Ind. Eng. Chem. Res. 1987,26, 1335-1339 solvents and SRC used in this study. We are also grateful to T. Chiba for his technical assistance.
Literature Cited
1335
Kline, E. A,; Harrison, M. E.; Farrnum, B. W. Prepr.-Am. Chem. SOC., Diu.Fuel Chem. 1982,27,18. Makabe, M.; Ouchi, K. Nippon Kagaku Kaishi 1980,867. Miller, R.L.; Silver, H. F. Energy Sources 1980,5,211. Miller, R. L.: Silver, H. F.; Hurtubise, R. J. Znd. Enp. Chem. Process Des. Deu. 1982,21,170. McMillen. D. F.: Malhotra. R.: Chann. S.: Ninenda. S. E. Proceedings of the International Conference-on Coay Science, Sydney, Oit 1985; p 91. Morita, Y., et al. J. Fuel SOC. Jpn. 1983,62,828. Murata, K.; Fukuju, Y. Symposium on Chemistry of Coal Liquefaction and Catalysis, Sapporo, 1985; p 1. Nagaishi, H.; Moritomi, H.; Sanada, Y.; Chiba, T . Proc. Conf. Coal Sci. Jpn, 21th 1984,21,23. Obara, T.; Yokono, T.; Sanada, Y. Fuel 1983,62, 813. Panvelker, S. V.; Ge, W.; Shah, Y. T.; Cronauer, D. C. Ind. Eng. Chem. Fundam. 1984,23,202. Raaen, V. F.; Roark, W. H. Fuel 1978,57,650. Shah, Y.T.; Golden, D.; Benson, S. J. Phys. Chem. 1977,81,1716. Shah, Y. T. Reaction Engineering in Direct Coal Liquefaction; Addison Wesley: London, 1981; p 162. Spencer, D. EPRI J . 1982,May, 31. Tagaya, H.; Chiba, K.; Sato, S.; Ito, K.; Sakurai, M. Nippon Kagaku Kuishi 1983,1172. Takeya, G.; Ito, M.; Suzuki, A.; Yokoyama, S. J . Fuel SOC.Jpn. 1964, 43,837. Uchino, H.; Yokoyama, S.; Satou, M.; Sanada, Y. J. Fuel SOC. Jpn. 1984,63,15. Whitehurst, D. D.; Mitchell, T. 0.;Farcasiu, M. Coal Liquefaction; Academic: New York, 1980; p 274. Yokono, T.; Marsh, H.; Yokono, M. Fuel 1981,60, 607. I
Aiura, M.; Masunaga, T.; Moriya, K.; Kageyama, Y. Fuel 1984,63, 1138. Bockrath, B. C.; Noceti, R. P. Prepr.-Am. Chem. SOC.,Diu. Fuel Chem. 1981,26,94. Brown, J. K.; Ladner, W. R. Fuel 1960,39,87. Chiba, K.; Tagaya, H.; Sato, S.; Watanabe, T. J . Fuel SOC. Jpn. 1984, 63,195. Chiba, K.;Tagaya, H.; Sato, S.; Ito, K. Fuel 1985a,64,68. Chiba, K.; Tagaya, H.; Yamauchi, T.; Tsukahara, Y. Chem. Lett. 1985b,945. Chiba, K.; Tagaya, H.; Sato, S.; Sugai, J.; Shibuya, Y.; Chiba, T. Conf. Coal Sci. Jpn., 22th 1985c,22,28. Chiba, K.; Tagaya, H.; Sato, S.; Ohgi, S. Bull. Yamagata Univ. 1986, 19,61. Curran, G. P.; Struck, R. T.; Gorin, E. Ind. Eng. Chem. Process Des. Deu. 1967,6, 166. Curtis, C. W.; Guin, J. A,; Jeng, J.; Tarrer, A. R. Fuel 1981,60, 677. Curtis, C. W.; Guin, J. A.; Kwan, K. C. Fuel 1984,63,1404. Curtis, C. W.; Guin, J. A.; Hale, M. A.; Smith, N. L. Fuel 1985,64,
461. Davies, G. 0.;Derbyshire, F. J.; Price, R. J. Inst. Fuel 1977,50,121. Furlong, L.E.; Effron, E.; Vernon, L. W.; Wilson, E. L. Chem. Eng. Prog. 1976,72,69. Guin, J. A.; Curtis, C. W.; Kwon, K. C. Fuel 1983,62,1412. Hellgeth, J. W.; Taylor, L. T. Fuel 1984,63,961. Kamiya, Y.;Ohta, H.; Mizuki, T.; Fukushima, A.; Aizawa, M. Proceedings of the International Conference on Coal Science, Pittsburgh, Aug 1983; p 195. Kimber, G. M. Proceedings of the International Conference on Coal Science, Sydney, Oct 1985; p 106.
