Commercial Steam Reforming Catalysts To Improve Biomass

2668

Ind. Eng. Chem. Res. 1998, 37, 2668-2680

Commercial Steam Reforming Catalysts To Improve Biomass Gasification with Steam-Oxygen Mixtures. 2. Catalytic Tar Removal Marıá P. Aznar,*,† Miguel A. Caballero,† Javier Gil,† Juan A. Martıń,† and Jose´ Corella‡ Chemical and Environmental Engineering Department, University of Saragossa, 50009 Saragossa, Spain, and Chemical Engineering Department, University “Complutense” of Madrid, 28040 Madrid, Spain

Eight different commercial catalysts, nickel based, for steam reforming of naphthas and of natural gas are tested in biomass gasification for hot gas cleanup and conditioning. They were manufactured by BASF AG, ICI-Katalco, UCI, and Haldor Topsøe a/s. The catalysts were tested in a slip flow after a biomass gasifier of fluidized bed type at small pilot-plant scale (10-20 kg of biomass/h). The gasifying agent used is steam-oxygen mixtures. A guard bed containing a calcined dolomite is used to decrease the tar content in the gas at the inlet of the catalytic bed. Main variables studied are catalyst type, bed temperature, H2O + O2 to biomass feed ratio, and time-on-stream. All catalysts for reforming of naphthas show to be very active and useful for tar removal and gas conditioning (in biomass gasification). 98% tar removal is easily obtained with space velocities of 14 000 h-1 (n.c.). No catalysts deactivation is found in 48 h-on-stream tests when the catalyst temperature is relatively high (780-830 °C). Using a simple first-order kinetic model for the overall tar removal reaction, apparent energies of activation (of around 58 kJ/mol) and preexponential factors are obtained for the most active catalysts. Introduction It is well-known how biomass gasification in fluidized bed produces a dirty raw gas that has to be cleaned of tar and particulates for most of its applications. Wet gas cleaning is usually not desired because it produces a flow of contaminated condensates of very difficult disposal and, besides, it cools the flue gas, lowering the overall thermal efficiency of the biomass-to-electricity process. Hot dry gas cleaning and upgrading is nowadays the best solution which, in turn, has two, at least, main ways, solutions or processes: using calcined dolomites (OCa‚OMg) or related materials and using steam reforming, nickel-based, catalysts. Calcined dolomites are cheap and useful materials, above 800 °C, for this application, but their use and study is here out of the scope of this paper which is concentrated on the use of nickel-based catalysts. Steam reforming catalysts have already proved their usefulness in biomass gasification for gas upgrading. They eliminate tars present in the raw flue gas by steam and dry (CO2) reforming reactions whose mechanism and reacting network are not well-known nowadays but which seem similar to the steam reforming of natural gas and of naphthas. The state-of-the-art and the institutions working or having worked in this process of hot gas cleaning with nickel catalysts have been reported previously (Aznar et al., 1992, 1993; Narvaéz et al., 1997; Aldeń et al., 1997; Caballero et al., 1997; Corella et al., 1998; Abatzoglou et al., 1998). In part 1 of this work (Caballero et al., 1997) it was shown how gas composition (H2, CO, CO2, CH4, C2, steam, ..., contents), heating value of the gas, gas yield, and * Author to whom correspondence should be addressed. † University of Saragossa. Fax: + 34 76 76 21 42. ‡ University “Complutense” of Madrid. Fax: + 34 1 394 41 64.

thermal efficiency of the gasification process were modified by a catalytic bed (nickel-based catalysts) located in a slip stream downstream of a fluidized-bed biomass gasifier. Now, in this part, it will be shown in detail how this catalytic bed eliminates the tars present in the flue gas and which variables affect this tar removal. This study is carried out at small pilot-plant scale with eight different commercial steam reforming catalysts. Some deactivation studies or lives of the catalysts are also presented. Experimental Section Pilot Plant Used. The present study has been carried out in a facility which can be considered a small pilot plant. It is based on a bubbling fluidized bed of 15-cm i.d. and 3.2-m height continuously fed with biomass near the bed bottom at flow rates of around 10 kg of biomass/h. It has been described previously (Aznar et al., 1995, 1996, 1997; Gil et al., 1997). For further comparisons on catalytic tar-removal activities under different flue gas atmospheres (compositions), it is clearly stated that the gasifying agent in this work has always been steam-O2 mixtures. Since there are thus three main reactants (steam, O2, and biomass), the following two ratios are needed to describe the reacting system at the gasifier inlet: (a) (H2O + O2)/biomass dry and ash-free, (kg/h)/(kg of daf/h), dimensionless. This gasifying ratio, which will be here referred to as GR, has an intrinsic meaning similar to the “equivalence ratio” used in biomass gasification with air. (b) (H2O/O2), (mol/h)/(mol/h), dimensionless. Gasifying with steam-O2 mixtures, gas quality, and product distribution at the gasifier exit have been studied in detail under very different process variables, and they have been reported previously (Gil et al., 1997). From that previous study, the intervals selected for the

