2252
Ind. Eng. Chem. Res. 1991,30,2252-2262
Glockner, P. W.; Keim, W.; Mason, R. F.; Bauer, S. R. Catalytic Oligomerization of Ethylene. German Patent 2,053,758,1971. Hirai, H.; Hiraki, K.; Noguchi, I.; Makishima, S. Electron Spin Resonance Study of Homogeneous Catalysts derived from n-Butyl Titanate and Triethyl Aluminum. J. Polym. Sci., Part A 1970, 8,147. Keim, W. Nickel Hydrides: Catalysis in Oligomerization and Polymerization Reactions of Olefins. Ann. N.Y. Acad. Sci. 1983,415, 191. Keim, W.; Peuckert, M. A New Nickel Complex for the Oligomerization of Ethylene. Organometallics 1983,2,594. Keim, W.; Behr, A.; Roper, M. Alkene and Alkyne Oligomerization, Cooligomerization and Telomerization Reactions. In Comprehensiue Organometallic Chemistry; Wilkinson, G., Stone, F. G. A,, Abel, E. W.; Eds.; Pergamon: Oxford, 1982;Vol. 8,Chapter 52,p 371. Kissin, Y. V. Oligomerization of Ethylene with a Homogeneous Sulfonated Nickel Ylide-Aluminum Alkoxide Catalyst. J.Polym. Sci., Polym. Chem. 1989,27,147. Langer, A. W. Linear Alpha-Olefins by Catalytic Oligomerization of Ethylene. J. Macromol. Sci., Chem. 1970,175. Lehmkuhl, H. Olefin Complexes of Organylmetals as Models for Organometallic Catalysts. Pure Appl. Chem. 1986,58,495. Lehmkuhl, H.; Rufinska, A.; Naydowski, C.; Danowski, F.; Bellenbaum, M.; Benn, R.; Schroth, G.; Mynott, R.; Pasynkiewicz, S. ($-Alkene) (VWyclopentadienyl Organyl Nickel) Complexes. Chem. Ber. 1984,117,3231. Novaro, 0.; Chow, S.; Magnouat, P. Mechanism of Oligomerization of a-Olefins with Ziegler-Natta Catalysts. J. Catal. 1976,41,91. Olive-Henrici, G.; Olive, S. The Active Species in Homogeneous Ziegler-Natta Catalysts for the Polymerization of Ethylene. Angew. Chem., Int. Ed. Engl. 1967,6,790. Olive-Henrici, G.; Olive, S. Olefin Polymerization with TransitionMetal Catalysts. Chem.-1ng.-Tech. 1971a,43,906.
Olive-Henrici,G.; Olive, S. Influence of Ligands on the Activity and Specificity of Soluble Transition Metal Catalysts. Angew. Chem., Int. Ed. Engl. 1971b, 10,105. Olive-Henrici, G.; Olive, S. Oligomerization of Ethylene with Soluble Transition-Metal Catalysts. Adv. Polym. Sci. 1974,15,1. Olive-Henrici, G.; Olive, S. Co-ordination and Catalysis. In Monographs in Modern Chemistry; Ebel, H. F., Series Ed.; Verlag Chemie: Weinheim, New York, 1977;Vol. 9,Chapter 8,p 186. Perrin, D. D.; Armarego, W. L.; Perrin, P. R. Purification of Laboratory Chemicals; Pergamon: London, 1978. Pillai, S. M.; Ravindranathan, M.; Sivaram, S. Dimerization of Ethylene and Propylene Catalyzed by Transition-Metal Complexes. Chem. Rev. 1986,86,353. Pillai, S . M.; Tembe, G. L.; Ravindranathan, M.; Sivaram, S. Dimerization of Ethylene to 1-Butene Catalyzed by the Titanium Alkoxide-TrialkylaluminumSystem. Ind. Eng. Chem. Res. 1988, 27,1971. Schulz, G. V. Highly Polymerized Compounds CXXII. The relation between Reaction Rate and Composition of the Reaction Products in Macro Polymerization Processes. 2.Phys. Chem. 1936,B30, 379. Schulz, G. V. The Kinetics of Chain Polymerization V. The Effect of Various Reaction Species on the Multimolecularity. 2.Phys. Chem. 1939,B43, 25. Shilov, A. E.; Dzhabiev, T. S.; Sabirova, R. D. Mechanism of Interaction between Triethyl Aluminum and Tetrabutyl Titanate and the Structure of Complexes formed. Kinet. Catal., Engl. Transl. 1964,5, 385. Van Zwet, H.; Bauer, S. R.; Keim, W. OligomerizingEthylene Using a Nickel containing Catalyst. German Patent 2,062,293,1971. Received for review January 3, 1991 Revised manuscript received May 1, 1991 Accepted June 10,1991
Steam Gasification of Cellulosic Wastes in a Fluidized Bed with Downstream Vessels Jos6 Corella,* Maria P. Aznar, Jesiis Delgado, and Elena Aldea Chemical and Environmental Engineering Department, University of Saragossa, 50009 Saragossa, Spain
In order to further use catalysts in a two- and three-stage process, advanced steam gasification of cellulosic wastes in a fluidized bed has been studied with and without a downstream high-temperature metallic filter and with and without a secondary vessel operating both empty and as a fluidized bed of calcined silica sand. Product distribution (gas, tar, char yields) and gas composition (H2,CO, CO,, CH4,C,) are given for all the operational variables studied: gasification temperatures, superficial velocity of the gas a t the gasifier inlet, residence time (space-time) of the gas in the gasifier, gasifier bed height, temperature of the filter chamber, state and temperature of the secondary vessel (empty and with a bed of sand), space-time and height of the second bed, size of the sand, and time on stream. The hot filter chamber, the silica sand in the second bed, and the carbonaceous solids formed are not inerts but they crack the produced tars, modifying the product distribution from the gasifier exit. Introduction Steam gasification of cellulosic wastes or of biomass in fluidized beds was studied in the early 1980s when it was thought this could be an alternative energy source. Nevertheless, the low price of oil and its derivatives, the disposal of cheap natural gas, and the healthy economies of the developed countries in the late 1980s meant that this process was not competitive as a source of energy. A lot of developing new technologies in this area were stopped by the absence of continuing funds. Only a few new processes for biomass were commercialized such as the one at Studvik Energy in Sweden (Rensfelt and Ekstrom, 1988),but many pilot installations were dismantled with
neither having gained experience nor having solved all the technical problems. This was the case with the steam gasification of cellulosic wastes in fluidized beds. This technology can have some future not only because of oil price increases in the future, but also because it can eliminate solid wastes with energy production by the generated gas; nevertheless this technology still has some technical problems. At the least, it could be improved and more advanced gaeifiers should be developed to make this technology useful. For instance, one of its problems is the primary product distribution. When steam is used as the gasifying medium, the gasifier exit flow is composed of H2, CO, COz,CHI, Cp,Cs,and higher hydrocarbons, together
0888-5885/91/2630-2252$02.50/0 0 1991 American Chemical Society
