Biomass Gasification with Air in Fluidized Bed. Hot Gas Cleanup with

Ind. Eng. Chem. Res. 2000, 39, 1143-1154

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Biomass Gasification with Air in Fluidized Bed. Hot Gas Cleanup with Selected Commercial and Full-Size Nickel-Based Catalysts Miguel A. Caballero,† Jose´ Corella,‡ Marıá-Pilar Aznar,*,† and Javier Gil† Chemical and Environmental Engineering Department, University of Saragossa, 50009 Saragossa, Spain, and Chemical Engineering Department, University “Complutense” of Madrid, 28040 Madrid, Spain

Three selected commercial, full-size steam-reforming catalysts for naphthas, BASF G1-50, ICI 46-1, and Topsøe R-67, are tested at pilot-scale level for hot gas cleanup in biomass gasification in a fluidized bed. Gas composition and tar content in the flue gas are measured before and after the catalytic bed. Variations of the catalytic bed in H2, CO, CO2, CH4, and H2O contents are reported for different operating conditions. Tar conversions and an apparent first-order kinetics constant for the overall tar removal reaction are calculated. Tar contents at the exit of the catalytic reactor as low as 10 mg/mn3 are obtained in a test of 50 h-on-stream without noticeable catalyst deactivation. Important variations in tar conversion with space time in the catalytic bed, with H2O/C* in the flue gas, and with the equivalence ratio in the upstream gasifier are observed. These results obtained at the pilot-scale level and with the use of full-sized commercial catalysts are an important forward step in demonstrating the technical feasibility of the overall biomass gasification process. Introduction Biomass gasification in fluidized bed produces a raw gas which usually has to be cleaned of tars for most of its advanced applications. Particulates, nitrogen compounds such as NH3 and CNH, and potassium in the flue gas also have to be taken into account and sometimes eliminated from such raw gas too. Hot gas cleaning is usually preferred over wet cleaning methods because it does not produce a toxic liquid waste stream and it has high overall thermal efficiency. Two different hot gas cleaning methods, approaches, and/or processes seem to be the most promising ones for future industrial applications. Such methods are based on using calcined dolomites (or related materials) or nickel-based catalysts. These two types of solids, in comparison, seem to be very different, but in fact their behaviors in tar removal are very similar:1 they both promote mainly steam- and dry- (CO2-) reforming reactions of tars toward H2 and CO. A good biomass gasification process has to be both technically and economically feasible. A lot of previous research led several institutions to the conclusion that the best process, for the moment, should be based on an optimized (according to its design and operation) gasifier followed by a nickel-based catalytic reactor. Such an optimized gasifier should generate a raw gas that is “quite clean” already, with a tar content below 2 g of tar/mn3, to avoid catalyst deactivation by coke in the downstream catalytic reactor. This relatively low tar content can be achieved, for instance, by in-bed (in the gasifier) use of dolomite.2,3 Having thus studied the gasifier optimization in previous papers (i.e., refs 2 and 3), we focused or concentrated this work only in the catalytic reactor, with nickel-based catalysts, downstream from an optimized fluidized-bed biomass gasifier. State-of-the-art methods on using nickel-based catalysts for hot gas cleaning in biomass gasification has * To whom correspondence should be addressed. Fax: + 34 976 76 21 42. E-mail: [email protected]. † University of Saragossa. ‡ University “Complutense” of Madrid. Fax: +34 91 394 4164.

been reviewed recently.4-9 There are two main approaches with these catalysts: to use monoliths, the well-known method followed among others by VTT in Finland,27 or to use commercial (nickel-based) steamreforming catalysts. This second approach is the one followed to date by the Universities of Madrid and Saragossa in Spain. A lot of data already exists on hot gas cleanup and conditioning concerning tar removal using commercial nickel-based catalysts in gasification with pure steam,4,10,11 with steam-O2 mixtures,6,7 or with air.5,8,9,12,13 The present work concerns only gasification with air, which seems to be the most available gasification technology nowadays. The process considered here is then biomass gasification with air in an optimized atmospheric fluidized bed followed by a fixed-bed catalytic reactor with a commercial steam-reforming catalyst. In previous works just on this matter5,8,9 a lot of operation variables as well as the selection/choice of the best available catalyst(s) were studied. For instance, from such previous works three commercial catalysts were selected for further testing: G1-50 from BASF AG, 46-1 from ICI-Katalco, and R-67 from Topsøe A/S. Selected results from different authors are shown in Table 1. It can be observed in this table how tar conversions higher than 90% have already been achieved by all authors, even with the conversions having relatively low gas residence times. The problem with tar elimination seems to have been solved thus, but in fact there are still some unknowns or unsolved problems. For instance, all the previous work shown in Table 1 was done with crushed catalysts. Crushed catalysts may be good for studying kinetics and/or intrinsic catalytic activities, but a commercial reactor with a commercial steam-reforming catalyst cannot use crushed catalysts. Commercial steam-reforming catalysts are rings that are relatively large and thus cannot be crushed for commercial application. The main difference and contribution of this work with respect to previous works is that here the commercial catalysts will be tested at the pilot-scale level in their full-size state.

