- Email: [email protected]
Catalytic gasification of biomass to produce hydrogen rich gas
ht. J. Hydrogen Energy, Vol. 23; No. I, PP. 551-557, 1998
Q 1998International Association for Hydrogen Energy Elsevier ScienceLtd All rights reserved. Printed in Great Britain PII: SO360-3199(97)00108-0 0360-3199198$19.00+0.00
Pergamon
CATALYTIC
GASIFICATION
S. RAPAGNA,” Dipartimento
OF BIOMASS RICH GAS
TO PRODUCE
HYDROGEN
N. JAND and P. U. FOSCOLO
di Chimica, Ingegneria Chimica e Materiali,
University
of L’Aquila,
67100 L’Aquila,
Italy
Abstract-Catalytic biomass steam gasification runs were performed in a bench scale plant consisting essentially of a fluidised bed gasifier and a secondary catalytic fixed bed reactor. This secondary reactor employed alternatively two different steam reforming catalysts and calcined dolomite. The operating conditions in the gasifier (temperature, biomass/steam ratio and biomass feed rate) were kept constant for all the runs (77O”C, 1 and 764 kg of biomass per hour per m3 of bed, respectively). The influence of the operating conditions in the catalytic converter on the production of gases, especially H,, was investigated over the temperature range of 660-83O”C, for Gas Hourly Space Velocities
(GHSV) in the range 9000-27,700 h-‘. About 2 m3 of dry gas (at ambient conditions) per kg of daf biomass were obtained by utilising the fresh catalyst at the highest temperature level, with more than 60% by volume being hydrogen. The lowest tar residue was 0.45 g/kg of daf biomass, which increased slightly over the three hours gasification time. Substantial carbon deposition was observed, mainly on the catalyst layers contacting the inlet gas. On the basis of these results, a process configuration suitable for industrial applications is discussed. 0 1998
International Association for Hydrogen Energy
INTRODUCTION The transformation of biomasses into hydrogen rich gas provides a competitive means for producing energy and chemicals from renewable resources [l]. Such a gas can be utilised in fuel cell units for electricity production at high efficiency levels (40% electrical efficiency and 85% overall thermal efficiency, including steam production). This application is dependent on tar elimination and CO conversion to levels less than 1% by volume by means of the water gas shift reaction [2]. Alternatively, the gas can be used as hydrogen source for refinery hydrotreating operations, for ammonia production [3], and, with suitable CO concentrations, for methanol and FisherTropsch synthesis. As reported extensively in the literature, biomass steam gasification (performed in the main in fluidised bed reactors) results in the conversion of carbonaceous materials to permanent gases (H,, CO, CO,, CH4, light hydrocarbons), char and tar. A schematic representation of the reaction steps involved in this process is given in Fig. 1. To render the gas amenable to commercial applications, tar components need to be reduced to a minimum (less than 1 g/m’ of gas at ambient conditions). This
* Author to whom all correspondence should be addressed.
helps to avoid corrosion problems and poisoning of the catalysts employed to perform the water gas shift reaction. Moreover, in this way more gas is obtained and the efficiency of the overall gasification process is increased. Much research has addressed this subject [414]; nickel based commercial steam reforming catalysts have been utilised together with, or as an alternative to, calcined dolomite. Different process configurations have also been examined in which the catalytically active substances are either placed directly in the gasifier, or in a secondary fixed or a fluidised bed reactor. A major problem has been recognised in the rapid loss of the catalytic activity due to the fouling by carbon build-up on the surface of the catalyst itself. Such deactivation can be partially overcame by utilising catalyst in the fluidised state and in a secondary reactor [4, 51;results obtained, however, show that the total gas yield under these conditions after several hours of operation is considerably lower than that at the start of the test: tar conversion is far for complete and heavy components are still present in the product gas. The kinetics of catalyst deactivation depend on many factors, such as catalyst type, bed temperature, gas residence time, steam/biomass ratio, catalyst particle size and, above all, the tar content of the raw gas. This in turn depends on the type of feedstocks [15], the gasification temperature, the steam/ biomass ratio and the biomass particle size. It is worth 551
S. RAPAGNA et al.
552
+ Hz, CO,COz, CH4, light and heavy hydrocarbons
biomass particle Fig. 1. Schematicrepresentationof the biomassgasificationprocess.
