Fresh Tar (from a Biomass Gasifier) Elimination over a Commercial

Ind. Eng. Chem. Res. 1997, 36, 317-327

317

Fresh Tar (from a Biomass Gasifier) Elimination over a Commercial Steam-Reforming Catalyst. Kinetics and Effect of Different Variables of Operation Ian Narva´ ez, Jose´ Corella,* and Alberto Orıó Department of Chemical Engineering, University “Complutense” of Madrid, 28040 Madrid, Spain

The upgrading of the raw gas from a biomass gasifier is studied with the commercial steamreforming BASF G1-25 S nickel-based catalyst. It is located downstream of the gasifier, a bubbling fluidized bed type in which air is used as gasifying agent. To increase the catalyst lifetime, a guard bed of a calcined dolomite at 800-850 °C is used. It decreases the throughput of tar entering the catalytic bed to amounts below 2 g tar/m3(NC). This work is focused only on the catalytic bed which easily decreases the tar content in the gas to only 1-2 mg/m3(NC). Variables studied include the particle diameter of the catalyst, time-on-stream, temperature of the catalytic bed, and gas and tar compositions. Both tar and gas compositions in the catalytic (Ni) reactor depend on the equivalence and H/C ratios existing in the gasifier and on the operating conditions of the guard bed of dolomite. A simple kinetic model is used to describe the overall tar elimination network. Its overall kinetic constant is used as index of the catalyst activity for tar elimination. Values of this overall kinetic constant are given for very different operating conditions. Introduction Gasification of biomass with air produces a dirty raw gas composed of a mixture of H2, CO, CO2, H2O, CH4, and light hydrocarbons, with some dust, tars, NH3, and other trace components. For most applications, this gas has to be cleaned of at least dust and tars. If the gas is going to be used for electrical production, burning it in gas turbines, for instance, to obtain high efficiencies the gas has to be cleaned at high temperatures, close to the ones at the gasifier exit (typically around 800 °C). This hot gas cleaning can be made by using filters (for dust) and downstream catalytic reactors (for tar). Two main lines and/or processes are nowadays emerging for these downstream reactors. One is the use of calcined dolomites or limestones, cheap materials, which are active for tar elimination at temperatures above 800 °C (i.e. Taralas et al., 1991; Corella et al., 1991a). The other one is the use of hydrocracking (Pedersen, 1994) or steam-reforming catalysts. The steam-reforming ones, which are nickel based, are the most studied in the world. They seem to be the most promising ones for future commercial applications, and this paper will focus only on these catalysts, while the use of dolomites and related materials for tar elimination is presented in other papers concurrent with this one (Delgado et al., 1996, 1997; Orıó et al., 1996). To our knowledge, these Ni catalysts were firstly placed in the same gasifier bed (Mudge et al., 1985; Rei et al., 1986). A process development pilot plant based on this concept was set up in the early eighties at the PNL of Richland, WA, but it was soon stopped and dismantled because of low catalyst lifetime. When the nickel catalyst is put in the same gasifier bed, it is surrounded by char which deactivates it, in addition to its simultaneous erosion if the gasifier bed is fluidized. So, using the catalyst in the same gasifier bed does not seem to be a useful idea. The use of the catalyst in a downstream vessel, in a two-stage process, then began to be studied by several institutions. For us, the pioneer institution in this area was Battelle’s Pacific Northwest Lab (PNL) in which Mudge, Baker, and co-workers carried out extensive work S0888-5885(96)00235-7 CCC: $14.00

(Mudge et al., 1985; Elliot and Baker, 1986; Baker et al., 1985, 1987; Mudge et al., 1988), which ended in the late eighties. Together with the work being carried out nowadays in the United States (Gebhard et al., 1994; Paisley and Gebhard, 1995; Kinoshita et al., 1995), the main effort in RTD in this area is being carried out in Europe. At least two groups in two European countries, Finland and Spain, are making extensive studies in catalytic hot gas cleaning and conditioning from biomass gasifiers. In the Technical Research Center of Finland (VTT), E. Kurkela, P. Simell, and co-workers have published several papers in the early nineties about the catalytic decomposition of tars from a fluidized bed biomass gasifier (Simell and Bredenberg, 1990; Leppala¨hti et al., 1991; Simell et al., 1992, 1994; Simell, 1996). Their valuable work is nevertheless obscured by the confidentiality that accompanies it. Due to the fact that the process is being developed for several commercial companies, several important details about the catalyst are omitted in such publications and remain thus to be known and published. Corella, Aznar, and co-workers started in the late eighties in Spain to study commercial nickel catalysts placed in downstream reactors (Corella et al., 1991a,b; Aznar, 1992, 1993). They used a small pilot plant facility similar the one used at PNL. Their work was designed for cleaning of hot gas coming from a biomass gasifier with steam. Such a gasifying agent, steam, will be the first important difference with respect to the present work in which the gasifying agent is air with some steam content. Their work could be summarized in two main conclusions: first, using downstream commercial nickel catalysts, tar conversions (eliminations) higher than 99.9% were easily obtained with very low space times (space velocities higher than 10 000 h-1); second, if the throughput of tars to the secondary reactor is high, that is to say, tar contents in the flue gas is higher than about 2-5 g/m3(NC), the catalyst deactivation is fast and the catalyst lifetime is low. The overall “catalytic gasification” process then becomes not feasible. Note that the Swedish MINO process also failed in the eighties by the catalyst deactivation. It was © 1997 American Chemical Society

