Biomass Gasification in Fluidized Bed at Pilot Scale with Steam

VOLUME 11, NUMBER 6

NOVEMBER/DECEMBER 1997

© Copyright 1997 American Chemical Society

Articles Biomass Gasification in Fluidized Bed at Pilot Scale with Steam-Oxygen Mixtures. Product Distribution for Very Different Operating Conditions Javier Gil,† Marıá P. Aznar,*,†,‡ Miguel A. Caballero,† Eva France´s,† and Jose´ Corella§ Chemical and Environmental Engineering Department, University of Saragossa, 50009 Saragossa, Spain, and Chemical Engineering Department, University “Complutense” of Madrid, 28040 Madrid, Spain Received December 27, 1996X

Biomass gasification in a fluidized bed with steam-O2 mixtures has been studied in detail at pilot plant scale. The gasifier used was 15 cm i.d. and 3.2 m high, and it was fed with pine wood chips at flow rates of 5-20 kg/h. Main operating variables studied were gasifier bed temperature (780-890 °C), steam to oxygen in the feeding ratio (2-3 mol/mol), and gasifying agent (H2O + O2) to biomass fed ratio (0.6-1.6 kg/kg daf). Product distribution here shown includes gas, tar and char yields, gas composition (H2, CO, CO2, CH4, steam, ...) and heating value, tar composition and content in the flue gas, gas heating value, apparent thermal efficiency, etc. Under good operating conditions the following gas is obtained: tar content of 5 g/Nm3, 30 vol % H2, heating value of 16.0 MJ/Nm3 (dry basis), gas yield of 1.2 Nm3 (dry basis)/kg biomass fed.

Introduction Biomass gasification is quite a well-known process that produces a valuable gas from a renewable energy, biomass, abundant in some scenarios. Biomass gasification can be classified or studied depending on the gasifier type (moving, fluidized bed, circulating, entrained, ...), operating pressure (in the gasifier), scale of processing, and also the main gasifying agent used. Besides some unusual gasifying agents, sometimes used at laboratory scale for very specific and academic purposes, the typical agents are air (with some moisture), steam, and steam + O2 mixtures. Depending on this gasifying agent, the raw gas produced at the gasifier exit has a different composition. The most used agent nowadays at demonstration or commercial scale is air with equivalence ratios of 0.2-0.3. By use of air, the †

University of Saragossa. Fax: +34 76 76 21 42. University “Complutense” of Madrid. Fax: +34 1 394 41 64. X Abstract published in Advance ACS Abstracts, September 15, 1997.

flue exit gas contains about 50 vol % N2 together with H2 (8-12 vol %), CO, CH4, C2, C3, CO2, H2O, and tars. This gas composition seems to be useful only for electricity production or heat generation. The exact gas composition at the gasifier exit in this process depends on a lot of operating variables, and it has been well studied by different authors, at different operating scales and with different gasifiers. For instance, Narvaéz et al.1 have recently reviewed and studied the effect of six operational variables on the product distribution and raw gas quality. Gasification with air being wellknown, will be thus absolutely out of the scope of this paper. Another interesting gasifying agent is pure steam. The product distribution is then quite different from the one obtained with air. It has been deeply studied among others by Antal,2,3 Corella et al.4,5 from 1984 till 1992,

‡ §

S0887-0624(96)00233-2 CCC: $14.00

(1) Narvaéz, I.; Orio, A.; Corella, J.; Aznar, M. P. Ind. Eng Chem.Res. 1996, 34, 2110-2120.