pointing out that the yield and composition of the primary product formed in the devolatilization step (Fig. 1) are related to the heating rate of the biomass particles: high heating rates produce more light gasesand less char and condensate [16]. Experiments performed by the authors with different biomass particle sizesin the range 2741090 pm, show that at the same temperature level (700-75O”Q the smallest particles produce about 0.5 m3 more gas per kg of biomass than is obtained with the biggest particles [ 171.For small particle sizesthe devolatilization process is mainly controlled by reaction kinetics; as the particle size increases,kinetic control gives way to heat transfer control. In this study catalytic gasification tests were performed with two nickel steam reforming catalysts (one spent and one fresh), as well as with calcined dolomite. The initial average diameter of the biomass particles was 1.1 mm. The aim of these experiments was to characterise the influence of operating parameters on the catalytic transformation of tar components into CO and H, in presence of steam, so as to obtain useful data for the design of industrial units. There is general consensus on the effectiveness of nickel catalysts for tar reforming, but suitable operating conditions, such as temperature and space time, need further assessment. The decision to test an exhaust nickel based catalyst has been dictated by economic considerations, which take into account the high cost of fresh catalyst and its life expectancy with the present processes[14]. APPARATUS AND EXPERIMENTAL
METHOD
Figure 2 provides a schematic representation of the bench scale facility utilised in this study. Its major components are: biomass feeding system, fluidised bed gasifier, catalytic fixed bed, gas cooling system, metering and analysing for the off-gases. The fluidised bed gasifier consists of an austenitic stainless steel cylindrical vessel of internal diameter 62 mm fitted with an alumina porous distributor plate, designed to allow for a good gas distribution at all temperatures. The bed is located in a cylindrical Carbolite Furnace provided with temperature and heating rate control systems. The temperatures in the reactor are measured by meansof two thermocouples, one immersed in the bed and the other located under the distributor. The bed inventory is sand particles (d, = 348 ,um, pp = 2640 kg/m3). Water for the generation of steam (the fluidising
gas) is fed to an electrically heated boiler by means of a peristaltic pump at a constant flow rate. The biomass feeding probe (of internal diameter 8 mm and cooled by means of air flowing through a jacket) is designed to deliver the biomass well inside the bubbling bed. The secondary catalytic fixed bed reactor (of internal diameter of 60 mm) is located after a ceramic candle filter and heated by means of a cylindrical electric furnace. The temperature is measured with a thermocouple located inside the reactor which is filled with catalyst: either one of the two nickel catalysts or calcined dolomite. Some further runs with calcined dolomite were performed by placing it directly in the gasifier, in place of the sand particles. Gas products were analysed by two on-line gas chromatographs: the first was fitted with a Carbosieve Sll column (1.5 m, 3 mm l.D) with helium as carrier gas, to detect Nz, COz, CH4, C,H, and C2H4(if any); the second was fitted with a Molecular Sieve 5A column (2 m, 4 mm I.D.) with argon as carrier gas, to detect mainly HZ, and also N,, CO and CH,. For each gasification run the permanent gas yield was measured by means of a volumetric gasmeter, after separation of the condensate (water and organic phases).The total quantity of tar was obtained by summing two contributions: the tar in the organic phase and the dissolved tar in the water phase. The organic phase of the condensate (a very minor quantity when the catalytic conversion was carried out) was separatedand weighted. The water soluble tar was quantified by means of TOC (total organic carbon) analysis with the assumption that naphthalene could be regarded to be the representative compound of the soluble tar. The size distribution and mean particle sizes of the catalyst utilised in the reactivity tests are reported in Table 1. To obtain information on the reduction temperatures for the two catalysts, a TPR (Temperature Programmed Reduction) test was performed. This consists of measuring the quantity of hydrogen consumed with 50 mg of catalyst heated from ambient to 900°C at the rate of lS”C/min, with a gas (3% H, in Argon) flowing at a rate of 0.05 l/min. The results obtained with this test are reported in Fig. 3. From these results the consumption of H, has been estimated to be 0.96 * 10m3and 3.8 * 10e31 for exhaust and fresh catalyst respectively. Runs were performed with the spent catalyst at temperatures of 75O”C-770°C with different quantities of catalyst (100, 200 and 330 g) supported between two
CATALYTIC
GASIFICATION
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nitrogen air water Fig. 2. Bench scale gasification plant. 1-fluidised bed reactor, 2-gas distributor, 3-wind box, &electric furnaces, 5-biomass feeder, 6--cyclone, 7--ceramic filter, g--cooling system, 9-gas flowmeter, IO-gas chromatographs, 1l-pump, 12-water man-
ometer, 13--catalytic fixed bed reactor.
Table 1. Size distribution
of the two catalysts
Spent catalyst Size, pm
% wt
Fresh catalyst % wt
> 1000 1000-850 850-710 710-600 60&500 500-425 < 425
0.39 30.62 31.16 16.33 17.02 4.19 0.29
0.1 27.86 30.20 17.16 20.40 4.2 0.18
705
719
Mean diameter, pm
results. After the maximum bed temperature was reached, the catalytic bed was left to cool down to the desired temperature level, the nitrogen-hydrogen feed was
TF’R of exhaust
--
catalyst
Q:!yyfJ--~;
layers of wool quartz in the reactor vessel. Before each run the catalyst heated to 850°C at a constant rate of 4”C/min in a 5% H2 in N2 atmosphere in order to enable the reduction to take place. The quantity of hydrogen utilised for the reduction process is well above that required by the catalyst, as determined from the TPR
Q
200
400
600
800
1000
Temperature, “C Fig. 3. Hydrogen consumption during the reduction of the two catalysts, as a function of temperature.
