Energy Fuels 2010, 24, 2707–2715 Published on Web 03/15/2010
: DOI:10.1021/ef901529d
Biomass Gasification: Catalytic Removal of Tars over Zeolites and Nickel Supported Zeolites Prashanth Reddy Buchireddy,† R. Mark Bricka,*,† Jose Rodriguez,‡ and William Holmes‡ †
Chemical Engineering, Mississippi State University, Starkville, Mississippi 39762, and ‡Mississippi State Chemical Laboratory, Mississippi State University, Starkville, Mississippi 39762 Received December 14, 2009. Revised Manuscript Received February 23, 2010
Tars have been identified as one of the major impurities associated with biomass gasification fuel gas. Tar buildup can cause blockages, plugging, corrosion, and catalyst deactivation, resulting in serious operational and maintenance problems. Therefore, tar removal is essential to ensure economic and effective fuel gas utilization. This study investigates the catalytic activity of zeolites and nickel supported zeolites for tar removal. Zeolites with varying pore sizes and acidity were tested, to evaluate the effect of pore size and acidity on tar removal. ZY had better catalytic activity due to its larger pore size compared with ZSM5. Catalytic activity of zeolites increased with an increase in the acidity. ZY-5.2, which is more acidic, had better naphthalene conversions compared with ZY-80. Impregnation of nickel on zeolites improved the activity significantly due to the steam reforming ability of nickel. Long-term catalytic activity tests were performed, the results of which showed that nickel-supported ZY-30 and ZY-80 had the best naphthalene conversions, with naphthalene conversions of greater than 99%, followed by nickel-supported ZY-5.2, SiO2/Al2O3, and chabazite. Also, very little loss in activity over a 97 h test period was noticed for nickelsupported ZY-80 and ZY-30, compared with nickel-supported ZY-5.2, SiO2/Al2O3, and chabazite. Characterization of calcined and spent catalysts was performed using X-ray diffraction, thermogravimetric analysis, and surface area analysis, to investigate the lower catalytic activity of nickel-supported ZY-5.2, SiO2/Al2O3, and chabazite. The decreased catalytic activity is attributed to the coke deposition and catalyst surface area.
Also, the overall efficiency of the gasification process in terms of energy could be improved by converting tars to fuel gases,5 and the carcinogenic nature of tars poses an environmental and health impact that may need to be reduced.6 Therefore, tar removal is essential to ensure economic and effective fuel gas utilization. Removal of tars can be classified into three general categories: (a) physical, (b) thermal, and (c) catalytic.2 However, catalytic tar removal has received much attention because tars can be cracked or reformed into gaseous components increasing the overall efficiency of the gasification process. Abu ElRub et al.7 have classified tar removal catalysts into nine groups: calcined rocks, olivine, clay minerals, ferrous metal oxides, char, fluid catalytic cracking catalysts, alkali metals, activated alumina, and transition-metal-based catalysts, particularly nickel (Ni). But, not all of the above-mentioned catalysts have been investigated extensively. The most widely researched catalysts were dolomites, alkali metals, and nickel. Extensive research8-13 has been performed on dolomites since they are inexpensive, abundantly available, and show good
1. Introduction An increase in the demand for energy worldwide has prompted concerns regarding the possible depletion of fossil fuel sources. Hence, attention has been focused on developing alternative energy sources. Biomass is one such energy source and is considered as a potential feedstock for sustainable energy production.1 Biomass gasification is one of the thermochemical routes of converting biomass to fuel gases (producer gas or synthetic gas). These gases can be used in the generation of electricity, production of transportation fuels and chemicals, hydrogen fuel production,2 etc.; however, these fuel gases contain many impurities, such as particulates, ammonia, hydrogen sulfide, hydrogen chloride, alkali metals, metals, and tars.3 Among these impurities, tars have been identified as a major concern within the gasification process. Tars are a complex mixture of polynuclear aromatic hydrocarbons (PAHs) having molecular weights higher than benzene. They range from a single-ring compound, phenol, to a six-ring compound, coronene.4 Tar buildup can cause blockages, plugging, corrosion, and catalyst deactivation, resulting in serious operational and maintenance problems.
(6) Sutton, D.; Kelleher, B.; Ross, J. R. H. Fuel Process. Technol. 2001, 73, 155–173. (7) Abu El-Rub, Z.; Bramer, E. A.; Brem, G. J. Ind. Eng. Chem. Res. 2004, 43, 6911–6919. (8) Orio, A.; Corella, J.; Narvaz, I. J. Ind. Eng. Chem. Res. 1997, 36, 3800–3808. (9) Simell, P. A.; Leppalahti, J. K.; Bredenberg, J. B. Fuel 1992, 71, 211–218. (10) Corella, J.; Aznar, M. P.; Gil, J.; Caballero, M. A. Energy Fuels 1999, 13 (6), 1122–1127. (11) Delgado, J.; Aznar, M. P.; Corella J. Ind. Eng. Chem. Res. 1996, 37, 3637–3643.