10.1021/ie990738t CCC: $19.00 © 2000 American Chemical Society Published on Web 04/13/2000

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Ind. Eng. Chem. Res., Vol. 39, No. 5, 2000

Table 1. Tar Conversions Obtained by Different Authors authors Baker et al., 198710

Leppa¨lahti et al., 199113

Simell et al., 199212

Kinoshita et al., 199511

Narvaez et al., 19975

Corella et al., 19999

catalyst

ICI 46-1

nickel “B”

UCI G-90B

BASF G1-25 S

gas residence time (s) dp (mm) temperature (°C) tar conversion

2-3 5-10 750 90

Engelhard Ni-0301 0.2-0.3 3 840 97.3

0.2-0.25 3 900 99.4

1.6 2 800 97.0

0.2-0.3 1.4 800 97.0

8 different catalysts 1-0.3 1.0-1.6 780-820 97-99.9

Pilot Plant, Catalytic Reactor, and Feedstock Feedstock. All the experimental equipment used in this work has already been described. As the feedstock, small pine wood chips were used. Their proximate and ultimate analysis can be found in the work of Gil et al.14 Pilot Plant. The pilot plant used has been fully described in ref 14. It was based on a bubbling and atmospheric fluidized bed, 15-cm inner diameter and 3.2-m height, that was continuously fed near the gas distributor plate with biomass flow rates of 8-15 kg/h. The gasifier operated was de-rated because of its location in the university campus, to avoid big flames and smokes which had big impacts in the surroundings. Equivalence ratios (ER) most often used were from 0.19 to 0.30. Main experimental conditions of all tests are given in Table 2. Catalytic Reactor(s). In a slip flow downstream from the gasifier there were two catalytic reactors both with 51-mm i.d. There were two catalytic reactors because, in former work, the nickel-based catalyst was placed in the second reactor only, the first one being used as a guard bed, with dolomite in it, to decrease the tar content in the flue gas (at the inlet of the second catalytic reactor) below 2 g/mn3. But it has been demonstrated in recent work15 that the dolomite placed in the gasifier bed shows activity (for tar elimination) similar to that of the dolomite placed downstream from the gasifier. So it was decided to move and locate the dolomite in the same gasifier bed, saving one reactor (the guard bed) and simplifying then the overall gasification-gas cleaning process. The former guard bed was not used anymore as a guard bed with dolomite; it was now used in this work for nickel catalysts. Two different catalytic reactors (connected in series) were used in this work thus. The first (former guard bed) and second catalytic reactors are completely described in the works of Pe´rez et al.16 and Caballero et al.,6 respectively. Both catalytic reactors were externally heated by ovens to operate them at different levels of temperature. The location and combination of catalysts used, including an “inert” material like silica sand, are shown in Table 3. It can be noticed in this table how some commercial (high and low temperature) CO-shift catalysts were tested too downstream from a nickel-based catalytic bed, but results with such CO-shift catalysts are out the scope of this paper. Details of the operation parameters and experimental conditions in the catalytic reactors are shown in the bottom part of Table 2. The temperature in the center and in the wall (inner side) of the first catalytic reactor and the pressure (above atmospheric) in the reactor for test AG-14 are shown just as an example in Figure 1. It can be appreciated in Figure 1 how these catalytic reactors were not rigorously isothermal, but a big effort was made to keep variations of temperature as low as possible (it can be remembered that working at the pilotscale level is not the same as working at the microscopic-, laboratory-, or small-scale level).

this work ICI 46-1 0.2-0.3 -7.0 to +14 840 99.8

Gas composition was measured before and after the catalytic reactor, and the results are shown in Figure 2 for some tests. Such measurements were made every half an hour. Values reported from now on gas composition will be the averaged ones (with respect to time-onstream). Tar Sampling and Analysis. Samples for tars were taken every hour before and after the catalytic reactor. Tar measurements are not standarized yet and all values concerning tar are relatives ones.17,18 The method used in this work has been described previously5 and it is similar to the one followed by VTT in Finland.12 Tar measured with this method will be called from now on tar* to indicate how this value is a relative measurement and that it is associated with the sampling and analysis methods followed in this work. Samples of tar were also simultaneously taken for their characterization by the SPA method developed by KTH in Stockholm.19 Tar composition was determined thus before and after the catalytic reactor. Tar composition at the gasifier exit depends on a lot of operation variables and the variation of the tar composition by the catalytic reactor is quite complex. A study of both data has been detailed in another paper.20 Just as an example, a typical tar composition (for ER ) 0.24) at the gasifier exit (inlet of the first catalytic reactor) is 28 wt % benzene, 15 wt % naphthalene, 13.6 wt % toluene, 7.0 wt % xylene, 7 wt % indene, 13 wt % tworing compounds, 9 wt % three- and four-ring compounds, and 7 wt % phenols. Catalysts Used. Three previously selected steamreforming catalysts, BASF G1-50, Topsøe R-67, and specially ICI 46-1 have been used. According to the space time wanted or required, the fixed bed had 8-12cm height for “high” space times and 3-6 cm for “low” space times. Because of the relatively large size of the commercial catalysts (rings of about 14-mm external diameter) with respect to the reactor inner diameter (51 mm), some channelling or bypassing could occur in the catalytic reactor. To avoid it to the maximum possible extent, all the commercial catalyst particles (rings) were cut carefully in half, and only ring ”halves” were used. When the reactor was filled up prior to each test, each particle of the catalyst (ring half, thus) was introduced and located carefully by hand to avoid further bypassing in the so-generated fixed bed. Even with these procedures (cutting into halves and careful handmade location of the catalyst in the bed), the authors consider and admit that a small amount of bypassing could still exist. Tar conversion could have been a little bit higher than the ones shown in this paper without such small amounts of bypassing. Depending on the type of arrangement used, the catalyst was placed in the first or in the second catalytic reactor. For instance, in run AG-4, the catalyst was located in the second reactor after a bed of silica sand (in the first catalytic reactor). In run AG-9, an “in equilibrium” FCC catalyst was placed in the gasifier