© 1997 American Chemical Society

1110 Energy & Fuels, Vol. 11, No. 6, 1997

Walawender et al.6,7 at Kansas State University, Mudge, Baker et al.8,9 at PNL in Washington, and Kinoshita and Wang.10 In steam gasification the flue gas from the gasifier does not contain N2 (except for a minor amount from purges) and the H2 content in such gas can be as high as 50-55 vol %.4 With a secondary and catalytic steam reformer, light hydrocarbons (CH4, C2, ...) and most tar can be converted into H2 and CO, increasing thus the H2 content in such gas to 70-75 vol %.11,12 This gas could realize new and promising applications, but steam gasification is endothermal and needs heat. This heat can be provided by burning in a separate chamber the char produced in the endothermic steam gasification and carrying such heat to the gasifier with a “medium” like silica sand. Several types of circulating systems were thus used to solve this problem,13-15 generating processes similar to the FCC for the cracking of heavy oil feedstocks. But such circulating units, well-known, designed, handled, and controlled in oil refineries, when applied to biomass gasification or cracking at small scale are, for us, quite complex and expensive. For instance, standpipes or transfer lines in such circulating systems offer some technical problems or difficulties in biomass processing. Although one Battelle-Columbus circulating demonstration unit, based on this concept, is being set up in Burlington, Vermont, similar but small pilot plants at Saragossa University, Spain14 and at TNEE, France15 were dismantled some years ago. Another solution for supplying heat in biomass gasification with steam is to add some amount of oxygen to the gasifier. Simultaneous exothermal reactions are then generated, together with the endothermal ones, and an autothermal gasification process can be achieved. The main problem is now the price of the oxygen. It is not cheap and the overall process could be not economically feasible. To our knowledge, it was the main reason for the failure of the big pilot in Clamecy (France) set up in the mid-80s by Creusot-Loire.16,17 Results gained with it belong nowadays to Stein-Industrie Inc. (GEC (2) Antal, M. J. In Fundamentals of Thermochemical Biomass Conversion; Overend, R. P., Milne, T. A., Mudge, L. K., Eds.; Elsevier Applied Sciences: London, 1983. (3) Antal, M. J. In Advances in Solar Energy; Boer, K. W., Duffie, J. A., Eds.; Plenum Press: New York, 1985; Vol. 2. (4) Corella, J.; Aznar, M. P.; Delgado, J.; Aldea, E. Ind. Eng Chem. Res. 1991, 30, 2252-2262. (5) Herguido, J.; Corella, J.; Gonza´lez-Saiz, J. Ind. Eng. Chem. Res. 1992, 31, 1274-1282. (6) Walawender, W. P.; Hoveland, D. A.; Fan, L. T. Ind. Eng Chem. Process Des. Dev. 1985, 24, 813-817. (7) Hoveland, D. A.; Walawender, W. P.; Fan, L. T. Ind. Eng Chem. Process Des. Dev. 1985, 24, 818-821. (8) Baker, E. G.; Mudge, L. K.; Brown, M. D. Ind. Eng Chem. Res. 1987, 26, 1335-1339. (9) Mudge, L. K.; Baker, E. G.; Brown, M. D.; Wilcox, W. A. In Research in Thermochemical Biomass Conversion; Bridgwater, A. V., Kuester, J. L., Eds.; Elsevier: London, 1988; pp 1141-1155. (10) Kinoshita, C. M.; Wang, W. P. Sol. Energy 1992, 49, 153-158. (11) Aznar, M. P.; Corella, J.; Delgado, J.; Lahoz, J. Ind. Eng. Chem. Res. 1993, 32, 1-10. (12) Aznar, M. P.; Corella, J.; Caballero, M. A.; Gil, J. Ind. Eng Chem. Res, in press. (13) Schiefelbein, G. F.; Stevens, D. J.; Geber, M. A. US Biomass Thermochemical conversion programme. Annual reports PNL-5445 (1984), PNL-5801 (1995); Pacific Northwest Laboratory: Richland, WA, 1984, 1995. (14) Herguido, J.; Corella, J.; Artal, G.; Garcıá-Bordeje´, J. E. In Biomass for Energy, Industry and Environment, 6th EC Conference; Grassi, G, Collina, A., Eds.; Elsevier Applied Sciences: London, 1992; pp 792-796. (15) Deglise, X. TNEE, France, European Patent 0 108 317 B1, 1986. (16) Chrysostome, G.; Lemasle, J. M. Rev. Gen. Therm. 1984, 275, 713-717. (17) Lesmasle, J. M.; Morcellin, M. ASCAB Report to EC (Brussels) for Contract No. EN3B-0108-F.

Gil et al.