554
S. RAPAGNA
switched off, and biomass gasification was started in the gasifier. With the fresh catalyst only one quantity has been tested (200 g), and the runs were carried out at three temperature levels: 665”C, 760°C and 830°C. The reduction procedure for the catalyst was again as described above. The tests with calcined dolomite were carried out by placing it either in the fixed bed or directly in the gasifier. The dolomite was first crushed and sieved to obtain a fraction with a mean particle size of 454 pm, and then calcined in air at 900°C for ten hours. The weight loss during calcination amounts to about 40% of the original weight: as a result the final density is about 1500 kg/m3. These values of particle size and density mean that the minimum fluidisation velocity fot the gasifier is of the same order as that obtained with the sand particles bed (E 0.085 m/s). Under these conditions the residence times of char and gas in the gasifier are comparable to those resulting from tests carried out with sand. RESULTS
AND
et al.
(Gas Hourly Space Velocity: gas flow rate at the exit of the catalytic reactor divided by the bulk volume of the catalyst bed) and the WHSV (Weight Hourly Space Velocity: mass flow rate of biomass fed to the gasifier, divided by the mass of catalyst in the reactor). The tar content of the gas leaving the catalytic bed (expressed as g of tar per kg of biomass) is reported in Figs 4 and 5, as a function of the gasification time, for exhaust and fresh catalyst respectively. The catalytic tests clearly show the gas yield to be considerably higher than from tests performed without catalyst (Tables 2 and 3): it effectively doubles when fresh catalyst is utilised. Moreover the gas composition obtained in this case is close to the equilibrium conditions for the water gas shift reaction (wgsr) as shown on Fig. 6, where the theoretical value has been calculated from
3%
DISCUSSION
The operating conditions in the gasifier were kept constant for all runs and are reported in Table 2. The total gas yields obtained with the two catalysts are reported in Table 3, together with the gas composition (a function of the catalytic bed temperature), the GHSV
Table 2. Operating conditions of the gasifier with the corresponding gas and tar yields Bed inventory, sand Biomass, almond shells Bed temperature Biomass feed rate Steam input rate Total gas yield Tar yield
dp = 348pm, pp = 2640 kg/m3 d, = 1.1 mm, pP = 1200 kg/m3 770°C 0.3 kg/h 0.3 kg/h 1 Nm3/kg of biomass 0.1 kg/kg of biomass
Table 3. Gas yields and composition
50
0
200
Gasification time, minutes Fig. 4. Tar yield expressed in g/kg of biomass in the produced gas, as a function of the gasification time for the spent catalyst. GHSV: (m) 27,700 h-l, (0) 14,400 h-l, (A) 9000 h-l.
obtained with the spent and the fresh catalysts, as a function of the catalytic bed temperature and quantities of catalyst utilised Fresh catalyst
Spent catalyst Temperature of the catalytic fixed bed WHSV, h-’ GHSV, h-’ Gas yield, normal m3 per kg of daf biomass Gas composition HZ co co2 ‘W
150
100
750 2.70 27,700
770 1.35 14,400
760 0.84 9000
665 1.35 13,850
760 1.35 15,650
830 1.35 18,000
1.35
1.40
1.50
1.80
1.88
1.98
in % by volume, 56.2 25.2 12.6 6.0
55.7 25.0 12.8 6.5
56.0 24.7 13.3 6.0
62.2 22.8 15.0
61.7 23.5 14.8
62.1 22.7 15.2
CATALYTIC
0
50
100
150
GASIFICATION
200
Gasification time, minutes Fig. 5. Tar yield expressed in g/kg of biomass in the produced gas, as a function of the gasification time for the fresh catalyst. GHSV: (m) 13,850h-‘, (0) 15,650h-l, (A) 18;OOO h-‘.
known values of the equilibrium constant for the wgsr and the experimental values of CO, CO, and H,O volumetric fractions. Figure 6 also shows that the hydrogen yield, although lower than for the corresponding active catalyst case, follows the same trend for operation without catalyst and with exhaust catalyst. Moreover with fresh catalyst, CH, is also completely
650
700
750
800
850
Temperature, “C Fig. 6. Ratio of the theoretical and experimental values of hydrogen present in the produced gas as a function of fluidised bed temperature: absence of catalysts (a), fresh catalyst (0) and spent catalyst (0).
OF BIOMASS
555
transformed into CO and H,, resulting in a hydrogen composition in the final product higher than 60% by volume. This value is compatible with the requirements imposed by the fuel cells on the feeding gas (i.e. Phosphoric Acid Fuel Cells), provided that a conventional CO shift operation is included in the process scheme. Figure 7 reports the total gas yield as a function of time, when dolomite is used, both in the gasifier itself instead of sand, and in a secondary fixed bed reactor. The continuous increase of gas yield with time for the case of dolomite in the gasifier can be explained on the basis that char accumulating in the bed is catalytically gasified in the presence of alkaline compounds. A major drawback of this procedure, however, is that it results in comminution of the bed inventory which gives rise to a very high production of fines-with attendant gas cleaning problems. On the other hand, the use of dolomite in a secondary reactor results in an increase in the overali gas yield over that reported in Table 2, and of the same order as that obtained with exhaust catalyst (although at a higher temperature level). The major problem encountered in the experimental tests with catalysts and dolomite in the fixed bed gas converter is related to the continuous increase of gas pressure drop due to carbon deposition especially on the layer firstly encountered by the gas. The formation of carbon in methane and naphtha steam reforming is well documented [18]. To avoid this phenomenon high steam/hydrocarbon ratios (in the range 3.44.5) are usually employed. In the case examined here, this ratio is much higher, so that the formation of carbon is not thermodynamically favoured. As a result the presence of carbon cannot be ascribed to the decomposition of CO by means of the Bouduard mech-
0
20
40
60
gasification time, minutes Fig. 7. Gas yield obtained by using dolomite directly in the gasifier (0) at temperature of 770°C and in the fixed bed reactor (0) at temperature of 81O’C.
