of a gasoline processor for hydrogen production. The average operating temperatures of each process stage are indicated.
tion are the main options in this context. The CO-PROX method is the most promising and economic approach: CO is oxidised with added air (CO + (1/2)O2 → CO2 ) with as low as possible concomitant loss of hydrogen by direct oxidation. Since LTWGS operates at temperature between 200 and 270 ◦ C and the fuel cell at 80 and 100 ◦ C, the available temperature operating range for the CO-PROX reactor is 80–270 ◦ C. The majority of catalysts for the CO-PROX have been studied for an intermediate temperature applications (150 ◦ C), in which case two coolers, upstream and downstream of the CO-PROX unit, are necessary (Fig. 1). The literature reports several studies on catalysts for the selective oxidation of CO, including noble metals (Pd, Ru, Rh, Pt) [5–8] and base metals (Co–Cu, Ni–Co–Fe, etc.) [8–10] generally supported on alumina. Pt-group-metals catalysts exhibit good activity and selectivity in a temperature range of 130–200 ◦ C. Activity in the order: Ru > Rh > Pt > Pd was observed [8]. The performances of the base-metal-based supported catalysts are less promising [8]. Although alumina is the most common support for noble metal-based catalysts, recent studies [11,12] have indicated that the use of zeolites as supports for platinum catalysts is promising to improve the catalyst selectivity, probably owing to the molecular sieve effects. In particular, Pt-A catalysts resulted the most selective among those tested (Pt-mordenite, Pt-X, Pt-A and Pt-Al2 O3 ) [11]. Since A zeolites are attractive also because of their low cost and high thermal stability [13], this paper describes the development of catalysts for preferential CO oxidation consisting of A zeolite-supported Pd, Ru, Pt. Particular attention was paid to the catalytic activity and selectivity in H2 , CO2 , H2 O rich stream as a function of temperature, in order to find a catalyst able to completely remove CO at a temperature similar to that of the outlet gas from LTWGS, as this would enable elimination of the intermediate cooler.
2. Experimental 2.1. Catalysts preparation and characterisation A-type zeolites with different pores size were purchased from Fluka: 3A-type zeolite (K12 [(AlO2 )12 (SiO4 )12 ] ×
H2 O), 4A-type zeolite (Na12 [(AlO2 )12 (SiO4 )12 ] × H2 O), 5A-type zeolite (Can Na12−2n [(AlO2 )12 (SiO4 )12 ] × H2 O), with pores of about 3, 4 and 5 Å diameter, respectively. All catalysts were prepared by impregnation [12,14,15]: a proper amount of precursor, to obtain 1 wt.% noble metal catalyst, was dissolved in 25 ml of distilled water; then 5 g of A-type zeolite were added and the mixture was stirred for 30 min at about 50 ◦ C. The mixture was dried in an oven at 90–100 ◦ C and, after grinding in an agate mortar, it was calcined in an electric oven in calm air at 500 ◦ C for 1 h. After calcination, the catalyst was reduced in H2 flow rate (50 N ml min−1 ) at 500 ◦ C for 2 h. Precursors for Pt, Pd and Ru-catalysts were: H2 PtCl6 (0.1327 g), Pd(NH3 )4 Cl2 (0.115 g) and Ru(NH3 )6 Cl3 (0.153 g) (all from Alfa), respectively. XRD analysis (Philips PW1710 apparatus equipped with a monochromator for the Cu K␣ radiation) was performed both on pure zeolites and on the prepared catalysts before and after the catalytic activity tests in order to verify the integrity of zeolite structure after exposition to water, present in the feed stream (see Section 2.2) and as a product of oxidation reactions. Impregnated catalysts were analysed by high-resolution transmission electron microscopy (HRTEM, Jeol JEM 2010 apparatus) to investigate the metal dispersion on supports. Temperature programmed CO2 desorption (TPD) tests were performed on zeolite 3A, 4A and 5A by a TPD/R/O apparatus (Thermoquest TPD/R/O 1100 analyser equipped with a Baltzer Quadstar 422 quadrupole mass spectrometer). CO2 saturation was performed by flowing 50 N cm3 min−1 CO2 for 1 h at 40 ◦ C, then a helium flow rate of 10 N cm3 min−1 was fed to the reactor while increasing the temperature at 5 ◦ C min−1 rate up to 450 ◦ C. 2.2. Reactor system and analytical methods Catalyst pellets were obtained by pressing at 125 MPa the powders into tablets, which were then crushed and sieved to produce 0.25–0.42 mm granules. A fixed-bed of about 2 cm in length, containing 0.15 g of catalyst pellets was enclosed in a glass tube (i.d.: 4 mm) and sandwiched between two glass–wool layers. The reactor was placed in a PID regulated oven, and a K-type thermocouple was inserted in the packed bed for oven regulation purposes.
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The feed stream flow rate (100 N cm3 min−1 ) contained 37 vol.% H2 , 18 vol.% CO2 , 0.5 vol.% CO, 5 vol.% H2 O, 1 vol.% O2 , 39.5 vol.% He. The desired composition and flow rate of each component (except water) were obtained by mass flow controllers, whereas a pressure transducer was used to check the pressure of inlet and outlet gases in order to check any possible undesired clogging of the catalyst fixed-bed. Water vapour was added by bubbling a He stream in a drechsel kept at a suitable operating temperature. The gas feedstock was delivered through a heated tube (above 70 ◦ C) to avoid H2 O condensation. The outlet gas stream was analysed through a Gas Chromatograph (Varian CP-3800) equipped with a thermal conductivity detector (TCD); the GC contained a “Poraplot Q” column (0.53 mm diameter by 30 m length), to separate CO2 and H2 O, and a “Molsieve 5A” column (0.53 mm diameter by 25 m length), to separate CO, H2 and O2 . The two columns were connected in series by a six-way valve. The columns were kept at 70 ◦ C and the sample injection was accomplished using helium as the carrier gas at a flow rate of about 2.8 ml/min. Detection limit of CO is 10 ppmV. The conversion of CO (ξ CO ) and O2 (ξO2 ), as well as the selectivity of CO oxidation (σ), determined in the 100–300 ◦ C range in the presence of excess hydrogen, were calculated as follows: [CO]out (1) ξCO = 1 − [CO]in [O2 ]out [O2 ]in
(2)
1 [CO]in − [CO]out . 2 [O2 ]in − [O2 ]out
(3)
ξO2 = 1 − and σ=
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3. Results and discussion A preliminary screening of the catalytic performance of all prepared catalysts, in terms of CO, O2 conversion and CO selectivity, was carried out. XRD analysis, performed on all zeolite-supported catalysts after the catalytic activity tests, demonstrates that the A zeolite structure is substantially preserved notwithstanding the water presence in the feed stream (5% H2 O). Although some authors proved that a negligible degradation of activity occurs on Pt-mordenite when water (up to 20%) was added to the feed stream [11], the resistance of A zeolites to high water amounts will be studied in further studies, as this is a very critical point for the practical application of A zeolite. The CO outlet concentration values for the three 1% Pd-zeolites are reported in Fig. 2. CO outlet concentrations decreased starting from about 150 ◦ C, reached the lowest values at about 180–200 ◦ C and increased again at temperature higher than 200 ◦ C. No catalyst reached total CO conversion: the lowest CO outlet concentration (about 930 ppmV) was obtained by 1% Pd-5A at 205 ◦ C. In contrast, the O2 conversion was complete in a temperature range of 200–300 ◦ C for all the catalysts. This means that the oxygen was consumed in the CO oxidation reaction: CO + 21 O2 → CO2
(4)
and also in the hydrogen oxidation reaction: H2 + 21 O2 → H2 O
(5)
which brings about a decrease in oxygen concentration and in power generation. The resulting selectivity σ reached values never higher than 20% for all the 1% Pd-catalysts. The inlet gas mixture
Fig. 2. CO outlet concentration vs. temperature for 1% Pd-3A, 1% Pd-4A and 1% Pd-5A catalysts with standard feed gas composition (5000 ppm CO, 1% O2 , 18% CO2 , 5% H2 O, 37% H2 and He as balance). Hourly space velocity: 67000 h−1 .