ALSTOHM group). Despite this negative experience in this process, the authors thought that biomass gasification with steam + O2 mixtures was interesting and scarcely studied. In fact, after studying the available results from the Clamecy plant,16 the authors thought that a part of that failure was due to a lack of knowledge of the effect of several process variables on the product distribution. As far as it is known, such a pilot of Clamecy was built up without enough previous experience in such a process: it was not technically optimized. It was decided thus to institute research and technical development (RTD) in this process and to direct a part of our RTD activities to obtain more knowledge in biomass gasification in fluidized bed with steam + O2 mixtures. IGT is also studying gasification with steam + O2 mixtures, developing its RENUGAS process, but such a process is for gasification under pressure,18 and besides, there is not much knowledge published about the detailed effect of its operation parameters. Therefore, this paper will be entirely devoted to knowing how this gasifying agent, steam + O2 mixtures, affects the product distribution under different operation conditions. Gasification with pure compounds (steam or oxygen) is out of the scope of this paper. Besides the data for gasification with pure steam,2-10 some data of gasification with pure oxygen can be found in ref 19. Results shown here encompass work done over 4 years (1992-1996) at the Universities of Saragossa and Madrid (Spain), which was financed by the JOULE2 and JOULE3 EC programs. During this period, around 40 experiments were performed. On average, each experiment required a month and six full time people working on it. That is to say, the manpower involved in the work presented in this paper has been about 240 man-months. An introductory part of this RTD activity was presented at the Banff conference.20 Now, the full product distribution, for all variables studied, in the whole RTD program, is presented here in detail. Experimental Section Pilot Plant Used. The pilot plant used in this work is shown in Figure 1. Some technical details of this facility have been presented in some recent international conferences.20-22 The gasifier is a fast fluidized bed of 15 cm i.d.and 3.2 m total height, continuously fed with biomass near the bed bottom. The feeding system has two hoppers with two locks and two screw feeders of 6 cm diameter. The first screw is a dosing device (with a slow, variable, and controlled speed). The second one has a very fast speed to avoid biomass pyrolysis (with subsequent plugging) in it. The feeding pipe has an external heat exchanger (near to the gasifier) to avoid biomass (18) Lau, F.; Carty, R. H.; Onischok, M.; Bain, R. L. Presented at the Strategic Benefits of Biomass Fuels Conference, Washington, DC, March 1993. (19) Guillory, J. L.; Goldbach, G. O. Proceedings of the ASME Solid Waste Processing Conference, Denver, CO, June 1-4, 1986; pp 195203. (20) Aznar, M. P.; Corella, J.; Gil, J. Martin, J. A; Caballero, M. A; Olivares, A.; Pe´rez, P.;. France´s, E. In Developments in Thermochemical Processing of Biomass, Proceedings of the Conference held in Banff, Canada, May 1996; Bridgwater, A. V., Boocock, D., Eds.; Blackie Academic and Professional: London, 1997; Vol. 2, pp 1194-1208. (21) Aznar, M. P.; Borque, J. A.; Campos, I. J.; Martin, J. A.; France´s, E.; Corella, J. In Biomass for Energy, Environment, Agriculture and Industry, 8th EC Conference in Vienna, Austria; Chartier, P. H., Beenackers, A. A. C. M., Eds.; Pergamon Press: London, 1995; Vol. 2, pp 1520-1527. (22) Aznar, M. P.; Gil, J.; Martıń, J. A.; Caballero, M. A.; Olivares, A.; Pe´rez, P.; France´s, E.; Corella, J. In New Catalyst for Clean Environment, 2nd VTT Symposium in Espoo; Maijanen, A., Hase, A., Eds.; VTT: Espoo, Finland, 1996; pp 169-176.

Biomass Gasification

Energy & Fuels, Vol. 11, No. 6, 1997 1111

Figure 1. Scheme of the pilot plant used. pyrolysis before it enters the gasifier. The feeding system is really basic and a key to the good functioning of the overall pilot plant. The system used here was developed in a previous study and during this period did not give major problems. Nevertheless, it has to be pointed out that the diameter of our screw feeder is the limiting factor for the size and shape of the biomass that can be fed. In our case, this size has to be below approximately 2 cm. The maximum flow rate of biomass fed was 20 kg/h. A bigger flow rate could be perhaps gasified in this pilot, but other limiting factors (the plant is located on a universitary campus located in turn in the center of a big city) did not allow us to use higher throughputs. Higher throughputs would imply use of big flames in the torch and large amounts of fumes in the stack. So this pilot usually worked derated, with biomass flow rates of 5-10 kg/h. At the bottom of the gasifier there is always a stationary bed of a fluidizing medium, silica sand, of -0.63 + 0.4 mm particle diameter (equivalent to a umf at gasifier conditions of 12 cm/s). An amount of 13 kg was used in each experiment, equivalent to a (bulk, fixed) bed height of 50 cm. The superficial gas velocity (gasifier conditions) at the bed inlet most often was 45 cm/s (it means a gas residence time in the gasifier bed, once fluidized, of 1.2-1.5 s). Temperature was measured in the gasifier bed and freeboard for points indicated in Figure 1. The temperature in the bed was varied from test to test. The temperature in the freeboard was usually 650 °C. After the gasifier there are three high-efficiency cyclones connected in series. The exit flow is cooled in a relatively big heat exchanger (vertical-multitubular), as Figure 1 shows, collecting the condensates in its bottom. The cool exit gas is measured for its flow rate and then is burnt out of the building and sent to a stack. Before the heat exchanger, a slip flow is taken from the flue gas and sent to two downstream reactors where its effective and exhaustive hot cleaning is studied by using different types of commercial catalysts, dolomites, and related materials. All this catalytic hot gas cleaning is out of the scope of this paper. It will be shown in forthcoming papers. For the stationary state, the temperature of the gasifier bed increases when (H2O + O2)/biomass increases and when H2O/ O2 decreases, which means an increase of the oxygen fed per