556
S. RAPAGNA et al.
anism, but is mainly due to the cracking and polymerisation reactions involving the heavy hydrocarbons. The addition of alkaline metals to the catalyst reduces the rate of carbon accumulation, but cannot eliminate it completely [4, 11, 181. This all points to a process configuration which should allow for frequent, if not continuous, catalyst regeneration. A further problem relates to the presence of sulphur in the biomass materials. The sulphur content can vary considerably with the type of feedstock and is usually very low; some sulphur is however always present as it is needed in the growing process of proteins. Sulphur is a poison for nickel catalysts. Evidence on catalyst deactivation by sulphur is not reported in the literature on catalytic treatment of the gas produced from biomass gasification; this may be due either to the low concentration of sulphur compounds in such gas, or to the fact that experience in this field is mainly confined to experimental studies over limited periods of time. Only experience with catalysts undergoing numerous regeneration cycles, as previously foreseen, will enable the importance of this phenomenon to be assessed. Finally in order to perform the catalytic reforming in industrial sizeunits, a major point to be considered relates to the implementation of an efficient and easy to operate system for supplying the heat requirement of the endothermic reactions involved. As the gasification process itself is also endothermic the thermal transfer problem is of great relevance for the overall plant feasibility. This problem is at present the subject of a joint research project financed by the EU (Contract JOR3-CT950037):
a dual fluidised bed configuration is employed with internal circulation of the bed material (including the reforming catalyst) between the gasification zone and a combustion zone which provides the necessary heat input for the biomass thermal conversion and tar reforming. Gas leaving the gasification section is prevented from contacting the combustion products, so as to maintain it free of nitrogen and with high calorific value [19]. This configuration also addresses the problem of catalyst deactivation: carbon deposited on the particle surfaces is burnt off in the combustion zone. CONCLUSIONS The experimental results show that fresh, nickel based
catalysts are extremely active for the elimination of CH, and tars: about 2 m3 of gas (at ambient conditions) are obtained per kg of biomass, with a H, content higher than 60% by volume. A gas of this specification is suitable
for industrial applications, both for highly efficient electricity production and as a feedstock for chemical synthesis. However, problems of coke formation and heat requirements do not favour fixed bed operation in industrial applications. A new configuration involving a dual action fluidised reactor has been proposed and is currently being studied to render the process industrially feasible. Acknowledgements-The
financial support of the Consiglio
Nazionale delle Ricerche and the EU under the Contract JOR3CT95-0037 is gratefully acknowledged.
REFERENCES 1. McKinley, K. R., Browne, S. H., Neill, D. R., Seki, A. and Takahashi, P. K., Hydrogen fuel from renewable resources. Enecqy Sources, 1990, 12, 105. 2. Cleghorn S., Ren X., Springer T., Wilson M., Zawodzinski C., Zawodzinski T. and Gottesfeld S.. Pem fuel cells for transportation and stationary power generation applications. In: Veziroglu T. N., Winter C. J., Baselt J. P. and Kreysa G. (Eds). Hydrogen Energy Progress, Proceedings of the 11th World Hydrogen Energy Conference, Stuttgart, 233 28 June 1996, 1637. 3. Production of Ammonia, European Fertiliser Manufactures Association; Booklet No 1, CU/E-96.22, 1, 1996. 4. Baker, E. G., Mudge, L. K. and Brown, M. D., Steam gasification of biomass with nickel secondary catalysts. 2nd. Eng. Chem. Res., 1987,26, 1335. 5. Aznar, M. P., Corella, J., Delgado, J. and Lahoz, J., Improved steam gasification of lignocellulosic residues in a fluidized bed with commercial steam reforming catalysts. Ind. Eng. Chem. Res., 1993,32, 1. 6. Kurkela, E., Stahlberg, P., Laatikainen, J. and Simell, P., Development of simplified IGCC-processes for biofuels: supporting gasification research at VTT. Bioresource Technology, 1993,46, 37. 7. Simell, P. A., Jukka, K., Leppalahti, J. and Bredenberg, J.