*To whom correspondence should be addressed. E-mail: bricka@ che.msstate.edu. (1) Devi, L.; Ptasinski, J. K.; Franns, J. J. G. J.; Paasen, S. V. B.; Burgman, P. C. A.; Kiel, J. H. A. Renewable Energy 2005, 30, 565–587. (2) Milne, T. A.; Abatzoglou, N.; Evans, R. J. NREL/TP-570-25357, NREL, Golden, CO, 1998. (3) Simell, P. A.; Bredenberg, J. B. Fuel 1990, 69, 1219–1225. (4) Bridgewater, A. V. Fuel 1995, 74, 631–653. (5) Simell, P.; Xahlberg, P. S.; Kurkela, E.; Albrecht, J.; Deutsch, S.; Sjostrom, K. Biomass Bioenergy 2000, 18 (3), 19–38. r 2010 American Chemical Society
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catalytic activity. Some of the problems reported with these materials are a decrease in mechanical strength over time,11,12 deactivation due to coke formation,11,13 and their inability to achieve the required low tar concentrations for many end use applications.14 The other group of catalyst materials investigated is the alkali metals. Alkali metals are typically added to the feed by dry mixing or wet impregnation methods. Although, they show good catalytic activity toward tar removal, several disadvantages exist. The alkali metal catalyst loses its activity due to particle agglomeration and volatilizes at high temperatures.7 Also, the recovery of the catalyst material is very difficult since it is usually added along with biomass. This makes the process expensive owing to separation of char and ash from alkali metals in addition to ash disposal issues.15 The most widely investigated catalysts for tar removal are nickel-supported catalysts. Nickel catalysts exhibit high activity toward steam reforming and dry reforming, aiding in complete tar removal and increasing the hydrogen and carbon monoxide content of the fuel gases.15 The activity of commercial nickel-supported catalysts has been reported to be 8-10 times higher than that of dolomites under similar operating conditions.16 Further, commercial steam reforming catalysts synthesized for steam reforming of heavy hydrocarbons were found to be more active for tar removal than steam reforming catalysts for light hydrocarbons.17 The main disadvantage of using the nickel catalysts is the loss of catalytic activity due to carbon fouling and sintering of nickel. Although the loss of activity of catalysts due to carbon deposition can be restored, sintering of nickel particles causes irreversible loss of activity.15 Dolomites and olivine catalysts also have a good activity toward tar removal, but their inability to achieve required low tar concentrations have prompted several researchers to investigate the use of nickel-supported dolomite and olivine catalysts.14,18-20 Wang et al. reported that 97% tar removal was achieved at an operating temperature of 750 °C and a space velocity of 12 000 h-1. Also, the activity of Ni/dolomite was comparable to commercial steam reforming catalysts.18 Swierczynski et al. investigated the use of Ni/olivine as a tar reforming catalyst. They reported that, at temperatures above 650 °C, 100% removal of toluene, which they used as a model compound, was achieved and coke formation was negligible.20 On the basis of these reported results, it can be concluded that Ni/dolomite and Ni/olivine catalysts have excellent activity for tar removal. In addition, Ni/dolomite and Ni/olivine catalysts are relatively cheaper compared with the common commercial steam reforming catalysts. Fluid catalytic cracking (FCC) catalysts are another group of materials that have the potential to eliminate tars from fuel
gas. Zeolites fall under this class of catalysts and are defined as “crystalline aluminosilicate materials with tetrahedral framework structure enclosing cavities occupied by cations and water molecules”.21 However, very limited research has been performed to evaluate zeolites for tar removal applications. Herguido et al.7 evaluated the use of spent FCC catalysts in equilibrium in a fluidized bed gasifier and found that tar was reduced from 78 g/m3 to 9 g/m3, with continuous regeneration of the catalyst. Dou et al. reported 1-methyl naphthalene conversions of more than 95% using zeolite Y at 550 °C and a gas hourly space velocity of 3000 h-1. However, their analysis of the product gas did not include the cracked compounds, such as lower aromatics or alkyl aromatics, which are considered tars.22 The advantages of zeolites over amorphous catalysts are related to their acidity, better thermal/hydrothermal stability, better resistance to nitrogen and sulfur compounds, tendency toward low coke formation, and easy regenerability.23 The other advantages with zeolites are their relatively low price and experience gained using these catalysts in FCC units, which offers better practical insight. However, the main disadvantage with these catalysts is the rapid deactivation caused by coke formation.7 Transition metal (group VIII) based catalysts are good steam and dry reforming catalysts. Nickel is the most widely used metal for steam reforming applications due to economic reasons and also has a relatively high activity compared with Co, Pt, Ru, and Rh.24,15 Utilizing the advantages of using zeolites and nickel metal mentioned above, the present study investigates the use of zeolites and nickel-supported zeolites as a potential tar removal catalyst. To facilitate this study, naphthalene was chosen as a model tar compound because it is one of the most stable tars, hence very difficult to remove.25,26 Nickel-supported zeolites were prepared, characterized, and tested for their efficiency in reforming tars. Also, experiments were conducted on SiO2/Al2O3 and Ni-supported SiO2/Al2O3 for comparison, since silica-alumina (SiO2/Al2O3) is a common catalyst support.
(12) Nanaez, I.; Orio, A.; Aznar, M. P.; Corella, J. J. Ind. Eng. Chem. Res. 1996, 35 (7), 2110–2120. (13) Simell, P. A.; Hakala, N.; Haario, H. E. J. Ind. Eng. Chem. Res. 1997, 36 (1), 42–51. (14) Wang, T. J; Chang, J.; Wu, C. Z.; Fu, Y.; Chen, Y. Biomass Bioenergy 2005, 28, 508–514. (15) Sutton, D.; Kelleher, B.; Ross, J. R. H. Fuel Process. Technol. 2001, 73, 155–173. (16) Olivares, A.; Aznar, M. P.; Caballero, M. A.; Gil, J.; Franes, E.; Corella, J. J. Ind. Eng. Chem. Res. 1997, 36, 5220–5226. (17) Aznar, M. P.; Caballero, M. A.; Gil, J.; Martin, J. A.; Corella, J. J. Ind. Eng. Chem. Res. 1998, 37, 2668–2680. (18) Wang, T.; Chang, J.; Lv, P. Energy Fuels 2005, 19, 22–27. (19) Zhang, R.; Wang, Y.; Brown, R. C. Energy Convers. Manage. 2007, 48, 68–77. (20) Swierczynski, D.; Libs, S.; Courson, C.; Kiennemann Appl. Catal., B: Environ. 2007, 74, 211–222.