Ind. Eng. Chem. Res., Vol. 30, No. 10, 1991 2253 with elutriated ash and char. Hydrocarbons condense at the installation exit together with the steam, giving a fraction soluble in water and another nonsoluble fraction called tars. These tars must always be entirely eliminated for many well-known results. The exit gas contains about 50% H2 and 20% CO; thus, it could be used to produce H2, to produce chemicals, or to be burnt as a medium heating value gas. However, its 4-10 vol% CHI content must be either eliminated (to produce a synthesis gas) or increased (to upgrade the heating value). There appears then a clear necessity to improve the steam gasification by modifying the product distribution. Together with other technical solutions such as O2 introduction, use of very high temperatures, etc. (Bridgwater, 19841, the use of catalysts could be a solution. Catalysts could be used for, at least, tar cracking, methanation, or steam reforming of the produced CHI and tars. These chemical processes and solutions have been widely studied and used in oil refineries and in coal industries, such as those reviewed by Juneja et al. (1987), but not in the thermochemical processing of solid wastes and of biomass. Thus, this and forthcoming papers will be devoted to the experimental study of the use of tar cracking, methanation, and CHI steam reforming catalysts in the steam gasification of cellulosic wastes in fluidized beds. The results using these catalysts will be shown in forthcoming papers. Here we will give only the product distributions obtained, in about 80 different runs, without catalysts in the same installation in which the catalyst will be later used. Concerning the important question of where to place the catalyst, at least three different sites or emplacements were envisaged: (a) impregnated in the solid waste; (b) placed in the same gasifier, mixed with or instead of the silica sand; and (c) placed in a secondary reactor in series with the gasifier. We thought that methanation, steam reforming, and some tar cracking catalysts had to be placed in a secondary reactor for at least four reasons: (1)Commercial methanation and steam reforming catalysts are made to be used in fixed beds, not in fluidized beds (like the gasifier). They are soft and easily eroded if fluidized. (2) The optimal temperature of these three processes (methanation, 350 OC; steam reforming, 750 "C; cracking with dolomites and limestones, 800-900 "C) is different from that of the gasification (750-800 "C). (3) The existing char in the gasifier bed would increase the rate of deactivation of the catalyst if placed in the same gasifier. (4) With two different reactors the second one can have a different diameter and therefore a different superficial gas velocity (u)and hydrodynamics. This fact allows the use of a particle size (and umf)for the catalyst different from that needed if the catalysts were placed in the Same gasifier bed. Thus, making the u/umf ratio in the second bed different from (less than) that in the gasifier bed, some advantages are obtained, such as a minor erosion and elutriation of the catalyst particles. For these four reasons we will use catalysts placed only in a secondary reactor downstream from the gasifier. Catalysts for tar cracking can be of, a t least, two different types: commercial and "in equilibrium- FCC catalysts (average diameter of 60-80 pm) and natural dolomites, magnesites, or limestones (Yeboah et al., 1980; Wen and Cain, 1984; Magne et al., 1990; SjBstrom et al., 1988). We are studying both types, but given the very low terminal velocities of the FCC catalyst particles (ut * 20 cm/s), they are quickly elutriated (Aznar et al., 1988a, 1989b, 1992; Corella et al., 1988a) and they would have to be used in a circulating system (Herguido et al., 1990).
Thus, in the present two-stage installation only dolomites, limestones, and magnesites will be used as cracking catalysts. We consider that the scale or the size of the gasifier and of the catalytic reactor is an important decision made at the outset that can influence not only further important results such as the product distribution and gas composition (Corella et al., 198813) but also future decisions about the technical and economical viability of the overall catalytic process. Between very small lab gasifiers (Le., 23-cm i.d.) and pilot-plant gasifiers (larger than 15-cm id.), we have had to choose an intermediate or bench-scale installation. For scaling up, this size has some advantages such as the continuous feeding of biomass, the continuous reforming of a real exit gas from a gasifier of wastes, and the continuous deactivation of the catalyst. Nevertheless, this scale requires the waste to be fed by the top of the gasifier, when feeding at the bottom would be much better: less tar produced (Corella et al., 1988a). However, the feeding at the bottom is only easily done in big fluidized bed gasifiers which in turn would need large quantities of catalysts that are outside our possibilities. In order to know the effects of the catalysts, to compare the product distribution obtained with and without catalysts, and to differentiate the catalytic activity from the effects (thermal cracking, for instance) of the installation, it was absolutely necessary to study the gasification in the same installation without catalysts. This study is the object of this first paper and it has three different parts: (steam) gasification (in a fluidized bed) (i) without a secondary catalytic reactor, (ii) with a high-temperature filter and an empty catalytic reactor, and (iii) with the catalytic reador filled with inert silica sand. Of course, the fluidized bed steam gasification of different types of celullwic wastes has been studied by several authors, and even their results have already been analyzed and compared (for instance, by Corella et al. (1988b, 1990)). Among the published papers on steam gasification of biomass in a single fluidized bed, the more relevant work with respect to the present one may be the study of Walawender, Fan and co-workers (i.e., Raman et al. 1980; Walawender et al., 1981, 1985; Hoveland et al., 1985; Singh et al., 1986). The gasification (or pyrolysis) with a secondary bed has also been studied in several countries: in Sweden by Alden et al. (1988) and Sjostrom et al. (19881, in France by Magne and Deglise (Le., Donnot et al., 1985; Magne et al., 19901, in Finland by Simell and son-Bredenberg (1990), and in the USA by Baker et al. at PNL, Richland, WA (i.e., Elliot and Baker, 1986; Baker et al., 1987; Mudge et al., 1988), and by Longwell and Peters at MIT (Ellig et al., 1985, Boroson et al. 1989a,b). Although, then, there are many previous good works on steam gasification of cellulosic wastes in fluidized beds, we nevertheless present here a lot of new information and data that will be the basis of comparison in forthcoming papers on catalytic, two-stage or advanced gasification.
Equipment Description A scheme of the whole installation when it already has the filter and the second vessel (reactor, sometimes) is shown in Figure 1. The gasifier is a 6-cm4.d. fluidized bed, 30 cm high, externally heated by a 4 kW furnace, followed by a disengaging zone, first conical 15 cm high and then cylindrical, 20-cm i.d., and 15 cm high. This freeboard zone was externally heated to diminish tar condensation, and it plays an important role (Hoveland et al., 1985) because in it (i) gas-phase (such as Shift) reactions occur and (ii) in the conical zone sometimes there is bridging of the char formed by the entering sawdust.