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Ind. Eng. Chem. Res., Vol. 37, No. 7, 1998 2669 Table 1. Some Main Experimental Conditions and Apparent Kinetic Constants for Tar Removal over Several Commercial Steam Reforming (Ni-Based) Catalysts kapp,tar run no.

catalyst

dp (mm)

5 6 7 8 9 10 11 12 13 14 15 16 20 21 22 23 24 25 26 27 28 29 30 31

UC C-11-9-062 UC C-11-9-062 UC C-11-9-062 BASF G25-1S nickel A Topsøe RKS-1 nickel Db nickel E nickel D silica sand nickel B nickel E nickel D nickel B nickel A nickel D nickel D nickel B nickel D nickel A nickel B silica sand nickel A nickel D

-0.8 + 0.2 -0.82 + 0.2 -0.8 + 0.2 -1.0 + 0.2 -1.0 + 0.2 -1.0 + 0.2 -1.0 + 0.2 -1.0 + 0.2 -1.0 + 0.2 -1.0 + 0.2 -1.0 + 0.2 -1.0 + 0.2 -1.0 + 0.2 -1.0 + 0.2 -1.0 + 0.2 -1.0 + 0.2 -1.0 + 0.2 -1.0 + 0.2 -1.0 + 0.2 -1.0 + 0.2 -1.0 + 0.2 -1.0 + 0.2 -1.0 + 0.2 -1.0 + 0.2

a

Tb (°C)

τ [(kg h)/ [m3 (Tb, wet)]]

SV [mn3 (wet)/ (m3 h)]

yH2Oa (vol %)

[mn3 (dry)/ (kg h)]

[mn3 (wet)/ (kg h)]

[m3 (Tb, wet)/ (kg h)]

750 725 735 785 800 785 825 770 760 820 790 800 670 630 640 804 780 775 830 790 780 820 800 780

0.127 0.122 0.022 0.031 0.025 0.036 0.033 0.037 0.032 0.080 0.035 0.037 0.025 0.028 0.027 0.021 0.019 0.023 0.020 0.027 0.025 0.029 0.035 0.034

3000 2700 5500 13800 11900 10500 9700 12700 10300 5300 9200 8400 13100 13400 14700 15000 16000 14000 15000 13100 13000 7500 12000 11500

39 45 31 41 46 43 57 55 50 45 40 42 46 47 45 40 44 53 56 45 53 42 52 55

5 5 9 33 22 19 7 12 15 3 18 21 26 29 25 51 44 23 27 28 29 4 28 22

8 9 12 55 41 33 16 27 29 7 30 36 47 56 45 85 79 49 62 52 62 7 42 35

30 34 46 214 175 129 64 103 111 29 118 143 164 183 149 349 333 188 249 201 238 17 132 98

Average steam content in flue gas, from mass balance. b Sintered.

gasifier operation in this work are

0.70 e GR e 1.60 2.0 e (H2O/O2) e 3.0 800 °C < gasifier bed temperature < 860 °C In this work the gasifier bed is silica sand only (without in-bed dolomite), with the corresponding char produced and existing under the stationary state. One test or experiment lasted on average 1.2 months (including test and feedstock preparation, analysis of the gas and tar samples, cleaning of the pilot plant, etc, ...). The running of the pilot plant during one test lasted on average 20 h (the plant was not able to work under reaction overnight because of its location in the University), of which about 10 h were under the stationary state (in all of its reactors, vessels, and equipment). Data reported here will refer thus only to stationary states. For catalyst deactivation studies, the plant was “stopped” overnight and operated the next day. Biomass Used. The biomass used as the feedstock in all tests has been pine (Pinus Pinaster) wood chips of -5.0 + 1.0 mm. This biomass was always partially dried until a moisture content of 10-12 wt % existed. This feedstock did not cause problems in the feeding system used in the pilot plant. Its physical and chemical characterization have been reported previously (i.e., Narvaez et al., 1996). Catalytic Reactor. The catalytic reactor was a downward fixed bed of 4.1-cm i.d. It was described in detail in part 1 of this work (Caballero et al., 1997). Four main variables have been studied in this work: catalyst type, temperature (Tb) of the catalytic bed, gas residence time expressed as space time (τ) or as space velocity (SV), and gas atmosphere composition which depends on the GR value used in the gasifier. Due to the type and size of the catalytic bed, radial and longitudinal temperature gradients were foreseen in it. For this