(21) Bhatia, S. Zeolite Catalysts: Principles and Applications; CRC Press: Boca Raton, FL, 1989. (22) Dou, B.; Gao, J.; Sha, X.; Baek, S. W. Appl. Thermal Eng. 2003, 23, 2229–2239. (23) Scherzer, J.; Gruia, A. J. Hydrocracking Science and Technology; Taylor & Francis, Inc.: Oxford, U. K., 1996. (24) Garcia, L.; Sanches, J. L.; Salvador, M. L.; Bilbao, R.; Arauzo, J. Developments in Thermochemical Biomass Conversion; Blackie Academic and Professonal: London, 1997; pp 1158-1169. (25) Devi, L.; Ptasinski, J. K.; Janssen, G. J. J. F. J. Ind. Eng. Chem. Res. 2005, 44, 9096–9104. (26) Coll, R.; Salvado, J.; Farriol, X.; Montane, D. Fuel Process. Technol. 2001, 74, 19–31. (27) Majdan, M.; Pikus, S.; Rzaczy nska, Z.; Iwan, M.; Maryuk, O.; Kwiatkowski, R.; Skrzypek, H. J. Mol. Sci. 2006, 791 (1-3), 53–60. (28) Nomura, M.; Akagi, K.; Murata, S.; Matsui, H. Catal. Today 1996, 29, 235–240.
2. Experimental Section 2.1. Catalyst Preparation. Commercially available zeolites were obtained from Zeolyst International, Pennsylvania, the properties of which are provided in Table 1.27,28 Chabazite was obtained from Zeox Mineral Materials Corporation, Tucson, Arizona. Nickel-supported zeolites were prepared by the wet impregnation technique. The desired amount of nickel nitrate (Ni(NO3)2 3 6H2O) was dissolved in deionized water, to which a predetermined weight of zeolite was added. This solution was continuously stirred and heated to 105 °C until all the water evaporated. The catalyst was then dried overnight in an oven at 110 °C. After drying, the catalyst was calcined in the presence of
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through a static mixer before entering the catalytic reactor to ensure proper mixing. All the tubing was heat traced and maintained at a temperature of 300 °C to prevent any condensation of naphthalene in the lines, preventing blockage. The gas exiting the catalytic reactor was passed through isopropyl alcohol to condense any unconverted tar compounds before being exhausted. Periodically, gas samples were collected in 1 L bags for additional analysis. 2.3. Sampling and Analysis. Unconverted naphthalene in the gas exit stream was measured by purging the exiting gas stream from the catalytic reactor through a series of impingers. The impinger setup was performed according to the guideline for sampling and analysis of tars and particulates in biomass producer gases.30 However, the impinger size was reduced to a volume capacity of 25 mL. The temperature of the impingers placed in the ice bath was maintained at -5 to -10 °C using an ice and salt mixture. The temperature of the gases leaving the last impinger was 10-15 °C. Purged isopropyl alcohol from the impingers was then collected and stored at 4 °C for further analysis. An Agilent 6890N gas chromatograph with flame ionization detector (GC-FID) was used to quantify the unconverted tars condensed in the isopropyl alcohol. A DB-5 (0.25 mm X 30 m) capillary column was used on the GC-FID and instrument parameters were set to separate and quantify aromatics (8 compounds) and PAHs (22 compounds), including naphthalene listed in the tar sampling guide.30 Samples were prepared with n-decane as an internal standard. The gas samples collected in the 1 L bags were analyzed using a GC (SRI 8610) equipped with a thermal conductivity detector. Gases H2, CO, N2, CO2, and CH4 were separated and quantified using a Cabosphere 80/100 (6 ft X 1/8”) stainless steel column (Alltech Associates Inc., U.S.A). 2.4. Experimental Conditions and Formulas. The effectiveness of catalytic tar removal using zeolites and nickel-supported zeolites was investigated in this study. Several experiments were performed to accomplish the above-mentioned objective. Different types of zeolites, ZSM5, ZY, Zβ, and chabazite were tested to evaluate their activity toward naphthalene removal. Also, the effect of zeolites’ acidity on naphthalene conversion was evaluated using ZY with different SiO2/Al2O3 ratios ranging from 5.2 to 80. These experiments were conducted at 750 °C, a steam to carbon ratio (S/C) of 5, a gas hourly space velocity (GHSV) of 12 800 h-1, and a naphthalene loading of 12 g/m3. Also, nickel-supported zeolites were tested for naphthalene conversion activity. Zeolites 7.5% Ni/ZY-5.2, 7.5% Ni/ZY-30, 7.5% Ni/ZY-80, and 7.5% Ni/ SiO2/Al2O3 were tested at temperatures ranging from 550 °C to 750 °C, a steam to carbon ratio (S/C) of 5, a gas hourly space velocity (GHSV) of 12 800 h-1, and a naphthalene loading of 12 g/m3. Although tar removal activity is more favorable at higher temperatures, our studies were limited to 750 °C considering the thermal stability of zeolites. Thermal stability of zeolites varies widely from 700 °C for lower silica zeolites to 1300 °C for highly silicaceous zeolites.31 Since activities of different zeolites were being evaluated in this study, the temperatures were limited to 750 °C. Naphthalene conversions, the steam to carbon ratio (S/C), and the gas hourly space velocity (GHSV) were calculated on the basis of the formulas provided below. Cin -Cout naphthalene conversion, % ¼ 100 ð1Þ Cin
Figure 1. Experimental setup for tar removal studies. Table 1. Properties of Catalysts catalyst
SiO2/Al2O3
pore size, A˚
surface area, m2/g
Si/Al Na-Chabazite (2) ZSM5 Zβ ZY-5.2 (3) ZY-30 ZY-80
28 4.0 24 25 5.2 30 80
4.1 (2) 5.5 (3) 6.7 (3) 8.0 (3) 8.0 8.0
381 520 425 680 750 750 780