2254 Ind. Eng. Chem. Res., Vol. 30, No. 10, 1991 1. Gasifier 2. Metallic filter 3. Catalytic bed 4. Waste feeder 5. Water pump 6. Condenser 7. Gas filter 8. Flowmeter 9. Gas sampling 10. Condensate
"2
Figure 1. Installation used.
The longitudinal profile of temperatures in the freeboard was measured. Usually it was at 500-600 "C, but sometimes, when external heating did not work, temperatures as low as 200-300 "C were detected. The heating rate of the fed sawdust was then lower and higher tar and char yields were obtained. The gas distributor plate is important; it is a plate with 9 bubble cups of 8-mm diameter each with 10 holes of 0.8-mm diameter. At the u/umfratios used here (from 1.5 to 5) this plate gives very small bubbles with good gas-solid contact. Blocking occurred at times, and periodically (every 4 months) the plate had to be mechanically cleaned. To preheat the fluidizing and gasifying gas, it was passed through a heated spiral pipe before reaching the gas distributor plate. The fluidizing gas was always steam. The operating pressure was always nearly (slightly above) atmospheric. The continuous feeding system of the biomass has special importance. Although several improvements were made, some runs had to be stopped by the plugging of this system. The metallic screw feeder has 3.5-cm i.d. and 25-cm length; it is inclined with an angle of 30". The screw feeder is connected to a variable-turning-speed motor (usually 14 rpm). To avoid bridging in the hooper, its walls have an angle of 80" (with respect to the horizontal) and it was continuously knocked. The downpipe or standpipe is 3-cm i.d., it was externally refrigerated with five quenching air flows, and it ends just at the vessel top (it does not extend into the freeboard). As cellulosic waste, pine (Pinus pinaster) sawdust was always used in these runs with a flow rate of 5-10 g/min (0.30-0.69 kg/h) and a particle size, initially, of -630 + 250 pm and, afterwards, of -2.0 + 1.0 mm. The selection and change of this size was made for two reasons: (i) to improve (increasing the size) the direct fall of the sawdust particles from the top of the vessel to the gasifier bed, without them being dragged by the exit gas flow and
pushed to the gasifier walls where they could stick, beginning a bridge that would collapse the run; (ii) to diminish the segregation of its char in the gasifier bed with the sand of -300 + 200 pm (and, sometimes,of -200 + 125 pm), according to the previous studies on segregation in these mixtures by Aznar et al. (1989a). The pine sawdust used had a composition (elemental analysis) of C, 41-45 wt %; H, 6.3-6.7 wt %; 0,49-53 wt % . Other properties were ash, 0.5-2 wt % ;moisture, 8-12 wt % fixed C, 12-13%; volatiles, 74-76 wt %; (LHV), 17.6-18.4 MJ/kg daf. When the installation had the downstream vessels, after the gasifier there was a porous metallic (Inconel) filter operating at 350-550 "C, depending on the run. Its nominal pore size was 10 pm, and its dimensions were 4 X 10 X 50 cm. On the filter surface a cake was always formed of a fine carbonaceous solid (particles less than 10 pm). This solid had a physical texture absolutely different from that of the char formed in the gasifier bed from the sawdust. We suppose it comes from the thermal cracking of the tars. The gas inlet in the filter chamber was tangential, of cyclonic type, so some fine char and sand dragged from the gasifier were collected at the bottom of this chamber and not on the filter surface. After the filter there was the secondary vessel in which the catalyst will be placed. The vessel had a lower part of 4.0-cm i.d. and 10-cm height followed by a conical zone (6.5-cm height) and an upper disengaging zone of 8.0-cm i.d. and 15.5-cm height. At the exit of this vessel there were a heat exchanger, a tar and water receiver, an effective demister (with glass wool), a continuous gas flowmeter (with temperature monitoring), a gas sampling device, and the exit gas pipe. Analysis The gasification products are gas (H2,CO, COZ, CHI, C2, and C3), liquids (and insoluble faction in the condensate
Ind. Eng. Chem. Res., Vol. 30, No. 10, 1991 2255 water and a soluble one), and solids (the char collected in the gasifier bed and in the freeboard walls and the carbonaceous solid in the filter cake). Periodically in each run, exit gas and liquid samples were collected to be analyzed. That gave us the evolution of the product distribution with time on stream. Usually six to seven samples were taken, every 10-15 min, by run. The sand and carbonaceous solids collected in several parts of the installation were calcined to determine their weight loss. This amount will be denoted char because the cake on the filter (it does not come from the sawdust fed) only was a small weight fraction of the overall carbonaceous solid collected in the installation. Gas samples were analyzed by conventional gas chromatography in a HP-5790 model. Hydrogen was determined with a molecular sieve 5 D column at 45 "C with nitrogen as carrier gas. CO, COz, CHI, Cz, and C3 were determined in a carbosieve SI1 column at 150 OC with helium as carrier gas. Some N2 was also detected in the exit gas. It comes from the two very small purges of nitrogen existing in the installation. The exit gas compositions shown in this work will be referred to the gas without NP. The determination of the organic content in the condensates was the most hard and time-consuming work, and even very disagreeable due to the odor and toxicity of the tars (mainly polyaromatic hydrocarbons (PAH)). Every sample of condensate was analyzed as follows: the sample is first filtered and the nonsoluble fraction remains in the cake. The cake is treated with acetone, which dissolves the tars. This acetone solution is then distilled at 80 "C giving a residue named nomluble tars (fraction nonsoluble in water). On the other hand, the condensate water with soluble organic components are analyzed in a high-temperature total organic carbon (TOC) analyzer (Dohrmann DC-90 model). This TOC determination gives the organics in milligrams per liter. To calculate the tar yield, an average molecular weight of 94 was supposed as Sjostrom et al. (1988). Liquid chromatography was used to determine tar composition (Aznar et al., 1989b), and phenol, cresol, and naphthalene were the dominant components in the tar from this noncatalytic gasification of wood. The TOC determination gives then that will be called water-soluble tars. Both fractions (soluble and nonsoluble) give the total tar content, yield, or production. This value will be the one given below. Tars were also analyzed by HPLC (liquid chromatography) in a Kontron apparatus with pumps (Model 4201, autosample (Model 460), and detector (Model 430). Precolumn and a (2-18 column were used. The detector was UV (A = 254 nm). Tars were dissolved in 30% H 2 0 and 70% CH,CN; the dissolution flow rate was 1.5 mL/min with a isocratic program. Chromatograms of the produced tar were similar to those shown by Ald6n et al. (1988). Due to our interest in eliminating these tars, chromatograms will only be shown when catalyst is used, because with catalyst in the secondary bed the chromatograms of the tar change a lot.