318 Ind. Eng. Chem. Res., Vol. 36, No. 2, 1997

attributed then to the effect of the sulfur present in such flue gas (Blackader and Rensfelt, 1984). To avoid a fast catalyst deactivation, by coke or by sulfur, is the key aspect in this catalytic gasification process. So, Corella et al. decided to use a guard bed before the catalytic one. The present work has thus been done using a guard bed, with a calcined dolomite, before that of nickel catalyst. This fact is the second main difference with respect to our previous work (Corella et al., 1995a). A three-stage gasification process is now used instead of the former two-stage one. This new process is more complex and expensive, but the improvement for the biomass gasification process can compensate its higher cost. Its economic feasibly is, of course, important, but it is out side of the scope of this work. Equipment Description The facility used in this work has three reactors connected in series, and it has been already described in a previous work (Narvaéz et al., 1996). In short, the first reactor is a biomass gasifier, bubbling fluidized bed type, of 6.0 cm i.d. and 60 cm height in its bottom part. It is followed by a freeboard of 12 cm i.d. and 20 cm height. It is externally heated by two ovens automatically controlled. The temperature is measured in the fluidized bed and in its freeboard. The fluidized bed has a stationary or fluidizing medium of 500 g of silica sand of -0.50 + 0.32 mm. The fluidizing and gasifying agent is air. Its flow rate is previously fixed in each run to get the desired value for the equivalence ratio (ER, air fed/stoichiometric air needed for total combustion). The u1/umf and ER values used in the gasifier were in the range of 1.1-4.0 and 0.20-0.45, respectively. Biomass was continuously fed by three screws to the bottom part of the gasifier, 4 cm above the gas distributor bubble cup. The correct operation of this feeding device has been a key factor in this experimentation. The third or last screw has a fast speed. It is externally cooled to avoid biomass pyrolysis in it which would produce tars and would oblige the stoppage of the feeding and the experiment. The effects of the type and shape of the feedstock on the feeding device required and on the product distribution obtained were previously studied (Herguido et al., 1992). No important differences on the product distribution were obtained with five different types of biomass. Given the fact, besides, of the small (5 cm i.d.) diameter of the biomass feeding pipe in this facility, we decided to use as biomass a representative and easyto-feed feedstock. Pine (pinus pinaster) sawdust between 4.0 and 0.8 mm was selected thus as the biomass to be gasified in this work. Its main characterization is proximate analysis (wt %): volatiles 74-76 fixed carbon 12-13 ash 0.5-1.2 moisture 8.5-16.5 elemental analysis (dry, ash free, wt %): carbon 50.0 hydrogen 5.7 oxygen 44.1 nitrogen 0.1-0.3 low heating value 18.0-18.4 MJ/kg daf

The moisture of the pine sawdust was carefully measured and varied from one experiment to another

Table 1. Some Examples of the Operating Conditions of the Catalytic Bed run number time on stream, h ER H/C O/C