Figure 2. Gasifier bed temperature vs gasifying agent-tobiomass ratio. (H2O/O2) ) 2 and 3 (mol/mol). kilogram of biomass (Figure 2). The gasifier has at its bottom an external oven for preheating the bed to the desired temperature and also as an external heat supply to maintain and control the temperature in the gasifier bed at the prefixed or desired one. The effect of such prefixed temperature in the control of the process or set point in shown in Figure 2. Increasing the difference between the set point and the gasifier real temperature means increasing the heat transfer from the oven to the bed. For the upper part of the gasifier, cyclones and connecting ducts also have several external heaters to get a relatively high temperature in the gasifier freeboard and to avoid condensations in the cyclones. The plant has feeding pipes and devices for steam, oxygen, air, and nitrogen. All these gases can be fed with known flow rates. The plant has a very advanced control. It was designed together with the help and expertise of the Process Control Institute (ICP) of the Spanish Council of Scientific Research (CSIC) in Cantoblanco, Madrid. The plant is exhaustively cleaned after each experiment. The feedstock is carefully prepared (collected, sieved, partially dried sometimes, etc.). Several samples of flue gas and condensates are taken in each experiment at different times-on-stream. These samples and also those for chars are analyzed afterward by several meth-

1112 Energy & Fuels, Vol. 11, No. 6, 1997 Table 1. Characteristics of Pinus pinaster Chips particle size (mm)

5(1

Proximate Analysis volatile matter (wt %) fixed carbon (wt %) ash (wt %) moisture (wt %)

74-76 12-13 0.5-1.2 10-14

Ultimate Elemental Analysis carbon (wt %) hydrogen (wt %) oxygen (wt %) nitrogen (wt %) sulfur (wt %)

50.0 5.7 44.1 0.1-0.3 0.03

ods. These analyses cause the time per experiment to be, on average, 1 month. As important as physicochemical results (like product distribution, yields, etc.) shown in the figures were several mechanical and engineering aspects of this pilot plant. These facts are not as easy to express, but they were very important. For instance, owing to the high temperatures of the gasifier under operation, the thermal stresses were important. This gasifier has a lot of joints (for cleaning, sampling, etc.) and seals that suffered a lot of mechanical and thermal stresses that, joined with the chemical attack (tars, for instance, are highly corrosive), caused some seals to sometimes suffer leakage of gas, leading to troubles such as small fires. These facts could perhaps not be considered as relevant for a scientific paper, but from the technical point of view and for a process engineer they are absolutely important and must be known and taken into account for future process development. Feedstock Used. The effect of the type of feedstock on the product distribution was experimentally studied in this pilot (in its former configuration) earlier. No big differences in product distribution were then found.5 So a feedstock that did not give problems in our feeding system was selected as the typical or representative one for these tests. Small pine (Pinus pinaster) chips was then selected as feedstock. Its main characteristics are shown in Table 1. Gas and Tar Sampling and Analyses. Gas and tar samples were periodically taken from the raw flue gas every 1 or 2 h. The sampling point was located after the cyclones, where it is shown in Figure 1. Gas, once sampled, was analyzed for major components (H2, CO, CO2, light hydrocarbons, N2, O2) by gas chromatography. The tar sampling device used has been previously described in detail.1 It is similar to the one used by VTT at Espoo (Finland), which, in turn, comes from a U.S. EPA norm. This tar sampling is not standardized yet in the world, and each institution is using its own method.23,27 So it is well-known and accepted that results concerning tar amounts and compositions could be somewhat different if taken in another institution,27 but the main results, trends, and conclusions here presented would remain the same and can be accepted worldwide. The tar sampling system here used has four impinger flasks of 150 cm3 each. The first two contain 300-400 cm3 in total of an emulsifier solution (water and a commercial soap which doesn’t form foams) to condense tar and water and to dissolve tar compounds in order to get only one liquid phase. The first two traps are placed in an ice bath and the second two (and empty) in another bath with dry (CO2) ice and acetone. There are a flow meter and a sucking pump in each sampling line. About 0.7 NL/min are drawn during 30-45 min through this system to get one sample. The overall solution is further mixed and slightly heated in a water bath to obtain a homogeneous sample. Total organic matter in this sample is then analyzed by a total organic carbon Dohrmann analyzer DC90, which is based on quantitation by a nondispersive infrared detector (NDIR) of the (23) BTG B.V. (Enschede, The Netherlands) reports to EC, DG XII, Brussels, for Contracts AIR-CT92-0294 and JOU2-CT93-0408 NOVEM. Utrecht (NL) Reports 9603 and 9608, January 1996.

Gil et al. amount of CO2 produced during the catalytic combustion at high temperature of the tar and emulsifier in the liquid sample. Depending on the range used (10-700 ppm and 1003500 ppm), a different amount (200 µL and 40 µL, respectively) of sample is injected. The standards of known carbon content used for the calibration of the different ranges (400 and 2000 ppm) are prepared by dissolving in water fixed amounts of potassium hydrogen phthalate, KHP (C8H5O4K). Since the solution initially inside the first two impinger flasks contains carbon coming from the soap, a blank sample is also taken and analyzed. The amount of carbon is later discounted in the samples. An averaged tar composition based on analyses made of tar and all components from majority (phenol) to minority (benzene) was used to convert TOC results (in mg C) to tar content (in mg tar), for which an averaged molecular weight for tar is needed.