Experimental Section To calculate the activity of the catalysts that will be used for the several reactions of the reacting network, it is necessary to know the product distribution from the steam gasification in this installation with two reactors in series but without catalyst. It is also necessary to know the effect of the several operational gasification variables and to select the best ones for further operation with catalysts. Therefore, the work without catalysts has been devised in three parts: gasifier without downstream filter or secondary vessel; gasifier with high-temperature filter and
1
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I
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GASIFICATION TEMPERATURE ("c)
Figure 2. Product distribution at the gasifier exit vs gasification temperature (steam/daf sawdust ratio = 0.80, T0,1 = 1.4 8).
70
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' 1
600
I
I
700 800 GASlFICATION TEMPERATURE yC)
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Figure 3. Gas composition at the gasifier exit at different gasification temperatures (steam/daf sawdust ratio = 0.80, ~ 0 =, 1~.4 8 ) .
empty secondary vessel; gasifier with filter and secondary vessel with a fluidized bed of an inert solid (silica sand).
Product Distribution with Gasifier Only, without Filter or Second Vessel Effect of the Gasification Temperature. The gasification temperature (TI)was studied between 660 and 810 "C. In these runs the following variables were maintained constant: sawdust flow rate = 7.1 f 0.8 g/min; sawdust (daf) flow rate = 6.2 f 0.7 g/min; steam/sawdust ratio (g/min/g/min) = 0.70 f 0.09; steam/sawdust (daf) ratio = 0.80 f 0.10; u , , ~= 15 cm/s; T ~ =, 1.4 ~ s; total pressure = 1.1 atm. Product distribution (gas, tar, and char yields) at different gasification temperatures is shown in Figure 2. These yields are the averaged values for the 2 h of the length of the run. As is already known, the gas yield increases and the char and the tar yields decrease with the rise in gasification temperature. Also, mass balances at the end of the run not shown here (Aznar, 1989) indicate that the steam reacted in the gasifier bed, or the steam conversion, increases with the gasifier temperature. The composition (main components) and low heating value (LHV) of the exit gas and the thermal effectiveness or energy recovery (defined as the percentage of the energy content in the sawdust fed that appears in the exit gas) without tars at different gasification temperatures are shown in Figures 3 and 4, respectively. LHV is always around 12-13 MJ/m3(NTP) (NTP = 1 atm, 20 "C) and energy recovery increases from 35% to 75% on increasing T from 660 to 810 "C. With these results, all further runs
2256 Ind. Eng. Chem. Res., Vol. 30, No. 10, 1991 80,
118
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5 600
700
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GASIFICATIONTEMPERATURE('C)
Figure 4. Low heating value of the gas at the gasifier exit and thermal efficiency vs gasification temperature.
will be made at gasification temperatures between 750 and 800 "C. When comparing this product distribution with that obtained in the same conditions in a fluidized bed gasifier of a bigger size (small pilot-plant level) and by other authors, important differences are seen. Apart from differences in gasification temperatures, steam/biomass ratios, type and moisture of the waste or biomass, etc., used by the different authors, it seems clear to us that the location of the feeding point also has an important effect on the product distribution. The very low density of the char, very quickly formed in feeding the biomass, makes this char segregate in the upper part of the bed (Aznar et al., 1989a). Then, when the feeding is by the top, as in Figure 1, at low superficial gas velocities such as the ones used here, the cellulosic waste produces a zone of char in the upper part of the bed. The waste then fed makes contact with this char at the top of the bed, and there it undergoes a heating rate lower than when the waste is fed at the bed bottom. In conclusion, more char and tar yields are obtained when the waste is fed at the top as in our installation. In any case, since the 80 runs shown here were made with the same type of feeding (at the top), these results are comparable among themselves. Remember, nevertheless, that this product distribution may change somewhat when applied to a bigger gasifier or to a different gasification process (Corella et al., 1988b). Effect of the Superficial Gas Velocity at the Ga) . the screwfeeder throughput sifier Inlet ( u ~ , ~Given (between 4 and 10 g/min), the diameter of the gasifier, and the steam/biomass (flow rates) ratio (around l.O), the superficial gas velocity at the gasifier inlet (u0,Jis fixed. We can vary uo,l between 10 and 25 cm/s. We have selected an interval for (uoJud) from 2 to 5 for the following reasons: (i) the longitudinal profile of the char in the bed, or the segregation index, is very much affected by the ( U ~ , ~ / U ∶ ) (ii) for (uo,l/ud) from 2 to 5 the segregation of the char is low if (dp sawdust/dp sand) > 3 (Aznar et al., 1989); (iii) the mean bubble diameter in the bed increases with (u0,,/ud) and the solid-gas contact is poorer. Thus, with uo,l being fixed, the umf of the char-sand mixture has to be about 5 cm/s. These considerations in turn made us select and use for our gasifier a silica sand of -300 + 200 km. The mixture at 50 vol 7% sand of this size with char (from the sawdust used) has an ud = 5 cm/s and a minimum complete fluidization velocity (ucf) = 9 cm/s (Aznar et al., 1988a,b;1992). Then, our small gasifier will operate at uo,l from 10 to 21 cm/s. Of course, in pilot and commercial gasifiers, since the throughput or flow rate of solids is much bigger, to maintain the biomass/steam ratio near unity, uo,l has to be also much bigger (of the
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4 10
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1 13
.
1 16
.
I
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.
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Figure 5. Product distribution in the gasifier vs superficial gas velocity at the gasifier inlet (760 "C).
order of 1-3 m/s) and the size of the sand also has to be bigger (1mm is a common choice). The other experimental conditions were as follows: gasification temperature = 760 " C ; steam/biomass ratio = 0.89 f 0.08 steam/biomass daf ratio = 1.02 f 0.09; biomass flow rate = 4.0-7.6 g/min; = from 4.0 to 8.5 f 1 cm; 70,1 = 0.40 f 0.05 s. Since all the variables are interrelated, to diminish uo,l (or steam flow rate) while maintaining the steam/sawdust (flow rates) ratio constant, it is necessary to diminish the sawdust flow rate to the gasifier. Then, at low uo,land/or biomass flow rate, less char is produced and accumulated in the bed and a lesser variation of the exit gas composition with time on stream is observed. The average (with respect to the time on stream) product distribution (gas, tar, and char yields) at different uo,lis shown in Figure 5. A strange result was obtained at low velocities (12 cm/s) that could be attributed to the fact that when uo,ldiminishes approaching ucf,the segregation index increases and more char is segregated in the upper part of the bed. For this reason the uo,lvalues selected for further runs were between 15 and 21 cm/s. In this interval, u0,+seems not to have any influence on the product distribution and, with the size selected for the sand and according to Aznar et al. (1989a),there is an acceptable mix in the bed between the char and the sand. Effect of the Residence Time of the Gas i n the Gasifier Bed. Since the reactant gas (steam) disappears somewhat in the gasifier and since the gas expansion in steam gasification is very important (eA from 2 to 41, we find it is necessary to use the space-time ( T ) for the gas instead of the mean residence time. For the gasifier (vessel 1)and at a time on stream = 0 (H= Ho,l),T ~ =, Ho,l/~o,l. ~ To know the effect of T ~ , several ~ , runs were made v w n g Ho,lin the following conditions: gasification temperature = 760 "C; steam/biomass ratio = 0.84 f 0.09; steam/ biomass daf ratio = 0.96 f 0.11; at 15 cm/s = from 4.5 to 30 (fl) cm; uo,l = 21 cm/s; T ~ =, from ~ 0.21 to 1.4 (*0.15) s. Product distribution (gas, tar, and char yields) and gas composition at the gasifier exit are shown in Figures 6 and 7, respectively. Tar yield seems to remain constant with bed height (or with T ~ , ~Figure ) , 6. We think this is due to the biomass feeding at the top. The biomass pyrolysis and tar production occurs then in the upper part of the bed. Tars go out with the upward gas flow without really entering into the bed. Thus, tar yield is not affected by the bed height. 70,1 has some effect on the gas composition, Figure 7. At low T ~ this , ~composition is near that obtained in fast pyrolysis (Corella et al., 1988b), indicating that gasification
Ind. Eng. Chem. Res., Vol. 30, NO. 10, 1991 2257
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Figure 6. Product distribution in the gasifier vs space-time of the gas in the gasifier bed (760 "C).