17 5.0 0.26 1.6 1.2

18 4.66 0.52 1.6 1.7

20 4.32 0.35 1.6 1.4

21 3.0 0.44 2.1 1.8

Secondary, Guard, or Tar-Cracking Bed of Calcined Dolomite calcined dolomite Seville Seville Seville Seville T2,c, °C 720 800 850 870 exit gas composition, dry vol % (inlet catalytic (Ni) reactor) H2 12.5 12.5 12.0 18.0 CO 15.0 15.0 13.0 13.0 CO2 15.0 15.5 12.4 20.0 CH4 2.8 2.8 2.5 2.5 C2H4 1.1 0.0 1.0 1.1 N2 53.6 54.2 59.1 56.6 Ctar,2, mg/nm3 1347 2140 585 1000 H2O 7.5 8.1 7.0 10.8 Steam-Reforming Bed: G1-25-S Catalyst catalyst G1-25-S G1-25-S G1-25-S W, g 200 200 200 T3,c, °C 650 650 720 L3, cm 11.0 11.0 11.0 τ3, s 0.09 0.09 0.12 τ3′, kg/m3(NC)/h 0.14 0.14 0.19 GHSV3(NC), h-1 10 000 10 000 10 300 dp mm (-5 + 2) (-5 + 2) (-3 + 1) umf, cm/s n.d. n.d. 130 u3, cm/s 104 104 114 exit gas composition, dry vol % H2 18.0 13.0 14.0 CO 17.5 14.5 14.0 CO2 12.0 11.5 11.0 CH4 2.8 2.4 0.7 C2H4 0.6 0.0 0.0 N2 57.1 58.6 51.3 Ctar,3, mg/m3(NC) 70 70 40 LHV3, mJ/m3(NC) 4.7 4.2 4.7 Ytar, g/kg daf fuel 0.05 0.11 0.04 Ygas, nm3/kg daf fuel 2.1 3.8 2.6 yH2Oe, vol % 6.9 6.8 9.1 X tar, % 94.8 96.7 88.0

G1-25-S 150 720 8.0 0.09 0.16 11 700 (-5 + 3) >200 95 18.0 13.0 16.0 1.9 0.0 51.1 30 4.3 0.02 3.0 8.8 97.0

one by addition of some water to the sawdust. There is some steam thus mixed with the air in the gasifying agent. This steam comes from the moisture of both biomass and air fed to the gasifier. The steam in the gasifier was thus another variable which was able to be varied, although in a short interval. The H/C and O/C ratios at the gasifier inlet depend on the flows of both biomass and air (together with their moistures) fed to the gasifier. Some examples of these ratios are shown in Table 1, together with some other main operating conditions, like steam content in the flue gas, for some representative experiments. After the gasifier there is an hot (500-600 °C) metallic filter and a bed of a calcined dolomite. This bed acts as a guard bed with respect to the downstream catalytic one. Some tar (typically around 90%) is eliminated or destroyed in this bed of dolomite. The tar content in the flue gas and the tar throughput sent to the catalytic (nickel) reactor is decreased with this guard bed to below 2 g tar/m3(NC), a value from which there could be catalyst deactivation by coke formation (from tar). The composition of the tar also changes in this bed of dolomite. Most of the remaining tar is naphthalene and PAHs. Several types of dolomites and related materials have been already tested under very different experimental conditions. The operating conditions of the fluidized bed gasifier were also varied. The effects of the operating conditions in the gasifier and in the secondary (or guard) bed of dolomite on the product

Ind. Eng. Chem. Res., Vol. 36, No. 2, 1997 319

Figure 1. Tar and gas sampling system.

distribution have been published in detail in other papers simultaneous to this one (Narvaéz et al., 1996; Orıó et al., 1996, 1997). So, no further details will be given here about these beds. The product distribution, including gas composition and tar content in the gas, at the exit of the gasifier and of the secondary or guard (dolomite) bed, are well-known in all experiments. That is to say, the gas composition and tar content at the inlet of the catalytic (nickel) bed are always known. Tar and Gas Sampling and Analysis Samples of gas and of condensates were taken at different times-on-stream in three different points in the installation. Two of them were located before and after the catalytic reactor. A detail of the sampling system is shown in Figure 1. The tar-sampling system was very similar to the one used in VTT (Leppa¨lahti et al., 1991). It is composed of an electrically heated pipe, to avoid condensation and plugging in it, a condenser, and four cold traps: the first two were in an ice bath and the last two in an acetone-CO2 ice bath. The volume of the gas passing through these cold traps to get one sample of condensate was of about 30 dm3 (gas flow rate of 0.060 m3(NC)/h during 30 min). The condensate so obtained was analyzed for their tar content by a total organic carbon analyzer, DC-90 Dohrmann model. An example of the tar content in the gas at the exit of the gasifier, guard bed, and catalytic (Ni) bed is shown in Figure 2. Note the clear and important decrease in this tar content in both guard and catalytic beds and how the gas from the catalytic bed is very clean (a few tens of mg/m3(NC)). The gas was sampled and analyzed in two different ways: on-line and off-line. On-line passing of the gas through an Rosemount apparatus enabled continuous measurement of H2, CO, and light hydrocarbons. The continuous measurement of the CO allowed instanta-

