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Ind. Eng. Chem. Res. 1997, 36, 42-51
Catalytic Decomposition of Gasification Gas Tar with Benzene as the Model Compound Pekka A. Simell,* Nina A. K. Hakala, and Heikki E. Haario Technical Research Centre of Finland, VTT Energy, P.O. Box 1601, FIN-02044 VTT, Finland
A. Outi I. Krause Department of Chemical Technology, Helsinki University of Technology, Kemistintie 1, FIN-02150 Espoo, Finland
Tar decomposition over dolomite catalyst in gasification conditions was modeled by benzene reaction with CO2. Kinetic studies were carried out at 1023-1173 K and ambient pressure in a plug flow reactor. Operation conditions without external or internal mass transfer limitations were used. Mechanistic models of the Langmuir-Hinshelwood type describing benzene decomposition were derived and tested. Experimental results could be best described by a kinetic equation where benzene single-site adsorption on dolomite was the rate-determining step and CO2 adsorption took place nondissociatively. Introduction Hot fuel gas cleaning is one of the most crucial steps in developing biomass gasification processes for power and heat production. Purification of biomass-derived gas from tars is essential due to possible plugging of downstream equipment like particulate filters and engine suction channels (Cahill et al., 1996). Quite extensive literature exists concerning tar decomposition from gasifier product gases. Nickel catalysts of various types have been found to be very effective catalysts for tar (Delgado et al., 1996; Ekstro¨m et al., 1985; Baker et al., 1987; Kinoshita et al., 1995; Simell et al., 1990) and as well for simultaneous ammonia removal in gasification conditions (Krishnan et al., 1988; Leppa¨lahti et al., 1991; Mojtahedi and Abbasian, 1995; Simell et al., 1996). Carbonate rock materials in turn have been reported to be active tardecomposing catalysts when used either in a secondary reactor (Alde´n et al., 1994; Aznar et al., 1989; Delgado et al., 1996; Ekstro¨m et al., 1985; Simell et al., 1992, 1995) or as additive bed materials in a gasifier (Narvaez et al., 1996; Yeboah et al., 1980). In these studies it has been demonstrated that almost complete tar conversions can be achieved with both materials at 900 °C and 0.1 MPa pressure when the catalyst is placed in a secondary reactor. However, the nickel catalysts are relatively expensive and hence require lifetimes in the range of years to be economically feasible (Simell et al., 1996). Consequently the alternative of using cheap bulk materials like dolomites or limestones becomes attractive. These materials can in principle be used in a separate reactor or directly as bed additives in a gasifier. Carbonate rock materials are highly active in the calcined state, which requires temperatures over 1073 K in typical gasification gas atmospheres (Alde´n et al., 1994; Aznar et al., 1989; Simell et al., 1995). The catalytic activity of carbonate rock has been reported to decrease in the order of limestone (CaO) > dolomite (CaO‚MgO) > MgO for steam reforming of naphthalene at 1000-1050 K (Betancur et al., 1995). Other authors * Author to whom correspondence should be addressed. E-mail:
[email protected]. S0888-5885(96)00323-5 CCC: $14.00
(Morita, 1978; Tralas et al., 1991) have observed different orders, which obviously depend on the reaction conditions applied as well as on the chemical and physical composition of the original materials (Simell et al., 1992). By carbonation the activity of Ca-containing materials decreases considerably (Donnot et al., 1991; Garcia and Huttinger, 1990; Simell et al., 1995). When tars decompose over carbonate rocks at high temperature, the formation of CO and H2 occur with simultaneous reaction of CO2 and H2O (Simell et al., 1994). Hence, the overall reactions of tars in catalytic hot gas cleaning can be described similarly to the schemes presented for steam reforming of hydrocarbons (Rostrup-Nielsen, 1984; Rostrup-Nielsen, 1973; Ridler and Twigg, 1989):