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% b Figure 7. Effect of the residence time of the gas in the gasifier bed on the composition of the exit gas (760"C, steam/sawdust daf ratio = 0.96).
occurs to a low extent. On increasing T ~or, Ho,l, ~ Hz and COz increases and CO decreases. This variation seems to indicate that the gas shift reaction (CO + H20 = COz + H,) is carried out in the bed or in ita freeboard. C H I and C, decrease with the bed height due, maybe, to the steam reforming reactions (CHI f 3Hz0 = CO, COz + 5H2). These results agree, for instance, with those of Donnot et al. (1985), who carried out these reactions in an empty reactor and in a reactor filled with sand and found that the rate of the above cited gas-phase reactions increases 2-4 times in a bed with sand. Boroson et al. (1989a,b) worked at a temperature much lower (400-600 "C)than ours and with a different apparatus (batch for the waste),
Figure 8. Effect of the filter and of the second vessel on the product distribution. Yields (kg/kg daf) with (Y)and without ( Y O ) filter and secondary vessel (TI= 750 "C,steam/daf sawdust ratio = 0.86, = 1.4 s, uo,l = 15 cm/s).
for which the comparison of our respective results can only be made qualitatively. In spite of this, we reach the same conclusion: the char formed in the bed and the carbonaceous deposita on the sand present in the bed enhance the conversion of the produced tars and the gas-phase reactions. The energy recovery or thermal efficiency of the gasification increases from 40 to 60%, and the LHV of the gas decreases from 16 to 12 MJ/m3(NTP) with the bed height (or with 70,~). The selected interval of T ~for, further ~ experimentation is 0.9-1.45 8.
Product Distribution with High-Temperature Filter and Empty Secondary Vessel Once the steam gasification was investigated in the gasifier alone, the filter and the second vessel were connected in series with the gasifier to get the complete installation shown in Figure 1. In this installation, with the second vessel being empty (without sand nor catalyst), runs were made at different temperatures in the filter chamber and in the second vessel (measured in its axis). The mass balance always closed in these runs between 85 and 100% of the input flows (Aznar, 1989). Product distribution (yields) now obtained compared with the one(s) obtained in the same gasification conditions but with the gasifier alone (without filter and secondary vessel) is given in Figure 8. Gas,tar, and char yields, Yi, now obtained and those obtained a t the gasifier exit (at 750 "C, uol = 15 cm/s, and HzO/biomassdaf ratio = 0.86), Yi", are shown in this figure. In the extended or enlarged installation, less tar and more gas yields are clearly obtained. This is due to the thermal cracking of the tar in the filter (500 "C) and in the secondary vessel (from 350 to 850 "C). According to Avidan and Shinnar (1990), below 450 "C the rate of thermal cracking is low for most hydrocarbons and above 600 OC this rate is fast for all hydrocarbons except methane. To verify the thermal cracking of tars in our installation, runs were made varying the temperature of the filter chamber between 360 and 573 "C. A clear increase in the exit gas yield is obtained at higher filter temperatures, Figure 9, indicating tar cracking. We think that this cracking is not only thermal but also catalytic. The coke, char, and/or activated carbon formed on the surface of the filter by thermal cracking has a slight catalytic cracking activity for the tar,as has been recently demonstrated by Boroson et al. (1989a). Also, the metal of the filter itself
2258 Ind. Eng. Chem. Res., Vol. 30, No. 10, 1991 100 -l
m
c
m
'0
T fllteh 550°C
2
a
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0
6o
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1
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9
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0,o
600
500
400
300
700
800
o f
I
1
I
I
0
5
10
15
20
HEIGHT SECOND BED (cm)
TEMPERATURE SECOND BED ('C)
Figure 9. Gas yield vs temperature of second vessel with sand (Hog = 10 cm) for two different temperatures in the filter chamber. 70 7
Figure 11. Second bed with silica sand effect of the height of this second bed on the product distribution.
i H 2 m
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(kg rrnd.h/Nms)
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300
.
, 400
.
l 500
.
l 600
.
l 700
.
l
.
800
10
l 900
TEMPERATURE SECOND VESSEL, EMPTY (OC)
Figure 10. Second vessel empty: effect of the temperature in this vessel on the final gas composition (Tl = 750 "C). Table I. Exit Gas Composition and Tar Concentration with and without Filter and Secondary Vessel (TI = 750 OC, u0,, = 15 cm/s) gas composition, tar ,VOI "/a concn, H, CO CH, CO," -, e/(mS(NTP)) 20 64-148 installation with filter and 52 23 6 secondary vessel (emptyY 16 104-195 installation without filter or 43 32 6 secondary vesselb
-
H20/daf = 0.67-0.84;T2 = 350-844 "C. bH20/daf = 0.86.
(Inconel) can have a cracking activity. All these effects give a lesser tar yield in the extended installation. The char yield is not much affected by the two new vessels because the char is mainly produced in the gasifier and remains in it, and in this char production the downstream vessels have no influence. The average (with time on stream) gas composition at the exit of the installation is shown in Figure 10. It does not vary significantly with the temperature of the secondary vessel in the interval here used. When this (average) composition is compared with the one obtained at the gasifier exit (Table I), a clear difference appears. Now H2 and COBcontents are higher and CO content is lower. Clearly this is due to the shift reaction (CO + H 2 0 = C02 + H2) in the (new) reaction volume created by the filter and the secondary vessel. The differences shown in Table I indicate that the results obtained with the gasifier alone cannot be taken as the basis for further comparison and study of the activity
o ! 0
I
I
I
5 10 15 HEIGHT SECOND BED (cm)
iD
Figure 12. Second bed with silica sand: tar concentration in the + 300 pm)vs bed height or space-time of the gas in the second bed (Tl = 780 O C , H,O/daf sawdust = 1.3-1.7, uo,l = 14.3-15.5 cm/s). gas from the second bed (at 800 O C , sand of -500
of the catalysts placed in the secondary vessel. At least, the results obtained with the whole installation will have to be the basis for calculating the conversion and gas composition modification produced by the catalyst when it is placed in the second vessel. In Table I tar concentration in the exit gas is also shown and these figures should be retained. For further comparison, when silica sand is used in the second vessel, the tar concentration at the second vessel inlet (CA~) taken as basis will be 75.8 g/m3(NTP). It corresponds to TI= 750 "C, H20/biomass daf = 0.84, uo,l = 15 cm/s, and the second reactor empty at 750 "C.