Figure 2. Example of the tar content in the gas at the exit of the three different reactors. Run no. 27. T1,c ) 800 °C; ER ) 0.32; H/C ) 2.1. Second bed: “Norte” calcined dolomite; T2,c ) 870 °C. Third bed: BASF G1-25 S catalyst; T3,c ) 800 °C.

neous detection of failures or problems in the biomass feeding system, like bridging in the hopper, which was solved by knocking it. The off-gas analyses was made by taking samples, as Figure 1 shows, at different times on stream and analyzing them in a HP gas chromatograph model 6890 equipped with Molecular Sieve 13X and Porapak Q columns. The on-line and off-line gas analyses gave similar results with only a maximum difference of two points in the volume percent. The gas composition presented here will be that obtained by gas chromatography. Two examples of gas composition before and after the catalytic reactor, corresponding to experiments 23 and 24, are shown in Figures 3 and 4. From these results one can observe how the H2 and CO contents in the flue gas increases in the catalytic (Ni) bed and how the CH4, C2H4, and CO2 decrease. It means a reacting network like (not adjusted):


Figure 3. Gas composition of the gas at the exit of the three different reactors. Run no. 23. T1,c ) 800 °C; ER ) 0.44; H/C ) 2.1. “Malaga” calcined dolomite; T2,c ) 845 °C. BASF G1-25 S catalyst; T3,c ) 750 °C.

{ } k

H2O 98

tars +

∑CnHm +

k′

CO2 98 CO + H2 + CH4 + ... ...

k′′

98

(1)

The exact amounts in which these components increase or decrease depend on the operating conditions of the catalytic bed. This variation in the gas composition through the catalytic bed is important to elaborate a tentative mechanism for this complex reacting network. The gas composition in the catalytic bed is also important for another purpose: the catalyst activity in this process is handled and compared by an apparent

kinetic constant (kapp) for tar elimination which in turn depends on the gas composition (Corella et al., 1995b, 1996). To correlate the values obtained for kapp with the gas composition, Figures 3 and 4 and corresponding ones for each experiment are very important. Besides the values shown in Figures 3 and 4 and similar ones, the steam content in the flue gas in the catalytic reactor has been also calculated because the steam is an important reactant in the catalytic removal of tar (Bonneau et al., 1991; Kuprez, 1992). The steam content in the gas is also important to compare the values for kapp obtained in this work of gasification with air with the ones simultaneously obtained by Aznar and co-workers (Aznar et al., 1996) for gasification with pure steam and with (steam + O2) mixtures.


Figure 4. Gas composition of the gas at the exit of the three different reactors. Run no. 24. T1,c ) 810 °C; ER ) 0.34; H/C ) 2.2. “Malaga” calcined dolomite; T2,c ) 820 °C. BASF G1-25 S catalyst; T3,c ) 800 °C.

Catalytic Bed and Catalyst Used The catalytic bed or third reactor has an internal diameter of 4.0 cm and an upward gas flow. The gas flow rate at its inlet, Q3,0, is mainly fixed by the operating conditions of the gasifier. On the other hand, the commercial catalyst used here was designed for fixed bed operation, and its fluidization would erode and deactivate it. Experiments were carried out thus in fixed bed conditions which implies that u3,0 < umf,3. To achieve this condition, and since u3,0 is fixed (by the gasifier conditions, diameter of the catalytic bed, and its temperature), umf for the catalyst has to be high enough. This fact forced use of the catalyst with particle diameters higher than 1.0 mm.

The catalytic reactor was externally heated by an electric oven. The temperature was measured in it with two thermocouples, in the center or axis of the bed and at the wall (inside). These thermocouples were moved along the bed height to obtain the profiles of temperature in it. The bed height was varied from one experiment to another one to get different space times, but it always was between 6 and 12 cm. To get the desired temperature level in the bed of the catalyst, a previous bed of an inert material was used. This inert was silica, small stones of about 5 mm diameter, previously calcined in an oven at 1000 °C. Typical profiles of temperatures in the center of the bed are shown in Figures 5 and 6. Experts in catalytic steam-reforming processes like Rostrup-Nielsen and


Figure 5. Example of the temperature profile in the axis of the catalytic reactor. Run no. 37.

Figure 7. Adsorption-desorption isotherm for the BASF G1-25 S catalyst.

Figure 6. Example of the temperature profile in the axis of the catalytic reactor. Run no. 22.

Figure 8. Pore volume distribution (by mercury porosimetry) for the BASF G1-25 S catalyst.