air at a heating rate of 5 °C/min until 550 °C was reached, and this temperature was maintained for 2 h. The calcined catalysts were palletized, crushed, and sieved to particle sizes between 20 and 30 mesh to avoid internal mass transfer limitations.29 Six different catalysts were prepared by the above-mentioned method: 7.5% Ni/ZY-5.2, 7.5% Ni/ZY-30, 7.5% Ni/ZY-80, 7.5% Ni/Si-Al, and 7.5% Ni/Chabazite. 2.2. Experimental Setup. A schematic representation of the experimental apparatus for naphthalene removal studies is presented in Figure 1. Experiments were conducted in a clear fused quartz compression tube with a 10.5 mm inside diameter and a length of 380 mm. The catalyst was held in place by quartz wool placed in the uniform temperature zone within the furnace. A type K thermocouple was placed at a distance of 10 mm from the catalyst bed to monitor the catalyst temperature. The temperature of the catalyst bed was monitored and recorded with a 4 channel data logging thermometer (Sper Scientific). The catalytic quartz reactor was heated in a tube furnace (Barnstead Thermolyne 2110), provided with a temperature controller. All the experiments were performed using syngas (CO, 20%; H2, 20%; CO2, 9%; CH, 3%; N2, 48%) as feed gas. This gas resulted in exposing the catalyst to a reducing environment. The flow of the syngas was controlled using a mass flow controller (Omega FMA 5508). Naphthalene was fed to the catalytic reactor using a stainless steel saturator that holds naphthalene. A naphthalene saturator was heated to a predetermined temperature using a furnace. Naphthalene vapors produced during heating were carried to the catalytic reactor by sparging preheated syngas through the naphthalene saturator. The syngas flow rate to the saturator was controlled by a mass flow controller (Omega FMA 5508). The amount of naphthalene input to the reactor was controlled either by varying the saturator temperature or syngas flow rate. The amount of water fed was controlled using a Hewlett-Packard 1050 high performance liquid chromatograph (HPLC) pump. Water enters the reactor through a heated capillary. Pressure gauges were attached to the naphthalene input line and before the reactor to monitor any blockages or clogs in the system. Naphthalene feed, steam, and syngas pass
S=C ¼ moles of water=½moles of carbonnaphthalene GHSV, h -1 ¼
Vo V
ð2Þ ð3Þ
(30) Paasen, S. V. B. V.; Neeft, J. P. A.; Knoef, H. A. M.; Buffinga, G. J.; Zielke, U.; Sjostrom, K.; Brage, C.; Hasler, P.; Simell, P. A.; Suomalainen, M.; Dorrington, M. A.; Thomas, L. Guideline for Sampling and Analysis of Tars and Particulates in Biomass Producer Gases, ECN-C-02-090, ECN, Petten, The Netherlands, 2002.
(29) Kinoshita, C. M.; Wang, Y; Zhou, J. C. J. Anal. Appl. Pyrol. 1994, 29, 169–81.
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Figure 2. Naphthalene conversion with zeolites (T, 750 °C; S/C, 5.0; GHSV, 12 800 h-1; naphthalene loading, 12 g/Nm3).
Figure 3. Variation of naphthalene conversion with SiO2/Al2O3 of zeolites (T, 750 °C; S/C, 5.0; GHSV, 12 800 h-1; naphthalene loading, 12 g/Nm3).
where Cin is the concentration of naphthalene entering the reactor, mg/min; Cout is the concentration of naphthalene exiting the reactor, mg/min; Vo is the volumetric flow rate of the gas mixture (CO, CO2, H2, CH4, N2) @ STP, m3/h-1; and V is the volume of the catalyst bed, m3.
Figure 2 shows that ZY-30 had the best activity toward naphthalene conversion followed by Zβ and ZSM5, with conversions of 32, 30, and 19%, respectively. This decrease in the activity can be attributed to the variation of pore sizes and surface area of the zeolites. The external surface area of zeolites is very low in comparison, approximately 1-2% of the internal pore surface.21 Hence, most of the available active sites are present in the zeolite pores. From Table 1, it can be seen that the pore size of ZSM5 is 5.5 A˚ compared with ZY-30, which has a pore size of 8 A˚. The size of the naphthalene molecule estimated by molecular orbital calculations was 4.9-6.8 A˚.32 Since the size of the naphthalene molecule is larger than the pore size of ZSM5, it is very difficult for the naphthalene molecule to diffuse into the pores of ZSM5. Hence, the naphthalene molecules would have restricted access to the active acidic sites available inside the zeolite pores, decreasing its activity. On the contrary, ZY-30 has a pore size of 8 A˚, which allows naphthalene to diffuse into the zeolite pores, increasing naphthalene conversion. Also, the surface area of ZSM5 (425 m2/g) is lower than that of Zβ (750 m2/g) and ZY-80 (750 m2/g). The lower surface area corresponds to a relatively lesser number of active sites. Hence, the activity of ZSM5 could have been lower than that of Zβ and ZY-80. 4.2. Effect of Zeolite’s Acidity on Naphthalene Conversion. The effect of zeolite acidity on naphthalene conversion is shown in Figure 3. An increase in the SiO2/Al2O3 ratio corresponds to a decrease in the acidity.23 Hence, ZY-5.2 with a SiO2/Al2O3 of 5.2 is more acidic compared to ZY-80 with a SiO2/Al2O3 of 80. Figure 3 shows that ZY-5.2 had the highest activity and SiO2/Al2O3 had the least activity toward naphthalene conversion. This could be due to the higher acidity of zeolites compared with amorphous SiO2/Al2O3. Zeolite catalysts are more active than amorphous catalysts due to their higher acidic strength both in terms of their strength and number of acid sites.23,33 Also, it can be seen that naphthalene conversion increased from 33 to 55%, with a decrease in SiO2/Al2O3 from 80 to 5.2 for ZY. This increase can once again be attributed to the higher acidity of ZY-5.2 compared with ZY-80. However, naphthalene conversion of Na-chabazite with a SiO2/Al2O3 of 4.0 was lower than ZY with a SiO2/Al2O3 of 80. This could be due to the presence of sodium, which neutralizes the acidic sites on chabazite.