Results and Discussion Products from biomass gasification are gas, condensates (water and tars, collected as liquids but in the gas phase in the gasifier), and char. Gas includes H2, CO, CO2, CH4, steam, and light hydrocarbons (C2, C3). The amounts of these compounds will be given as yields or as volumetric contents in the gas phase. In this process there are three main reactants: biomass, steam, and oxygen. To describe the reacting system, two ratios have been selected:

-(H2O + O2)/biomass (here referred as GR or R), (kg fed/h)/(kg daf fed/h) -H2O/O2, mol/mol These two ratios are enough to know what is fed into the gasifier. Besides those two parameters, three other variables have been studied: temperature in the bed of the gasifier (Tb or Tbed), temperature in the gasifier freeboard, and addition of some amount of dolomite into the gasifier bed (in the gasifier bed, mixed with the silica sand) and mixed with the biomass fed into the gasifier. The effect of this latest variable will be shown in a forthcoming paper.24 In this paper, feedstock will be pure biomass (without dolomite) and the stationary bed or fluidizing solid will be only silica sand. The effect of the gas residence time or space-velocity in the gasifier bed had been studied previously.4 Gas Composition. H2 Content in the Flue Exit Gas. The hydrogen content in the raw flue exit gas is shown in Figure 3 for different temperatures in the gasifier bed and H2O/O2 and gasifying agent-to-biomass ratios GR. H2 content in the raw gas ranges between 14 and 30 vol %. This H2 content decreases when GR increases or when H2O/O2 decreases because as the O2 fed increases, more H2 is burnt in the gasifier and less H2 is found at the exit. In biomass gasification with pure steam the H2 content increased with the H2O/biomass ratio.4,5 Now the (H2O + O2)/biomass ratio includes O2, and it clearly changes the trend of the H2 produced. The influence of the temperature of the gasifier bed on the H2 amount in the flue gas is shown in Figure 4. The increase of the H2 content with Tbed is well understood because when Tbed is increased, the rate of the biomass and tar cracking and reforming reactions (24) Aznar, M. P.; Martin, J. A.; Caballero, M. A.; Corella, J. Ind. Eng. Chem. Res., submitted.


Figure 3. Hydrogen content in the gas at the gasifier exit at different temperatures, H2O/O2, and gasifying agent-to-biomass ratios.

Figure 4. Hydrogen content in the gas at the gasifier exit vs gasifier (bed) temperature. (H2O/O2) ) 2 and 3 (mol/mol).

Figure 5. Carbon monoxide content in the gas at the gasifier exit at different temperatures, H2O/O2, and gasifying agentto-biomass ratios.

also increases, and such cracking and reforming reactions produce hydrogen. When the H2O/O2 ratio increases, the amount of H2 produced also increases, as Figure 4 shows. CO Content. The CO in the exit gas is quite high when compared with the 12-22 vol % obtained in biomass gasification with air.1 Now the CO content in the raw gas ranges between 30 and 50 vol %, dry basis, as Figure 5 shows. The variation of the amount of CO in the gas phase with the (H2O + O2)/biomass and H2O/ O2 ratios is shown in Figure 5 and its dependency with Tbed in Figure 6. When (H2O + O2)/biomass and Tbed are increased, the CO content clearly decreases. When the H2O/O2 ratio is decreased, the CO content also decreases. The effect of both ratios can be easily


Figure 6. Carbon monoxide content in the gas at the gasifier exit vs gasifier (bed) temperature. (H2O/O2) ) 2 and 3 (mol/ mol).

Figure 7. Carbon dioxide content in the gas at the gasifier exit at different temperatures, H2O/O2, and gasifying agentto-biomass ratios.

Figure 8. Carbon dioxide content in the gas at the gasifier exit vs gasifier (bed) temperature. (H2O/O2) ) 2 and 3 (mol/ mol).

explained because higher amounts of O2 burn CO to give CO2. The effect of Tbed is not easy to understand. One possible explanation could be that, if its equilibrium is not reached, the rate of the shift reaction (CO + H2O S CO2 + H2) increases with Tbed and it causes the CO in the gas to decrease at the gasifier exit. CO2 Content. The CO2 content in the raw product gas is shown in Figure 7 for different (H2O + O2)/biomass and (H2O/O2) ratios and in Figure 8 for different temperatures in the gasifier bed. The amount of CO2 in such a gas is between 14 and 37 vol % dry basis. When GR or (H2O + O2)/biomass ratio is increased and (H2O/O2) is decreased, the CO2 produced increases


Gil et al.

Figure 9. Methane content in the gas at the gasifier exit at different temperatures, H2O/O2, and gasifying agent-to-biomass ratios.

Figure 11. Light hydrocarbons (C2) content in the gas at the gasifier exit at different temperatures, H2O/O2, and gasifying agent-to-biomass ratios.