B., Catalytic purification of tarry fuel gas with carbonate rocks and ferrous materials. Fuel, 1992,71, 211. 8. Kinoshita, C. M., Wang, Y. and Zhou, J., Effect of reformer conditions on catalytic reforming of biomass-gasification tars. Ind Ena. Chem. Res., 1995,34. 2949. 9. Gebhard, S C., Wang, D., dverend, R. P. and Paisley, M. A., Catalytic conditioning of synthesis gas produced by biomass gasification. Biomass and Bioenergy, 1994,7, 307. 10. Magne P., Donnot A. and Deglise X., Dolomite as a catalyst of tar cracking: comparison with industrial catalysts. In: Grassi G., Gosse G. and dos Santos G. (Eds). Biomass for Energy and Industry, 5th EC Conference, Lisbon, 9-13 October 1989,259O. 11. Arauzo, J., Radlein, D., Piskorz, J. and Scott, D. S., Catalytic pyrogasification of biomass. Evaluation of modified nickel catalysts. Ind Eng. Chem. Res., 1997,36, 67. 12. Vassilatos, V., Taralas, G., Sjostrom, K. and Bjornbom, E., Catalytic cracking of tar in biomass pyrolysis gas in the presence of calcined dolomite. The Can. J of Chem. Eng., 1992,70, 1008. 13. Delgado, J., Aznar, M. P. and Corella, J., Biomass gasification with steam in fluidized bed: effectiveness of CaO, MgO, and CaO-MgO for hot raw gas cleaning. Znd. Eng. Chem. Res., 1997,36, 1535. 14. Lammers G., Orio A. and BeenackersA. A. C. M., Catalytic tar removal from biomass producer gas with in situ catalyst regeneration. In: Chartier P., Ferrer0 G. L., Henius U. M., Hultberg S., Sachau J. and Wiinblad M. (Eds). Biomass for Energy and the Environment, Proceedings of the 9th European Bioenergy Conference, Copenhagen, 24-27 June 1996, 1416. 15, Herguido, J., Corella, J. and Gonzalez-S&z, J., Steam gasi-
fication of lignocellulosic residues in a fluidised bed at a small pilot plant. Effect of the type of feedstock. Ind. Eng. Chem. Res., 1992,31, 1274. 16. Di Blasi, C., Kinetic and heat transfer control in the slow and floash pyrolysis of solids. Ind. Eng. Chem. Res., 1996, 35, 37.
CATALYTIC
GASIFICATION
17. Rapagna S. and Latif A., Steam gasification of almond shells in a fluidised bed reactor: the influence of temperature and particle size on product yield and distribution. Biomass and Bioenergy, in press. 18. Twigg M. V., Catalyst Handbook, 2nd edition, Wolfe Publishing, 1989.
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19. Foscolo P. U., Production of hydrogen-rich gas by biomass gasification: application to small scale, fuel cell electricity generation in rural areas. EU Annual Report, Contract JOR3-CZ95-0037, January 1997.
Q 1998International Association for Hydrogen Energy Elsevier ScienceLtd All rights reserved. Printed in Great Britain PII: SO360-3199(97)00108-0 0360-3199198$19.00+0.00
Pergamon
CATALYTIC
GASIFICATION
S. RAPAGNA,” Dipartimento
OF BIOMASS RICH GAS
TO PRODUCE
HYDROGEN
N. JAND and P. U. FOSCOLO
di Chimica, Ingegneria Chimica e Materiali,
University
of L’Aquila,
67100 L’Aquila,
Italy
Abstract-Catalytic biomass steam gasification runs were performed in a bench scale plant consisting essentially of a fluidised bed gasifier and a secondary catalytic fixed bed reactor. This secondary reactor employed alternatively two different steam reforming catalysts and calcined dolomite. The operating conditions in the gasifier (temperature, biomass/steam ratio and biomass feed rate) were kept constant for all the runs (77O”C, 1 and 764 kg of biomass per hour per m3 of bed, respectively). The influence of the operating conditions in the catalytic converter on the production of gases, especially H,, was investigated over the temperature range of 660-83O”C, for Gas Hourly Space Velocities
(GHSV) in the range 9000-27,700 h-‘. About 2 m3 of dry gas (at ambient conditions) per kg of daf biomass were obtained by utilising the fresh catalyst at the highest temperature level, with more than 60% by volume being hydrogen. The lowest tar residue was 0.45 g/kg of daf biomass, which increased slightly over the three hours gasification time. Substantial carbon deposition was observed, mainly on the catalyst layers contacting the inlet gas. On the basis of these results, a process configuration suitable for industrial applications is discussed. 0 1998
International Association for Hydrogen Energy
INTRODUCTION The transformation of biomasses into hydrogen rich gas provides a competitive means for producing energy and chemicals from renewable resources [l]. Such a gas can be utilised in fuel cell units for electricity production at high efficiency levels (40% electrical efficiency and 85% overall thermal efficiency, including steam production). This application is dependent on tar elimination and CO conversion to levels less than 1% by volume by means of the water gas shift reaction [2]. Alternatively, the gas can be used as hydrogen source for refinery hydrotreating operations, for ammonia production [3], and, with suitable CO concentrations, for methanol and FisherTropsch synthesis. As reported extensively in the literature, biomass steam gasification (performed in the main in fluidised bed reactors) results in the conversion of carbonaceous materials to permanent gases (H,, CO, CO,, CH4, light hydrocarbons), char and tar. A schematic representation of the reaction steps involved in this process is given in Fig. 1. To render the gas amenable to commercial applications, tar components need to be reduced to a minimum (less than 1 g/m’ of gas at ambient conditions). This
* Author to whom all correspondence should be addressed.