Receiued for review January 14, 1986 Revised manuscript received February 18, 1987 Accepted April 11, 1987
Steam Gasification of Biomass with Nickel Secondary Catalysts Eddie G. Baker,* Lyle K. Mudge, and Michael D. Brown Pacific Northwest Laboratory,' Richland, Washington 99352
Nickel secondary catalysts are effective for increasing the gas yield from steam gasification of biomass by converting tars and other hydrocarbons t o gas. Coke buildup on the catalyst can cause loss of catalyst effectiveness in a very short time. Bench-scale gasification tests showed that coke buildup can be prevented if the catalyst is used in a separate fluid bed reactor. When the catalyst was used in the primary gasification reactor or in a separate fixed bed reactor, rapid catalyst deactivation occurred. Gasification or pyrolysis of biomass produces a wide spectrum of products including gases, char, light oils, and tar. Pacific Northwest Laboratory (PNL) has been exploring the use of catalysts to produce gases from biomass. In particular, supported nickel catalysts were found to be effective in producing high yields of synthesis gas from the hot, raw gas produced by steam gasification of biomass (Mitchell et al., 1980). Several other investigators have obtained similar results (Tanaka et al., 1984; Yokoyama et al., 1983; Ekstrom et al., 1985; Corte et al., 1985). We refer to these catalysts as secondary to differentiate them from catalysts such as alkali carbonates which catalyze pyrolysis and gasification reactions. In tests in a small fixed bed laboratory reactor (10-12 g/h wood feed) several catalysts were found that retained their activity over a long period of time without regeneration. One catalyst was tested for over lo00 h and was still active when the test was terminated (Baker et al., 1983;
Mudge, et al., 1985). These tests also determined the optimum conditions for synthesis gas production to be 750 "C and a stream/wood mass ratio of about 0.8. When these same catalysts were tested in a 1 tpd (ton/day) fluidized bed process research unit (PRU), they quickly lost their activity. Fouling due to coke buildup was the cause of catalyst deactivation (Baker et al., 1984; Mudge et al., 1985). In the PRU, the catalyst was used in a fluidizable form in the gasifier vessel. A bench-scale gasification unit (BSG) was built to resolve the apparent discrepancy in results between the laboratory and PRU tests. The BSG has provision for placing the catalyst in the fluidized bed gasification vessel similar to the PRU or downstream in a separate catalytic reactor to more closely simulate the original laboratory studies.
+Operated for the U.S. Department of Energy by Battelle Memorial Institute under Contract DE-AC06-76 RLO 1830.
in steam gasification of biomass. With a nickel secondary catalyst, the yields of methane and light hydrocarbon gases
0888-5885187 f 2626-I335$0I.50/0
Background Table I shows the most important secondary reactions
0 1987 American Chemical Society
1336 Ind. Eng. Chem. Res., Vol. 26, No. 7 , 1987 Table I. Secondary Reactions in Steam Gasification of Biomass tar cracking tar C* + C,H, + gas (1) methane reforming CH, + HZO Q 3Hz + CO (2) light hydrocarbon reforming C,H, + H,O CO + H, (3) CO + HzO s H2 + COS water gas shift reaction (4) CnH2, + Hz CnHzn+2 (5) hydrogenation tar reforming tar + H 2 0 CO + H, (6) carbon-steam reaction C*" + HzO CO + H, (7)
-
-
-+
O C *
= if a catalyst is present, this represents coke on the cata-
lyst.