Figure 10. Methane content in the gas at the gasifier exit vs gasifier (bed) temperature. (H2O/O2) ) 2 and 3 (mol/mol).

Figure 12. Light hydrocarbons (C2) content in the gas at the gasifier exit vs gasifier (bed) temperature. (H2O/O2) ) 2 and 3 (mol/mol).

because, of course, increasing the amount of O2 fed the combustion reaction increases with the subsequent formation of CO2. Somewhat more difficult to explain is why the CO2 content at the gasifier exit decreases with Tbed, as Figure 8 shows. To our knowledge, it can be due to the fact that the rates of the dry (with CO2) reforming reactions of CH4, light hydrocarbons, tars, and biomass increase with temperature. All of them consume CO2, thus decreasing its amount at the gasifier exit. The CO-shift reaction mentioned above would produce simultaneously CO2, but results shown in Figure 8 seem to indicate that in the gasifier, when Tbed is increased, the CO2-consuming reactions would be more important than the shift reaction. No other explanation is found for the overall result concerning CO2, shown in Figure 8. CH4 and Light Hydrocarbons in the Gas Phase. The CH4 in the exit gas, under the experimental conditions used here, ranges between 5 and 7.5 vol %, dry basis (Figures 9 and 10) and the C2 compounds between 2.2 and 3.8 vol % (Figures 11 and 12). The trends shown in these figures are similar for CH4 and for C2 hydrocarbons. So they will be discussed together. The content in CH4 and C2 hydrocarbons in the exit gas decreases when the (H2O + O2)/biomass is increased or when H2O/O2 decreases, as Figures 9 and 11 indicate, because O2 burns and H2O reforms (destroys) in-bed some amounts of CH4 and C2. These two reactions (O2burning and H2O-reforming) would also explain the effect of Tbed shown in Figures 10 and 12. When Tbed is increased, the rate of several reactions (combustion,

Figure 13. (H2/CO) at the gasifier exit at different temperatures, H2O/O2, and gasifying agent-to-biomass ratios.

steam, and dry reforming, cracking, ...) increases, which eliminates CH4 and light hydrocarbons from the flue gas. H2/CO and CO/CO2 Ratios in the Flue Exit Gas. The H2/CO ratio in the exit gas is important for further possible end uses of this gas. It is shown in Figures 13 and 14 for several values of GR, H2O/O2, and Tbed. This H2/CO ratio ranges between 0.28 and 0.75, which is very low for methanol or ammonia production. Nevertheless, it can be said in advance (it will be shown in a subsequent paper) that this ratio in this process can be easily increased to 2.0 with a downstream catalytic bed. By the explanations given in previous paragraphs concerning H2 and CO2 production, the H2/CO ratio


Figure 14. (H2/CO) at the gasifier exit vs gasifier (bed) temperature. (H2O/O2) ) 2 and 3 (mol/mol).

Figure 15. (CO/CO2) at the gasifier exit at different temperatures, H2O/O2, and gasifying agent-to-biomass ratios.

Figure 16. (CO/CO2) at the gasifier exit vs gasifier (bed) temperature. (H2O/O2) ) 2 and 3 (mol/mol).

increases with Tbed and H2O/O2 (Figure 14) and decreases with increasing GR [(H2O + O2)/biomass] (Figure 13). The CO/CO2 ratio obviously depends on the CO and CO2 formed whose amounts were previously discussed. This CO/CO2 ratio is above 1.0 in this process and its detailed values are shown in Figures 15 and 16. The CO/CO2 ratio increases with Tbed and with the H2O/O2 ratio (Figure 16) and decreases when GR [H2O + O2)/biomass] is decreased (Figure 15). Lower Heating Value of the Raw Gas. The lower heating value (LHV) of the raw gas depends on its composition, which in turn depends on the H2O/O2 and GR ratios as has been clearly shown in proceedings paragraphs. The LHV values are shown in Figure 17


Figure 17. Low heating value of the dry gas at the gasifier exit vs gasifying agent-to-biomass ratio. (H2O/O2) ) 2 and 3 (mol/mol).

Figure 18. Steam content in the flue gas at the gasifier exit vs gasifying agent-to-biomass ratios. (H2O/O2) ) 2 and 3 (mol/ mol).