helps to avoid corrosion problems and poisoning of the catalysts employed to perform the water gas shift reaction. Moreover, in this way more gas is obtained and the efficiency of the overall gasification process is increased. Much research has addressed this subject [414]; nickel based commercial steam reforming catalysts have been utilised together with, or as an alternative to, calcined dolomite. Different process configurations have also been examined in which the catalytically active substances are either placed directly in the gasifier, or in a secondary fixed or a fluidised bed reactor. A major problem has been recognised in the rapid loss of the catalytic activity due to the fouling by carbon build-up on the surface of the catalyst itself. Such deactivation can be partially overcame by utilising catalyst in the fluidised state and in a secondary reactor [4, 51;results obtained, however, show that the total gas yield under these conditions after several hours of operation is considerably lower than that at the start of the test: tar conversion is far for complete and heavy components are still present in the product gas. The kinetics of catalyst deactivation depend on many factors, such as catalyst type, bed temperature, gas residence time, steam/biomass ratio, catalyst particle size and, above all, the tar content of the raw gas. This in turn depends on the type of feedstocks [15], the gasification temperature, the steam/ biomass ratio and the biomass particle size. It is worth 551
S. RAPAGNA et al.
552
+ Hz, CO,COz, CH4, light and heavy hydrocarbons
biomass particle Fig. 1. Schematicrepresentationof the biomassgasificationprocess.
pointing out that the yield and composition of the primary product formed in the devolatilization step (Fig. 1) are related to the heating rate of the biomass particles: high heating rates produce more light gasesand less char and condensate [16]. Experiments performed by the authors with different biomass particle sizesin the range 2741090 pm, show that at the same temperature level (700-75O”Q the smallest particles produce about 0.5 m3 more gas per kg of biomass than is obtained with the biggest particles [ 171.For small particle sizesthe devolatilization process is mainly controlled by reaction kinetics; as the particle size increases,kinetic control gives way to heat transfer control. In this study catalytic gasification tests were performed with two nickel steam reforming catalysts (one spent and one fresh), as well as with calcined dolomite. The initial average diameter of the biomass particles was 1.1 mm. The aim of these experiments was to characterise the influence of operating parameters on the catalytic transformation of tar components into CO and H, in presence of steam, so as to obtain useful data for the design of industrial units. There is general consensus on the effectiveness of nickel catalysts for tar reforming, but suitable operating conditions, such as temperature and space time, need further assessment. The decision to test an exhaust nickel based catalyst has been dictated by economic considerations, which take into account the high cost of fresh catalyst and its life expectancy with the present processes[14]. APPARATUS AND EXPERIMENTAL
METHOD
Figure 2 provides a schematic representation of the bench scale facility utilised in this study. Its major components are: biomass feeding system, fluidised bed gasifier, catalytic fixed bed, gas cooling system, metering and analysing for the off-gases. The fluidised bed gasifier consists of an austenitic stainless steel cylindrical vessel of internal diameter 62 mm fitted with an alumina porous distributor plate, designed to allow for a good gas distribution at all temperatures. The bed is located in a cylindrical Carbolite Furnace provided with temperature and heating rate control systems. The temperatures in the reactor are measured by meansof two thermocouples, one immersed in the bed and the other located under the distributor. The bed inventory is sand particles (d, = 348 ,um, pp = 2640 kg/m3). Water for the generation of steam (the fluidising
gas) is fed to an electrically heated boiler by means of a peristaltic pump at a constant flow rate. The biomass feeding probe (of internal diameter 8 mm and cooled by means of air flowing through a jacket) is designed to deliver the biomass well inside the bubbling bed. The secondary catalytic fixed bed reactor (of internal diameter of 60 mm) is located after a ceramic candle filter and heated by means of a cylindrical electric furnace. The temperature is measured with a thermocouple located inside the reactor which is filled with catalyst: either one of the two nickel catalysts or calcined dolomite. Some further runs with calcined dolomite were performed by placing it directly in the gasifier, in place of the sand particles. Gas products were analysed by two on-line gas chromatographs: the first was fitted with a Carbosieve Sll column (1.5 m, 3 mm l.D) with helium as carrier gas, to detect Nz, COz, CH4, C,H, and C2H4(if any); the second was fitted with a Molecular Sieve 5A column (2 m, 4 mm I.D.) with argon as carrier gas, to detect mainly HZ, and also N,, CO and CH,. For each gasification run the permanent gas yield was measured by means of a volumetric gasmeter, after separation of the condensate (water and organic phases).The total quantity of tar was obtained by summing two contributions: the tar in the organic phase and the dissolved tar in the water phase. The organic phase of the condensate (a very minor quantity when the catalytic conversion was carried out) was separatedand weighted. The water soluble tar was quantified by means of TOC (total organic carbon) analysis with the assumption that naphthalene could be regarded to be the representative compound of the soluble tar. The size distribution and mean particle sizes of the catalyst utilised in the reactivity tests are reported in Table 1. To obtain information on the reduction temperatures for the two catalysts, a TPR (Temperature Programmed Reduction) test was performed. This consists of measuring the quantity of hydrogen consumed with 50 mg of catalyst heated from ambient to 900°C at the rate of lS”C/min, with a gas (3% H, in Argon) flowing at a rate of 0.05 l/min. The results obtained with this test are reported in Fig. 3. From these results the consumption of H, has been estimated to be 0.96 * 10m3and 3.8 * 10e31 for exhaust and fresh catalyst respectively. Runs were performed with the spent catalyst at temperatures of 75O”C-770°C with different quantities of catalyst (100, 200 and 330 g) supported between two
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GASIFICATION
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I
J
nitrogen air water Fig. 2. Bench scale gasification plant. 1-fluidised bed reactor, 2-gas distributor, 3-wind box, &electric furnaces, 5-biomass feeder, 6--cyclone, 7--ceramic filter, g--cooling system, 9-gas flowmeter, IO-gas chromatographs, 1l-pump, 12-water man-
ometer, 13--catalytic fixed bed reactor.