Ind. Eng. Chem. Res., Vol. 39, No. 5, 2000 1145 Table 2. Experimental Conditions test no. AG-1

AG-2

AG-3

AG-4

AG-5

Feeding ER biomass flow rate (kg/h) moisture (% wet basis) (not calc.) dolomite/ biomass (g/kg of daf) dp dolomite (mm)

0.24 11.4 10 30

0.26 10.8 10 30

0.21 10.9 10 30

0.28 10.7 10 30

0.19 9.8 10 30

-0.63 to +0.40

-0.63 to +0.40

-0.63 to +0.40

-0.63 to +0.40

-0.63 to +0.40

total weight (kg) % dolomite (%) Tbed (°C) T“freeboard” (°C) τ0 (s) u0 (cm/s) u0/umf sand u0/umf dolomite

15.7 23 820 690 1.28 54 5.6 6.5

14.5 25 800 690 1.15 55 5.7 6.6

15.9 22 790 700 1.64 42 4.4 5.1

14.5 15 820 690 1.13 56 5.8 6.7

12.4 27 800 640 1.50 36 3.8 4.3

position w/ respect to the gasifier catalyst weight (g) dp (mm) Hbed (cm) Tbed,center (°C) u0 (cm/s) τaveraged (kg‚h/mTb,wet3) τ0 SV (mn,wet3/h‚m3) Fbed (kg/L) H2Oinlet (vol %) H2O/C* (mol/atom‚g of C) Ctars*,inlet (mg/mn3)

second reactor

Catalytic Bed second reactor second reactor

third reactor

second reactor

BASF G1-50 122 -6.4 to +5.0 8 830 53 0.043 0.15 6500 1.11 14.1 2.2 3400

TOPSOE R-67 115 -6.4 to +5.0 8 835 43 0.040 0.19 6600 1.07 16.8 2.8 2375

ICI 46-1 115 -14 to +7.0 11 840 43 0.052 0.24 3800 0.78 7.5 1.7 275

ICI 46-1 115 -14 to +7.0 12 840 54 0.045 0.22 3800 0.71 17.6 3.1 7910

Gasifier

TOPSOE R-67 115 -6.4 to +5.0 8 840 37 0.053 0.22 4700 1.07 13.5 2.3 5277 test no.

AG-6

AG-7

AG-8

AG-9

AG-10

Feeding ER biomass flow rate (kg/h) moisture (% wet basis) (not calc.) dolomite/ biomass (g/kg of daf) dp dolomite (mm)

0.19 10.0 11 30

0.20 9.1 7 30

0.22 10.5 7 30

0.27 8.8 11 51 (FCC)

0.29 8.2 10 30

-0.63 to +0.40

-0.63 to +0.40

-0.63 to +0.40

0.069 (FCC)

-0.63 to +0.40

total weight (kg) % dolomite (%) initial dp dolomite (mm) Tbed (°C) T”freeboard” (°C) τ0 (s) u0 (cm/s) u0/umf sand u0/umf dolomite

12.6 28 -1.0 to +0.0.63 800 600 1.49 37 3.9 4.5

Gasifier 12.7 29 -1.0 to +0.63 810 640 1.53 36 3.8 4.3

14.0 21 -1.0 to +0.4 820 680 1.30 47 4.9 5.7

16.0 FCC FCC 820 680 1.52 46 4.8 5.5

15.3 28 -1.0 to +0.4 820 680 1.40 48 5.0 5.8

position w/ respect to the gasifier catalyst weight (g) dp (mm) Hbed (cm) Tbed,center (°C) u0 (cm/s) τaveraged (kg‚h/mTb,wet3) τ0 (s) SV (mn,wet3/h‚m3) Fbed (kg/L) H2Oinlet (vol %) H2O/C* (mol/atom‚g of C) Ctars*,inlet (mg/mn3)

second reactor

Catalytic Bed second reactor

second reactor

third reactor

second reactor

ICI 46-1 115 -14 to +7.0 12 830 52 0.043 0.23 4000 0.71 14.2 2.2 7043

ICI 46-1 115 -14 to +7.0 12 850 38 0.056 0.32 3000 0.71 13.1 2.5 7310

ICI 46-1 115 -14 to +7.0 12 830 45 0.053 0.27 3300 0.71 11.1 2.3 4764

ICI 46-1 115 -14 to +7.0 12 840 61 0.050 0.20 3400 0.71 6.4 1.1 580

ICI 46-1 116 -14 to +7.0 12 840 49 0.050 0.24 3600 0.72 17.2 3.7 1660

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Table 2. Continued test no. AG-11 ER caudal biomass (kg/h) moisture (% wet basis) (not calc.) dolomite/ biomass (g/kg of daf) dp dolomite (mm)