Table 11. ProDerties of Secondarv Catalvst catalyst source NCM" G-90Cb composition, wt % active metals 9.5% Ni 15% Ni 4.25% CuO 9.25% MOOS support SiOz-Al2O3 70-76% A1203 58% CaO ~
~
ICI-46-1'
~~~
BET surface area, m2/g metal surface area, m2!g (based on H, chemisorption) catalyst size, U.S. sieve primary fluid bed secondary fluid bed secondary fixed bed
205
3-15
16.5% Ni (21% NiO) 14% SiO, 29% A1203 13% MgO 13% CaO 7% KZO 11.0
5
-40 to +70 -40 to +70
W. R. Grace. "nited
Table 111. Typical Analysis of Feedstocks and Char and Tar Byproducts feedstocks pine flakes bagasse char tar elemental anal. C 50.3 44.7 85.0 85.0 H 6.4 5.9 1.0 6.2 0 42.7 44.4 14.0 8.8 N 0.2 S 0.03 ash 0.6 4.8 HHV Btu f lb 8770 7870 kJf kg 20400 18300 major components, benzene, toluene, GC/MS xylene, phenol, cresol, indene, methylindene, naphthalene, methylnaphthalene, acenaphthylene, acenaphthene, phenanthrene
-40 to +70 -40 to +70 to
f4
-20 to +70
in.,
-1 to +0.5 cm
Catalyst. 'Imperial Chemical Industries
are reduced (reactions 2 and 3), the shift reaction (reaction 4) is driven to equilibrium, and the tar yield is reduced to very low levels. Reaction 6 represents the ultimate fate of tar in the presence at nickel catalysts, but direct reaction of steam and tar over the nickel metal has not been proven. Reaction 6 may also be represented as the sum of reactions 1, 3, and 7. The exact mechanism by which the tar is converted is not known but may be a combination of both. Catalytic deactivation was found to be due to coke buildup on the catalyst. Although carbon formation is not favored thermodynamically a t the reactor conditions, carbon accumulation depends on the relative rates of the carbon forming reaction (1)and carbon removing reaction (7). The catalyst support was found to have a significant effect on carbon buildup on the catalyst, apparently by affecting the balance between reactions 1and 7. In general the less acidic supports reduced carbon buildup and deactivation (Baker and Mudge, 1984).
Materials The properties of three different catalysts tested are shown in Table 11. NCM is a catalyst made especially for this project by W. R. Grace and consists of nickel, copper, and molybdenum salts impregnated on a proprietary high surface area support. G90C is a commercial steam reforming catalyst from United Catalysts. These were two of the most effective catalysts in laboratory studies. ICI-46-1 is also a commercial nickel-based steam reforming catalyst with a different support than G-90C. NCM was available only in -40-mesh to +70-mesh spheres and was tested in the primary gasifier fluid bed and in a secondary fluid bed reactor. Catalysts G90C and
ICI-46-1 were both 1.2-2.0-cm rings when received. These were broken to 0.5-1.0-cm chunks for tests in a fixed bed secondary reactor and ground to a fine mesh for the fluidized bed tests. Catalyst G90C was used in all three configurations, and ICI-46-1was used only in the secondary fluid bed reactor. Typical analyses of the feedstocks used in the tests are shown in Table 111. Pine flakes were used in all of the tests in the BSG. Both wood and bagasse were used in the PRU tests. Table I11 also includes a typical elemental composition for the char produced in the gasification vessel. The composition of the tar produced in the gasifier vessel with no catalyst present is also shown in Table 111. The tar is primarily aromatic, with single-ring phenolics being the most common components. The residual tar left after catalytic processing is also aromatic but has no phenolics. Benzene and naphthalene are the most common components. More detailed analyses of the tars produced in the BSG are given by Elliott and Baker (1986).