Table 1. Size distribution
of the two catalysts
Spent catalyst Size, pm
% wt
Fresh catalyst % wt
> 1000 1000-850 850-710 710-600 60&500 500-425 < 425
0.39 30.62 31.16 16.33 17.02 4.19 0.29
0.1 27.86 30.20 17.16 20.40 4.2 0.18
705
719
Mean diameter, pm
results. After the maximum bed temperature was reached, the catalytic bed was left to cool down to the desired temperature level, the nitrogen-hydrogen feed was
TF’R of exhaust
--
catalyst
Q:!yyfJ--~;
layers of wool quartz in the reactor vessel. Before each run the catalyst heated to 850°C at a constant rate of 4”C/min in a 5% H2 in N2 atmosphere in order to enable the reduction to take place. The quantity of hydrogen utilised for the reduction process is well above that required by the catalyst, as determined from the TPR
Q
200
400
600
800
1000
Temperature, “C Fig. 3. Hydrogen consumption during the reduction of the two catalysts, as a function of temperature.
554
S. RAPAGNA
switched off, and biomass gasification was started in the gasifier. With the fresh catalyst only one quantity has been tested (200 g), and the runs were carried out at three temperature levels: 665”C, 760°C and 830°C. The reduction procedure for the catalyst was again as described above. The tests with calcined dolomite were carried out by placing it either in the fixed bed or directly in the gasifier. The dolomite was first crushed and sieved to obtain a fraction with a mean particle size of 454 pm, and then calcined in air at 900°C for ten hours. The weight loss during calcination amounts to about 40% of the original weight: as a result the final density is about 1500 kg/m3. These values of particle size and density mean that the minimum fluidisation velocity fot the gasifier is of the same order as that obtained with the sand particles bed (E 0.085 m/s). Under these conditions the residence times of char and gas in the gasifier are comparable to those resulting from tests carried out with sand. RESULTS
AND
et al.
(Gas Hourly Space Velocity: gas flow rate at the exit of the catalytic reactor divided by the bulk volume of the catalyst bed) and the WHSV (Weight Hourly Space Velocity: mass flow rate of biomass fed to the gasifier, divided by the mass of catalyst in the reactor). The tar content of the gas leaving the catalytic bed (expressed as g of tar per kg of biomass) is reported in Figs 4 and 5, as a function of the gasification time, for exhaust and fresh catalyst respectively. The catalytic tests clearly show the gas yield to be considerably higher than from tests performed without catalyst (Tables 2 and 3): it effectively doubles when fresh catalyst is utilised. Moreover the gas composition obtained in this case is close to the equilibrium conditions for the water gas shift reaction (wgsr) as shown on Fig. 6, where the theoretical value has been calculated from
3%
DISCUSSION
The operating conditions in the gasifier were kept constant for all runs and are reported in Table 2. The total gas yields obtained with the two catalysts are reported in Table 3, together with the gas composition (a function of the catalytic bed temperature), the GHSV
Table 2. Operating conditions of the gasifier with the corresponding gas and tar yields Bed inventory, sand Biomass, almond shells Bed temperature Biomass feed rate Steam input rate Total gas yield Tar yield
dp = 348pm, pp = 2640 kg/m3 d, = 1.1 mm, pP = 1200 kg/m3 770°C 0.3 kg/h 0.3 kg/h 1 Nm3/kg of biomass 0.1 kg/kg of biomass
Table 3. Gas yields and composition
50
0
200
Gasification time, minutes Fig. 4. Tar yield expressed in g/kg of biomass in the produced gas, as a function of the gasification time for the spent catalyst. GHSV: (m) 27,700 h-l, (0) 14,400 h-l, (A) 9000 h-l.
obtained with the spent and the fresh catalysts, as a function of the catalytic bed temperature and quantities of catalyst utilised Fresh catalyst
Spent catalyst Temperature of the catalytic fixed bed WHSV, h-’ GHSV, h-’ Gas yield, normal m3 per kg of daf biomass Gas composition HZ co co2 ‘W
150
100
750 2.70 27,700
770 1.35 14,400
760 0.84 9000
665 1.35 13,850
760 1.35 15,650
830 1.35 18,000
1.35
1.40
1.50
1.80
1.88
1.98
in % by volume, 56.2 25.2 12.6 6.0
55.7 25.0 12.8 6.5
56.0 24.7 13.3 6.0
62.2 22.8 15.0
61.7 23.5 14.8
62.1 22.7 15.2
CATALYTIC
0
50
100
150
GASIFICATION
200
Gasification time, minutes Fig. 5. Tar yield expressed in g/kg of biomass in the produced gas, as a function of the gasification time for the fresh catalyst. GHSV: (m) 13,850h-‘, (0) 15,650h-l, (A) 18;OOO h-‘.
known values of the equilibrium constant for the wgsr and the experimental values of CO, CO, and H,O volumetric fractions. Figure 6 also shows that the hydrogen yield, although lower than for the corresponding active catalyst case, follows the same trend for operation without catalyst and with exhaust catalyst. Moreover with fresh catalyst, CH, is also completely
650
700
750
800
850
Temperature, “C Fig. 6. Ratio of the theoretical and experimental values of hydrogen present in the produced gas as a function of fluidised bed temperature: absence of catalysts (a), fresh catalyst (0) and spent catalyst (0).