3. Catalyst Characterization Catalyst samples were characterized by powder X-ray diffraction on a Rigaku Ultima III X-ray Diffractometer using Cu KR (λ = 0.154 nm, 40 kV, 44 mA) radiation to determine the nickel phase. The diffractographs were scanned in the range of 2θ between 10° and 70°, with a sampling width of 0.02° and scan speed of 4°/min. Transmission electron microscopy (TEM) was conducted on a JEOL model JEM100 C XII (80 kV). Catalyst specimens were prepared by dispersing the powdered samples in isopropyl alcohol. Drops of alcohol mixed with ground catalyst were transferred onto a Cu grid and dried at room temperature. Microscopy analysis was performed on the specimens prepared to determine the size of nickel particles. Surface area measurements of the support and prepared catalysts was obtained from nitrogen adsorption isotherms at 77 K, using a Quantachrome AUTOSORB-I instrument. Samples were outgassed for 4 h at 180 °C prior to making these measurements. Thermogravimetric analysis was performed on a Versa Therm HS Thermogravimetric Analyzer (TGA) manufactured by Thermo Cahn Inc. (Waltham, MA), to determine the amount of coke deposited on the catalyst. 4. Results and Discussions 4.1. Catalytic Activity of Zeolites. Naphthalene conversion studies were performed with different zeolites at 750 °C, a steam to carbon ratio (S/C) of 5, a gas hourly space velocity (GHSV) of 12 800 h-1, and a naphthalene loading of 12 g/m3. Several zeolites were tested, the properties of which are presented in Table 1. Figure 2, shows the activity of the zeolites toward naphthalene conversion in comparison with thermal cracking. The catalytic activity of the zeolites is attributed to the acidic nature of zeolites. The acidic nature of zeolites aids in the naphthalene conversion via cracking reaction as shown in eq 4. C10 H8 w C þ gas þ lower hydrocarbons
ð4Þ
(32) Isoda, T.; Maemoto, S.; Kusakabe, K.; Morooka, S. Energy Fuels 1999, 13, 617–623. (33) Rosemary, S. Molecular Sieves; 2nd ed.; Springer-Verlag: New York, LLC, 1998.
(31) Auerbach, S. M.; Carrado, K. A.; Dutta, P. K. Handbook of Zeolite Science and Technology; Taylor & Francis, Inc.: Oxford, U. K., 2003.
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Figure 4. Post run catalyst samples.
Figure 6. Variation of product gas composition with catalyst, initial gas composition: H2, 20%; CO, 20%; CO2, 9%; CH4, 3% (T, 750 °C; S/C, 5.0; GHSV, 12 800 h-1; naphthalene loading, 12 g/Nm3). Table 2. Coke Analysis coke deposited, %
Figure 5. Variation of naphthalene conversion with zeolites and nickel supported zeolites (T, 750 °C; S/C, 5.0; GHSV, 12 800 h-1; naphthalene loading, 12 g/Nm3).
ð5Þ
C10 H8 þ 10CO2 w 20CO þ 4H2
ð6Þ
650° C
750° C
8.93% 7.74% 8.43% 2.30%
2.62% 1.21% 0.86% 0.72%
pared with Ni/ZY-80. Table 2 shows the amount of coke deposited on different catalysts at 750 °C. It can be seen that coke deposited increases from 0.7 to 2.62% with an increase in the acidity of the Ni zeolites. Formation of coke deactivates the catalyst by poisoning the active nickel sites and acidic sites and blocking the pores of the zeolites.21,23 Hence, the activity of Ni zeolites toward naphthalene conversion decreased with an increase in the acidity. The other reason for the decrease in activity of Ni zeolites with an increase in acidity could be due to thermal degradation and/or formation of inactive nickel aluminates. Catalyst deactivation by thermal degradation is a result of the loss of catalytic surface area and/or nickel particle growth, which occurs at higher temperatures (>500 °C). Thermal stability of zeolites tends to increase with an increase in the SiO2/ Al2O3.33 Therefore, an increasing order in thermal stability of zeolites tested is Ni/ZY-80 > Ni/ZY-30 > Ni/ZY-5.2. Since Ni/ZY-80 is thermally more stable than Ni/ZY-5.2, the activity of the former could be higher toward naphthalene conversion. Also, transformation of the active nickel phase to an inactive nickel aluminate phase could contribute to a loss of catalyst activity. Since, the alumina content of Ni/ZY5.2 is higher than that of Ni/ZY-80, the probability of formation of an inactive nickel aluminate phase is greater. This might be another reason for the lower catalytic activity of Ni/ZY-5.2 compared with Ni/ZY-80. Variation of exit gas composition with the different zeolites tested is presented in Figure 6. These gas samples were collected at the end of the experimental run. These exit gas composition variations are due to several thermodynamically favorable reactions occurring in series and/or parallel. Figure 6 shows that the H2 and CO content is higher for Ni/ zeolites (H2, 21-21.3%; CO, 20.5-21.2) compared with the zeolites (H2, 19.5-20.5%; CO, 19-19.9%). This increase in the H2 and CO contents is attributed to the presence of nickel metal on the zeolites. Since Ni is a good steam and dry reforming catalyst, naphthalene conversions increase for
An increase in the acidity of zeolites increases the activity of the catalyst while decreasing its resistance toward coke formation. Figure 4 shows a photograph of post-run catalyst samples for ZY-80, ZY-30, and ZY-5.2. The intensity of darkness increased from ZY-80 to ZY-5.2. This could be due to an increase in coke formation with an increase in the acidity of the catalyst. 4.3. Catalytic Activity of Nickel-Supported Zeolites. Experiments were performed on nickel-supported ZY with varying SiO2/Al2O3 and were compared with ZY to evaluate the catalytic activity toward naphthalene conversion, the results of which are presented in Figure 5. As shown in the figure, nickel-supported catalysts had a significant improvement in activity toward naphthalene conversion for all the zeolites tested. Naphthalene conversion improved significantly from 13 to 87% and 33 to 98% for SiO2/Al2O3 and ZY-80, respectively, with the impregnation of nickel. The improved performance of nickel-supported zeolites is attributed to the steam and dry reforming activity of impregnated nickel as shown in eqs 5 and 6. C10 H8 þ 10H2 O w 10CO þ 14H2
catalyst 7.5% Ni-ZY 5.2 7.5% Ni-ZY 30 7.5% Ni-ZY 80 7.5% Ni-SiO2/Al2O3
Figure 5, also shows that the activity of nickel-supported zeolites increases with a decrease in the zeolites’ acidity. Naphthalene conversions reported were 78 and 99% for Ni/ZY-5.2 and Ni/ZY-80, respectively. In contrast, the activity of zeolites decreased with a decrease in the acidity. This anomaly could be attributed to the coke formation on the catalyst surface. Coke formation increases with an increase in acidity, as discussed earlier. Since Ni/ZY-5.2 is more acidic, the ability of coke formation is greater com2711
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Figure 8. Effect of temperature on gas composition (catalyst, 7.5% Ni/ZY-80; S/C, 5.0; GHSV, 12 800 h-1; naphthalene loading, 12 g/ Nm3).