Feeding 0.24 9.2 10 30

0.35 8.3 15 30 -0.63 to +0.40

total weight (kg) % dolomite (%) initial dp dolomite (mm) Tbed (°C) T”freeboard” (°C) τ0 (s) u0 (cm/s) u0/umf sand u0/umf dolomite position w/ respect to the gasifier catalyst weight (g) dp (mm) Hbed (cm) Tbed,center (°C) u0 (cm/s) τaveraged (kg‚h/mTb,wet3) τ0 (s) SV (mn,wet3/h‚m3) Fbed (kg/L) H2Oinlet (vol %) H2O/C* (mol/atom‚g of C) Ctars*,inlet (mg/mn3)

AG-12

-0.63 to +0.40 Gasifier 14.6 26 -1.0 to +0.4 800 660 1.52 42 4.4 5.1

14.3 24 -1.0 to +0.4 840 690 1.13 55 5.7 6.6

second reactor ICI 46-1 116 -14 to +7.0 12 840 45 0.054 0.27 3300 0.72 7.3 1.2 970

third reactor ICI 46-1 86 -14 to +7.0 10 820 46 0.037 0.22 4100 0.64 3.8 0.4 530

Catalytic Bed second third reactor reactor ICI 46-1 ICI 46-1 57 57 -14 to +7.0 -14 to +7.0 6 6 830 860 82 81 0.014 0.018 0.07 0.07 12000 12000 0.71 0.71 12 9 1.6 1.4 1600 270

AG-13

AG-14

0.24 8.3 10 30

0.24 8 10 30

-0.63 to +0.40

-0.63 to +0.40

15.1 26 -1.0 to +0.4 830 680 1.69 39 4.1 4.7 second reactor ICI 46-1 30 -14 to +7.0 3 820 102 0.062 0.029 30000 0.74 13 1.5 700

third reactor ICI 46-1 39 -14 to +7.0 3.5 840 80 0.0095 0.044 23000 0.83 8 1.4 190

17 30 -1.0 to +0.4 840 680 1.61 46 4.2 5.1 second reactor ICI 46-1 30 -14 to +7.0 3 830 79 0.0085 0.038 20000 0.74 13 1.6 700

third reactor ICI 46-1 39 -14 to +7.0 3.5 840 57 0.013 0.061 14000 0.83 8 1.5 140

Table 3. Location and Combinations of the Catalysts Used (Gasifying Agent: Air) gasifier bed

first cat. bed (second reactor)

silica sand 3 diff. nickel + dolomite catalysts silica sand + dolomite silica sand + FCC silica sand + dolomite silica sand + dolomite

second cat. bed (third reactor)

test code

silica sand

3 diff. CO-shift AG-1-AG-3 catalysts AG-5-AG-8 AG-10 and AG-11 1 nickel catalyst AG-4 and AG-12

dolomite

1 nickel catalyst AG-9a and AG-9b

1 nickel catalyst 1 nickel catalyst AG-13 and AG-14 silica sand

silica sand

AG-15

bed, a bed of dolomite was used in the first catalytic reactor, and the Ni catalyst was located in the second catalytic reactor. In runs and AG-14, parts a and b, both reactors contained Ni-catalysts beds (Table 3).

Figure 1. Temperatures and pressure (above atmospheric) in the catalytic bed with time-on-stream (test AG-14).

Both catalytic reactors have the same diameter (51 mm) and similar heights. The first catalytic reactor has upward gas flow and the second one has downward gas flow. The reference temperature, expressed as Tbed, is always the measured one in the center (in height and diameter) of the bed. Although the length of the thermocouple sheath is fixed in each reactor, the length of the 1.5-mm diameter movable thermocouple itself (located in the sheath) was selected and rigorously measured to make sure that the end of the thermocouple (measuring point) was placed in the center of the catalytic bed. In others words, the end of the thermocouple was capable of being changed from test to test, depending on the catalytic bed height.

Results The gas composition, gas yield, and tar* were measured at the catalytic bed inlet and exit. Tar* measurements were expressed as tar* content in the flue gas and, from it, tar* conversion by the catalytic bed (Xtar*). These values are indicated in Table 4 for all tests. Let us show how all these variables are modified by the catalytic reactor. Because the three catalysts used (ICI 46-1, BASF G1-50, and Topsøe R-67) proved to have activity similar to tar* elimination,9 the results obtained with these three catalysts will be shown simultaneously. Error theory21,22 was also taken into account in all measurements to check if the differences in data were in the interval of error or not.

Ind. Eng. Chem. Res., Vol. 39, No. 5, 2000 1147

Figure 2. Gas composition at the catalytic bed inlet and exit with time-on-stream (tests AG-5-AG-8. Tb ) 840 °C; τ ) 0.050 kg‚h/ mTb,wet3; ER ) 0.19-0.22. +, N2; 0, H2; b, CO; ×, CO2; 3, CH4; ], C2Hn].