Product Distribution with Silica Sand in the Second Vessel Beds of calcined (at 900 OC)silica sand (from the beach) of two different sizes were placed in the second vessel, and the temperature and height (space-time) of this second bed were studied. Effect of the Second Bed Height or Space-Time. The second bed height (H0,2)has been investigated in two different conditions (note that in both conditions uo,2= 60 cm/s > umf,the second bed then being fluidized): (a) TI = 780 "C, H20/daf sawdust = 1.3-1.7, T2 = 800,O C , sand in second bed = -500 + 300 pm; (b) TI= 750 "C, H20/daf sawdust = 0.9-1.3, T2 = 745 OC,sand in second bed = -300 + 200 pm.
Ind. Eng. Chem. Res., Vol. 30,No. 10,1991 2259
60-
3
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8
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TEMPERATURE SECOND BED ("C) 7n
'"
4
o ! 0
b
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I
I
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I
5
10
15
20
Figure 14. Effect of the temperature of the second bed (bed height = 10 cm, sand = -500 + 300 rm)in the product distribution at the installation exit. (TI= 780 "C,HaO/daf sawdust = 1.3-1.7,u , , ~= 14.3-15.5 cm/s).
HEIGHT SECOND BED
Figure 13. Second bed with silica sand: effect of the second bed height on the gas composition at the exit of the installation, for two different gasification conditions: (a) T2= 800 O C , dp arena = -500 + 300 pm, Tl = 780 O C , H20/daf sawdust = 1.3-1.7 u , , ~= 14.3-15.5 cm/s; (b) T2= 745 O C , dp arena = -300 + 200 pm, TI= 750 O C , H20/daf sawdust = 0.9-1.3, 14.9-17.4 cm/s.
Product distributions for the conditions are shown in Figure 11. Gas and char yields seem to increase slightly with and tar yield seems to decrease. This tar yield variation is shown with more detail in Figure 12. Tar concentration in the exit gas (in the operational conditions given in the caption of the figure) is clearly affected by H0,2, diminishing when bed height increases. Therefore, silica sand (with the carbonaceous deposit formed on it by tar cracking) at 800 "C is active and cracks the tars, and this fact has to be remembered for further experimentation. Regarding the gas composition at the exit of the second bed,this composition remains unchanged with Ho,z (Figure 13). This gas composition, of course, depends on other operation conditions such as gasification temperature but not with the second bed height (in the interval of 3-15 cm), as Figure 13 shows. Effect of Temperature of Second Bed ( Tz).With the second bed filled with silica sand of -500 + 300 pm and at a height (H0,2) of 10 cm, runs have been made at different temperatures in the second bed (T2), ranging from 600 to 900 "C. In these runs TI = 780 "C, = 530 O C , H,O/sawdust daf = 1.3-1.7,and uo,l = 14.3-15.5 cm/s. Yields of tar, char, and gas at the end of the run and at the exit (for the gas and tar) are shown in Figure 14. Gas yield increases a little with T2, and tar yield clearly diminishes. This means that the sand or the carbonaceous deposit formed on its surface, or both, has an appreciable cracking activity for tars above -650 "C. This is seen in more detail in Figure 15,in which the tar concentration in the exit gas clearly diminishes with TP To compare with further results using a catalyst, it should be kept in mind
500
600
700
800
900
1000
TEMPERATURE SECOND BED ("C)
Figure 15. Effect of the temperature of the second bed on the tar concentration in the final or exit gas. 70
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2260 Ind. Eng. Chem. Res., Vol. 30, No. 10, 1991 70 7
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600°C 740°C
Kinetics of the Tar Cracking on the Sand Comparing the results obtained with silica sand in the second vessel (Figure 15) with the ones obtained without silica sand (Table I), it is easily deduced that the silica sand cracks the tar (A) at temperatures of 600-900 "C. This (thermal and catalytic) tar cracking rate can be fitted to the following simple kinetic equation:
m
o ! 0
I
100
I
200 TIME ON STREAM (min)
I 300
Figure 17. Second bed with silica sand (-500 + 300 pm, H0,* = 10 cm): variation of the tar concentration in the exit gas along the time on stream (parameter T2, tar concentration at bed inlet = 75.8 g/(N m3).
observed that the composition is the same (with, perhaps, a little increase in H2 and a diminution in CO, by the gas shift reaction). In conclusion: in the interval of 350-900 "C, and for mean residence times (space-times) for the gas of about 0.2 s, gas composition (H2,CO, COP,CH,, ...) does not vary with the temperature of the second vessel and with the presence of silica sand, but the tar content in the gas clearly varies. Effect of the Time on Stream. We say in advance here that the deactivation of the catalyst in the second bed is going to be the most important problem and one of the main interests of our research. This deactivation is going to be detected and measured by the variation with time on stream of the exit gas composition and of the product distribution (gas and tar yields) at the exit of the installation. Thus, we found it was necessary to know if the gas composition and the tar yield change at the exit with time on stream when there is only sand in the second bed, without a catalyst. For this purpose, samples of gas and condensates were taken and analyzed periodically during every run. In Figure 17 an example of the variation with time on stream of the tar in the exit gas is shown. The most common result is that the exit gas composition, tar and CHI included, remains constant with time on stream (during the 4 h of the length of these runs). Nevertheless it is necessary to say that sometimes important variations in the exit tar concentration are obtained in the same run, maintaining constant the waste and steam input flow rates, temperatures, etc., as can be seen in Figure 17. A slight diminution with time on stream of the tar concentration in the exit gas can occur due to the fact that the char and the carbonaceous deposits (coke) continuously formed and accumulated on the sand can progressively crack the tars coming from the gasifier. A progressive increase in the tar concentration in the exit gas is sometimes detected. This fact occurs when there is an important increase of char in the upper part of the gasifier bed or when there is bridging in the conical zone of the gasifier. Then, in that zone, the char agglomerates, the temperature decreases, and the incoming sawdust or waste does not enter into the bed but makes contact with a char at a low temperature. Then, the heating rate and final temperature of the sawdust or waste are lower and more tar is produced. Therefore, this increase (with time on stream) in the tar production is clearly detected at the installation exit. When this bad functioning on the run was detected, the results from such run were neglected.
