Hydrogen Production by Biomass Gasification with Steam−O2

Energy & Fuels 2006, 20, 1305-1309

1305

Hydrogen Production by Biomass Gasification with Steam-O2 Mixtures Followed by a Catalytic Steam Reformer and a CO-Shift System Maria P. Aznar, Miguel A. Caballero, Jose´ Corella,*,† Gregorio Molina,† and Jose´ M. Toledo† Chemical and EnVironmental Engineering Department, CPS, 3 Maria de Luna st., UniVersity of Saragossa, 50018 Saragossa, Spain, and Department of Chemical Engineering, UniVersity “Complutense” of Madrid, 28040 Madrid, Spain ReceiVed December 21, 2005. ReVised Manuscript ReceiVed February 10, 2006

This paper studies the effect of adding a CO-shift reactor downstream from a fluidized-bed biomass gasifier and a steam-reforming catalytic bed. The upstream gasifier was of small pilot-plant scale, 10 (kg of biomass)/ h, so the downstream catalytic reactors, steam-reformer and CO-shift, operated under a real gasification gas. The main gasifying agent used was H2O-O2 mixtures. Some results are also reported for gasification with air with some steam. The CO-shift catalytic system used had one high- (HT) and one low-temperature (LT) reactor. Commercial catalysts were used. CO conversions (eliminations) higher than 90% and a H2 content as high as 73 vol %, dry basis, have been obtained by the CO-shift system. The CO conversion and the increase (up to 14 vol %) of H2 content correlated well with the molar steam/CO ratio in the gasification gas at the inlet of the HT reactor.

Introduction Some authors working on biomass gasification believe that a H2-rich gas could be produced by biomass gasification. This would be useful for some advanced applications of the produced gasification gas. Biomass has, on average, only 6 wt % of hydrogen, which would make it, in principle, of not much interest for H2 production. Nevertheless, some extra H2 can come from the cracking of H2O if the biomass is also used as a fuel to supply the high amount of heat needed to crack the molecule of H2O. In fact, a raw gasification gas with 48-55 vol %, dry basis, H2 was generated by biomass gasification with pure steam in a fluidized bed.1,2 Since this raw gasification gas contains some CH4, tar, and light hydrocarbons, when a bed of calcined dolomite (CaO‚MgO), limestone (CaO), and/or magnesite (MgO) was used downstream from the gasifier, the H2 content in the gasification gas was increased to 50-58 vol %, dry basis.3 Given that these minerals have only a small steam-reforming activity for tar and CH4 elimination, some tar and CH4 still remains in the gasification gas after the bed of dolomite or limestone.3 When another more active (nickel-based) steamreforming catalyst is added to the two upstream reactors [biomass gasifier + downstream bed of dolomite], the H2 content in the gasification gas was increased to 55-63 vol %, dry basis.4 * Corresponding author. E-mail: [email protected]. † University “Complutense” of Madrid. (1) Corella, J.; Aznar, M. P.; Delgado, J.; Aldea, E. Steam gasification of cellulosic waste in a fluidized bed with downstream vessels. Ind. Eng. Chem. Res. 1991, 30 (10), 2252-2262. (2) Herguido, J.; Corella, J.; Go´nzalez-Saiz, J. Steam gasification of lignocellulosic residues in a fluidized bed at a small pilot scale. Effect of the type of feedstock. Ind. Eng. Chem. Res. 1992, 31 (5), 1274-1282. (3) Delgado, J.; Aznar, M. P.; Corella, J. Biomass gasification with steam in fluidized bed: Effectiveness of CaO, MgO, and CaO-MgO for hot raw gas cleaning. Ind. Eng. Chem. Res. 1997, 36 (5), 1535-1543.