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Fig. 3. CO outlet concentration vs. temperature for 1% Ru-3A, 1% Ru-4A and 1% Ru-5A catalysts with standard feed gas composition. Hourly space velocity: 67,000 h−1 .
used in the catalytic tests contained 5000 ppmV CO and 1 vol.% oxygen which is four-fold the stoichiometric value needed for the reaction (1). As a consequence, when both CO and O2 conversions are equal to 1, it is not possible to obtain a selectivity value higher than 25%. As far as the 1% Ru-supported catalysts are concerned, Fig. 3 shows that 1% Ru-3A and 1% Ru-4A catalysts gave rather high CO outlet concentration values in the whole temperature range: the lowest CO outlet concentrations were 1650 ppmV at 165 ◦ C, and 1170 ppmV at 148 ◦ C for 1% Ru-3A and 1% Ru-4A catalysts, respectively. The 1% Ru-5A catalyst showed a particular trend: it reached 700 ppmV of CO outlet concentration at about 125 ◦ C, then it increased up to 2850 ppmV at 184 ◦ C and decreased again to 10 ppmV (the detection limit of CO) at 250 ◦ C. The complete CO conversion observed at high temperature is quite strange but can be explained considering the CH4 formation according to the reaction: CO + 3 H2 → CH4 + H2 O
(6)
that brings about a consumption of CO. Methane was in fact detected in the exhaust gases above 180 ◦ C and its amount was equal to about 0.5% in the temperature range between 240 and 280 ◦ C. The complete CO conversion in the same temperature range could confirm the occurrence of CO methanation (reaction (6)) rather than CO2 methanation. Accordingly, some authors [6] demonstrated that CO methanation is favoured in CO/CO2 /H2 mixture rather than CO2 methanation. Some methanation level was also detected in this temperature range for the 1% Ru-3A and the 1% Ru-4A catalysts. These results are in agreement with literature data [6,7], reporting CH4 formation on Ru-based catalysts above 180 ◦ C.
The O2 conversion was complete at temperatures higher than 165, 148, 125 ◦ C for 1% Ru-3A, 1% Ru-4A, 1% Ru-5A catalysts, respectively, and the resulting selectivity values were, once again, not higher than 20% for all the catalysts. Similar experiments carried out on 1% Pt-zeolite catalysts produced quite different results. Fig. 4 shows that the 1% Pt-3A and the 1% Pt-5A catalysts yielded total CO conversion (CO outlet concentration ≤10 ppmV) in the range of 165–220 ◦ C, whereas the lowest value of CO outlet concentration achieved by 1% Pt-4A catalyst was 100 ppmV at 165 ◦ C. The 1% Pt-3A catalyst kept the complete CO conversion in a wide range of temperature: from 185 to 220 ◦ C; in contrast, at this high temperature the CO outlet concentrations for 1% Pt-4A and 1% Pt-5A were 1300 and 675 ppmV, respectively. All catalysts gave a selectivity of 25%, the maximum selectivity reachable when ξ CO and ξO2 are equal to 1, at the temperature of the highest CO-conversion. In good accordance with the CO-conversion data, the selectivity of 1% Pt-3A remained as high as about 25% also at very high temperatures (up to 285 ◦ C), whereas the other two catalysts showed selectivity value lower than 20% at the same temperature. The decrease of CO-conversion (due to the high CO outlet concentration data) observed at high temperature for 1% Pt-4A and 1% Pt-5A catalysts seems to be mainly a consequence of the loss of selectivity: at high temperature the amount of O2 available for CO oxidation is determined by the amount of O2 consumed for the simultaneously occurring H2 oxidation (reaction 5) [13]. However, the observed decrease in CO conversion with temperature might also be related to the reverse water gas shift reaction (RWGS): CO2 + H2 → CO + H2 O (7) limiting the CO conversion at high temperatures [16].
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Fig. 4. CO outlet concentration vs. temperature for 1% Pt-3A, 1% Pt-4A and 1% Pt-5A catalysts with standard feed gas composition. Hourly space velocity: 67,000 h−1 .