Description of Experimental System The bench-scale gasification system was designed for evaluation of catalysts for biomass gasification. It has a nominal wood feed capacity of 1.8 kg/h. The system contains two reactors in series, a biomass pyrolysis/gasification reactor followed by a secondary catalytic reactor for gas-phase reactions. A schematic for the system is shown in Figure 1. Wood is screw fed from a hopper to the bottom of the pyrolysis reactor where it is entrained into the reactor by superheated steam. The pyrolysis reactor consists of a fluidized bed section 7.8-cm i.d. and 60-cm high followed by a disengaging zone 90-cm high and 12.8-cm i.d. The fluidized bed section contains the catalyst or an inert material either silica sand or bauxite. Gases and some char leave the top of the reactor and pass through sintered metal filters to remove char. Char is also vacuumed from the top of the bed about every 15 rnin to maintain the bed level at a constant depth. Bed material removed from the gasifier through the vacuum system with the char is fed back to the bed through the wood feed system. Hot gas from the filters passes through the secondary catalyst reactor. The secondary fluidized bed reactor is a duplicate of the primary pyrolysis reactor except it contains a bubble-cap gas distributor. Actual fluidized
Ind. Eng. Chem. Res., Vol. 26, No. 7, 1987 1337 Secondary
Table V. Typical Results with a Fresh Nickel Catalyst and with No Catalyst fresh Ni catalyst no catalyst dry gas composition, vol % 60 41 H,
caI8l"tlc
1
23
16 4 10 28
2.6 2.1
1.2 0.8
90 10
80 13 7
16 1
co'
3
Candensale Tar
Reactor Dral"
Figure 1. Flow schematic for bench-scale gasification system. Table IV. Tyoical Conditions for Catalyst Tests wood feed rate, g/m 12-16 bed depth, cm steam rate, g/m 11-13 primary 30 temp, "C secondary 20-40 primary 750 catalyst loading, g 600-700 secondary 750 char residence time, m 30-40 pressure, kPa primary 101 secondary 101
catalyst depth is about 30 cm. A secondary fixed bed reactor was also tested. It is 150-cm long by 5.1-cm i.d. Catalyst depth is about 40 cm. Following the secondary reactor is a condenser/scrubber where water and tar are removed. The cool, clean gas is analyzed with a Carle I11 H gas chromatograph and then vented. During each test a small slipstream of hot gas is cooled and passed through a scrubber containing methylene chloride to remove tars and oils. The gas flow is measured and then vented. The resulting solution was analyzed by using a Perkin-Elmer 3920 gas chromatograph with a 3% OV-17 column.
gas production, m3/kga total synthesis gas (H2 + CO) C conversion, wt 70 gas solid liquidb
aVolume of dry gas per unit weight of dry wood measured at atmospheric pressure and 20 "C. bTars and water-soluble organics.
----
-*-
0 C 0
0
2
4
6
8
Catalyst Age (Time
Bench Scale G90C NCM Ni/AIz03
10
on
PRU(Wood) PRU (Bagasse)
12
14
16
Stream). hours
Figure 2. Catalyst deactivation in the primary reactor. 25
Results Catalyst tests were conducted in three modes: (1)catalyst in the fluid bed gasifier, ( 2 ) catalyst in a secondary fixed bed reactor, and (3) catalyst in a secondary fluid bed reactor. Three tests were made in the first mode, four in the second mode, and seven in the third mode. The tests ranged in length from 4 to 50 h depending on how well the catalyst maintained activity. The majority of the tests were made a t the same conditions which are shown in Table IV. At these conditions, the gas residence time in the primary vessel is about 30 s. Gas residence time in the secondary fluidized bed reactor is about 20 s and about 6s in the secondary fixed bed reactor. Catalyst loading was the same in both the fluidized bed and the fixed bed secondary vessels, 600-700 g. This corresponds to a gas residence time in the catalyst bed of about 2-3 s. In all of the tests (in all three modes) with the catalysts shown in Table 11, catalyst activity was initially high. The gas consisted primarily of H2, CO, COz, and a trace of methane. No tars or oils were present initially. Typical start-of-run data are shown in Table V. Several uncatalyzed tests were made to provide a comparison, and results are also shown in Table V. The gas production numbers in Table V can be compared to a maximum theoretical yield of synthesis gas from biomass calculated assuming 100% carbon conversion to CO and COz. With a H,/CO ratio of 2.0, the maximum theoretical gas yield is about 2.4 m3/kg; with a H2/C0 ratio of 3.0 (higher steam to wood ratio), the maximum theoretical yield is about 2.8 m3/kg. The yield of synthesis gas and the appearance of tar in the
s
10 15t
3 Catalyst Age (Time
on Stream). hours
Figure 3. Carbon accumulation on catalysts in the primary reactor.
condensate from the gasifier were used as indicators of a loss of catalyst activity.