Tars could also disappear in the bed by steam reforming reactions. Since there is a large excess of steam in our bed, its concentration is supposed to remain constant in the bed; then both reactions (cracking and steam reforming) are included in k. In our second bed the type of flow is very near the piston flow (Levenspiel's or dispersion modulus < 0.005). The differential mass balance for the tar is thus QOCAO dXA = (-FA) d W, XA being tar conversion by cracking and W weight of sand. For CAo,the previously obtained value (with the second reactor empty and at Tl = 750 "C) of 75.8 g/m3(NTP) has been used. Furthermore, tar cracking in the bed is important for the process but it does not significantly alter the total flow rate, Qo = Q, which is the same supposing that tA (in the second bed) = 0 and Q is a well-known value in each experiment. Thus, with eq 1 and the mass balance (plug flow):
or
To choose the best value for n, the values of k for different n values have been calculated and compared, Figure 18. The best fitting is for n = 1, and this value will be retained. To verify if in the second bed the assumption of t A = 0 is correct, we use the runs in which Qo and CAoare was varied). The mass balance constant (only 70,2or HOt2 with n = 1 and t A # 0 is
In Figure 18 the deviation or error for several values of is shown. The best adjustment is obtained for t A = 0, and this value is retained for further analysis. In conclusion, eq 1 with n = 1 and t A = 0 fits the data well. k, defined by eq 1with n = 1, was calculated in runs in which T2was varied from 600 to 900 "C. Representing k (m3(NTP)/(kgh))according to Arrhenius law (Figure 19), a very clear difference is obtained at high and low temperatures. A t T < 740 "C: tA
ko = 2.70 m3(NTP)/(kg.h)
E = 1.4 kcal/mol = 5.9 kJ/mol and at T > 740 "C: ko = 2520 m3(NTP)/(kg.h)
E = 15.1 kcal/mol = 63.1 kJ/mol The very low activation energy obtained a t low temperatures, 10 times lower than the corresponding one at high temperatures, clearly indicates that there are two competitive tar cracking mechanisms, or that tars behave
Ind. Eng. Chem. Res., Vol. 30, No. 10, 1991 2261 50
I
J 40-
2Y
8
Acknowledgment
30-
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20-
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!
I
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1 CRACKING ORDER
0
a 21
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1.0 1,5 2,O EXPANSION FACTOR ( EA)
0.5
b
2,5
Figure 18. Checking (a) reaction order and (b) expansion factor for tar cracking in the second bed of sand.
-c
Y
I
-0,o 0,8
Let us finally repeat that the data given in this work, which correspond to about 80 different and "hard" runs, will serve in further papers as the basis to improve the product distributions shown here by using catalysts in the second bed.
I
I
I
I
099
190
1,1
192
lOW/T(K)
Figure 19. Arrhenius representation for the tar cracking over silica sand.
has two different lumps: one lump will correspond to molecules "easily" cracked (nonstable thermal fraction) and the other would correspond to the stable thermal fraction. From our experimental results, at 600-740 "C the nonstable thermal fraction would be about 50% of the tar (obtained at 750 "C). Boronson et al. (1989) for their tars obtained a thermal (vapor-phase) conversion of 0-30% in the range of 400-600 "C. They also proposed a fraction of tar being thermally stable. Alden et al. (1988) also found that some tar components were easily cracked but others, like naphthalene, were very stable at high temperatures and very difficult to crack. Finally, the tar concentration in the exit gas has been predicted with eq 1 for several conditions in the second bed. A good fit was obtained, which gives us confidence in this kinetic equation.
This work has been made under the financial support of the shared contract between the EEC, DGXII, Project No. EN3B-0103-E, and the DGICYT (Madrid), Project No. PB88-078. The TOC determinations by FORET at La Zaida (Zaragoza, Spain) were very appreciated. Fruitful discussions with G . Taralas (RIT, Stockolm) and P. Mange (Universiti de Nancy) are recognized. The aid of J. Lahoz, M. P. Martinez, E. Agorreta, and N. Cebrian in some runs of this work was also very much appreciated.
Nomenclature CA, CAO= tar concentration at the exit and at the inlet of the second bed, g/m3(NTP) dp = averaged particle diameter, mm E = energy of activation for tar cracking over sand, kcal/mol Q, Qo = gas flow rate at the exit and at the inlet of the second bed, m3(NTP)/h k = cracking kinetic constant ko = preexponential factor of the Arrhenius equation, (m3(NTP))"/(h.kg of sand)-(gof tar)"-') H0,', Ho,z= bed height at t = 0 in the first and in the second bed, cm mA = mass of tar, g n = reaction order for tar cracking, eq 1 rA = rate of disappearance of the tar over silica sand, g/(kg of sand-h) t = time, h T1,T2 = temperature in the gasifier bed and in the second vessel (in its axis), respectively, "C u = superficial gas velocity, cm/s u0,',u0,2= superficial gas velocity at the inlet of the first bed (gasifier) and of the second (catalytic) vessel, cm/s umf= minimum fluidization superficial gas velocity, cm/s XA = tar conversion in the second bed of sand Y , Yo = yield, at the end of run (for the char) or at the exit of the second vessel (for the gas and tar) or of the gasifier, respectively, kg/kg daf W = weight of sand in the second bed, kg EA = expansion factor, defined as Q = Qo(l + e A X A ) , adim. T ~= ,space-time ~ for the gas in the first (gasifier)bed, defined as ~ 0 , 1 / ~ 0 , 1s, T ~ =, space-time ~ for the gas in the second bed, defined as H0,2/u0,2, s
Registry NO.COZ, 124-38-9; CHI, 74-82-8 CHfiHz, 74-85-1; CHSCH3, 74-84-0; H C z C H , 74-86-2.
Literature Cited AldBn, H.; Espenb, B.-G.;Ftensfelt, E. Conversion of Tar in Pyrolysis Gas from Wood Using a Fixed Dolomite Bed. In Thermochemical Processing of Biomass; Bridgwater, A. V., Kuester, J. L., Eds.; Elsevier: London, 1988; pp 987-1001. Avidan, A. A.; Shinnar, R. Development of Catalytic Cracking Technology. A Lesson in Chemical Reactor Design. Ind. Eng. Chem. Res. 1990,29 (6), 931-942. Aznar, M. P. Catalytic Steam Gasification of Biomass in Fluidized Bed. Ph.D. Thesis, Department of Chemical Engineering University of Saragossa, Spain, 1989. Aznar, M. P.; Gracia, F. A.; Corella, J. The ud and Q of Mixtures of Agricultural and Forestry Wastes with a Secondary and Fluidizing Solid. Part I. An. Quim. 1988a, 84, 379-385. Aznar, M. P.; Gracia, F. A.; Corella, J. The ud and uCfof Mixtures of Agricultural and Forestry Wastes with a Secondary and Fluidizing Solid. Part 11. An. Quim. 198813, 84, 386-396.