To check the occurrence of reaction (7), the inlet gas mixture was fed without CO to the Pt-catalysts and the CO appearance was measured. Fig. 5 shows the CO outlet concentrations obtained with the three Pt-catalysts using the standard inlet gas mixture (full symbols) and the CO-free inlet gas mixture (empty symbols); the resulting difference ( CCO ) between these two outlet CO concentrations is also reported. With CO-free feedstock, the RWGS reaction is very limited for the 1% Pt-3A and the 1% Pt-5A catalysts and evident for the 1% Pt-4A one. By comparing these CO outlet concentrations with those obtained from the activity tests (standard feed mixture), in the range of high temperature where O2 conversion is practically equal to 1, the very low loss of selectivity observed for the 1% Pt-3A catalyst seems to be mainly due to the RWGS reaction, whereas the substantial loss of selectivity obtained for 1% Pt-4A and 1% Pt-5A catalysts seems to be caused by both reaction (7) (RWGS) and reaction (5) (H2 oxidation). However, for 1% Pt-5A catalyst, the contribution of the RWGS reaction to the outlet CO concentration can be considered negligible, as compared to that of H2 oxidation. The CCO values, reported for each catalyst and compared in the Fig. 5d, represent the CO amount that cannot be oxidised because the oxygen has been completely consumed. In summary, CCO remains very low for the 1% Pt-3A catalyst, it is progressively increasing with temperature for the 1% Pt-4A catalyst and presents high values in all the considered temperature range for the 1% Pt-5A catalyst. Pt dispersion of Pt-based catalysts was evaluated by HRTEM analysis. TEM micrographs of 1% Pt-3A, 1% Pt-4A and 1% Pt-5A catalysts are shown in Fig. 6A–C, respectively. All the catalysts show a rather poor homogeneous Pt dispersion with large Pt clusters on preferential zones of
the zeolite support. The average Pt crystal size varies from a few nanometers (2–3 nm) to some tens nanometers (20–25 up to 50 nm) for all catalysts. These data are in good agreement with X-ray diffraction investigation, as the peaks of metallic platinum (JCPDS card 04-0802) were well visible in the diffraction pattern of the Pt-catalysts, suggesting the presence of large Pt crystallites. The platinum crystal size of each Pt-catalyst was also estimated by the measurement of the full width half maximum (FWHM) of the platinum lines in the diffraction pattern according to Debye–Scherrer equation [17]. The resulting average platinum particle size was in the range of 21–26 nm; so almost all the employed Pt would be outside the zeolite pores. This seems to be a positive feature for the Pt-A zeolite catalysts prepared by wet impregnation, at least if compared with those prepared by cation-exchange of the A zeolites in an aqueous solution of Pt(NH3 )4 Cl2 ·H2 O for which a considerable fraction of the employed platinum would be inside the zeolite pores: those latter, in fact, gave worse catalytic performance. The superior performance of the 1% Pt-3A catalyst, as regards the RWGS inhibition might be due to the structure of the 3A zeolite. Its pore size (3 Å) is close to the diameter of CO2 molecule (2,5 Å), which could render difficult the CO2 admittance into 1% Pt-3A pores, in contrast with the 4A and 5A-zeolites pores (4 and 5 Å, respectively). Also the different basicity of the zeolites 3A, 4A and 5A, related in turn to the different nature of their counter ions (K+ , Na+ and Ca++ , respectively) may have an influence: it this could explain a different interaction of supports with CO2 and a consequent different outcome of the RWGS reaction. Temperature programmed CO2 desorption tests, performed on the 3A, 4A and 5A zeolites (Fig. 7) show indeed a different interaction with CO2 . The zeolite 3A shows a small
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Fig. 5. CO outlet concentration in catalytic activity tests carried out on (a) 1% Pt-3A, (b) 1% Pt-4A and (c) 1% Pt-5A catalysts; (d) CCO . Full symbols: standard feedstock. Empty symbols: CO-free feedstock. (×, +, ) symbols: CCO for 1% Pt-3A, 1% Pt-4A, 1% Pt-5A catalysts, respectively. Hourly space velocity: 67,000 h−1 .
peak at about 80 ◦ C, whereas the zeolite 4A desorbs a much higher CO2 amount at about 100 ◦ C. Instead, the zeolite 5A shows two small CO2 desorption peaks at about 90 and 230 ◦ C. These results are in good agreement with the heats of adsorptions for CO2 in the A zeolites determined from
IR isotherms by other authors [18]: CO2 presented increasing values of heat of adsorption versus K+ , of zeolite 3A, Na+ of zeolite 4A and Ca2+ of zeolite 5A. Accordingly, the CO2 desorption from zeolite 3A, 4A and 5A occurs at rising temperature (Fig. 7). The two CO2 desoption peaks
Fig. 6. HRTEM micrographs of Pt-catalysts: (A) 1% Pt-3A (600,000×); (B) 1% Pt-4A (300,000×); (C) 1% Pt-5A (80,000×).
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Fig. 7. Temperature programmed CO2 desorption for zeolite 3A, zeolite 4A and zeolite 5A.
observed for zeolite 5A could probable be due to the presence of both Na+ (peak at about 90 ◦ C) and Ca2+ (peak at about 230 ◦ C) cations, the latter can be considered stronger adsorption sites because of their bivalent charge. Similar experiments performed on the Pt-A zeolite catalysts showed the same results, in that the metallic particles seem not to be involved in the CO2 interaction. This means that also the possible presence of traces of chlorine from chlorine-containing Pt-precursor does not influence the CO2 interaction on zeolites. The limited interaction of the zeolite 3A with CO2 , whatever the origin, either geometrical or chemical, could account for the higher selectivity and the total CO-conversion (up to 10 ppmV) in a wide temperature range (185–220 ◦ C) of the 1% Pt-3A catalyst. The low CO2 concentration in the neighbourhood of the catalytic active sites, due to its low affinity with the support, could be responsible for: (i) the CO preferential oxidation (reaction (4)) enhancement; (ii) the RWGS (reaction (7)) inhibition. By the same token, the strong interaction between CO2 and zeolite 4A could explain the high contribution of the RWGS (reaction (7)) to the substantial loss of selectivity of the 1% Pt-4A catalyst. Zeolite 5A shows, as a whole, a limited interaction with CO2 , that could account for the negligible occurrence of the RWGS with the 1% Pt-5A catalyst, as for zeolite 3A. As these results regard the interaction between zeolites A and CO2 , the RWGS (reaction (7)) should have similar occurrence on Pd/Ru/Pt-A zeolites catalysts, independently of the employed noble metal. In contrast, the CO outlet concentrations observed at high temperature (>180 ◦ C) show different trends if Pd/Ru/Pt-3A, Pd/Ru/Pt-4A, and Pd/Ru/Pt-5A catalytic performances are compared (Figs. 2, 3 and 4). This fact, however, is not conclusive, because not the sole RWGS
but also the H2 oxidation (reaction (5)) and—limitedly to Ru-based catalysts—the methanation (reaction (6)) bring about the final measured CO outlet concentration. We experimentally verified that with CO-free feed mixture the RWGS reaction is very limited for not only for the 1% Pt-3A catalyst but also for the 1% Pd-3A and 1% Ru-3A ones: this means that the RWGS outcome seems to be directly related to the support structure. In contrast, the H2 oxidation can depend on the nature and the structure of both the active phase (i.e. the noble metal) and the support: its outcome, in fact, is very different for each tested catalyst. Further investigations have been performed on the promising 1% Pt-3A catalyst evaluating the effect of oxygen amount and hourly space velocity variation on its activity and selectivity. Different experiments were performed by varying O2 concentration in the feed flow rate: 1, 0.75, 0.5 vol.% O2 were used, with corresponding λ values of 4, 3, 2 (λ = 2[O2 ]/[CO]), respectively. Fig. 8a shows the CO-conversion values versus temperature obtained in the different experimental conditions: CO conversion practically equal to 1 was achieved at 185 ◦ C with λ = 3 and 4, whereas the maximum CO conversion at the same temperature with λ = 2 was equal to 0.6. This means that the catalyst is not so selective to completely oxidise CO even when an oxygen amount twice higher than the stoichiometric value is employed. As a consequence, the highest selectivity (Fig. 8b) was reached with λ = 3: it was equal to about 33%, which is the maximum value reachable at the tested operating conditions. It is important to note, however, that with λ = 3 the temperature range of total CO-conversion was more limited than with λ = 4.