for several H2O/O2 and GR ratios. They vary between 11 and 16 MJ/Nm3 dry gas. When the H2O/O2 ratio decreases, the LHV decreases (Figure 17) because more O2 is fed, which burns in situ some gas components. When GR (for a given H2O/O2 ratio) is increased, the LHV decreases by the same reason (more O2 is fed). Steam Content. The steam content in the raw gas is an important variable. For instance, if the tar present in such a gas is going to be destroyed with a downstream bed of a calcined dolomite or a nickel-based steamreforming catalyst, the effectiveness of such tar destruction depends on the steam content in the gas phase because it is an important reactant for tar removal reactions. On the other hand, not much steam should be present in the raw gas because it would imply a lost of energetic efficiency (heat is given to the overall gasification process to evaporate the water and produce such steam). The steam content in the gas phase is quite high in this process. It varies between 32 and 60 vol % (Figure 18). It increases on increasing the H2O/ O2 and GR ratios as Figure 19 also shows. The steam content in the gas at the gasifier exit decreases somewhat with Tbed (Figure 18). It can be easily understood because steam is one main reactant and its overall rate of disappearance (by several reactions) increases with temperature. Gas Yield. The gas yield, very important datum in a gasification process, varies between 0.85 and 1.2 Nm3 (dry gas)/kg biomass fed, as Figure 20 shows. These values are low because they are referred to as dry basis


Gil et al.

Figure 19. Steam content in the flue gas at the gasifier exit vs gasifier temperature. (H2O/O2) ) 2 and 3 (mol/mol).

Figure 22. Apparent thermal efficiency of the gasifier vs gasifier (bed) temperature. (H2O/O2) ) 2 and 3 (mol/mol).

Figure 20. Gas yield at the gasifier exit at different temperatures, H2O/O2, and gasifying agent-to-biomass ratios.

Figure 23. Apparent thermal efficiency of the gasifier at different temperatures, H2O/O2, and gasifying agent-to-biomass ratios.

Apparent Thermal Efficiency. An apparent thermal efficiency is defined as energy in the gas to energy in the biomass fed. It is calculated as

LHV (MJ/Nm3, dry gas) × gas yield (Nm3, dry gas/kg biomass daf fed) × 100 LHV of feed (MJ/kg biomass daf fed)

Figure 21. Gas yield at the gasifier exit vs gasifier (bed) temperature. (H2O/O2) ) 2 and 3 (mol/mol).

(and in this process most of the gasifying agent, or flue gas at the inlet, is steam that is not taken into account in the gas yield so defined). The gas yield, dry basis, increases with the (H2O + O2)/biomass ratio as Figure 20 indicates for several circumstances. The gas yield also increases with Tbed as Figure 21 shows for two H2O/O2 ratios. Trends shown in Figures 20 and 21 are easily admitted. Notice that in Figure 21 nothing is said about the heat supplied for the gasification. Higher H2O/O2 ratios give higher gas yields, but such a fact requires more external heat supply (from an external oven) or working with lower temperatures in the bed as Figure 21 indicates (when H2O/O2 ) 3, the temperature in the gasifier bed working at autothermal conditions is 785-830 °C; when H2O/ O2 ) 2, Tbed is 830-890 °C).

In our pilot plant the approximate thermal efficiency ranges between 60 and 97%. This apparent thermal efficiency increases with Tbed as Figure 22 shows. Higher temperatures in the gasifier bed allow higher thermal efficiencies. Notice from Figure 22 how (with the main external oven off) with (H2O/O2) ) 2 a higher Tbed is obtained than when (H2O/O2) ) 3, this fact being due to the higher amount of O2 fed in the first case when (H2O/O2) ) 2. On the other hand, this thermal efficiency decreases with GR, as Figure 23 indicates, for several values of Tbed and of H2O/O2. For the stationary state (after a few hours from the starting up period) the gasifier operated most of the time with the external oven off. Only a preheater to vaporize the fed water was on. That is to say, for the stationary state not much external heat was required to get the desired value for Tbed. It means that the apparent thermal efficiency shown in Figures 22 and 23 is very close to a real thermal efficiency. Of course, this gasification plant and process is near autothermal thanks to the heat generated by the O2 added to the feed. Its cost would remain its main drawback.


Figure 24. Char produced (and collected in the cyclones) vs gasifying agent-to-biomass ratio. (H2O/O2) ) 2 and 3 (mol/mol).

Figure 25. Char produced (and collected in the cyclones) vs gasifier (bed) temperature. (H2O/O2) ) 2 and 3 (mol/mol).

Char Yield. Char is another main product from biomass gasification, and its yield should be as small as possible to get a good gasification process. Char was collected in every experiment/test in two points: in the gasifier bed (mixed with the sand) and in the vessels located at the bottom of the three cyclones. Since the hydrodynamics in all experiments were always nearly the same, their influence on the char yield was always the same; nearly all char produced was elutriated out the gasifier because of its very low density. The amount collected in the cyclones was always much bigger than the amount collected in the gasifier bed. For this reason, the char yield shown here means the char collected in the cyclones. The figures given here for char yield would be thus somewhat lower than the true yields. Char yield decreases with the gasifying agent-tobiomass ratio as Figure 24 shows. To get a char yield less than 10%, that is to say to get a good gasification, the value for GR has to be higher than 1.0. The bed temperature, in the level used here, does not influence much the char yield, as Figure 25 shows for several values of GR. About the effect of the H2O/O2 ratio, values of 2 for this ratio produce more char than values of 3, as it is deduced from Figures 24 and 25. Tar Content in the Raw Gas. Tar Composition. Since the authors are not expert in determining tar composition, some samples of the condesates were sent to NREL in Golden, CO, and the KTH in Stockholm, Sweden. At the NREL our tar was characterized by molecular beam mass spectrometry (MBMS), a novel and promis-


Figure 26. Average spectrum for a typical sample of our condensates at the gasifier exit characterized by molecular beam mass spectrometry (MBMS) at NREL.