Test with Catalyst in Primary Reactor (Gasifier). These tests were made to confirm results obtained previously in the 1ton/day fluidized bed PRU. In the PRU, the catalysts tested were a Ni/A1203hydrocracking catalyst supplied by W. R. Grace (Mudge et al., 1981) and the same NCM catalyst tested in the bench-scale tests (Mudge et al. 1983). Both lost activity rapidly in the PRU as shown in Figure 2. Similar deactivation also occurred in the bench-scale tests with both NCM and G90C. Loss of catalyst activity is apparently due to fouling by buildup of carbon which blocks access to the catalyst pores. Figure 3 shows carbon buildup in the bench-scale tests and in PRU tests with bagasse. Some catalyst is carried out of the gasifier with the product gas which probably accelerates the apparent loss
1338 Ind. Eng. Chem. Res., Vol. 26, No. 7, 1987 2w
Table VI. Typical Results with “Steady-State”Catalyst gas composition, vol 70 gas production, m3/ kg” H2 55 total 1.9 synthesis gas (H2+ Co) 1.5 COZ 14 C conversion, w t % CZH,, C*H, 1 gas 88 CHI 6 solid 10 liquidb 2 co 24
--1 ~
-@-?
Bench Scale Laboratory G90C NCM
-J
Volume of dry gas per unit weight of dry wood measured a t 20 “C and atmospheric pressure. *Tars and water-soluble organics.
mz loot
0 75
0
1
I
I
I
I
2
4
6
8
10
12 i
$4
G
G90C
I’
NCM
Catalyst Age (Time on Stream) hours
I
Figure 4. Catalyst deactivation in the secondary fixed bed reactor.
.
-
e
I
00
8
4
16
12
20
24
28
32
36
Catalyst Age (Time on Stream), hours
Figure 6. Carbon accumulation on catalysts in the secondary fluidized bed reactor. 1
25
C
NCM ’ IC1 4 6 1
1 00
0 75(
0
GSOC
N i l = A1203
N I Cu Mo/SiOz A1203 NI S i 0 2 A1203 CaO M g O K 2 0
I
1
I
I
I
I
I
4
8
12
16
20
24
28
1 32
Catalyst Age ITime on Stream) hours
Figure 5. Catalyst activity in the secondary fluidized bed reactor.
of catalyst activity. However, the large amount of carbon on the catalyst remaining in the bed indicates it has little or no activity and even if the bed were still full there would be little overall improvement in catalyst effectiveness. Tests w i t h Catalyst i n a Secondary Fixed Bed Reactor. One extended test was made with G90C in a secondary fixed bed reactor to compare to previous results obtained in laboratory gasification tests. The initial loss of activity is similar to that seen in the laboratory as shown in Figure 4. After 15-20 h in the laboratory tests, catalyst activity stabilized at a synthesis gas yield of 1.25-1.30 m3/kg wood, and no further deactivation was noted. In the bench-scale tests, G-90C continued to deactivate down to a synthesis gas yield of about 1.0 m3/kg. After 16 h of exposure, the catalyst had 6 wt % carbon on it compared to less than 1 wt 70after 40 h of exposure in the laboratory unit. After about 1200 h of exposure in the laboratory unit, NCM had only 4 wt % carbon. Tests w i t h Catalysts in a Secondary Fluid Bed Reactor. The NCM catalyst was available only in -40mesh to +70-mesh spheres so it was tested in a secondary fluidized bed reactor. The initial activity of NCM was somewhat higher in the secondary fluidized bed. Some activity was lost in the first 8-10 h, but after that no further loss of activity was noted during 34 h of exposure. Synthesis gas yield leveled off at about 1.5 m3/kg, 20% higher than achieved in laboratory gasification tests. This test was so encouraging that two other catalysts, G90C and ICI-46-1, were also tested in this mode. Results with G90C and ICI-46-1 were both similar to NCM. Results of these three tests are shown in Figure 5 . All three catalysts reached a “steady-state’’ activity with a synthesis gas yield of about 1.5 m3/kg. Table VI shows typical results during this steady-state period. These can be compared to the results with a fresh catalyst and with no catalyst shown previously in Table V.
z
Non-Catalytic Threshold
v)
05
I
i
I
I
I
I
----
Primary Vessel Secondary Fixed Bed
15
Secondary Fluid Bed
.-e-.