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Fig. 8. Effect of λ (CO/O2 ratio obtained by varying O2 concentration in the standard feed gas) on CO conversion and selectivity vs. temperature for 1% Pt-3A catalyst. Hourly space velocity: 67,200 h−1 .
Fig. 9. Effect of hourly space velocity on CO conversion as a function of temperature for 1% Pt-3A catalyst with standard feed gas composition.
4. Conclusions Since the high selectivity is a key-factor for the catalyst application, 1% Pt-3A was selected to be used in its best operating conditions (λ = 3) for the further catalytic activity measurements. In Fig. 9, the CO conversion of the 1% Pt-3A catalyst is plotted versus temperature at different volumetric hourly space velocities: 67,000, 134,000, 268,000, 536,000 h−1 (VHSV, calculated as the ratio of the standard volumetric flow rate (STP, 25 ◦ C, 1 atm) and the empty reactor volume occupied by the catalyst bed [19]). Complete CO conversion was reached for all the VHSV values with a corresponding temperature increasing with VHSV: from 185 to 205 ◦ C, 224 and 243 ◦ C for VHVS equal to 67,000, 134,000, 268,000 and 536,000 h−1 , respectively. The diffusional resistances, therefore, are not the rate limiting step as the process is mostly under the kinetic control. An outlet CO concentration of 10 ppmV is kept up to 264 ◦ C at the space velocity of 536,000 h−1 . In contrast, the selectivity was not influenced by the VHVS variation: for all the considered conditions, it remained at about 33% at the complete CO-conversion temperature. The catalytic performance of 1% Pt-3A catalyst at the higher VHVS of 536,000 h−1 , because of its complete CO conversion at 264 ◦ C, makes practicable the actual elimination of the heat exchanger downstream the LTWGS unit (270 ◦ C, Fig. 1).
A variety of catalytic materials, noble metals (Pd, Ru, Pt) supported on zeolites, were tested in a laboratory test reactor fed with a stream flow rate containing CO, CO2 , H2 , O2 , H2 O, in order to identify the best catalyst to remove CO, a poison for the PEM fuel cell electrodes. The catalysts were prepared by impregnation method and characterized by XRD and HRTEM analyses. The Pd and Ru-supported catalysts did not reach a complete CO-conversion, whereas the Pt-catalysts showed the highest CO-conversion and selectivity. In particular the 1% Pt-3A catalyst kept the complete CO-conversion in a wide temperature range, showing the highest selectivity for the CO oxidation with minimal involvement in side reactions, such as H2 oxidation and RWGS reaction. The RWGS outcome seems to be directly related to the support structure, whereas the H2 oxidation can depend on the nature and the structure of both the active phase (i.e. the noble metal) and the support. With the most promising material, the 1% Pt-3A catalyst, the CO/O2 ratio (λ) and the hourly space velocity were varied. The best catalytic performance was observed at λ = 3; the decrease of the contact time brought about an increase of the temperature corresponding to the complete CO-conversion. The complete CO-conversion and the maximum attainable selectivity observed for the
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1% Pt-3A catalyst at a temperature of 264 ◦ C (λ = 3 and VHVS = 536,000 h−1 ) could make practicable the suppression of the heat exchanger downstream the LTWGS unit.
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Development of A zeolites-supported noble-metal catalysts for CO preferential oxidation: H2 gas purification for fuel cell Ilaria Rosso∗ , Camilla Galletti, Guido Saracco, Edoardo Garrone, Vito Specchia Dipartimento di Scienza dei Materiali e Ingegneria Chimica, Politecnico di Torino, Corso Duca degli Abruzzi, 24, 10129 Turin, Italy Received 27 April 2003; received in revised form 7 August 2003; accepted 20 October 2003
Abstract Even traces of CO in the hydrogen-rich feed gas to proton exchange membrane fuel cells (PEMFC) poison the platinum anode electrode and dramatically decrease the power output. In this work, a variety of catalytic materials consisting of noble metals supported on A zeolites were synthesised, characterised and tested under realistic conditions in the quest of a catalyst for the removal of CO via the CO preferential oxidation (CO-PROX) reaction. Pt, Pd and Ru-based catalysts, prepared by wet impregnation and characterised by XRD and HRTEM, were investigated in a fixed bed reactor, by determining CO conversion and selectivity through the outlet concentrations of CO and O2 . In contrast to supported Pd and Ru catalysts, Pt-catalysts showed complete CO-conversion and a comparatively high selectivity. The 1% Pt-3A catalyst showed the best performance: it kept the complete CO-conversion in a wide temperature range, showing the highest selectivity for CO oxidation with minimal involvement in side reactions, such as H2 oxidation and RWGS reaction. Experimental data proved that the RWGS outcome is directly related to the support structure. The rather high temperature (≈260 ◦ C) at which complete CO conversion is achieved by the 1% Pt-3A catalyst enables to locate the CO-PROX unit immediately after the low temperature water-gas shift unit of the fuel processor converting hydrocarbons into hydrogen-rich gas. © 2003 Elsevier B.V. All rights reserved. Keywords: Pt/Pd/Ru-A zeolite; CO preferential oxidation; Return water gas shift
1. Introduction Because of the perspectives of a significantly higher efficiency and of almost no emission of pollutants, proton exchange membrane fuel cells (PEMFC) have been extensively studied in the last two decades for many applications and especially for low emissions vehicles [1,2]. Pure hydrogen is the ideal fuel for the PEMFC. Currently, however, there is no available technology for the safe and economical storage of enough hydrogen to give to a PEMFC powered vehicle an acceptable driving range. On-demand generation of hydrogen by reforming methanol or other liquid hydrocarbon fuels in a fuel processor is a viable means of providing the needed hydrogen and the driving range. To this purpose, a complex system has to be developed including feeding, production and purification of the gas rate flowing to the fuel cells. The development of a complete on-board fuel processor for the production of clean hydrogen from commercial gaso∗ Corresponding author. Tel.: +39-011-5644710; fax: +39-011-5644699. E-mail address: [email protected] (I. Rosso).