2262 Ind. Eng. Chem.
Res.,Vol. 30, No. 10,1991
Aznar, M. P.; Gracia, J. A,; Corella, J. Segregation in the Fluidization of Agricultural and Forestry Wastes. Longitudinal Profiles. An. Quim. 1989a,85, 100-108. Aznar, M. A,; Delgado, J.; Corella, J.; Lahoz, J. Steam gasification of biomass in fluidized bed with a secondary bed. In Pyrolysis and Gasification; Ferrero, G. L., et al., Eds.; Elsevier: London, 1989b;pp 629-634. Aznar, M. A,; Gracia, J. A.; Corella, J. The ud and uCfof Mixtures of Agricultural and Forestry Wastes with a Secondary and Fluidizing Solid. Int. Chem. Eng. 1992,in press. Baker, E. G.; Mudge, L. K.; Brown, M. D. Steam Gasification of Biomass with Nickel Secondary Catalysts. Ind. Eng. Chem. Res. 1987,26,1335-1339. Boroson, M. L.; Howard, J. B.; Longwell, J. P.; Peters, W. A. Heterogeneous Cracking of Wood Pyrolysis Tars over Fresh Wood Char Surfaces. Energy Fuels 1989a,3,735-740. Boroson, M. L.; Howard, J. B.; Longwell, J. P.; Peters, W. A. AIChE J. 1989b,35, 120-128. Bridgwater, A. V., Ed. Thermochemical Processing of Biomass; Butterworths: London, 1984. Corella, J.; Herguido, J.; Gonzdez-Saiz, J.; Alday, F. J.; RodriguezTrujillo, J. L. In Research in Thermochemical Biomass Conversion; Bridgwater, A. V., Kuester, J. L., Eds.; Elsevier: London, 1988a;pp 754-765. Corella, J.; Herguido, J.; Alday, F. J. Pyrolysis and Steam Gasification of Biomass in Fluidized Beds. Influence of the Type and Location of the Biomass Feeding Point on the Product Distribution. In Research in Thermochemical Biomass Conversion; Bridgwater, A. V., Kuester, J. L., Eds.; Elsevier: London, 1988b; pp 384-398. Corella, J.; Herguido, J.; Gonzcilez-Saiz, J. Steam Gasification of Biomass in Fluidized Bed. Effect of the Type of Feedstock. In Pyrolysis and Gasification; Ferrero, G. L., et al., Eds.; Elsevier: London, 1990; pp 618-623. Donnot, A.; Reningovolo, J.; Magne, P.; Deglise, X. Flash Pyrolysis of Tar from the Pyrolysis of Pine Bark. J. Anal. Appl. Pyrolysis 1985,8,401-414. Ellig, D. L.; Lal, C. L.; Mead, D. W.; Longwell, J. P.; Peters, W. A.; Pyrolysis of Volatile Aromatic Hydrocarbons and n-Heptane over Calcium Oxide and Quartz. Ind. Eng. Chem. Process Des. Dev. 1985,24,1080-1087. Elliot, D. C.; Baker, E. G. The Effect of Catalysis on Wood-Gasification Tar Composition. Biomass 1986,9, 195-203. Herguido, J.; Rodriguez-Trujillo, J. L.; Corella, J. Gasification of Biomass with Tar Cracking Catalyst in a Circulating Multisolid Fluid Bed Pilot Plant. In Biomass for Energy and Industry;
Grassi, G., et al., Eds.; Elsevier: London, 1990;pp 2793-2797. Hoveland, D. A.; Walawender, W. P.; Fan, L. T. Steam Gasification of Pure Cellulose. 2. Elevated Freeboard Temperature. Ind. Eng. Chem. Process Des. Dev. 1985,24,818-821. Juneja, M. N.;Mazundar, A.; Biswas, D. K.; Rao, S. K. Fuel Sci. Technol. 1987,6(2)45-60. Magne, P.; Donnot, A.; Deglise, X. Dolomite as a Catalyst for Tar Cracking. Comparison with Industrial Catalyst. In Biomass for Energy and Industry; Grassi, G., et al., Eds.; Elsevier: London, 1990; pp 2590-2594. Mudge, L. K.; Baker, E. G.; Brown, M. D.; Wilcox, W. A. Catalytic Destruction of Tars in Biomass-Derived Gases. In Research in ThermochemicalBiomass Conversion;Bridgwater, A. V.,Kuester, J. L., Eds.; Elsevier: London, 1988; pp 1141-1155. Raman, K. P.; Walawender, W. P.; Fan, L. T. Gasification of Manure in a Fluidized Bed Reactor. The Effect of Temperature. Ind. Eng. Chem. Process Des. Deu. 1980,10,623-629. Rensfelt, E.; Ekstram, C. Fuel Gas From Municipal Waste in an Integrated Circulating Fluid-Bed GasificationlGas-Cleaning Process. In Energy from Biomass and Wastes XII; I G T New Orleans, 1988. Simell, P. A.;son-Bredenberg, J. B. Catalytic purification of tarry fuel gas. Fuel 1990,69,1219-1225. Singh, S. K.; Walawender, W. P.; Fan, L. T.; Geyer, W. A. Steam Gasification of Cottonwood (Branches) in a Fluidized Bed Wood Fiber Sci. 1986,18 (2),327-344. Sjbtrom, K.; Taralas, G.; Liinaki, L. In Thermochemical Processing of Biomass; Bridgwater, A. V., Kuester, J. L., Eds.; Elsevier: London, 1988;pp 974-986. Walawender, W. P.; Ganesan, S.; Fan, L. T. Steam Gasification of Manure in a Fluid Bed. Influence of Limestone as a Bed Additive. In Symposium Papers from Energy from Biomass and Wastes V, Lake Buena Vista, FL; Mass, D. L., Ed.; IGT Chicago, 1981;pp 517-527. Walawender, W. P.; Hoveland, D. A.; Fan, L. T. Steam Gasification of Pure Cellulose. 1. Uniform Temperature Profile. Ind. Eng. Chem. Process Des. Dev. 1985,24,813-817. Wen, W. Y.; Cain, E. Catalytic Pyrolysis of a Coal Tar in a Fixed-Bed Reactor. Ind. Eng. Chem. Process Des. Dev. 1984,23,623-637. Yeboah, Y. D.; Longwell, J. P.; Howard, J. B.; Peters, W. A. Effect of Calcined Dolomite on the Fluidized Bed Pyrolysis of Coal. Ind. Eng. Chem. Process Des. Dev. 1980,19,646-653. Receiued for review November 26, 1990 Revised manuscript received May 16, 1991 Accepted May 28, 1991