0
5
10
15
20
25
30
Catalyst Age (Time on Stream), hours
Figure 7. Summary of catalyst deactivation and carbon accumulation in different operating modes.
Coke buildup on the catalysts in the secondary fluid bed reactor was dramatically reduced compared to catalysts in the primary reactor and in a secondary fixed bed reactor. Coke on the catalyst leveled off at about 3 w t % for G90C and 1 wt % for NCM as shown in Figure 6. Discussion Previous laboratory tests confirmed that coke formation resulted from decomposition of tars and light, unsaturated hydrocarbon gases. The coke blocked access to the catalyst pores, causing a loss of catalyst activity (Baker and Mudge, 1984). The primary finding of these bench-scale tests is that the placement of a secondary catalyst (primary bed vs. secondary bed) and the mode of contact (fluid bed vs. fixed bed) has a dramatic effect on catalyst deactivation. Figure 7 summarizes the rate of coke formation and catalyst deactivation for all these systems. The differences in the results in Figure 7 must be explained in terms of carbon-forming and carbon-removing reactions (1)and ( 7 ) in Table I. Since the conditions for gasification of carbon from the catalyst (eq 7) are nearly identical in the primary gasification reactor and in the
Ind. Eng. Chem. Res., Vol. 26, No. 7, 1987 1339 secondary fluid bed reactor, the difference in the results between these two systems must be due to reaction (1). We previously theorized that the rapid coke formation on catalysts in the primary gasification reactor is due to secondary tar reactions which produce coke, gases, and secondary tar (Baker and Mudge, 1984). Placing the catalyst in a secondary vessel allows the primary tar to react uncatalyzed in the primary vessel. The catalyst in the secondary vessel then only has to treat the resulting secondary tar. The results of the bench-scale gasification tests are in accordance with this theory. The results of the bench-scale tests show that carbon on the catalyst peaks at about 8-10 h in the secondary fluid bed. At this point, the rate of coke gasification is apparently equal to the rate of coke formation, and no further accumulation of carbon occurs. Assuming that processing conditions do not change, it appears that this steady-state catalyst activity can be maintained for an extended period of time without regeneration. This is similar to the results achieved in the laboratory gasifier tests. Heat- or mass-transfer limitations are one possible cause of the relatively poor performance of the secondary fixed bed compared to the secondary fluid bed. Due to the strongly endothermic nature of the reactions that take place on the catalyst, poor heat transfer in the fixed bed reactor could result in the catalyst pellet being significantly cooler than the measured bulk gas temperature. Computer modeling efforts (Brown et al., 1986; Mudge et al., 1987) showed pellet temperatures as much as 100 "C below that of the bulk gas. Increased carbon deposition could then be the result of dropping to near the critical temperature (-600 "C) below which carbon is thermodynamically stable in the system and gasification of carbon on the catalyst (reaction 7) is not possible. Even if gasification of carbon on the catalyst is still thermodynamically favorable, the lower temperatures reduce the rate of carbon gasification on the catalyst pellet which may result in a net accumulation of coke on the catalyst. In the small laboratory fixed bed reactor, radiation from the hot tube wall apparently