0926-3373/$ – see front matter © 2003 Elsevier B.V. All rights reserved. doi:10.1016/j.apcatb.2003.10.016
line is the objective of an industrial-type EU project, PROFUEL, which involves, besides our group, several partners from the automotive and catalyst manufacturing industries [3]. The fuel processor is schematically described in Fig. 1. It consists of: a dosing system, an autothermal reformer, a high-temperature water-gas-shift (HTWGS) reactor, a desulphuriser, a low-temperature water-gas-shift (LTWGS) reactor and a CO preferential oxidation unit (CO-PROX). All primary fuel processing reactions, which take place in the reformer, produce substantial amount of carbon monoxide, which is, unfortunately, incompatible with direct feeding of PEMFC stacks, as CO strongly chemisorbs on the platinum-based electrodes and deactivates them. In order to reduce drastically the CO concentration and increase the hydrogen concentration, the water-gas-shift reactors (CO + H2 O → CO2 + H2 ) are placed immediately after the reformer. This reaction reduces the CO content of the hydrogen-rich gas stream to about 0.5–1% (5000–10,000 ppm) [4], still far above the 10 ppm tolerance limit of a typical anode catalyst, so that a third process unit is required. Palladium-based membrane purification, catalytic methanation and catalytic CO preferential oxida-
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Fig. 1. Schematic description of a gasoline processor for hydrogen production. The average operating temperatures of each process stage are indicated.
tion are the main options in this context. The CO-PROX method is the most promising and economic approach: CO is oxidised with added air (CO + (1/2)O2 → CO2 ) with as low as possible concomitant loss of hydrogen by direct oxidation. Since LTWGS operates at temperature between 200 and 270 ◦ C and the fuel cell at 80 and 100 ◦ C, the available temperature operating range for the CO-PROX reactor is 80–270 ◦ C. The majority of catalysts for the CO-PROX have been studied for an intermediate temperature applications (150 ◦ C), in which case two coolers, upstream and downstream of the CO-PROX unit, are necessary (Fig. 1). The literature reports several studies on catalysts for the selective oxidation of CO, including noble metals (Pd, Ru, Rh, Pt) [5–8] and base metals (Co–Cu, Ni–Co–Fe, etc.) [8–10] generally supported on alumina. Pt-group-metals catalysts exhibit good activity and selectivity in a temperature range of 130–200 ◦ C. Activity in the order: Ru > Rh > Pt > Pd was observed [8]. The performances of the base-metal-based supported catalysts are less promising [8]. Although alumina is the most common support for noble metal-based catalysts, recent studies [11,12] have indicated that the use of zeolites as supports for platinum catalysts is promising to improve the catalyst selectivity, probably owing to the molecular sieve effects. In particular, Pt-A catalysts resulted the most selective among those tested (Pt-mordenite, Pt-X, Pt-A and Pt-Al2 O3 ) [11]. Since A zeolites are attractive also because of their low cost and high thermal stability [13], this paper describes the development of catalysts for preferential CO oxidation consisting of A zeolite-supported Pd, Ru, Pt. Particular attention was paid to the catalytic activity and selectivity in H2 , CO2 , H2 O rich stream as a function of temperature, in order to find a catalyst able to completely remove CO at a temperature similar to that of the outlet gas from LTWGS, as this would enable elimination of the intermediate cooler.
2. Experimental 2.1. Catalysts preparation and characterisation A-type zeolites with different pores size were purchased from Fluka: 3A-type zeolite (K12 [(AlO2 )12 (SiO4 )12 ] ×
H2 O), 4A-type zeolite (Na12 [(AlO2 )12 (SiO4 )12 ] × H2 O), 5A-type zeolite (Can Na12−2n [(AlO2 )12 (SiO4 )12 ] × H2 O), with pores of about 3, 4 and 5 Å diameter, respectively. All catalysts were prepared by impregnation [12,14,15]: a proper amount of precursor, to obtain 1 wt.% noble metal catalyst, was dissolved in 25 ml of distilled water; then 5 g of A-type zeolite were added and the mixture was stirred for 30 min at about 50 ◦ C. The mixture was dried in an oven at 90–100 ◦ C and, after grinding in an agate mortar, it was calcined in an electric oven in calm air at 500 ◦ C for 1 h. After calcination, the catalyst was reduced in H2 flow rate (50 N ml min−1 ) at 500 ◦ C for 2 h. Precursors for Pt, Pd and Ru-catalysts were: H2 PtCl6 (0.1327 g), Pd(NH3 )4 Cl2 (0.115 g) and Ru(NH3 )6 Cl3 (0.153 g) (all from Alfa), respectively. XRD analysis (Philips PW1710 apparatus equipped with a monochromator for the Cu K␣ radiation) was performed both on pure zeolites and on the prepared catalysts before and after the catalytic activity tests in order to verify the integrity of zeolite structure after exposition to water, present in the feed stream (see Section 2.2) and as a product of oxidation reactions. Impregnated catalysts were analysed by high-resolution transmission electron microscopy (HRTEM, Jeol JEM 2010 apparatus) to investigate the metal dispersion on supports. Temperature programmed CO2 desorption (TPD) tests were performed on zeolite 3A, 4A and 5A by a TPD/R/O apparatus (Thermoquest TPD/R/O 1100 analyser equipped with a Baltzer Quadstar 422 quadrupole mass spectrometer). CO2 saturation was performed by flowing 50 N cm3 min−1 CO2 for 1 h at 40 ◦ C, then a helium flow rate of 10 N cm3 min−1 was fed to the reactor while increasing the temperature at 5 ◦ C min−1 rate up to 450 ◦ C. 2.2. Reactor system and analytical methods Catalyst pellets were obtained by pressing at 125 MPa the powders into tablets, which were then crushed and sieved to produce 0.25–0.42 mm granules. A fixed-bed of about 2 cm in length, containing 0.15 g of catalyst pellets was enclosed in a glass tube (i.d.: 4 mm) and sandwiched between two glass–wool layers. The reactor was placed in a PID regulated oven, and a K-type thermocouple was inserted in the packed bed for oven regulation purposes.