reason temperatures in the bed were measured with three thermocouples (two in the bed axis and one in the bed wall, inner side). The temperature used here as reference (Tb) is the average (arithmetic mean) between the bed axis and the wall (inner side) at 5 cm (roughly the middle of the bed height) from the bed inlet. Several research programs (such as AIR2, JOULE2, JOULE3) of the European Union (DGXII) are financing projects on catalytic hot gas cleaning (since the present one) in which several Institutions from different countries are participating and have to compare their results. In periodic and internal meetings among the partners in these projects, it has been made clear how important (for comparison purposes) it is to define in the same way the variables to be further compared. This is the reason for the units selected and used for τ and SV in this work. The activity of the catalysts for the overall tar elimination will be given by a kinetic constant (kapp,tar). It is based on or deduced from a first-order kinetic equation and a single and overall reaction of tar elimination (Narvaéz et al., 1997). The units of this index (see Table 1, for instance) are also very important for further comparisons. Catalysts Used. Eight different commercial steam reforming catalysts have been tested in this work. They come from the four main manufactures of this type of catalyst in the world: BASF AG, ICI-Katalco, Topsøe A/S, and United Catalyst Inc. (UC). Four of these catalysts are for steam reforming of light hydrocarbons (natural gas), and four are for reforming of “heavier hydrocarbons” (naphthas). Their physical (pore volume and distribution, BET surface area, ...) and chemical characterization have been shown in detail in part 1 of this work (Caballero et al., 1997). Due to secrecy agreements, some catalysts have to be referred as A, B, C, and D (having been well characterized and described in part 1 of this work). Commercial steam reforming catalysts are suministrated as rings (with one or several holes) of about 1-cm

2670 Ind. Eng. Chem. Res., Vol. 37, No. 7, 1998

Figure 1. Tar content in the flue gas in three different locations for run 28 using the nickel B catalyst.

height and 1-cm o.d. Since the catalytic reactor had 4.1cm i.d., to avoid big wall effects, these catalysts were all crushed and sieved until a size of -1.0 + 0.2 mm (which, was the size always used in this work) was reached. Crushing of these catalysts is not good, of course, mainly for the impregnated catalysts, but the authors had no other solution at hand. Two other facts concerning the particle size and shape of the catalysts used have to be taken into account: (1) The effectiveness factors of the BASF G1-25S catalyst, for the same reaction, was determined experimentally previously for different particle sizes (Narvaéz et al., 1997). From that work it was deduced that for particle diameters bigger than around 0.3 mm the internal diffusion starts to have influence on the overall process (in the apparent activation energy, for instance). To compare the intrinsic chemical activity of the catalysts, the size of -1.0 + 0.2 mm was selected. In this interval of particle sizes the internal diffusion does not control fully but it starts to have some influence (Narvaéz et al., 1997) and it will have to be taken into account (in the values obtained for the activation energy, for instance). (2) In some gasification processes the flue gas coming to the catalytic bed can contain some particles (char, silica sand elutriated from the gasifier, etc., ...). In this case a honeycomb or monolith catalyst is recommended, but this shape it is not available in the market yet for steam-reforming catalysts. So, although monoliths will be tested when available, only (crushed) commercial rings were tested here and a filter had to be used before the catalytic bed. Some main experimental conditions for the catalytic bed are shown in the first columns of Table 1. A noticeable fact is that the steam content (yH2O) in the flue gas in the catalytic reactor is quite high (clearly higher than the steam content when biomass gasification is carried out with air). The (H2O/C*) ratio was always high, always higher than 2.0 which is the