CnHm + nH2O f nCO + (n + m/2)H2
(1)
-∆H°(1173 K) ) -753 kJ mol-1 (for benzene) CO + H2O T CO2 + H2
-∆H°(1173 K) ) 33 kJ mol-1 (2)
The steam reforming of hydrocarbons with CaO-based materials has been studied by various authors. Garcia and Hu¨ttinger (1990) have studied the kinetics of steam gasification of naphthalene on CaO at 923-1073 K and found that the reaction rate was of first order with respect to hydrocarbon and strongly inhibited by CO2, which was a reaction product. The rate equation of the Langmuir-Hinselwood type was derived assuming the gasification of intermediate surface carbon to be the rate-determining step. Similarly Morita (1978) has suggested in his studied on steam reforming of residual oils over dolomite at 1173 K that the steam gasification of carbon intermediate is the rate-determining step. However, it is also possible that the hydrocarbons react in gasification gas atmosphere with CO2 which is one of the main constituents in the gas (about 15 vol %) (Rostrup-Nielsen and Bak Hansen, 1993; Alde´n et al., 1994). Thermodynamically this “dry” reforming reaction (eq 3) is slightly more favorable (∆G° ) -502 kJ mol-1 for benzene) than the steam-reforming reaction © 1997 American Chemical Society
Ind. Eng. Chem. Res., Vol. 36, No. 1, 1997 43
(∆G° ) -490 kJ mol-1 for benzene) in the operation conditions of hot gas cleaning (1173 K).
CnHm + nCO2 f 2nCO + (m/2)H2
(3)
-∆H°(1173 K) ) -948 kJ mol-1 (for benzene) Hence, carbon dioxide can affect tar reforming over CaO-based materials by forming surface carbonates and by being the oxidating reagent. There is, however, scarce information available on the reactions of CO2 with higher hydrocarbons over carbonate rocks or other basic oxides. Krylov et al. (1994, 1995) have studied the CO2-reforming kinetics of methane, C2, C3, C4, and C7 hydrocarbons over Mn, Cr, and Ni oxides presenting rate equations of the Langmuir-Hinselwood type for the oxidation reactions. However, kinetic parameters and reaction conditions for the derived correlations were not published. CO2-reforming kinetics of methane over group VIII metals and metal oxides has been studied more widely. Richardson and Paripatyadar (1990) found the rate equations of the Langmuir-Hinshelwood type to fit most adequately their data obtained with Rh catalyst at 873-973 K. Rostrup-Nielsen and Bak Hansen (1993) concluded that their mechanism derived for steam reforming can without drastic impacts also be applied for dry reforming by replacing H2O with CO2 in a study dealing with Ni, Ru, Rh, Pd, Ir, and Pt catalysts at 673-923 K. However, the dry reforming rate was slower and was explained to be due to adsorption and dissociation of CO2 on the catalyst and subsequent adsorption of CO, which inhibits the reaction. An alternative approach was presented by Bodrov and Appelbaum (1966) who presented dry reforming kinetics with the same expression as that used for steam reforming in a work with Ni-foil catalyst at 1073-1173 K. Their assumption was that steam was the actual oxidating component and that it was formed by the shift reaction 2 from CO2 prior to the reforming reaction. Activity of some oxide catalysts containing alkalineearth metal oxides has been observed by Ruckenstein and Hu (1995) to decrease in the order NiO/CaO . NiO/ MgO > NiO/SrO > NiO/BaO at 1063 K. The high activity of the CaO-supported catalyst was explained to be due to the insolubility of the NiO and CaO resulting in high free NiO surface area. To improve understanding of hot gas cleaning chemistry in gasification processes, it is obvious that the effects of CO2 on tar compounds should be studied in more detail. In this work we interpret the kinetics of the CO2-reforming reaction of a tar model compound at a temperature range relevant to hot gas cleaning conditions. Benzene was chosen for the tar model compound because it is the main constituent of hightemperature tar and also because it represents a stable aromatic structure apparent to this type of tar rich in polyaromatics. Experimental Section In the experiments the effects of temperature, space time, and reagent concentrations on the rate of benzene decomposition were measured. Apparatus and Test Procedure. The apparatus comprised a fixed-bed tube reactor equipped with a gasmixing system and gas sampling and analysis devices. The reactor was made of quartz (i.d. 9 mm, length 350 mm), and it had a quartz grid, where the bed material was placed and a thermocouple pocket (o.d. 4 mm). The