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The feed stream flow rate (100 N cm3 min−1 ) contained 37 vol.% H2 , 18 vol.% CO2 , 0.5 vol.% CO, 5 vol.% H2 O, 1 vol.% O2 , 39.5 vol.% He. The desired composition and flow rate of each component (except water) were obtained by mass flow controllers, whereas a pressure transducer was used to check the pressure of inlet and outlet gases in order to check any possible undesired clogging of the catalyst fixed-bed. Water vapour was added by bubbling a He stream in a drechsel kept at a suitable operating temperature. The gas feedstock was delivered through a heated tube (above 70 ◦ C) to avoid H2 O condensation. The outlet gas stream was analysed through a Gas Chromatograph (Varian CP-3800) equipped with a thermal conductivity detector (TCD); the GC contained a “Poraplot Q” column (0.53 mm diameter by 30 m length), to separate CO2 and H2 O, and a “Molsieve 5A” column (0.53 mm diameter by 25 m length), to separate CO, H2 and O2 . The two columns were connected in series by a six-way valve. The columns were kept at 70 ◦ C and the sample injection was accomplished using helium as the carrier gas at a flow rate of about 2.8 ml/min. Detection limit of CO is 10 ppmV. The conversion of CO (ξ CO ) and O2 (ξO2 ), as well as the selectivity of CO oxidation (σ), determined in the 100–300 ◦ C range in the presence of excess hydrogen, were calculated as follows: [CO]out (1) ξCO = 1 − [CO]in [O2 ]out [O2 ]in
(2)
1 [CO]in − [CO]out . 2 [O2 ]in − [O2 ]out
(3)
ξO2 = 1 − and σ=
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3. Results and discussion A preliminary screening of the catalytic performance of all prepared catalysts, in terms of CO, O2 conversion and CO selectivity, was carried out. XRD analysis, performed on all zeolite-supported catalysts after the catalytic activity tests, demonstrates that the A zeolite structure is substantially preserved notwithstanding the water presence in the feed stream (5% H2 O). Although some authors proved that a negligible degradation of activity occurs on Pt-mordenite when water (up to 20%) was added to the feed stream [11], the resistance of A zeolites to high water amounts will be studied in further studies, as this is a very critical point for the practical application of A zeolite. The CO outlet concentration values for the three 1% Pd-zeolites are reported in Fig. 2. CO outlet concentrations decreased starting from about 150 ◦ C, reached the lowest values at about 180–200 ◦ C and increased again at temperature higher than 200 ◦ C. No catalyst reached total CO conversion: the lowest CO outlet concentration (about 930 ppmV) was obtained by 1% Pd-5A at 205 ◦ C. In contrast, the O2 conversion was complete in a temperature range of 200–300 ◦ C for all the catalysts. This means that the oxygen was consumed in the CO oxidation reaction: CO + 21 O2 → CO2
(4)
and also in the hydrogen oxidation reaction: H2 + 21 O2 → H2 O
(5)
which brings about a decrease in oxygen concentration and in power generation. The resulting selectivity σ reached values never higher than 20% for all the 1% Pd-catalysts. The inlet gas mixture
Fig. 2. CO outlet concentration vs. temperature for 1% Pd-3A, 1% Pd-4A and 1% Pd-5A catalysts with standard feed gas composition (5000 ppm CO, 1% O2 , 18% CO2 , 5% H2 O, 37% H2 and He as balance). Hourly space velocity: 67000 h−1 .
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Fig. 3. CO outlet concentration vs. temperature for 1% Ru-3A, 1% Ru-4A and 1% Ru-5A catalysts with standard feed gas composition. Hourly space velocity: 67,000 h−1 .
used in the catalytic tests contained 5000 ppmV CO and 1 vol.% oxygen which is four-fold the stoichiometric value needed for the reaction (1). As a consequence, when both CO and O2 conversions are equal to 1, it is not possible to obtain a selectivity value higher than 25%. As far as the 1% Ru-supported catalysts are concerned, Fig. 3 shows that 1% Ru-3A and 1% Ru-4A catalysts gave rather high CO outlet concentration values in the whole temperature range: the lowest CO outlet concentrations were 1650 ppmV at 165 ◦ C, and 1170 ppmV at 148 ◦ C for 1% Ru-3A and 1% Ru-4A catalysts, respectively. The 1% Ru-5A catalyst showed a particular trend: it reached 700 ppmV of CO outlet concentration at about 125 ◦ C, then it increased up to 2850 ppmV at 184 ◦ C and decreased again to 10 ppmV (the detection limit of CO) at 250 ◦ C. The complete CO conversion observed at high temperature is quite strange but can be explained considering the CH4 formation according to the reaction: CO + 3 H2 → CH4 + H2 O
(6)
that brings about a consumption of CO. Methane was in fact detected in the exhaust gases above 180 ◦ C and its amount was equal to about 0.5% in the temperature range between 240 and 280 ◦ C. The complete CO conversion in the same temperature range could confirm the occurrence of CO methanation (reaction (6)) rather than CO2 methanation. Accordingly, some authors [6] demonstrated that CO methanation is favoured in CO/CO2 /H2 mixture rather than CO2 methanation. Some methanation level was also detected in this temperature range for the 1% Ru-3A and the 1% Ru-4A catalysts. These results are in agreement with literature data [6,7], reporting CH4 formation on Ru-based catalysts above 180 ◦ C.