The high endothermicity of the process of gasification with pure steam may be solved in two different ways: (1) By using a dual fluidized-bed circulating system (DFBCS). This was the solution adopted first by Prasad and Kuester5 and BattellesColumbus;6 afterward, in the late eighties, by Herguido and co-workers;7,8 and more recently, in the late nineties, by Hofbauer and co-workers.9 These DFBCSs have proved their technical feasibility, but if several more reactors had to be added after those circulating systems, the resulting overall process would become too complex and difficult to operate. (2) Another solution is to provide some heat for the steam gasification by adding some oxygen to the steam; that is to say, to gasify with [H2O + O2] mixtures. In this case, the H2 content in the raw gasification gas can reach 30 vol %, dry basis, when (4) Aznar, M. P.; Corella, J.; Delgado, J.; Lahoz, J. Improved steam gasification of lignocellulosic residues in a fluidized bed with commercial steam reforming catalysts. Ind. Eng. Chem. Res. 1993, 32, 1-10. (5) Prasad, B. V. R. K.; Kuester, J. Process analysis of a dual fluidized bed biomass gasification system. Ind. Eng. Chem. Res. 1988, 27 (2), 304310. (6) Paisley, M. A.; Overend, R. P. The SilvaGas process from future energy resourcessA commercialization success. In Proceedings of the 12th European Conference on Biomass for Energy, Industry and Climate Protection, Amsterdam, The Netherlands, June 17-21, 2002; pp 975-978. (7) Herguido, J.; Rodrı´guez-Trujillo, J. L.; Corella, J. Gasification of biomass with tar cracking catalysts in a circulating multisolid fluidized bed pilot plant. In Biomass for Energy and Industry, Proceedings of the 5th International Conference on Biomass, Lisbon, Portugal, Oct 9-13, 1989; Grassi, G., Gosse, G., dos Santos, G., Eds.; Elsevier Applied Science Publishers: London, 1990; Vol 2, pp 2793-2797. (8) Herguido, J.; Corella, J.; Artal, G.; Garcı´a-Bordeje´, J. E. Results with a multisolid circulating dual fluid bed pilot plant for the advanced steam gasification of biomass. In Biomass for Energy, Industry and EnVironment, Proceedings of the 6th International Conference on Biomass, Athens, Greece, April 22-26, 1991; Grassi, G., Colina, A., Zibeta, H., Eds.; Elsevier Applied Science Publishers: London, 1992; pp 792-796. (9) Pfeifer, C.; Rauch, R.; Hofbauer, H. In-bed catalytic tar reduction in a dual fluidized bed biomass steam gasifier. Ind. Eng. Chem. Res. 2004, 43 (7), 1634-1640.

10.1021/ef050428p CCC: $33.50 © 2006 American Chemical Society Published on Web 03/16/2006

1306 Energy & Fuels, Vol. 20, No. 3, 2006

Aznar et al.

Table 1. H2 Contents (Vol %, Dry Basis) Obtained in the Gasification Gas at the Exit of Different Reactors, for Two Different Gasifying Agents fasifying agent at the gasifiers exit in-bed/silica sand in-bed/silica sand + dolomite gasifier + bed of dolomite gasifier + bed of dolomite + bed of nickel-based catalyst