minimum value used in commercial steam reformers (Basini and Piovesan, 1998). Steam reforming of hydrocarbons such as tars can generate coke on the catalyst surface, which can deactivate the catalyst. Catalyst deactivation occurs when the throughput of tars coming to the catalytic bed is high (Aznar et al., 1993). These authors have given the limit of about 2 g of tars/mn3 to avoid in this process catalyst deactivation by coke. Since the tar content in the raw gas usually has a value higher than this limit (Gil et al., 1997), it is of basic importance to decrease this tar content in the flue gas below the said limit of 2 g of tars /mn3. This fact can be achieved in several ways. The one used in this work was to use a guard bed of calcined dolomite before the catalytic bed. This guard bed decreases the tar content in the flue gas until about 0.5-1 g/mn3 (Figure 1). This tar content can now be easily processed (removed) by the catalytic bed. Tar Definition, Sampling, Analysis, and Characterization. For comparison purposes tar definition and measurement (sampling and analysis) has been standardized very recently, as Corella and Milne et al. among others claimed (Corella, 1996; Milne et al., 1997). On March 17-19, 1998, organized by IEA, Thermal Gasification of Biomass Task, the European Commission (DGXVII), and U.S. DOE, there was a meeting in Brussels in which a protocol was accepted for tar definition and measurement in the field of biomass gasification. It will be published in June 1998. From that date tar should have the same meaning and value worldwide but the present work was made before such date, and the tar content and even the type of tar mentioned in this paper can possibly be somewhat different from the ones from other authors or gasifiers (Corella, 1996; Milne et al., 1997). Since, the tar sampling method used here was similar to the “most developed” one (as the one used by VTT in Finland) shown in the said tar protocol meeting held in Brussels, these authors believe that amounts and conclusions in

Ind. Eng. Chem. Res., Vol. 37, No. 7, 1998 2671

Figure 2. Tar conversion in the catalytic bed vs temperature [s, nickel D][0.019 < τ < 0.037 (kg h)/m3 (Tb, wet); H2O/O2 ) 2 and 3 mol/mol].

this paper concerning “tar” can be applied to other gasification plants. Although there has been an international agreement concerning tar as a whole lump, there is still a clear need for advancement in the definition of other sublumps related to tar. For instance, in catalytic gas cleaning (this work) it seems clear how there are some “easy-to-destroy (convert)” or “soft” tars and some other “difficult-to-destroy” or “hard” tars. Although these concepts and/or sublumps are not standardized yet, we will have to use them in this work. In this work the flue gas was periodically sampled (for gas and for condensates) before and after the catalytic bed and before and after the guard bed. Tars were measured in the condensates under the method previously described by Narvaéz et al. (1996), which is very similar to the standard method recently agreed upon. The tar composition varies in the guard bed and in the catalytic bed. Samples of tars were sent to NREL (Golden, CO) and to KTH (Stockholm, Sweden) to be characterized by molecular beam mass spectrometry (Evans and Milne, 1997) and by solid-phase adsorption-desorption (Brage et al., 1997), respectively. Although this characterization is not fully accomplished yet for all operation parameters, there are already useful data available in our plant on this subject. It can be said thus that the bed of calcined dolomite mostly eliminates/destroys the “soft” tars (the tars that can be easily destroyed such as phenol derivatives) and only the hard tars (naphthalene and similar PAHs) come now to the catalytic bed (Caballero et al., 1997). This fact has to be taken into account when comparing kinetic constants for tar removal obtained with nickel catalysts and with calcined dolomites: the reactant (tar) is not the same (does not have the same composition) in both beds.

Results In most of the figures in this paper (Figures 2-5, ...) results will be indicated only by a point (×, O, 0, ...) although they should be given by “not-small” rectangles (as in the simultaneous paper of Orio et al. (1997), for instance). In fact, in the catalytic bed there are temperature differences of 10-25 °C, and in tar sampling and analysis there are some errors too. Senent’s errors theory (Senent, 1962), based on statistical analysis, has always been taken into account in this work, but points instead of rectangles are used in the figures in this paper because (i) they make such figures clearer and (ii) using rectangles instead of points do not add more information and does not modify the conclusions. Statistics and errors theory have been taken into account thus, and points shown in Figures 2-5, ..., will have to be understood as the center of “the rectangle” in each measurement. Tar Conversions (in the Catalytic Bed). From the measured tar contents in the flue gas at the catalytic bed inlet and exit, tar conversion (Xtar) in the catalytic bed was calculated for all catalysts under well-known experimental conditions. This overall tar conversion includes all catalytic reactions (for tar removal) as well as thermal reactions. Effect of the Bed Temperature. The effect of the bed temperature is shown in Figure 2 for space times (τ) of 0.019-0.037 kg of catalyst‚h/m3 flue gas (Tb, wet). All points correspond to H2O/O2 ratios (in the feeding system) of 3.0, except two points shown in Figure 2. Figure 2 shows how tar conversion increases with temperature in this process. In nearly all experiments tar conversions higher than 96% were achieved. Effect of the Gasifying Ratio. The gasifying ratio [GR, or (H2O + O2)/biomass fed] used in the gasifier has an effect on tar conversion. Tar conversion depends on the gasifying ratio (GR) used in the upstream gasifier, as Figures 2 and 3 clearly demonstrate. The two lines