Twigg have indicated how these temperature profiles are of basic importance in the reforming of natural gas and of naphtas (Rostrup-Nielsen, 1984; Twigg, 1989). These profiles could also be very important in this new process of catalytic steam reforming of tars. So, they were always measured. As our fixed bed catalytic reactor is thus not isothermal, the temperature of reference for the bed of catalyst (T3,c) shown in figures and tables in this paper will be the one measured in its axis, at the middle of its height. Other important operating conditions for the catalytic bed are shown in Table 1 for some representative experiments. Notice in Table 1 how (i) the steam content in the catalytic bed is between 7 and 11 vol %, and (ii) the H2O/C (from tars) ratio is more than 2.2 (limit in commercial steam reformers to avoid coke formation on the catalyst surface). The catalyst used was a BASF commercial one used for steam reforming at high temperatures of natural gas. Its denomination is G1-25S. It contains 15 wt % NiO over R-Al2O3 (BASF, 1994). Its adsorption isotherm is shown in Figure 7 indicating this catalyst is not porous and has quite a low BET surface area (2.0 m2/ g). Its total (up to 10 µm) pore volume is 0.20 cm3/g, and its pore distribution is shown in Figure 8 with pore diameters in the range of 100-2000 nm (1000 Å to 2 mm). It is received in big rings of about 20 × 20 mm and crushed and sieved until the desired particle size was obtained before its use. Once put in the catalytic bed, it was always activated in situ before its use by

reducing it in a flow of a mixture of hydrogen (20%) and nitrogen at 820 °C for 4 h. Kinetic Model Used The activity of the catalyst for “tar removal” under very different conditions will be here expressed by an apparent kinetic constant for such reaction, kapp. Corella and co-workers have already presented (Corella et al., 1996) a tentative reacting network for the catalytic “tar removal” from the flue gas which includes reactions of steam and dry (CO2) reforming of the tar. Our model uses the following and simplest kinetic equation n -rtar ) kappCtar

(2)

m′ m′ + k′yCO + k′′ + ... kapp ) kyH 2O 2

(3)

with n ) 1, and

But notice that

k′, k′′, k′′′ ) f(tar reactivity in the third or catalytic bed) ) f ′(tar composition in the third bed) (4) and


(tar composition... (and yH2O, yCO2, ...) in the third bed) ) f ′′(experimental conditions in the gasifier, in the guard bed and other ones not known yet) (5) This fact means that for a given catalyst and operating conditions in the third or catalytic bed

kapp ) f ′′′(experimental conditions in the gasifier and in the guard bed) (6) Exact composition of the fresh or realistic tar existing in the catalytic bed is very difficult to be known. Samples of tars in the condensates before and after the catalytic bed have been kindly analyzed for us by TPS AB and KTH in Stockholm and by NREL in Golden, CO. Nevertheless we cannot say yet the exact composition of the fresh tar in the catalytic reactor because fresh tar is not stable. When cooling, condensing, trapping, and analyzing it after several days of its sampling, it has polymerized and its composition has changed quite a bit. k, k′, k′′, ... depend on the realistic fresh tar composition existing in the catalytic reactor and can not be related with composition of the aged or “old” tar analyzed. The fresh tar composition is not well-known for us yet. So, our approach says that the fresh tar composition in the catalytic bed depends on how it was produced in the gasifier and on what happened with it in the guard bed. Tar reactivity in the catalytic bed could be thus related with the product distribution at the gasifier exit which, in turn, depends on the equivalence (ER), H/C, and O/C ratios there used (Narvaéz et al., 1996). Besides, the guard bed eliminates about 90% of the tars (the most reactive ones) and modifies the tar composition. So, the experimental conditions in the guard bed affect the values of k, k′, k′′, ... and therefore of kapp. Such experimental conditions will have always to be stated (said or known). The fresh and reactant tar being a complex lump, its kinetics of disappearance can follow an order higher than one, as several authors have demonstrated for situations similar to this one (i.e. Corella et al., 1985; Astarita and Sandler, 1991). The assumption of n ) 1 in eq 2 cannot be true and might thus be revised in the future. Nevertheless, n ) 1 has been selected in this work for the following two reasons: (i) Most institutions working in Europe in catalytic hot gas cleaning from biomass gasifiers have agreed upon using n ) 1 to further compare their results under the same basis (Corella, 1995c). (ii) The tar coming to the third or catalytic (Ni) reactor is a lump much less complex than the one at the gasifier exit. The guard (dolomite) bed eliminates about 90% of the tar and “simplifies” its composition. At the exit of the bed of dolomite (inlet of the nickel bed), most tar is composed of naphtalene. So, the tar lump in the nickel bed is not as complex as at the gasifier exit. Working under conditions very close to plug flow, a mass balance for tar and eq 2 (with n ) 1) gives

kapp ) -ln(1 - Xtar)/τ′ ) SV[-ln(1 - Xtar)]