only silica sand is used as the in-bed material,10,11 or 45 vol % H2, dry basis, when there is 40 wt % of calcined dolomite in the gasifier bed12 which acts as an in-bed steam-reforming catalyst. However, if a downstream bed of calcined dolomite is added, then this H2 content increases to 50-52 vol %, dry basis.13 When another further bed of a steam-reforming nickelbased catalyst is added, the H2 content in the gasification gas can be increased to 52-59 vol % H2, dry basis.14,15 Since these H2 contents are the basis of the work here reported, they are summarized in Table 1. The H2 content in the gasification gas can still be increased by adding a CO-shift reactor to the above said processes and/ or schemes. The CO-shift catalysts increase the H2 content in the flowing gas by the exothermic reaction CO + H2O S CO2 + H2. Although the CO-shift process is well-known in the chemical industry, when these authors started to present results in 1998 concerning this approach,16 it was the first time, to the authors’ knowledge, that CO-shift reactors were connected to biomass gasifiers. The resulting overall biomass gasification process was new at the time. Further, Brown and co-workers17,18 at University of Iowa also connected two CO-shift reactors downstream from a biomass fluidized-bed gasifier followed by a guard bed of dolomite and a catalytic steam-reforming bed. With these four catalytic reactors after their biomass gasifier, they obtained a H2 content of 60.8 vol %, N2-free basis, when the biomass gasification was made with pure steam,17 and 26.7 vol % H2, dry basis, when the gasification was made with air.18 Currently, some more authors (i.e., refs 19 and 20) are working on this issue, because this approach offers new perspectives and some more future for biomass gasification. (10) Aznar, M. P.; Corella, J.; Gil, J.; Martı´n, J. A.; Caballero, M. A.; Olivares, A.; Pe´rez, P.; France´s, E. Biomass gasification with steam and oxygen mixtures at pilot scale and with catalytic gas upgrading. Part I: Performance of the gasifier. In DeVelopments in Thermochemical Biomass ConVersion; Bridgwater, A. V., Boocock D. G. B., Eds.; Blackie Academic & Professional: London, 1997; Vol. 2, pp 1194-1208 (ISBN: 0-75140350-4). (11) Gil, J.; Aznar, M. P.; Caballero, M. A.; Frances, E.; Corella, J. Biomass gasification in fluidized bed at pilot scale with steam-oxygen mixtures. Product distribution for very different operating conditions. Energy Fuels 1997, 11 (6), 1109-1118. (12) Olivares, A.; Aznar, M. P.; Caballero, M. A.; Gil, J.; Frances, E.; Corella, J. Biomass gasification: Produced gas upgrading by in-bed use of dolomite. Ind. Eng. Chem. Res. 1997, 36 (12), 5220-5226. (13) Perez, P.; Aznar, M. P.; Caballero, M. A.; Gil, J.; Martin, J. A.; Corella, J. Hot gas cleaning and upgrading with a calcined dolomite located downstream a biomass fluidized bed gasifier operating with steam-oxygen mixtures. Energy Fuels 1997, 11 (6), 1194-1203. (14) Caballero, M. A.; Aznar, M. P.; Gil, J.; Martı´n, J. A.; France´s, E.; Corella, J. Commercial steam reforming catalysts to improve biomass gasification with steam-oxygen mixtures. 1. Hot gas upgrading by the catalytic reactor. Ind. Eng. Chem. Res. 1997, 36 (12), 5227-5239. (15) Aznar, M. P.; Caballero, M. A.; Gil, J.; Martı´n, J. A.; Corella, J. Commercial steam reforming catalysts to improve biomass gasification with steam-oxygen mixtures. 2. Catalytic tar removal. Ind. Eng. Chem. Res. 1998, 37 (7), 2668-2680. (16) Caballero, M. A.; Aznar, M. P.; Gil, J.; Martı´n, J. A.; Corella, J. CO-shift catalytic beds after a biomass gasifier and a steam-reforming catalytic reactor to get new and interesting exit gas compositions. In Biomass for Energy and industry, proceedings of the 10th European Conference and Technology Exhibition, Wu¨rzburg, Germany, June 1998; Kopetz, H., Weber, T., Palz, W., Chartier, P., Ferrero, G. L., Eds.; C.A.R.M.E.N.: Rimpar, Germany, 1998; pp 1789-1793.

pure H2O

H2O + O2

48-55 (refs 1 and 2)

30 (refs 10 and 11) 45 (ref 12) 50-52 (ref 13) 52-59 (refs 14 and 15)

50-58 (ref 3) 55-63 (ref 4)

This paper addresses the effect of locating one or two COshift catalytic beds downstream from a two-step gasification process: fluidized-bed gasifier with in-bed dolomite + reactor with a steam-reforming nickel-based catalyst. The approach, system, and/or process here studied is, therefore, very similar to that recently published by Brown and co-workers.17,18 The main difference between both systems is that, in this work, a guard bed of dolomite is not used downstream from the gasifier, as we used in the past.21 Instead, we locate the dolomite in the same fluidized-bed gasifier, which provides very similar results22 and avoids the use of one reactor, the guard bed, in the complex overall process. So, the scheme of the process studied here is the following:

The main results presented in this work were obtained using steam-O2 mixtures as the gasifying agent. Nevertheless, since the authors currently operate the gasifier with air (+ some H2O from the moistures of the biomass and of the air), the CO-shift reactor was also operated with air (+ H2O) in the upstream gasifier. Some results will, therefore, also be shown here using air (+ some H2O) as the gasifying agent. Experimental Facility Used Gasifier. The gasifier used was a bubbling fluidized bed of 15 cm i.d. and 3.2 m height, continuously fed with biomass near the bed bottom. The feeding system had two hoppers with two locks and two screw feeders of 6 cm diameter. This gasifier usually worked with biomass flow rates of ∼10 kg/h. The ingasifier-bed material was a mixture of 70-80 wt % silica sand and 20-30 wt % calcined dolomite. Silica sand + calcined dolomite (13 kg) were used in each test, equivalent to a (bulk, fixed) bed height of 50 cm. The superficial gas velocity (gasifier conditions) at the bed inlet typically used was 45 cm/s. Some (17) Brown, R. C.; Sadaka, S. S.; Norton, G.; Xu, M.; Bown, N. Thermochemical production of hydrogen from biomass. In Proceedings of the 2nd World Conference on Biomass for Energy, Industry and Climate Protection, Rome, May 10-14, 2004; Van Swaaij, W. P. M., Fja¨llstro¨m, T., Helm, P., Grasi, A., Eds.; ETA-Florence: Florence, Italy, 2004; pp 1144-1147 (ISBN: 88-89407-01-2). (18) Zhang, R.; Brown, R. C.; Suby, A. Thermochemical generation of hydrogen from Switchgrass. Energy Fuels 2004, 18 (1), 251-256. (19) Ederer, H. J.; Fritsch, T.; Henrich, E.; Mas, C. Water-gas-shiftreaction using MOS2 catalyst following biomass gasification. In Proceedings of the 2nd World Conference on Biomass for Energy, Industry and Climate Protection, Rome, May 10-14, 2004; Van Swaaij, W. P. M., Fja¨llstro¨m, T., Helm, P., Grasi, A., Eds.; ETA-Florence: Florence, Italy, 2004; pp 918920 (ISBN: 88-89407-01-2). (20) Effendi, A.; Hellgardt, K.; Zhang, Z.-G.; Yoshida, T. Optimising H2 production from model biogas via combined steam reforming and CO shift reactions. Fuel 2005, 84, 869-874. (21) Narvaez, I.; Orio, A.; Aznar, M. P.; Corella, J. Biomass gasification with air in an atmospheric bubbling fluidized bed. Effect of six operational variables on the quality of the produced raw gas. Ind. Eng. Chem. Res. 1996, 35 (7), 2110-2120. (22) Corella, J.; Aznar, M. P.; Gil, J.; Caballero, M. A. Biomass gasification in fluidized bed: Where to locate the dolomite to improve gasification? Energy Fuels 1999, 13 (6), 1122-1127.

Hydrogen Production by Biomass Gasification

other details concerning the gasifier can be found in refs 10 and 11. It must be pointed out that the total tar content in the raw gasification gas under these experimental conditions was relatively low, of the order of only 2 g/Nm3.23 This is a key limit to get a good performance, high life, of the catalytic reactors located downstream from the gasifier.15 Feedstock. An easy-to-feed and easy-to-gasify feedstock was selected for the research presented in this paper: small pine (Pinus pinaster) wood chips. This type of biomass has a content of sulfur and of nitrogen of only a very few ppms. Its full characterization, including its detailed elemental analysis, was published previously.10,11 It may be repeated here that its H2 content was 5.7 wt %. A 2-3 wt % of the overall mass flow rate was calcined dolomite. It was continuously fed, mixed with the biomass, to replace the dolomite eroded and carried out of the gasifier by elutriation. Catalytic Reactors. After the gasifier and before the heat exchanger, a slip flow was taken from the gasification raw gas and sent to several catalytic reactors connected in series, as indicated in refs 10 and 11. The first catalytic reactor contained a commercial steam-reforming nickel-based catalyst to eliminate most of the tar present in the gasification gas, correspondently increasing the H2 content in it. At the exit of this reactor (inlet of the high-temperature (HT) shift reactor) the tar content was only 10-30 mg/Nm3 (for more details on the gas composition in that location, see ref 15). This low tar content did not deactivate the downstream CO-shift catalysts, at least in the few hours on stream of each test. Downstream from the catalytic steam-reforming reactor, there were one or two reactors with commercial CO-shift catalysts. Two types of tests were made: with only one CO-shift catalyst and with a combination of high- and low-temperature catalysts. When two shift reactors were used, the first one was a hightemperature (HT) reactor and the second one was a lowtemperature (LT) reactor. The CO-shift catalytic reactors had 4.1 cm i.d. and 63 cm total height. Before the test itself, the shift reactors were externally heated by two ovens of 1.5 kW to get the desired level of temperature in each bed. When the whole gasification plant reached steady state, these external ovens were switched off. So, both CO-shift beds operated under near-adiabatic conditions, which is a well-known and/or recommended operation for CO-shift catalytic reactors (i.e., ref 24). Between the HT and the LT beds, there was some cooling to decrease by 200-250 °C the temperature of the gas before entering the LT reactor. In some of the latest tests, some preheated steam was also introduced between the steamreforming and the HT CO-shift catalytic reactor to increase the H2O/CO ratio in the flowing gas. Because of the size of the CO-shift catalytic reactors, temperature gradients could be important. So, three thermocouples were placed in each bed, two in the axis and one in the wall, inner side. Of course, temperatures in the shift reactors were carefully controlled during each test to avoid catalyst sintering (which occurs at 350 °C with the low-temperature catalyst and at 550 °C with the high-temperature one). To keep constant the H2O/CO ratio and the gas composition at the inlet of the CO-shift reactor was not difficult in this facility, because the upstream gasifier operated without troubles. (23) Corella, J.; Li, G.; Toledo, J. M. Experimental Conditions to get less than 2 g tar/Nm3 in a Fluidized Bed Biomass Gasifier. Presented at the 14th European Conference on Biomass for Energy Industry and Climate Protection, Paris, France, Oct 17-21, 2005. (24) Rase, H. F. Case study 105: The CO-shift reactor. Chemical Reactor Design for Process Plants; John Wiley: New York, 1977; Vol. 2.