Figure 3. Tar conversion in the catalytic bed for different gas atmospheres [770 < Tb < 800 °C; 0.019 < τ < 0.037 (kg h)/m3 (Tb, wet)].

Figure 4. Tar concentration in the flue gas at the exit of the catalytic bed for different GR values and catalysts [770 < Tb < 800 °C; 0.019 < τ < 0.037 (kg h)/m3 (Tb, wet)].

drawn in Figure 2 correspond to the nickel D catalyst, which showed to be one of the most active and most used (in this work) catalysts. Results shown in Figures 2 and 3 indicate that tars produced at high (>0.95) GR values are more difficult to convert (destroy) than tars generated at low ( 0.95.

Table 3. Values for the Apparent Activation Energy and Preexponential Factor for the Overall Tar Removal Reaction on Several Catalysts (under Some Internal Diffusion Influence) gasifying agent catalyst Eapp, kJ/mol kapp,0, m3 (Tb, wet)/ (kg h) a

steam + O2 (this work) nickel D 58 ( 18 250 000

air (Narvaéz et al., 1997) BASF G1-25S 72 ( 12 143 000

steam (Delgado et al., 1997) calcined dolomite 97 ( 14 (1.2-1.4) × 106 a

mn3 (wet)/(kg h).

Table 4. Kinetic Constants (at t ) 0) for Methane Removal over Several Commercial Steam Reforming (Ni-Based) Catalysts kCH4 mn3 mn3 mn3 (dry)/(kg h) (wet)/(kg h) (Tb, wet)/(kg h)

run no.

catalyst

5 6 7 8 9 10 11 12 13 14 15 16 20 21 22 23 24 25 26 27 28 29 30 31

UC C-11-9-062 UC C-11-9-062 UC C-11-9-062 BASF G25-1S nickel A Topsøe RKS-1 nickel Da nickel E nickel D silica sand nickel B nickel E nickel D nickel B nickel A nickel D nickel D nickel B nickel D nickel A nickel B silica sand nickel A nickel D

a

4 5 6 21 13 38 3 29 38 0.12 41 11 12 13 14 37 21 30 38 19 25 0.12 22 40

7 9 9 35 25 66 7 65 76 0.26 68 19 23 25 25 62 38 63 86 34 54 0.24 31 80

26 33 31 116 96 245 26 226 295 1 248 76 75 81 80 242 127 224 323 125 208 1 120 320

Sintered.

Table 5. Values for the Apparent Activation Energy and Preexponential Factor for the Overall Methane Removal Reaction for Different Catalysts Tested

Figure 9. Effect of the space time and gasifying agent-to-biomass ratio on the apparent kinetic constant for tar removal over different commercial steam-reforming catalysts [760 < Tb < 810 °C].

means around 6-8 men-months, there are only a few points for some catalysts. It is difficult then to draw a line in this figure for some catalysts. After careful analysis and averaging of the results obtained for the different catalysts, the best possible fitting for all catalysts gave an apparent activation energy (Eapp) of 58 kJ/mol. This (averaged) value for Eapp (with some error, of course) was the one selected for all the catalysts tested. The preexponential factors (kapp,0) of the Arrhenius equation (for this selected value of Eapp), calculated for all catalysts, are shown Table 2. For the same Eapp value, kapp,0 is a good index for comparing catalytic activities. Table 2 gives thus the ranking found in this work for the catalysts, according to their activities for tar elimination. Values of Eapp and kapp,0 obtained in this work for the most active catalyst, nickel D, are compared in Table 3 with other ones previously obtained for the same overall reaction (tar removal) on different catalysts in gasifica-

catalyst

Eapp (kJ/mol)

k′app,0 [m3 (Tb, wet)/(kg h)]