(7)

kapp so calculated will be the index used in this paper to describe the catalyst activity for tar removal. Effect of the Particle Diameter of the Catalyst To know if the internal diffusion plays a role in this reaction, several experiments were made under the

Figure 9. Effect of the particle diameter on the apparent kinetic constant for tar* elimination. BASF G1-25 S catalyst (*tar obtained at ER ) 0.28-0.34, H/C ) 1.8-2.1, T2,c ) 850 °C).

Figure 10. Thiele modulus vs particle diameter of the BASF G125 S catalyst for the overall tar* elimination reaction. (*tar obtained at ER ) 0.28-0.34, H/C ) 1.8-2.1, T2,c ) 850 °C).

same experimental conditions varying the particle diameter (dp) of the catalyst. The commercial catalyst was crushed and sieved until different size intervals were obtained. This method is questionable because some particles so obtained can have more nickel than other ones but (i) we did not have another method for getting particles with different dp, and (ii) due to the relatively high amount of catalyst (100-300 g) used in each experiment, the differences in nickel content between particles, statistically, would not be very important. The study was carried out at two levels of temperature in the third reactor, 750 and 800 °C, with particle sizes between 1.3 (-1.6 + 1.0) and 3.6 (-4.0 + 3.2) mm. kapp was calculated with eq 7 for all experiments with different particle diameters. The results are presented in Figure 9. These results indicate how dp has an important influence on this process under the said experimental conditions. The Thiele modulus (φs) and the effectiveness factor were now calculated with well-established methods (Satterfield, 1970). The intrinsic value of kapp, without internal diffusion, does not need to be known because (i) φs can be calculated having different values of kapp at different dp and (ii) extrapolation to dp ) 0 is difficult in Figure 9. The Thiele modulus calculated for several particle diameters, at two temperatures in the catalytic bed, are shown in Figure 10. The φs-dp relationship for a calcined dolomite at 850 °C (Delgado et al., 1997),


Figure 11. Effectiveness factor vs Thiele modulus for the tar* elimination reaction over the BASF G1-25 S catalyst. (*tar obtained at ER ) 0.28-0.34, H/C ) 1.8-2.1, T2,c ) 850 °C).

for the same reaction and process, is also shown in Figure 10 for comparison purposes. Note that the line for dolomite in Figure 10 is close to the ones for the nickel catalyst because it was obtained at a higher temperature (850 °C) than the ones used for the Ni catalyst (750 and 800 °C) and, of course, reactivity (or intrinsic kinetic constant, involved in φs) increases with temperature. The calculated effectiveness factor (η) is shown in Figure 11. From this figure the following important conclusions emerge: (1) Our work is made under some internal diffusion control (near the transition region). The effectiveness factors are not very low, but they are clearly less than 1.0. (2) To get η ) 1.0, in this reaction and over this catalyst, dp < 0.3-0.4 mm should be used, as it is deduced from Figures 10 and 11. (3) Using this catalyst with its original commercial size and shape, η would be around 0.1, a very low value. It would indicate that for a future commercial application a new shape and size should be used: a small size one (which would imply fluidized bed operation, which in turn would imply a new and “harder” catalyst), or a monolith or honeycomb. Effect of the Time-on-Stream The biomass gasification process with catalytic hot gas cleaning here studied can be feasible at commercial scale if the catalyst has a long life. The previous and published data without using a guard bed indicated that the catalyst deactivated very soon. Using now a guard bed, tar conversion in the catalytic (Ni) bed was measured at different times-on-stream. kapp was then calculated, with eq 7, at different times-on-stream. Results at two temperature levels are shown in Figure 12. At low temperatures (650-750 °C) some deactivation appears after 2-3 h on stream, but at high (g800 °C) temperature no deactivation is detected in 4-11 h. Of course, these times-on-stream are low to be used or extrapolated to a future commercial operation. Long term runs under the best experimental conditions found here will be made in the future. When results shown in Figure 12 are compared with similar ones previously obtained without a guard bed (Aznar et al., 1993), one can see how now the catalyst life has increased a lot. There have been two clear

Figure 12. Effect of the time-on-stream on the apparent kinetic constant for tar* elimination over the BASF G1-25 S catalyst. (*tar obtained at ER ) 0.32, H/C ) 2.1).