The O2 conversion was complete at temperatures higher than 165, 148, 125 ◦ C for 1% Ru-3A, 1% Ru-4A, 1% Ru-5A catalysts, respectively, and the resulting selectivity values were, once again, not higher than 20% for all the catalysts. Similar experiments carried out on 1% Pt-zeolite catalysts produced quite different results. Fig. 4 shows that the 1% Pt-3A and the 1% Pt-5A catalysts yielded total CO conversion (CO outlet concentration ≤10 ppmV) in the range of 165–220 ◦ C, whereas the lowest value of CO outlet concentration achieved by 1% Pt-4A catalyst was 100 ppmV at 165 ◦ C. The 1% Pt-3A catalyst kept the complete CO conversion in a wide range of temperature: from 185 to 220 ◦ C; in contrast, at this high temperature the CO outlet concentrations for 1% Pt-4A and 1% Pt-5A were 1300 and 675 ppmV, respectively. All catalysts gave a selectivity of 25%, the maximum selectivity reachable when ξ CO and ξO2 are equal to 1, at the temperature of the highest CO-conversion. In good accordance with the CO-conversion data, the selectivity of 1% Pt-3A remained as high as about 25% also at very high temperatures (up to 285 ◦ C), whereas the other two catalysts showed selectivity value lower than 20% at the same temperature. The decrease of CO-conversion (due to the high CO outlet concentration data) observed at high temperature for 1% Pt-4A and 1% Pt-5A catalysts seems to be mainly a consequence of the loss of selectivity: at high temperature the amount of O2 available for CO oxidation is determined by the amount of O2 consumed for the simultaneously occurring H2 oxidation (reaction 5) [13]. However, the observed decrease in CO conversion with temperature might also be related to the reverse water gas shift reaction (RWGS): CO2 + H2 → CO + H2 O (7) limiting the CO conversion at high temperatures [16].
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Fig. 4. CO outlet concentration vs. temperature for 1% Pt-3A, 1% Pt-4A and 1% Pt-5A catalysts with standard feed gas composition. Hourly space velocity: 67,000 h−1 .
To check the occurrence of reaction (7), the inlet gas mixture was fed without CO to the Pt-catalysts and the CO appearance was measured. Fig. 5 shows the CO outlet concentrations obtained with the three Pt-catalysts using the standard inlet gas mixture (full symbols) and the CO-free inlet gas mixture (empty symbols); the resulting difference ( CCO ) between these two outlet CO concentrations is also reported. With CO-free feedstock, the RWGS reaction is very limited for the 1% Pt-3A and the 1% Pt-5A catalysts and evident for the 1% Pt-4A one. By comparing these CO outlet concentrations with those obtained from the activity tests (standard feed mixture), in the range of high temperature where O2 conversion is practically equal to 1, the very low loss of selectivity observed for the 1% Pt-3A catalyst seems to be mainly due to the RWGS reaction, whereas the substantial loss of selectivity obtained for 1% Pt-4A and 1% Pt-5A catalysts seems to be caused by both reaction (7) (RWGS) and reaction (5) (H2 oxidation). However, for 1% Pt-5A catalyst, the contribution of the RWGS reaction to the outlet CO concentration can be considered negligible, as compared to that of H2 oxidation. The CCO values, reported for each catalyst and compared in the Fig. 5d, represent the CO amount that cannot be oxidised because the oxygen has been completely consumed. In summary, CCO remains very low for the 1% Pt-3A catalyst, it is progressively increasing with temperature for the 1% Pt-4A catalyst and presents high values in all the considered temperature range for the 1% Pt-5A catalyst. Pt dispersion of Pt-based catalysts was evaluated by HRTEM analysis. TEM micrographs of 1% Pt-3A, 1% Pt-4A and 1% Pt-5A catalysts are shown in Fig. 6A–C, respectively. All the catalysts show a rather poor homogeneous Pt dispersion with large Pt clusters on preferential zones of
the zeolite support. The average Pt crystal size varies from a few nanometers (2–3 nm) to some tens nanometers (20–25 up to 50 nm) for all catalysts. These data are in good agreement with X-ray diffraction investigation, as the peaks of metallic platinum (JCPDS card 04-0802) were well visible in the diffraction pattern of the Pt-catalysts, suggesting the presence of large Pt crystallites. The platinum crystal size of each Pt-catalyst was also estimated by the measurement of the full width half maximum (FWHM) of the platinum lines in the diffraction pattern according to Debye–Scherrer equation [17]. The resulting average platinum particle size was in the range of 21–26 nm; so almost all the employed Pt would be outside the zeolite pores. This seems to be a positive feature for the Pt-A zeolite catalysts prepared by wet impregnation, at least if compared with those prepared by cation-exchange of the A zeolites in an aqueous solution of Pt(NH3 )4 Cl2 ·H2 O for which a considerable fraction of the employed platinum would be inside the zeolite pores: those latter, in fact, gave worse catalytic performance. The superior performance of the 1% Pt-3A catalyst, as regards the RWGS inhibition might be due to the structure of the 3A zeolite. Its pore size (3 Å) is close to the diameter of CO2 molecule (2,5 Å), which could render difficult the CO2 admittance into 1% Pt-3A pores, in contrast with the 4A and 5A-zeolites pores (4 and 5 Å, respectively). Also the different basicity of the zeolites 3A, 4A and 5A, related in turn to the different nature of their counter ions (K+ , Na+ and Ca++ , respectively) may have an influence: it this could explain a different interaction of supports with CO2 and a consequent different outcome of the RWGS reaction. Temperature programmed CO2 desorption tests, performed on the 3A, 4A and 5A zeolites (Fig. 7) show indeed a different interaction with CO2 . The zeolite 3A shows a small
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Fig. 5. CO outlet concentration in catalytic activity tests carried out on (a) 1% Pt-3A, (b) 1% Pt-4A and (c) 1% Pt-5A catalysts; (d) CCO . Full symbols: standard feedstock. Empty symbols: CO-free feedstock. (×, +, ) symbols: CCO for 1% Pt-3A, 1% Pt-4A, 1% Pt-5A catalysts, respectively. Hourly space velocity: 67,000 h−1 .
peak at about 80 ◦ C, whereas the zeolite 4A desorbs a much higher CO2 amount at about 100 ◦ C. Instead, the zeolite 5A shows two small CO2 desorption peaks at about 90 and 230 ◦ C. These results are in good agreement with the heats of adsorptions for CO2 in the A zeolites determined from
IR isotherms by other authors [18]: CO2 presented increasing values of heat of adsorption versus K+ , of zeolite 3A, Na+ of zeolite 4A and Ca2+ of zeolite 5A. Accordingly, the CO2 desorption from zeolite 3A, 4A and 5A occurs at rising temperature (Fig. 7). The two CO2 desoption peaks
Fig. 6. HRTEM micrographs of Pt-catalysts: (A) 1% Pt-3A (600,000×); (B) 1% Pt-4A (300,000×); (C) 1% Pt-5A (80,000×).