BASF G1-25-S nickel E nickel D nickel B nickel A

62 ( 18 62 ( 18 62 ( 18 62 ( 20 30 ( 18

456 000 342 000 334 000 300 000 101 000

tion with air. The lowest value for Eapp found in this work (58 kJ/mol compared to 72 and 97 kJ/mol; Table 3) would indicate that nickel D catalyst is more active than the BASF G-25S catalyst used by Narvaéz et al. (and than a calcined dolomite, of course). Comparison of catalytic activities for tar removal is presented in Figure 9 in another way. kapp,tar is given there at different values of τ and GR. Figure 9 shows the differences between catalysts and how nickel D is the most active in our process. Catalytic Activity for the Simultaneous CH4 Removal For some applications of the produced gas, such as its application in fuel cells, methane and other light hydrocarbons present in the flue gas should be also removed from the flue gas. Methane conversions on these catalysts were shown in part 1 of this work (Caballero et al., 1997). Using now a first-order (for methane) kinetic model similar to the one given by eq 1, a kinetic constant for methane removal (kCH4) can also be calculated by

kCH4 ) [-ln(1 - XCH4)]/τ

(4)

kCH4 values obtained in this work for different catalysts and with eq 4 are shown in Table 4. These kCH4


Figure 10. Arrhenius plot for the methane removal reaction (at t ) 0) over the commercial steam-reforming catalysts used in this work [s, nickel D; - -, nickel B; -‚‚-, nickel A; ‚‚‚, UC].

Figure 11. Effect of the time-on-stream on the apparent kinetic constant for tar elimination over the nickel A, B, and D catalysts at high temperature (780-830 °C) [0.019 < τ < 0.037 (kg h)/m3 (Tb, wet)].

values are plotted in Figure 10 according to the Arrhenius equation. Some lines are drawn in this figure for some catalysts. After a careful and similar analysis to the tar removal reaction, an apparent energy of activation (E′app) for methane removal of 62 kJ/mol was selected for all catalysts. Using now the Arrhenius equation, the preexponential factors were found. They are shown in Table 5. This table shows the ranking of catalysts for methane removal.

Catalyst Life As was explained before, several tests were made for the same catalyst in several days to get data on the catalysts activity vs time-on-stream. So, tests of 45 h-on-stream under the stationary state were made for the most active (regarding tar removal) catalysts. The activity of some catalysts for tar removal, given by their kapp,tar , is shown in Figures 11 and 12 at two different temperature levels. At relatively high (780-830 °C)


Figure 12. Effect of the time-on-stream on the apparent kinetic constant for tar elimination over the nickel A, B, D, and E catalysts at low temperature (670-770 °C) [0.019 < τ < 0.037 (kg h)/m3 (Tb, wet)].

Figure 13. Methane conversion vs time-on-stream for different catalysts at a relatively high temperature (780-830 °C) in the catalytic bed.

temperature there is not a noticeable deactivation (Figure 11), but at low temperature (670-770 °C) the deactivation is important (Figure 12). This fact can be explained in the following way: at high temperatures (780-830 °C) the overall rate of coke-removal (from the catalyst surface) reactions by steam and dry (CO2) gasification is higher than the rate of the coke-forming reactions. So, no buildup of coke appears on the catalysts surface and there is not deactivation in short (45 h) periods of time. At low temperatures the coke

formation rate would be higher than the coke removal one, and a net buildup of coke would appear, generating a catalyst deactivation. Of course, to find out if the catalyst is deactivated by the few parts per million of H2S present in the flue gas, long-term tests would be required but they are, at present, out of the realm possibility of these authors. Methane conversion in the catalyst bed was simultaneously measured at different times-on-stream. This is shown in Figure 13. Some catalysts lose very soon


(in about 20 h) some activity for this reaction. So, there would be a selective deactivation: Some catalysts become deactivated for the methane-reforming reaction, but they maintain their activity for the tar-reforming reactions. This fact can be further investigated in a future work.

invoice of samples of catalysts by BASF AG (Ludwigshafen, Germany), ICI-Katalko (Billingham, Cleveland, U.K.), Haldor Topsøe A/S (Lyngby, Denmark), and United Catalyst Inc. (Louisville, KY) is also gratefully acknowledged. Nomenclature