Energy & Fuels, Vol. 20, No. 3, 2006 1307

It was, therefore, easy to avoid variations with time-on-stream of temperature in shift reactors. Two important operation parameters, temperatures and hourly space velocity, in the two CO-shift beds were as follows. catalyst HT LT

SV (Nm3, wet gas/(m3 h))

T (°C) inlet

exit

350 200

480 260

1300-2700 4600-5100

CO-Shift Catalysts. Commercial CO-shift catalysts were provided by ICI Katalco (U.K.) and BASF AG (Germany). They were tested with their commercial sizes and shapes. High-Temperature (HT) CO-Shift Catalyst. Two different commercial HT catalysts were used: ICI 15-5 and BASF K611. Their composition is a mixture of iron and chromium oxides and is similar to that used by Brown and co-workers18 The catalysts were used in their commercial shape, pellets 5.4 mm diameter × 3.6 mm high for the ICI catalyst and tablets of 6 × 6 mm for the BASF one. These catalysts require an initial reduction of the iron oxide (Fe2O3) to magnetite (Fe3O4). The essential features of the reduction are that steam must be present and that the CO content of the reducing gas should be less than a limiting value, which could overheat the catalyst. On a wetgas basis, CO should typically be 90 vol % have been obtained. (iv) To get CO conversions >90%, high (g2) steam/CO ratios are needed at the inlet of the HT shift reactor. (v) From the experience gained in this research, the authors believe that an improved design and operation of the CO-shift catalytic reactors, with an optimized temperature profile, could still improve the results presented here. In fact, the thermodynamic equilibrium constant at 260 °C (exit temperature of the LT shift reactor) is 75, and the similar ratio calculated with experimental results is only 15, indicating that there are still some possible further improvements. (vi) With such a low tar content (on average, in these tests, 20 (mg of tar)/Nm3) in the gasification gas entering the HT CO-shift catalytic reactor, there was no deactivation of the shift catalysts after 10 hours on stream. (vii) The overall gasification process (advanced fluidizedbed gasifier + steam-reforming bed + two CO-shift reactors) is complex and could result with the produced H2 having a high cost. (viii) Nevertheless, the overall gasification process described in this work was operated without major technical problems. For these authors, the process is technically feasible.

(4)

experimental

Comparing these two values (75 and 15), it may be deduced that the gas composition obtained was, therefore, still somewhat far from equilibrium. It indicates that still better results could be obtained from a further and improved experimentation.

Acknowledgment. We are grateful to Javier Gil and JuanAntonio Martı´n for their help in performing some experiments. We are also grateful to BASF AG and ICI-Katalco for providing samples of their catalysts. EF050428P