Figure 7. Effect of temperature on naphthalene conversion (S/C, 5.0; GHSV, 12 800 h-1; naphthalene loading, 12 g/Nm3).
reactions occurring in series or simultaneously might have contributed to the changes in exit gas compositions. As shown in the figure, at 550 °C, a decrease in the H2 and CO content with a corresponding increase in CH4 and CO2 from the initial values is observed. This variation could be due to a series of thermodynamically favorable reactions taking place simultaneously. At lower temperatures (550 °C), methanation is thermodynamically favored, as shown in eq 12,15 increasing the methane content of the exit gas stream, while consuming CO and H2. We attribute this increase in CO2 content to the water gas shift activity, as shown in eq 9. As the temperature increases from 550 to 750 °C, the H2 and CO content increases, with a decrease in CH4 and CO2. This decrease in methane could be due to methane reforming, as shown in eqs 7 and 8. Also, a reverse water gas shift reaction, which is more favorable above 675 °C, could have contributed to an increase in CO, with a corresponding decrease in CO2 content.20
Ni/zeolites as shown in Figure 5. This increase in naphthalene conversions via eqs 5 and 6 might have increased the CO and H2 content of the exit gas. Also, a decrease in the methane and CO2 compositions is observed for Ni/zeolites compared with those for zeolites. This decrease could be attributed to the methane reforming and reverse water gas shift activity, as shown in eqs 7, 8, and 9. CH4 þ H2 O S CO þ 3H2
ð7Þ
CH4 þ CO2 S 2CO þ 2H2
ð8Þ
CO þ H2 O S CO2 þ H2
ð9Þ
4.4. Effect of Catalyst Bed Temperature. Experiments were conducted on nickel-supported zeolites in the temperature range of 550-750 °C, with a S/C ratio of 5.0, a GHSV of 12 800 h-1, and a naphthalene loading of 12 g/m3, to evaluate the effect of temperature on naphthalene conversion. Figure 7 shows the effect of temperature on the different zeolites tested. As shown in the figure, naphthalene conversion increases with an increase in temperature from 550 to 750 °C with the catalysts tested. An increase in the conversion from 55 to 76% and 84 to 99% was noticed with Ni/ZY-5.2 and Ni/ZY-80, respectively. This increase in the conversion is attributed to the steam reforming, which is thermodynamically favorable at higher temperatures. Also, poisoning of the active sites due to coke formation on the catalyst has an influence on the naphthalene conversions. As shown in Table 2, the amount of coke deposited on the catalyst decreased with an increase in temperature from 650 to 750 °C for all the catalysts tested. This decrease in the coke formation with an increase in temperature could be due to the water gas and boudouard reactions, as shown in eqs 10 and 11. The reaction of coke with water and CO2 is thermodynamically more favorable at higher temperatures, producing CO and H2.34,35 C þ CO2 S 2CO
ð10Þ
C þ H2 O S CO þ H2
ð11Þ
CO þ 2H2 S CH4
ð12Þ
In addition to methanation, methane reforming, water gas shift, and reverse water gas shift reactions might have been involved and contribute to the variations in the gas compositions, reactions 1-5 also might have contributed to the variations in the exit gas composition. 4.5. Estimation of Apparent Activation Energy. Apparent activation energy was calculated for naphthalene conversion assuming a first order reaction model. This assumption was made by several researchers and has been widely accepted.36-39 The rate of removal of naphthalene (-rnap) can be approximated by the following equation (eq 13). ð13Þ -rnap ¼ Kappnap Cnap The first order apparent rate constant (Kappnap) under plug flow conditions can be calculated using eq 14, which was derived from an integral plug flow reactor model. ½ -lnð1 -XÞ ð14Þ Kappnap ¼ τ (36) El-Rub, A.; Bramer, E. A.; Brem, G. Proceedings of Expert Meeting on Pyrolysis and Gasification of Biomass and Waste, France, 2002; pp 337-346. (37) Jess, A. Chem. Eng. Possessing 1996, 35, 487–494. (38) Devi, L. Catalytic Removal of Biomass Tars; Olivine as Prospective in-Bed Catalyst for Fluidized-Bed Biomass Gasifiers, Ph.D. thesis, Technical University of Eindhoven, Eindhoven, The Netherlands, 2005. (39) El-Rub, A. Biomass char as in-situ catalyst for tar removal in gasification systems, Ph.D. thesis, Twente University, Enschede, The Netherlands, 2008.
Figure 8 presents the effect of temperature on gas composition variation for the experiment performed with 7.5% Ni/ZY-80 catalyst. Several thermodynamically favorable (34) Basu, P.; Basu, B. Combustion and Gasification in Fluidized Beds; Taylor & Francis, Inc.: Oxford, U. K., 2006. (35) Tomita, A. Catal. Surveys Jpn 2001, 5 (1), 17–24.