Gas Composition. The H2 content in the flue gas increases (∆H2) by 4-15 vol % (dry basis), as shown in Figure 3. This increase clearly decreases when the ER used in the upstream gasifier increases. Figure 3 also shows how ∆H2 is minor when the catalyst was placed in the secondary catalytic reactor (downstream from the first one) and how silica sand had a low catalytic activity in this case (in which, it has to be remembered, the “soft tars” are destroyed in the upstream gasifier by the inbed use of dolomite). CO Content. The CO content increases (∆CO) by 3-5 vol % (dry basis) by the catalytic reactor (Figure 3). This ∆CO increases slightly with ER increasing in the upstream gasifier. CH4 and C2Hn Contents. The CH4 and C2Hn contents decrease a little by the catalytic reactor, from 0.5 to 2.5 vol % in the case of CH4 and from 1.0 to 1.8 vol % in the case of C2Hn (Figure 4). These increments are progressively smaller when ER increases in the gasifier (Figure 4). Because experiments were made at different space times in the catalytic reactor, it is possible to know the effect of such a variable in CH4 conversion. Such variation is shown in Figure 5. CO2 and H2O Contents. The CO2 and H2O contents decrease by the catalytic reactor, as shown in Figure 6. Such a decrease was expected because the used catalysts are for steam reforming. Steam is a reactant, and thus, it is consumed in the catalytic bed. But not only does steam act as a reactant but so does CO2. CO2 decreases from 2 to 4.5 vol % (dry basis) and H2O from 2 to 10 vol %. The decrease in H2O content in the flue gas is higher than the decrease in CO2.

Figure 3. Increase of the H2 and CO contents in the flue gas by the catalytic reactor for different ER values (*, catalyst in the third reactor).

Figure 4. Decrease of the CH4 and C2Hn contents in the flue gas by the catalytic reactor for different ER values (*, catalyst in the third reactor).

The gas compositions at the gasifier exit, after the

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Table 4. Main Experimental Results test no. AG-1

AG-2

AG-3

AG-4

AG-5

AG-6

13.6 18.9 14.5 4.4 1.7 7 2.04 1900 6.1

15.5 21.1 15.5 4.9 1.7 12 1.8 7900 7.0

13.4 22.5 13.4 5.1 1.8 14 1.8 7000 7.1

21.3 22.3 12.0 3.3 0 4 2.2 10 6.3 96 64

29.8 25.4 9.6 3.5 0.08 9 2.0 27 7.7 99 126

27.1 27.3 10.6 4.0 0.2 11 1.9 21 7.9 99 135

AG-10

AG-11

AG-12

10.5 19.5 14.7 4.8 1.7 15 2.1 1660 6.0

12.6 16.4 14.5 4.2 1.2 7 2.5 970 5.6

15.7 20.1 11.8 5.6 1.7 4 1.9 530 7.3

17.7 23.2 12.3 3.9 0.26 10 2.2 32 6.4 98 79

16.5 21.6 10.7 3.4 0.3 2 2.55 63 5.9 93 51

22.5 25.4 7.4 1.6 0.6 2 2.24 3 6.6 99 141

AG-14a

AG-14b

AG-14b

18 22.9 10.6 5.1 0.5 8 1.79 190 3.9

14.9 19.6 12.7 7.1 1.3 13 1.43 700 7.4

18 22.9 10.6 5.1 0.5 8 1.43 140 6.9

21 24.6 9.5 3.4 0.3 6 1.82 3 6.8 85 200

18 22.9 10.6 5.1 0.5 8 1.73 140 6.9 80 190

21 24.6 9.5 3.4 0.3 6 1.52 3 6.8 90 170

Catalytic Bed Inlet gas composition H2 (vol %, dry basis) CO (vol %, dry basis) CO2 (vol %, dry basis) CH4 (vol %, dry basis) C2Hn (vol %, dry basis) H2O (vol %, wet basis) Ygas (mn3/kg of daf) tar* content (mg/mn3) LHV (MJ/mn3)

13.1 22.0 13.4 5.2 1.8 14 2.07 3400 7.0

13.9 19.6 12.7 5.3 1.7 15 2.06 2400 6.7

16.2 18.63 13.9 5 1.7 13 2.03 5200 6.7

Catalytic Bed Exit gas composition H2 (vol %, dry basis) CO (vol %, dry basis) CO2 (vol %, dry basis) CH4 (vol %, dry basis) C2Hn (vol %, dry basis) H2O (vol %, wet basis) Y gas (mn3/kg of daf) tar* content (mg/mn3) LHV (MJ/mn3) Xtars* (%) kapp,tars* (mTb,wet3/kg‚h)

21.4 26.8 10 3.5 0.05 5 2.3 50 7.0 98 97

24.4 22.9 10.2 3.2 0.05 3 2.4 98 6.7 96 80

AG-7

AG-8

23.4 19.2 12.5 4.3 0.7 12 2.3 282 6.9 95 56 test no. AG-9

Catalytic Bed Inlet gas composition H2 (vol %, dry basis) CO (vol %, dry basis) CO2 (vol %, dry basis) CH4 (vol %, dry basis) C2Hn (vol %, dry basis) H2O (vol %,wet basis) Ygas (mn3/kg of daf) tar* content (mg/mn3) LHV (MJ/mn3)