OF BIOMASS
555
transformed into CO and H,, resulting in a hydrogen composition in the final product higher than 60% by volume. This value is compatible with the requirements imposed by the fuel cells on the feeding gas (i.e. Phosphoric Acid Fuel Cells), provided that a conventional CO shift operation is included in the process scheme. Figure 7 reports the total gas yield as a function of time, when dolomite is used, both in the gasifier itself instead of sand, and in a secondary fixed bed reactor. The continuous increase of gas yield with time for the case of dolomite in the gasifier can be explained on the basis that char accumulating in the bed is catalytically gasified in the presence of alkaline compounds. A major drawback of this procedure, however, is that it results in comminution of the bed inventory which gives rise to a very high production of fines-with attendant gas cleaning problems. On the other hand, the use of dolomite in a secondary reactor results in an increase in the overali gas yield over that reported in Table 2, and of the same order as that obtained with exhaust catalyst (although at a higher temperature level). The major problem encountered in the experimental tests with catalysts and dolomite in the fixed bed gas converter is related to the continuous increase of gas pressure drop due to carbon deposition especially on the layer firstly encountered by the gas. The formation of carbon in methane and naphtha steam reforming is well documented [18]. To avoid this phenomenon high steam/hydrocarbon ratios (in the range 3.44.5) are usually employed. In the case examined here, this ratio is much higher, so that the formation of carbon is not thermodynamically favoured. As a result the presence of carbon cannot be ascribed to the decomposition of CO by means of the Bouduard mech-
0
20
40
60
gasification time, minutes Fig. 7. Gas yield obtained by using dolomite directly in the gasifier (0) at temperature of 770°C and in the fixed bed reactor (0) at temperature of 81O’C.
556
S. RAPAGNA et al.
anism, but is mainly due to the cracking and polymerisation reactions involving the heavy hydrocarbons. The addition of alkaline metals to the catalyst reduces the rate of carbon accumulation, but cannot eliminate it completely [4, 11, 181. This all points to a process configuration which should allow for frequent, if not continuous, catalyst regeneration. A further problem relates to the presence of sulphur in the biomass materials. The sulphur content can vary considerably with the type of feedstock and is usually very low; some sulphur is however always present as it is needed in the growing process of proteins. Sulphur is a poison for nickel catalysts. Evidence on catalyst deactivation by sulphur is not reported in the literature on catalytic treatment of the gas produced from biomass gasification; this may be due either to the low concentration of sulphur compounds in such gas, or to the fact that experience in this field is mainly confined to experimental studies over limited periods of time. Only experience with catalysts undergoing numerous regeneration cycles, as previously foreseen, will enable the importance of this phenomenon to be assessed. Finally in order to perform the catalytic reforming in industrial sizeunits, a major point to be considered relates to the implementation of an efficient and easy to operate system for supplying the heat requirement of the endothermic reactions involved. As the gasification process itself is also endothermic the thermal transfer problem is of great relevance for the overall plant feasibility. This problem is at present the subject of a joint research project financed by the EU (Contract JOR3-CT950037):
a dual fluidised bed configuration is employed with internal circulation of the bed material (including the reforming catalyst) between the gasification zone and a combustion zone which provides the necessary heat input for the biomass thermal conversion and tar reforming. Gas leaving the gasification section is prevented from contacting the combustion products, so as to maintain it free of nitrogen and with high calorific value [19]. This configuration also addresses the problem of catalyst deactivation: carbon deposited on the particle surfaces is burnt off in the combustion zone. CONCLUSIONS The experimental results show that fresh, nickel based
catalysts are extremely active for the elimination of CH, and tars: about 2 m3 of gas (at ambient conditions) are obtained per kg of biomass, with a H, content higher than 60% by volume. A gas of this specification is suitable
for industrial applications, both for highly efficient electricity production and as a feedstock for chemical synthesis. However, problems of coke formation and heat requirements do not favour fixed bed operation in industrial applications. A new configuration involving a dual action fluidised reactor has been proposed and is currently being studied to render the process industrially feasible. Acknowledgements-The
financial support of the Consiglio
Nazionale delle Ricerche and the EU under the Contract JOR3CT95-0037 is gratefully acknowledged.
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17. Rapagna S. and Latif A., Steam gasification of almond shells in a fluidised bed reactor: the influence of temperature and particle size on product yield and distribution. Biomass and Bioenergy, in press. 18. Twigg M. V., Catalyst Handbook, 2nd edition, Wolfe Publishing, 1989.
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