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Fig. 7. Temperature programmed CO2 desorption for zeolite 3A, zeolite 4A and zeolite 5A.
observed for zeolite 5A could probable be due to the presence of both Na+ (peak at about 90 ◦ C) and Ca2+ (peak at about 230 ◦ C) cations, the latter can be considered stronger adsorption sites because of their bivalent charge. Similar experiments performed on the Pt-A zeolite catalysts showed the same results, in that the metallic particles seem not to be involved in the CO2 interaction. This means that also the possible presence of traces of chlorine from chlorine-containing Pt-precursor does not influence the CO2 interaction on zeolites. The limited interaction of the zeolite 3A with CO2 , whatever the origin, either geometrical or chemical, could account for the higher selectivity and the total CO-conversion (up to 10 ppmV) in a wide temperature range (185–220 ◦ C) of the 1% Pt-3A catalyst. The low CO2 concentration in the neighbourhood of the catalytic active sites, due to its low affinity with the support, could be responsible for: (i) the CO preferential oxidation (reaction (4)) enhancement; (ii) the RWGS (reaction (7)) inhibition. By the same token, the strong interaction between CO2 and zeolite 4A could explain the high contribution of the RWGS (reaction (7)) to the substantial loss of selectivity of the 1% Pt-4A catalyst. Zeolite 5A shows, as a whole, a limited interaction with CO2 , that could account for the negligible occurrence of the RWGS with the 1% Pt-5A catalyst, as for zeolite 3A. As these results regard the interaction between zeolites A and CO2 , the RWGS (reaction (7)) should have similar occurrence on Pd/Ru/Pt-A zeolites catalysts, independently of the employed noble metal. In contrast, the CO outlet concentrations observed at high temperature (>180 ◦ C) show different trends if Pd/Ru/Pt-3A, Pd/Ru/Pt-4A, and Pd/Ru/Pt-5A catalytic performances are compared (Figs. 2, 3 and 4). This fact, however, is not conclusive, because not the sole RWGS
but also the H2 oxidation (reaction (5)) and—limitedly to Ru-based catalysts—the methanation (reaction (6)) bring about the final measured CO outlet concentration. We experimentally verified that with CO-free feed mixture the RWGS reaction is very limited for not only for the 1% Pt-3A catalyst but also for the 1% Pd-3A and 1% Ru-3A ones: this means that the RWGS outcome seems to be directly related to the support structure. In contrast, the H2 oxidation can depend on the nature and the structure of both the active phase (i.e. the noble metal) and the support: its outcome, in fact, is very different for each tested catalyst. Further investigations have been performed on the promising 1% Pt-3A catalyst evaluating the effect of oxygen amount and hourly space velocity variation on its activity and selectivity. Different experiments were performed by varying O2 concentration in the feed flow rate: 1, 0.75, 0.5 vol.% O2 were used, with corresponding λ values of 4, 3, 2 (λ = 2[O2 ]/[CO]), respectively. Fig. 8a shows the CO-conversion values versus temperature obtained in the different experimental conditions: CO conversion practically equal to 1 was achieved at 185 ◦ C with λ = 3 and 4, whereas the maximum CO conversion at the same temperature with λ = 2 was equal to 0.6. This means that the catalyst is not so selective to completely oxidise CO even when an oxygen amount twice higher than the stoichiometric value is employed. As a consequence, the highest selectivity (Fig. 8b) was reached with λ = 3: it was equal to about 33%, which is the maximum value reachable at the tested operating conditions. It is important to note, however, that with λ = 3 the temperature range of total CO-conversion was more limited than with λ = 4.
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Fig. 8. Effect of λ (CO/O2 ratio obtained by varying O2 concentration in the standard feed gas) on CO conversion and selectivity vs. temperature for 1% Pt-3A catalyst. Hourly space velocity: 67,200 h−1 .
Fig. 9. Effect of hourly space velocity on CO conversion as a function of temperature for 1% Pt-3A catalyst with standard feed gas composition.
4. Conclusions Since the high selectivity is a key-factor for the catalyst application, 1% Pt-3A was selected to be used in its best operating conditions (λ = 3) for the further catalytic activity measurements. In Fig. 9, the CO conversion of the 1% Pt-3A catalyst is plotted versus temperature at different volumetric hourly space velocities: 67,000, 134,000, 268,000, 536,000 h−1 (VHSV, calculated as the ratio of the standard volumetric flow rate (STP, 25 ◦ C, 1 atm) and the empty reactor volume occupied by the catalyst bed [19]). Complete CO conversion was reached for all the VHSV values with a corresponding temperature increasing with VHSV: from 185 to 205 ◦ C, 224 and 243 ◦ C for VHVS equal to 67,000, 134,000, 268,000 and 536,000 h−1 , respectively. The diffusional resistances, therefore, are not the rate limiting step as the process is mostly under the kinetic control. An outlet CO concentration of 10 ppmV is kept up to 264 ◦ C at the space velocity of 536,000 h−1 . In contrast, the selectivity was not influenced by the VHVS variation: for all the considered conditions, it remained at about 33% at the complete CO-conversion temperature. The catalytic performance of 1% Pt-3A catalyst at the higher VHVS of 536,000 h−1 , because of its complete CO conversion at 264 ◦ C, makes practicable the actual elimination of the heat exchanger downstream the LTWGS unit (270 ◦ C, Fig. 1).
A variety of catalytic materials, noble metals (Pd, Ru, Pt) supported on zeolites, were tested in a laboratory test reactor fed with a stream flow rate containing CO, CO2 , H2 , O2 , H2 O, in order to identify the best catalyst to remove CO, a poison for the PEM fuel cell electrodes. The catalysts were prepared by impregnation method and characterized by XRD and HRTEM analyses. The Pd and Ru-supported catalysts did not reach a complete CO-conversion, whereas the Pt-catalysts showed the highest CO-conversion and selectivity. In particular the 1% Pt-3A catalyst kept the complete CO-conversion in a wide temperature range, showing the highest selectivity for the CO oxidation with minimal involvement in side reactions, such as H2 oxidation and RWGS reaction. The RWGS outcome seems to be directly related to the support structure, whereas the H2 oxidation can depend on the nature and the structure of both the active phase (i.e. the noble metal) and the support. With the most promising material, the 1% Pt-3A catalyst, the CO/O2 ratio (λ) and the hourly space velocity were varied. The best catalytic performance was observed at λ = 3; the decrease of the contact time brought about an increase of the temperature corresponding to the complete CO-conversion. The complete CO-conversion and the maximum attainable selectivity observed for the
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1% Pt-3A catalyst at a temperature of 264 ◦ C (λ = 3 and VHVS = 536,000 h−1 ) could make practicable the suppression of the heat exchanger downstream the LTWGS unit.
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