14.6 22.4 13.6 4.7 1.7 13 1.8 7310 6.9

12.5 21.8 13.8 4.4 1.7 11 1.9 4764 6.5

12.9 19.9 14.8 4.9 1.7 6 2.2 580 6.5

Catalytic Bed Exit gas composition H2 (vol %, dry basis) CO (vol %, dry basis) CO2 (vol %, dry basis) CH4 (vol %, dry basis) C2Hn (vol %, dry basis) H2O (vol %, wet basis) Ygas (mn3/kg of daf) tar* content (mg/mn3) LHV (MJ/mn3) Xtars* (%) kapp,tars* (mTb,wet3/kg‚h)

27.9 25.9 10.7 3.7 0.05 10 2.03 12 7.6 99 113

23.8 25.1 11.0 3.6 0.26 6 2.05 70 7.2 98 80

22.8 25.2 11.2 3.5 0.1 5 2.33 11 7.0 98 79 test no.

AG-13

AG-13

AG-14a

Catalytic Bed Inlet gas composition H2 (vol %, dry basis) CO (vol %, dry basis) CO2 (vol %, dry basis) CH4 (vol %, dry basis) C2Hn (vol %, dry basis) H2O(vol %, wet basis) Ygas (mn,dry gas3/kg of daf) tar* content (mg/mn3) LHV (MJ/mn3)

14.1 18.3 15.4 4.7 1.9 12 2.17 1600 6.7

18.2 21.6 13.0 4.1 1.4 9 2.21 270 7.0

14.9 19.6 12.7 7.1 1.3 13 1.82 700 7.4 Catalytic Bed Exit

gas composition H2 (vol %, dry basis) CO (vol %, dry basis) CO2 (vol %, dry basis) CH4(vol %, dry basis) C2Hn (vol %, dry basis) H2O (vol %, wet basis) Ygas (mn3/kg of daf) tar* content (mg/mn3) LHV (MJ/mn3) Xtars* (%) kapp,tars* (mTb,wet3/kg h)

18.2 21.6 13.0 4.1 1.4 9 2.21 270 7.0 86 140

21.2 24.1 10.6 3.2 0.7 4 2.34 2 6.9 91 130

18 22.9 10.6 5.1 0.5 8 2.09 190 6.9 73 210

Ind. Eng. Chem. Res., Vol. 39, No. 5, 2000 1149 Table 5. Gas Composition at the Gasifier Exit and after the Second Reactor (Catalytic Bed): Experimental Ones and Corresponding to Equilibrium for 0.20 < ER < 0.26 (a) 0.0062 < τ < 0.014 kg‚h/mTb,wet3 experimental, exit of (vol %, wet basis) H2 CO CO2 CH4 H2O

gasifier 14-15 18-20 13-15 4-7 12-13

catalyst bed 18 22-25 11-13 4-5 8-9

equilibrium (840 °C) 27-29 29-32 6-8 0.014-0.026 7-8

(b) 0.043 < τ < 0.056 kg‚h/mTb,wet3 experimental, exit of (vol %, wet basis) Figure 5. CH4 conversion in the catalytic bed at different space times in it, for three different catalysts and for silica sand.

H2 CO CO2 CH4 H2O

gasifier 13-15 19-22 12-14 4-5 13-14

catalyst bed 21-26 19-26 10-12 3-4 6-12

equilibrium (840 °C) 27-29 29-32 6-8 0.014-0.026 6-8

Figure 7. Low heating value of the gas at the catalytic bed inlet and exit (0.037 < τ < 0.056 kg‚h/mTb,wet3).

Figure 6. Decrease in the CO2 and H2O contents in the flue gas by the catalytic reactor for different ER values (*, catalyst in the third reactor).

catalytic reactor, and for chemical equilibrium are shown in Table 5. Lowest Heating Value (LHV) of the Flue Gas. Once the gas composition is known, the LHV is easily calculated at the inlet and exit of the catalytic reactor. Such values are shown in Figure 7. LHV increases some (by 0.5 MJ/mn3, on average) by the catalytic reactor, and this increase (of LHV) is progressively lower upon ER increasing in the upstream gasifier. Gas Yield. The gas yields at the inlet and exit of the catalytic reactor are shown in Figure 8. The gas yield increases by the catalytic reactor by 0.2 mn3/kg of biomass daf (on average). This increase in gas yield seems to decrease at relatively high values of ER.

Figure 8. Gas yield from biomass at the catalytic bed inlet and exit (Tb ) 840 °C; 0.040 < τ < 0.054 kg‚h/mTb,wet3.

Tar Content. The tar* content at the inlet and exit of the catalytic reactor in a “relatively long-term” test is shown in Figure 9. Because of the location of the pilot plant, which was not allowed to work at night, such a long-term test became, in fact, four different tests (AG5-AG8) performed in consecutive days. The decrease in tar* content (averaged with respect to time-on-stream) for the three catalysts used is shown

1150


Figure 9. Tar* content in the flue gas at the catalytic bed inlet and exit for tests AG-5-AG-8.