Figure 13. Comparison of the tar conversion obtained by several authors (using nickel catalysts) at different operating temperatures.

improvements: (i) to work (in the nickel bed) above 800 °C, instead of 680-750 °C, as before and (ii) to use a guard bed. Now (in this work) the throughput or loading of tar to the catalytic (Ni) bed is below 2-3 g tar/m3(NC) instead of 20-50 g tar/m3(NC) as happened previously (Aznar et al., 1993). With these two improvements, catalyst life has clearly increased a lot. Notice that the low throughput of tar to the catalytic (Ni) bed could also be achieved, instead of using a guard bed, by a good design and operation of the gasifier, which would include in-bed use of dolomite, secondary air, and high temperatures in the freeboard. Effect of the Temperature on the Catalytic Bed The temperature of the catalytic bed, T3, has a clear effect on the catalyst activity and life, as it has been just shown. This variable has been thus studied in detail. There is some discussion about the best temperature interval to be used. The tar conversion published by several authors, using different space velocities, is compared in Figure 13. Excepting researchers from VTT, which use up to 900 °C, most people work below 820 °C. Several experiments were made, using the BASF G125S catalyst with dp ) -1.6 + 1.0 mm, at different temperatures. kapp values for tar removal, calculated with eq 7, are shown at different temperatures in Figure


Figure 14. Effect of temperature of the catalytic bed on the apparent kinetic constant for tar* elimination over the BASF G125 S catalyst. (*tar obtained at ER ) 0.28-0.30, H/C ) 1.8-2.1, with the guard bed at T2,c ) 850 °C).

Figure 15. Arhenius representation for the k3,app values over BASF G1-25 S catalyst. (Tar obtained at ER ) 0.28-0.30, H/C ) 1.8-2.1.) Table 2. Values for the Apparent Activation Energy and Preexponential Factor for the Overall Tar Removal Reaction this work

over a calcined dolomite

tar tar generated tar generated generated in a gasifier with in a gasifier with in a gasifier steam (Delgado air (Orıó with air et al., 1997) et al., 1996, 1997) Eapp, kJ/mol kapp,0, m3(NC), dry/(kg h)

72 ( 12a 143 000a

84 ( 6 2600 ( 700b

97 ( 14 (1.2-1.4) × 106 c

a Under some internal diffusion control. b Under big internal diffusion control. c m3, wet/(kg h).

14. The temperature used in this Figure, T3,c, and in next ones is a “reference temperature” as was previously indicated. The values of kapp are represented in Figure 15 according to the Arrhenius law. The values of Eapp and kapp,0 obtained from this Figure are shown in Table 2. Note the high value obtained for the preexponential factor, indicating an high catalytic activity. Notice also how these values, according to Figures 9-11, were obtained under some internal diffusion control. The true value for Eapp would be somewhat higher that the one indicated in Table 2. The values for Eapp and kapp,0 obtained over a calcined dolomite for the removal of tar

Figure 16. Effect of T3,c on the concentration of CO, CO2, and light hydrocarbons at the exit of the catalytic reactor. BASF G125 S catalyst; dp ) -1.6 + 1.0 mm (gasifier: ER ) 0.32, H/C ) 2.1).

Figure 17. Effect of the equivalence ratio on the apparent kinetic constant for tar* removal over theBASF G1-25 S catalyst (*tar passed through a guard bed of dolomite at T2,c ) 850 °C).

generated in biomass gasification with steam (Delgado et al., 1997) and with air (Orıó et al., 1996, 1997) are also indicated in Table 2. The temperature of the catalytic reactor not only has influence in the tar removal reaction but also in all the reactions existing in the reacting network. As an overall effect, the gas composition at the exit of the catalytic reactor changes with T3,c as Figure 16 shows. The CO2 content in the exit gas decreases with T3,c indicating that CO2 (and not only the steam) also reacts with the tar present in the flue gas. Note that the amount of light hydrocarbons in the gas remains nearly unchanged. This fact is an index that the catalyst reforms (destroys) better the tars than the small molecules of the light hydrocarbons present in the flue gas. Effect of the Tar and Gas Compositions on the Catalytic Activity for Tar Removal When kapp (in the catalytic (Ni) bed) is related to the equivalence ratio (ER) in the gasifier, a relationship appears, as Figure 17 shows. It confirms the approach given by eqs 3-6. In a previous paper (Narvaéz et al., 1996) the important effect of ER on the product distribution at the gasifier exit (in the same experimental facility) was shown. It was there demonstrated how ER determines the composition of the gas and tar produced