Table 1. Experimental Conditions in the Three Test Series Performed test series temperature, K Wcat, 10-3 kg benzene, ppmv CO2, vol % 1 2 3
823-1173 823-1173 833-1100
0.05-0.50 0.07-0.26 0.20-0.21
50-400 90-100 450-500
10 0.06 0.6-7
catalyst was mixed in a 1:1 volume ratio with silicon carbide to ensure isothermal conditions in the catalyst bed. The reactor was packed so that first a layer of silicon carbide (3.5 × 10-4 kg) was placed on the grid and then the catalyst mixture layer and finally a layer of silicon carbide (2.5 × 10-3 kg) to improve heat transfer to the reacting gas. The reactor was placed inside a three-zone furnace. The catalyst bed temperature profile was measured in each experiment with a K-type thermocouple, which was moved in the thermocouple pocket. The temperature of the catalyst bed was kept within 1 K from the mean temperature. The reactor was of downflow type, and it was operated in differential and integral modes. The feed gas flow (0.06-0.18 mn3/h) was controlled by mass flow controllers. In the gas mixing apparatus both CO2 and benzene/N2 mixture were added to the N2 flow to form the feed gas. The pressure drop over the catalyst bed ranged from 10 to 60 kPa depending on the gas flow rate. This was taken into account in the kinetic parameter estimates. Three series of test runs were performed, the first one with a large excess of CO2 to benzene, the second with a stoichiometric amount of CO2, and the third with CO2 varying between the stoichiometric and large excess amounts. In the case of excess CO2, a content of 10 vol % was used at temperature of 800 °C and above and 5 vol % at 750 °C to prevent carbonation of the dolomite. Experimental conditions are summarized in Table 1. The test runs were started by calcinating and stabilizing the dolomite catalyst. When the reactor was connected to the gas-mixing system, a nitrogen flow of 0.03 mn3/h through the reactor was started. The calcination was carried out in situ at 1073 K for 1 h under the nitrogen flow. After the calcination, the flows of all reagent gases were started and the reactor was heated to the required temperature. The catalyst was then allowed to stabilize for 1.5-2 h before test runs were started. The weight of the original uncalcined sample was used as the basis in the kinetic analysis of the results. Analytical Methods. A Hewlett-Packard gas chromatograph 5790 A series equipped with a flame ionization detector was used for benzene analysis. A HP-1 (0.53 mm × 30 m) capillary column was used. The existence of other hydrocarbons was checked by adsorbing them in dichloromethane in gas washing bottles and by detecting them from the solvent. The contents of the permanent gas components (H2, O2, CO2, CO, CH4, and C2 hydrocarbons) were analyzed for gas samples taken into gas-sampling bulbs. The methods used for higher hydrocarbon and gas analyses have been described in detail by Simell and Bredenberg (1990). The contents of benzene, CO, CO2, and H2O were also measured continuously with a GASMET FT-IR analyzer in some test runs. Reagents and Catalyst. The materials used were of Finnish dolomite from the Kalkkimaa quarry and of silicon carbide produced by Arendal Smeltwerk AB. Their respective chemical compositions are presented in Table 2. The BET surface area of calcined dolomite
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Ind. Eng. Chem. Res., Vol. 36, No. 1, 1997
Table 2. Chemical Composition of the Bed Materials (wt %) dolomite SiC
Ca
Mg
Fe
Ni
Ti
Mn
Al
Si
Cl
S
O + res
19.0
11.0
1.5 0.1