maintained the catalyst and the bulk gas temperatures near isothermal. This produced results similar to the bench-scale secondary fluid bed reactor. Conclusions Use of a nickel secondary catalyst can dramatically increase the yield of synthesis gas (H, CO) from steam gasification of biomass. In previous tests in a PRU, tar cracking on the catalyst resulted in rapid deactivation due to carbon fouling. Tests with a bench-scale gasifier showed that carbon fouling and deactivation are a function of placement of the catalyst (primary bed vs. secondary bed) and the mode of contact (fluid bed vs. fixed bed). Three catalysts were found which retain their activity for an extended period of time without regeneration in a sec-
+
ondary fluid bed reactor. Catalyst deactivation is rapid when the catalyst is placed in the primary gasifier or in a secondary fixed bed reactor. Acknowledgment We thank Gary F. Schiefelbein and Mark A. Gerber of the Biomass Program Office at PNL for their support. We are particularly indebted to Wayne A. Wilcox who was responsible for construction and operation of the benchscale gasifier. This study was funded by the Biofuels and Municipal Waste Technology Division of the U S . Department of Energy. Registry No. Ni, 7440-02-0; Cu, 7440-50-8; Mo, 7439-98-7; C&a, 71-43-2; CcHbCH3, 108-88-3; CH,C,H*CHs, 1330-20-7; CGHbOH, 108-95-2; CH,CBH4OH, 1319-77-3; Hz, 1333-74-0; HzC=CHCH,, 115-07-1; H3CCHZCH3, 74-98-6; COZ, 124-38-9; CzC=CH,, 74-85-1; H&CH3,74-84-0; CHI, 74-82-8; CO, 630-08-0; C, 7440-44-0; indene, 95-13-6; methylindene, 29036-25-7; naphthalene, 91-20-3; methylnaphthalene, 1321-94-4; acenaphthylene, 208-96-8; acenaphthene, 83-32-9; phenanthrene, 85-01-8.
Literature Cited Baker, E. G.; Mudge, L. K. J . Anal. Appl. Pyrolysis 1984, 6, 285. Baker, E. G.; Mitchell, D. H.; Mudge, L. K.; Brown, M. D. Energy Prog. 1983, 3(4), 226. Baker, E. G.; Mudge, L. K.; Brown, M. D. Chem. Eng. Prog. 1984, 80(12), 43. Brown, M. D.; Baker, E. G.; Mudge, L. K. "Deactivation of Catalysts in Biomass Gasification", presented at the 1986 AlCHE Summer Meeting in Boston, Aug. 25-27, 1986; PNL-SA-13931. Corte, P.; Lacoste, C.; Traverse, J. P. J. Anal. Appl. Pyrolysis 1985, 7(4), 323. Ekstrom, C.; Lindman, N.; Petterson, R. In Fundamentals of Thermochemical Biomass Conversion; Overend, T., Milne, T., Mudge, L. K. Eds.; Elsevier Applied Science: London. 1985. Elliott, D. C.; Baker, E. G. Biomass 1986, 9, 195-203. Mitchell, D. H.; Mudge, L. K.; Robertus, R. J.; Weber, S. L.; Sealock, L. J., Jr. Chem. Eng. Prog. 1980, 76(9), 53. Mudge, L. K.; Weber, S. L.; Mitchell, D. H.; Sealock, L. J., Jr.; Robertus, R. J. "Investigations on Catalyzed Steam Gasification of Biomass", Report PNL-3695, 1981; Pacific Northwest Laboratory, Richland, WA. Mudge, L. K.; Baker, E. G.; Mitchell, D. H.; Robertus, R. J.; Brown, M. D. "Catalytic Gasification Studies in a Pressurized Fluid Bed Unit, Report PNL-4594, 1983; Pacific Northwest Laboratory, Richland, WA. Mudge, L. K.; Baker, E. G.; Mitchell, D. H.; Brown, M. D. Trans. ASME J.,Solar Energy Eng. 1985, 107(2), 88. Mudge, L. K.; E. G.; Baker, E. G.; Brown, M. D.; Wilcox, W. A. "Bench Scale Studies on Gasification of Biomass in the Presence of Catalysts", Report PNL-5699, 1987, Pacific Northwest Laboratory, Richland, WA. Tanaka, Y.; Yamaguchi, T.; Yamasaki, K.; Ueno, A.; Kotera, Y. Ind. Eng. Chem. Prod. Res. Dev. 1984,23(2) 225. Yokoyama, S.; Ogi, T.; Koguchi, K.; Nakamura, E. Chem. Lett. 1983, 151.
Received for review January 21, 1986 Accepted April 18, 1987