Figure 18. Effect of the H/C ratio on the apparent kinetic constant at two different temperatures. BASF G1-25 S catalyst (ER ) 0.28-0.34; T2,c ) 850 °C).

in the gasifier. It is thus easy to understand now how ER affects the kapp value in the downstream reactor, as Figure 17 clearly shows. In further comparisons of the kapp values, the ER value will thus always have to be indicated or known. The steam present in the flue gas is an important reactant in the tar removal network, eq 1. Its content in the flue gas depends in this process on the moisture of the biomass and air fed to the gasifier, values known here. This steam could be indicated as the (H2O/C) ratio fed to the gasifier (there are not other feeding points in this work), but we preferred here to handle it as the H/C ratio which is calculated in each experiment. Some examples of the H/C value are indicated in Table 1. When representing kapp in the catalytic (Ni) reactor against the H/C ratio, Figure 18, it is observed how there is also a relationship between these two parameters. kapp increases with H/C, which can be easily understood. When further kapp values are compared, H/C will have also to be indicated. Our final conclusion is thus that both equivalence and H/C ratios have an important influence on the activity (kapp) of the downstream catalyst for the tar removal reaction. Acknowledgment This work has been carried out under the EU, DGXII, Agroindustry Programme Project No. AIR2-CT931436. The authors thank the European Commission for its financial support. Discussions with Dr. Martyn Twigg on the use of commercial steam-reforming catalyst for this new process have been very fruitful for us. Tar composition analysis made at KTH and TPS AB, Sweden, and by NREL, Golden, CO, and the invoice of samples of catalysts by BASF GmbH, Ludwigshafen, are also recognized. Thanks are also given to Dr. M. P. Aznar of University of Faragossa for her continuous assistance and help and to Mr. Jose Prieto (Pepo) of the JCP-CSIC (Madrid) for his advanced control system setup in the pilot plant used here. Nomenclature Ctar ) tar concentration (mg/m3(NC)) dp ) particle diameter of the catalyst (mm) Eapp ) apparent activation energy for the overall reaction of tar elimination (kJ/mol) ER ) equivalence ratio at the gasifier inlet

GHSV ) gas hourly space velocity in the catalytic reactor, defined as Q3,e /V3 (h-1) H/C ) hydrogen to carbon ratio at the gasifier inlet, adim kapp ) apparent kinetic constant for tar conversion over the catalyst (m3(NC)/kg h) kapp,0 ) preexponential factor in the Arrhenius eq (m3(NC)/ (kg h)) k, k′, k′′ ) kinetic constants for reactions involved in the overall tar elimination network, given by eq 1 (m3(NC)/ (kg h)) L ) height of the catalyst bed (cm) LHV ) low heating value of the gas (mJ/m3(NC)) mB,0 ) mass flow rate of biomass fed to the gasifier (kg daf/h) m1,0 ) mass flow rate of air to the gasifier (kg/h) m2,0 ) crude gas mass flow rate to the secondary reactor (kg/h) O/C ) oxygen to carbon ratio at the gasifier inlet, adim P ) absolute pressure (bar) Q ) gas flow rate (nm3/h) rtar ) rate of tar removal (mgtar/(kg h)) SV ) space velocity in the catalytic reactor (nm3/(h kg)) T ) temperature (°C) umf ) minimum fluidization velocity (cm/s) u ) superficial gas velocity, at Ti and Pi (cm/s) V ) volume of the bed of catalyst (m3) Xtar ) tar conversion (adim) yH2O, yCO2 ) content of steam and CO2 in the flue gas (adim) W ) weight of catalyst (kg) Ygas3 ) gas yield from the catalytic reactor (m3(NC)/kg biomass daf) Ytar3 ) tar yield from the catalytic reactor (kg/kg biomass daf) Greek Symbols τ ) space time in the catalytic reactor (s) τ′ ) space time, defined as W/Q (kg catalyst h/m3(NC)) η ) effectiveness factor φs ) Thiele modulus for spheres, defined as (dp/6) (kapp,intrinsicF/De)1/2 (adim) Subscripts 1 ) gasifier 2 ) second or guard bed, bed of dolomite 3 ) third or catalytic reactor i ) first, second or third reactor o ) inlet e ) exit c ) bed center or axis w ) wall

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Received for review April 24, 1996 Revised manuscript received October 24, 1996 Accepted November 8, 1996X IE960235C

X Abstract published in Advance ACS Abstracts, January 1, 1997.

Fresh Tar (from a Biomass Gasifier) Elimination over a Commercial

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