Figure 27. Tar compounds distribution at the gasifier exit (run 25, R ) 1.08, H2O/O2 ) 3, Tb ) 825 °C).

ing approach for molecular characterization of tars developed and used at NREL by Evans and Milne.25 The average spectrum for a typical sample of condensates from our gasifier is shown in Figure 26. Phenol94 and cresol108 are the main detected compounds in our tar. Some samples of our condensates were also analyzed at the KTH.26 An example of the results there obtained is shown in Figure 27. These results correspond to experiment 25 made at Tbed ) 825 °C, GR ) 1.08, and (H2O/O2) ) 3. Two types of tars were detected: phenolics and nonphenolic aromatics. In the first fraction, phenol and cresols are the most abundant. In the second fraction, naphthalene, indene, and toluene are most abundant. Notice how even these two excellent laboratories were not able to identify an important fraction of the tars (many peaks in Figure 26 and fractions called unknown in Figure 27). Tar Content. The tar content in the raw gas clearly decreases with the gasifying-to-biomass ratio as shown in Figure 28. To get tar contents less than 10 g/Nm3, values for GR greater than 1.0-1.1 have to be used, as Figure 27 indicates. The effect of the gasifier bed temperature on the tar content is shown in Figure 29. For (H2O/O2) ) 3, when Tbed is increased, the tar content decreases. This effect is more important at low (1.2) values of GR the effect of Tbed does not seem important, as Figure 29 shows. Results shown in Figures 28 and 29 indicate how, in the gasifier used here (working without a secondary flow inlet, with relatively low freeboard temperatures, and with only silica sand, without dolomite, as the stationary bed), it is very difficult to get tar contents below 5-10 g/Nm3, dry basis. This level of tar is, for us, a “quite” clean gas but not a clean “enough” gas. Several applications of this raw gas could need its upgrading or cleaning for tars. Such upgrading and cleaning of the gas has been already studied in this pilot and will be the objective of forthcoming papers. Conclusions A pilot plant for fluidized bed gasification of biomass is being operated at the University of Saragossa (Spain). Although other types of biomass and gasifying agents are sometimes used, it was operated only for this work with pine wood chips (5-20 kg/h) and H2O/O2 mixtures. Three parameters have been exhaustively studied: gasifying agent (H2O + O2) to biomass ratio, steam-tooxygen fed ratio, and temperature of the gasifier bed. For (H2O/O2) ) 3 the gasifier used here works as an autothermal reactor. For gasification with steam and (28) Olivares, A.; Aznar, M. P.; Caballero, M. A.; Gil, J.; Corella, J. Making a Business from Biomass, Proceedings of the 3rd Biomass Conference of the Americas, Montreal, Canada, August 1997; Vol. 1, p 765 and addendum.

About the tar produced, its main components are phenol, cresol, naphthalene, indene, and toluene. The tar content in the raw gas varied between 2 and 50 g/Nm3, dry gas. Although a “good gasification” can be understood to depend on the main objectives or interests, the authors consider that the best and recommended operating condition for the process here studied are the following.

gasifier bed temperature: between 800 and 860 °C (H2O + O2)/biomass: from 0.8 to 1.2 (kg/ kg daf) (H2O/O2): around 3.0 (mol/mol) gas residence time (in the gasifier bed): around 2 s With these conditions a “quite” clean gas (tar contents of around 5 g/Nm3) is obtained. To get a flue hot gas with a lesser tar content (or a higher H2 content), dolomite in bed and/or catalytic beds downstream should be used.24,28 Acknowledgment. This work has been carried out thanks to the financial support of the European Commission, JOULE 2 and JOULE 3 programs, Contracts No. JOU2-CT93-0399 and JOR3-CT95-0053. Thanks are given to Jose´ Prieto (Pepo) et al. of the ICP of the CSIC (Madrid) for their help in developing and setting up the advanced control system used in the pilot plant. We also like to express our gratitude to Dr. K. Sjo¨strom et al. from the KTH of Stockholm (Sweden) and to Dr. Bob Evans from the NREL for their valuable analyses of our condensates. Glossary GR Tb, Tbed umf

gasifying ratio, steam and oxygen-to-biomass ratio, kg/kg daf temperature in the bed of the gasifier, °C minimum fluidization superficial gas velocity, cm/s EF9602335

Biomass Gasification in Fluidized Bed at Pilot Scale with Steam

Recommend Documents