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Table 3. Comparison of Kinetic Data for Decomposition of Naphthalene over Various Catalysts catalyst
temperature, °C
reaction medium
residence time, s
activation energy, KJ/mol
frequency factor
Ni/ZY-80 Ni/ZY-30 Ni/ZY-5.2 Ni/Si-Al
550-750 550-750 550-750 550-750
H2O þ CO2 þ CO þ H2 H2O þ CO2 þ CO þ H2 H2O þ CO2 þ CO þ H2 H2O þ CO2 þ CO þ H2
0.3 0.3 0.3 0.3
31.8 18.4 20.1 14.4
3235 309 257.1 162.2
Figure 9. Estimation of apparent activation energy for different catalysts by Arrhenius plot (S/C, 5; naphthalene loading, 12 g/m3; GHSV, 12 800 h-1).
Figure 10. Naphthalene conversion with time on stream for catalysts tested (T, 750 °C; S/C, 5; naphthalene loading, 12 g/m3; GHSV, 12 800 h-1).
Temperature dependence on an apparent rate constant for naphthalene conversion was determined by using the Arrhenius equation as shown below (eq 15), from which the apparent activation energy (Ea) and frequency factor (Koappnap) can be calculated.
Table 4. Surface Area of Fresh and Spent Catalysts (BET)
Kappnap ¼ Koappnap expð -Ea=RTÞ
surface area, m2/g
ð15Þ 3
where Cnap is the concentration of naphthalene, mol/m ; X is naphthalene conversion; τ is the gas residence time, s; R is the universal gas constant, (8.314 10-3, KJ/mol K); and T is the temperature, K. Tests were performed in the range of 550-750 °C, to determine the activation energy with a naphthalene loading of 12 g/m3, over different Ni/zeolites. Activation energy and frequency factors were calculated from the Arrhenius plot shown in Figure 9 and are presented in Table 3. Activation energies for different catalysts varied between 14.4 and 31.8 kJ/mol, and the frequency factors ranged from 160 to 3200. It should also be noted that the lower activation energy values could also be due to the external mass transfer limitations. 4.6. Long-Term Catalytic Activity. Tests were conducted with different catalysts for a period of 97 h to evaluate the long-term catalytic activity for naphthalene conversion. The results of these experiments are presented in Figure 10. From the figure, it can be seen that Ni/ZY-80 had the best activity, with naphthalene conversions of 99.7%, while Ni/Chabazite had the least naphthalene conversion of 34.6% at the end of 97 h. As shown in the figure, the decreasing order of activity of the catalysts tested was Ni/ZY-80 ≈ Ni/ZY-30 > Ni/ZY5.2 > Ni/Si-Al > Ni/Chabazite. The lower activity of Ni/ ZY-5.2 could be due to its high acidic nature compared with other zeolites, as discussed in section 4.3. Also, the surface areas of Ni/ZY-5.2 and Ni/Si-Al were 481 and 401 m2/g, respectively, as shown in Table 4. A higher surface area of the catalyst allows for better dispersion of metals on its surface, increasing the overall active metal surface area.40 Hence, the
catalyst
calcined
spent
% decrease
7.5% Ni/ZY-5.2 7.5% Ni/ZY-30 7.5% Ni/ZY-80 7.5% Ni/Chabazite 7.5% Ni/Si-Al
481 646 673 283 401
265.8 459.1 487.1 2.25 190.9
44.7 28.9 27.6 99.2 52.3
relatively lower surface area of Ni/ZY-5.2 and Ni/Si-Al could have resulted in a decreased activity toward naphthalene removal. The activity of Ni/ZY-30 was similar to that of Ni/ZY-80, with naphthalene conversions of 99.3%. Also, Ni/ZY-80 and Ni/ZY-30 were very stable with very little loss of activity and had naphthalene conversions of above 99%, over the 97 h time period tested. In contrast, a drop in the activity of Ni/ ZY-5.2, Ni/Chabazite, and Ni/Si-Al was observed, as shown in Figure 10. Naphthalene conversions dropped from 97 to 88.8% and 94 to 70% for Ni/ZY-5.2 and Ni/Si-Al, respectively. The drop in activity was very high with Ni/ Chabazite, with naphthalene conversions decreasing from 98 to 34%. The drop in the activity of the catalysts can be contributed to the formation of coke on the catalyst, a loss in surface area and sintering of nickel due to thermal degradation, and/or formation of NiAl2O4 spinel. The gas composition variation with time on stream for Ni/ ZY-30 is presented in Figure 11. There was a decrease in the H2 and CO contents from 27 to 23% and 22.2 to 20.8%, respectively, with a corresponding increase in CH4 and CO2 contents within the first 10 h, as shown in the figure. After 10 h, the gas composition was consistent for the rest of the experiment, with higher than initial values of H2 and CO. This variation in gas composition could be due to methane reforming, a reverse water gas shift, and steam and dry reforming activity of the catalyst, as discussed in section 4.4. However, the initial decrease in CO and H2 with a corresponding increase in CH4 and CO2 is possibly due to a decrease in the catalytic activity and selectivity for methane reforming and reverse water gas shift activity.
(40) Wen, Y. W.; Cain, E. J. Ind. Eng. Chem. Proc. Des. 1984, 23 (4), 627–637.
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Figure 12. Thermogravimetric analysis of spent catalysts.