Figure 10. Tar* content in the flue gas at the catalytic bed inlet and exit for different values of ER in the gasifier (Tb ) 840 °C; 0.014 < τ < 0.056 kg‚h/mTb,wet3;*, catalyst in the third reactor).

in Figure 10 for different values of ER in the upstream gasifier. It can be clearly observed how from 4 g of tar/ mn3 the tar* content decreases well below 1 g of tar*/ mn3. Because the bottom part of Figure 10 does not give enough detailed information, which is very important in this work, the bottom part (tar* content at the exit of the catalytic bed) is amplified and shown then in Figure 11. Now, it is clearly seen how the tar* content in the flue gas at the catalytic reactor exit in most tests was below 50 mg of tar*/mn3. This fact had already been found5,9 but had not for commercial catalysts at their full size and shape. This result is, maybe, the key result in this work: “tar* can be easily removed with targeted or selected commercial steam-reforming catalysts”. Effect of Different Operation Parameters on Tar* Conversion From the tar* content at the inlet and exit, tar* conversion (Xtar*) was calculated and related to the main operation parameters as follows: Effect of the Gas Residence Time (in the Catalytic Reactor). Different tests were made at different gas residence times or space times (τ). Variation of Xtar* with τ at 840 °C is shown in Figure 12. It is observed how, at a very low space time of 0.02 kg‚h/mTb,wet3, tar* conversion is 90%. At τ ) 0.06 kg‚h/mTb,wet3, tar* conversion is 99%. It is surprising how fast the tar* decomposition reaction is over these nickel-based catalysts at 840 °C.

Figure 11. (Bottom part of Figure 10 with more detail) Tar content in the flue gas at the exit of the catalytic bed (Tb ) 840 °C; 0.014 < τ < 0.056 kg‚h/mTb,wet3; *,catalyst in the third reactor).

Figure 12. Tar* conversion at different gas residence times in the catalytic reactor (0.19 < ER < 0.35; Tb ) 840 °C; dp(catalyst) ) -14 to +7 mm; dp(silica sand) ) -1.6 to +1.0 mm).

To know to which extent such tar* elimination at 840 °C is by catalytic or by thermal reactions, some tests were done with calcined silica sand of dp ) -1.6 to +1.0 mm and the results are shown in Figure 12. It is observed how even silica sand has relative activity for tar* elimination at such high temperatures, but how tar* conversions never exceed 50% with silica sand. The remaining 50% tars* (or “hardest tars”) are destroyed only by catalytic reactions. Thermal reactions on silica sand destroy (eliminate) only tars* which could be called “easy-to-destroy tars”.20 Effect of ER. The value of ER used in the upstream gasifier has an influence on the downstream tar* conversion as shown in Figure 13. When ER is increased, the tar* content in the flue gas at the gasifier exit decreases,2,5 but such tars (the ones produced at high ER values) would be more difficult to be catalytically converted/destroyed.20,23 The results shown in Figure 13 confirm the importance of ER on the refractoriness of tars to be catalytically converted. Effect of H2O/C*. The H2O/C* ratio in the flue gas, C* being the atom‚g of C in the flue gas (CH4, C2Hn, tars*) to be reformed, is of deep importance in catalytic steam reforming of natural gas and of heavier hydrocarbons.24 This ratio also proved to be quite important in the steam reforming of tars* on crushed catalysts.9 Now, in this work, it is also important, as Figure 14


Figure 13. Tar* conversion in the catalytic reactor for different values of ER in the gasifier (0.018 < τ < 0.056 kg‚h/mTb,wet3; Tb ) 840 °C; dp ) -14 to +7 mm).

Figure 14. Tar* conversion in the catalytic bed for different H2O/ C* ratios in the flue gas at the inlet of the catalytic bed (0.018 < τ < 0.056 kg‚h/mTb,wet3; Tb ) 840 °C; dp ) -14 to +7 mm; 0.19 < ER < 0.35).

shows. Tar* conversion increases somewhat with H2O/ C* (a sharper increase cannot be observed in this figure because tar conversions were next to equilibrium, 100%). Reacting Network All the data shown to this point indicates that the reacting network in the catalytic reactor is the following one:

{ }

H2O tar* + CO f H2 + CO + CH4 + C2Hn 2

{ }

H2 O CH4, C2Hn + CO f H2 + CO 2 CO + H2O / CO2 + H2 This reacting network explains and/or fits data shown in Figures 3-14. It is not new. It was already known. What is new in this work is to what extent these reactions are carried out with some selected and fullsize commercial catalysts. Apparent Kinetic Constant for Tar* Elimination The easiest kinetic model for tar removal (single and first-order reaction with respect to tar*) has been

Figure 15. Apparent and first-order kinetic constant for overall tar* elimination for different values of ER in the gasifier (0.018 < τ < 0.056 kg‚h/mTb,wet3; Tb ) 840 °C; dp ) -14 to +7 mm).