Conclusions The most important conclusions are, in short, as follows: 1. Commercial steam reforming catalysts for naphthas are more active for tar removal (in a flue gas coming from a biomass gasifier) than the commercial steam reforming catalysts for light hydrocarbons (natural gas, methane). 2. The effectiveness of these catalysts in this process (that is to say, the activity for tar removal) is so high that no important differences appear between the four commercial catalysts for reforming of naphthas tested in this work. The differences in activity between the catalysts tested are shown in Table 2. 3. Experimental variables with effects on the tar removal conversion are temperature (of the catalytic bed), space time or space velocity, catalyst particle size, and gas atmosphere composition, which, in turn, depends on the gasifying ratio used in the upstream gasifier. 4. Using a single overall reaction with a first-order kinetic equation, an apparent activation energy of 58 kJ/mol for the overall tar removal reaction fits the experimental results. 5. Using only one lump (tar) and a single first-order kinetic equation is not a good enough model. The kinetic constant of such model depends thus on several parameters such as space time (Figure 6) and flue gas composition or gasifying ratio (Figure 7). These two facts (not correct for a kinetic constant) cannot be explained with such a model. A model with at least two sublumps (like soft and hard) for tars is needed to explain the results obtained here. 6. At “relatively high” (780-830 °C) temperatures there is no catalyst deactivation in a test of 45 h-onstream under the stationary state. Long-term (more than 200 h, at least) tests are needed to check the feasibility of these catalysts in a biomass gasification process at a commercial scale. 7. It is confirmed in this work that the tar content in the flue gas coming to the catalytic bed has to be low enough (the authors propose 2 g of tars/ mn3 as the limit) to avoid catalyst deactivation by coke. Acknowledgment This work was carried out under the JOULE III Program of the EU DG-XII, Project No. JOR3-CT950053. The authors thank the European Commission for its financial support. Work was also done under the Spanish DGES-financed Project No. PB96-0743. Discussions on the use of commercial steam reforming catalysts for this process with Dr. Martyn V. Twigg of Johnson Matthey in Royston (England), with Mr. Harald Sha¨fer and Roland Spahl of BASF AG, Ludwigshafen (Germany), and with P. E. Hojlund Nielsen of Haldor Topsøe A/S of Lyngby (Denmark) have been very fruitful for us. Tar composition analyses made by Dr. K. Sjo¨stro¨m and colleagues at KTH (Sweden) and by R. Evans at NREL, Golden, CO (USA), are recognized. The

Ctar ) tar content in the flue gas (mg/mn3) daf ) dry, ash free dp ) particle diameter of the catalyst (mm) Eapp ) apparent activation energy for tar removal (kJ/mol) E′app ) apparent activation energy for methane removal (kJ/mol) H2O/C* ) steam to carbon-to-be-reformed (CH4 + C2 + ... + tars) ratio in the flue gas at the inlet of the catalytic reactor, mol of H2O/atomic g of C GR ) gasification ratio at the gasifier inlet, defined as [(H2O + O2)/biomass daf fed], dimensionless, (kg/h)/(kg/ h) kapp,tar ) apparent kinetic constant for tar removal (m3 (Tb, wet)/(kg‚h)) kCH4 ) apparent kinetic constant for methane removal (m3 (Tb, wet) /(kg‚h)) kapp,0 ) preexponential factor for kapp,tar (m3 (Tb, wet)/(kg‚h)) k′app,0 ) preexponential factor for kCH4 (m3 (Tb, wet)/(kg‚h)) Q ) gas flow rate (m3 (Tb,wet)/h) rtar ) rate of tar elimination (mg of tar/kg of catalyst‚h) SV ) space velocity in the catalytic reactor [mn3 wet/(h‚m3)] t ) time on stream (h) Tb or Tbed ) temperature measured in the center of the catalytic bed (°C) W ) weight of catalyst (kg) XCH4 ) methane conversion (dimensionless) Xtar ) tar conversion (dimensionless) yH2O ) steam content in the flue gas (dimensionless) Greek Symbols τ ) space time, defined as W/Q [(kg h)/[m3 (Tb, wet)]]

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Received for review September 18, 1997 Revised manuscript received April 24, 1998 Accepted April 24, 1998 IE9706727

Commercial Steam Reforming Catalysts To Improve Biomass

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