Figure 11. Gas composition variation with time on stream, initial gas composition: H2, 19.5%; CO, 20%; CO2, 10%; CH4, 3%, (T, 750 °C; GHSV, 12 800 h-1; S/C, 5; catalyst, 7.5% Ni/ZY-30; naphthalene loading, 12 g/m3). Table 5. Amount of Coke Deposited on the Spent Catalyst (TGA) catalyst
coke deposited, %
7.5% Ni/ZY-5.2 7.5% Ni/ZY-30 7.5% Ni/ZY-80 7.5% Ni/Chabazite 7.5% Ni/Si-Al
12.82 10.35 6.18 3.21
A thermogravimetic analysis was performed to determine the amount of coke deposited on the spent catalyst samples, the results of which are provided in Table 5 and Figure 12. The amount of coke deposited increased with an increase in the catalyst acidity. Ni/ZY-5.2 had the maximum amount of coke deposited with 12.8%, followed by Ni/ZY-30 with 10.35%. The presence of coke deactivates the catalyst by covering the active sites, as discussed in section 4.3. Hence, naphthalene conversion was maximum with Ni/ZY-80, which had the least amount of coke deposited followed by Ni/ZY-30 and Ni/ZY-5.2, respectively. In contrast, the activities of Ni/Si-Al and Ni/Chabazite were lower than that of the Ni zeolites, although the amount of coke deposited was lower at 3.2% and 0.4%, respectively. This could be due to several reasons, which will be discussed later. A comparison of coke deposited shown in Tables 2 and 5 shows that the amount of coke deposited increased with an increase in time on stream. The amount of coke deposited on Ni/ZY5.2 increased from 2.5 to 12.8% with an increase in time on stream from 2.5 to 97 h. A similar trend was observed with the other zeolites, which suggests the possibility of coke being deposited over time. The weight of the catalyst drops in the temperature range of 550-800 °C, as shown in Figure 12, which corresponds to the oxidation of coke in the presence of air. Since the coke is being oxidized in the temperature range of 550-800 °C, the coke is assumed to be in the form of graphitic carbon.20 The surface area of the calcined and spent catalysts was obtained from BET measurement, the results of which are presented in Table 4. The decrease in the surface area of Ni/ ZY-5.2 was 44.7%, followed by those of Ni/ZY-30 and Ni/ ZY-80. The thermal stability of zeolites increases with an increase in SiO2/Al2O3.33,41 Hence, Ni/ZY-80 with a SiO2/ Al2O3 of 80 was thermally more stable with a decrease in the surface area of 27% versus that of Ni/ZY-5.2. The surface
Figure 13. X-ray diffraction patterns of calcined Ni/ZY-5.2, Ni/ZY30, and Ni/ZY-80.
area of Ni/ZY-5.2 reduced by 44.7%. The decrease in surface area of the spent catalyst is a result of partial collapse of the internal pore structure of zeolites. This decrease in surface area will result in a reduced availability of acidic sites for reaction. Also, the accessibility of reactant molecules to nickel species encaged in the pores of the catalysts is reduced due to structural collapse. This will result in a reduced activity of catalysts.21,23 The decrease in surface area of Ni/Si-Al was greater than that of Ni/ZY at 52%. This decrease is attributed to the lower thermal stability of amorphous catalysts compared to zeolites,23 which could have resulted in lower naphthalene conversions, as shown in Figure 10. In contrast, the surface area of spent Ni/Chabazite decreased by 99.2%, which is attributed to the total structural collapse. This anomaly resulted in a reduced naphthalene conversion for Ni/Chabazite, which decreased steadily from 98 to 34% over time. X-ray diffraction was performed on calcined and spent catalyst samples, the diffractograms of which are presented in Figures 13 and 14. Figure 13 presents the diffraction patterns for calcined Ni/zeolites. Three peaks were identified at 2θ values of 37.2, 42.3, and 63°, which correspond to the NiO peaks. However, the peaks were less pronounced for Ni/ ZY-30. Also, the NiO peaks for Ni/ZY-5.2 were indistinguishable from the noise. This could be due to the interaction of nickel oxide with silica and alumina at high calcination temperatures that led to phase transformations. However, the transformed phases could not be distinctly observed due to the lower nickel concentrations. Further investigation needs to be done to explore the reasons behind the phase changes. Two distinct peaks were observed at 2θ values of 44.3 and 51.6°, as shown in Figure 14, which shows the XRD
(41) Flanigen, E. M.; Bekkum, H.; Jansen, J. C. Introduction to Zeolite Science and Practice, 2nd ed.; Elsevier Science and Technology: New York, 1991, pp 156.
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Nickel-supported zeolites have a very high tar removal potential. Zeolite Y had the better activity toward naphthalene conversions compared with Zβ and ZSM5. The higher activity of zeolite Y was attributed to its relatively larger pore size compared with Zβ and ZSM5. ZY-5.2 had better activity toward naphthalene conversion compared with ZY-30 and ZY-80. The higher activity of ZY-5.2 was attributed to its relatively higher acidic nature. Nickel-supported zeolites had a very high activity toward naphthalene removal compared with zeolites. Ni/ZY-80 had the best activity among the catalysts tested, with conversions of greater than 99.5. The superior activity of Ni/ZY-80 was attributed to the activity of Ni and the acidic nature of zeolite support. Long-term activity tests showed that Ni/ZY-80 and Ni/ZY30 achieved naphthalene conversions of greater than 99% with very little loss of activity at the end of the 97 h test duration. The activities of Ni/ZY-5.2, Ni/Si-Al, and Ni/Chabazite decreased over time, and the decrease in activity was associated with coke deposition and a loss of active surface area.
Figure 14. X-ray diffraction patterns of spent Ni/ZY-5.2, Ni/ZY30, and Ni/ZY-80.
patterns of spent catalysts. These two peaks correspond to metallic Ni species. The phase change from NiO to Ni for a spent catalyst is due to the reduction of NiO species during the experiment, which was carried out under a H2 and CO atmosphere.18,19,42-44 Conclusions Several zeolites and nickel-supported zeolites were examined for their potential to remove tars from gasification producer gas. The following conclusions are presented on the basis of the results obtained.
Acknowledgment. This work was supported by the Sustainable Energy Research Center (SERC) at Mississippi State University, funded under the Department of Energy. The authors would like to thank Dr. Mark White, Dr. Hossein Toghiani, and Dr. Amit Gujar for their technical input and assistance throughout the course of this study. The authors also thank Eugene Columbus and James Wooten for providing access to necessary instrumentation for this study. In closing, the authors would like to acknowledge Jared Jones and John Blakely for providing laboratory assistance.
(42) Moronta, A.; Iwasa, N.; Fujita, S.; Shimokawabe, M. Clays Clay Miner. 2005, 53 (6), 622–630. (43) Wang, S.; Zhu, H. Y.; Lu, G. Q. J. Colloid Interfac. Sci. 1998, 204, 128–134. (44) Laosiripojana, N.; Sutthisripok, W.; Assabumrungrat Chem. Eng. J. 2005, 112, 13–22.
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