Figure 16. Apparent and first-order kinetic constant for different H2O/C* ratios in the flue gas at the inlet of the catalytic reactor (0.018 < τ < 0.056 kg‚h/mTb,wet3; Tb ) 840 °C; dp ) -14 to +7 mm; 0.19 < ER < 0.35).

discussed in other works.20,25,26 It is useful for a quick comparison of the results obtained by different institutions, or with different catalysts, or under different operating conditions, such as the simplest first-order kinetic model that provides an apparent kinetic constant (kapp) for the overall tar* removal (kapp,tar*) which can be easily calculated by

kapp, tar* ) [-ln(1 - Xtar*)]/τ

(1)

This parameter is shown in Figure 15 for the three catalysts under different ER values in the upstream gasifier. kapp,tar* values for these full-size catalysts range between 50 and 130 mTb,wet3/kg‚h. It is checked again how upon increasing ER, kapp,tar* decreases, and how kapp,tar* increases somewhat with the H2O/C* ratio, as Figure 16 shows for two intervals of ER. It is important now to compare the values for kapp,tar* obtained here with the those previously obtained for the same (or similar) catalysts under the same operation parameters but with the catalysts being crushed in such previous works. This comparison is shown in Table 6. Now, with the catalyst in its full size, the activity parameter (kapp,tar*) is clearly lower than that of when the catalyst was crushed. This fact is obvious, but what is new in this work is to what extent kapp,tar* is decreased

1152


Table 6. Representative Values for the Apparent Kinetic Constants for Overal Tar Removal (in Biomass Gasification with Air) Obtained by Different Authors author Narvaez et al., catalyst space time (kg‚h/mTb,wet3) dp (mm) temperature(°C) ER H2O/C* kapp,tars* (mTb,wet3/kg h)

19975

BASF G1-25 S 0.025 -1.6 to +1 800 0.28 3.9 194

this work

Corella et al., 19999 ICI 46-1 0.014 -1.6 to +1 845 0.23 2.3 294

ICI 46-1 0.014 -14 to +7.0 840 0.24 1.6 130

BASF G1-50 0.043 -6.4 to +5 830 0.24 1.9 97

The ICI 46-1 catalyst was used at 840 °C and with ER from 0.19 to 0.22 (notice, with these relatively low ER values the tar yield or tar content is the maximum one.2,3,23 In other words, when work was performed in the catalytic bed under relatively bad conditions, no deactivation was detected. Of course, a test of 50 h-onstream in a pilot plant under the stationary state is not enough to design a commercial plant, but it was the maximum for these authors and/or plant. Several months-on-stream are absolutely out of the rhealm of possibility of this research and development which, one should remember, was undertaken at a university. The last step for checking the technical feasibility of this gas cleaning process should be undertaken by an industrial company. Conclusions Figure 17. Apparent and first-order kinetic constant for a “long term” pilot-scale test, at the stationary state, using the ICI 46-1 catalyst (Tb,averaged ) 840 °C; 0.045 < τ < 0.056 kg‚h/mTb,wet3; dp ) -14 to +7 mm).

by using the catalyst in its full size. In the case of the ICI 46-1 catalyst, under the same conditions (indicated in Table 6), the value of kapp,tar* is 130 instead of 294 mTb,wet3/kg‚h). The design of future commercial catalytic beds for tar* removal will have to take into account this important difference in the values for such a kinetic constant. Full-size commercial catalysts provide kapp,tar* values different from the ones obtained with crushed catalysts, and this well-known fact is quantified here. Lifetime of the Catalysts Once proven, the high activity of these commercial catalysts for tar removal and gas composition upgrading, they have to show a long lifetime too. Deactivation could be due, in this process, to three main causes: particulates, or dust, sulfur, and coke formation (from tar, mainly). In this pilot plant the content in particulate in the produced flue gas is low because of the three inseries high-efficiency cyclones located at the gasifier exit. The sulfur content is very low in biomass and in the thus-produced gas. Coke formation from tar is important when the tar content in the flue gas is above 2 g/mn3,4,7,9 which is not the case in which in-gasifierbed use of dolomite and an optimized operation of the gasifier generate a raw gas with a low tar content.2 Besides, coke formed on the catalyst surface would be continuously gasified by the H2O, H2, and CO2 present in the flue gas at these relatively high temperatures (as the one, 840 °C, used in this work).12 As a result of all the above conditions, the results shown in Figure 17 are not surprising. A relatively longterm test (50 h under stationary state) was made by working 4 days in four continued tests (AG-5-AG-8).

Optimization of the gasifier (both its design and operation) is definitely of big importance for the technical feasibility of the downstream catalytic hot gas cleanup. The necessary low tar content in the flue gas at the catalytic reactor inlet was only obtained by an optimized gasifier. In this way, the gasifier used in this work under the following and selected main variables,

ER moisture of the biomass dolomite content in the gasifier bed bed temperature

0.19-0.35 8-10 wet basis, % 20-30 wt % 800-820 °C

produced a gas with the following main properties: tar* content LHV H2O (vol %, wet basis) H2O/C*

1-8 g/mn3, dry gas 6-7 MJ/mn3 7-16 0.5-3 mol/atom‚g of C*

With this gas at the inlet of the catalytic reactor and, under the following conditions in it, temperature (bed center) space time dp catalyst

830-840 °C 0.0062-0.056 kg‚h/mTb,wet3 -14 to +7 mm

the main components in the flue gas and some of its properties vary in this way (by the catalytic bed): H2 content CO content CO2 content CH4 content H2O content LHV Ygas tars* content

increases from increases from decreases from decreases from decreases from increases from increases from decreases from

11-15 16-23 12-15 4-7 7-16 6-7 1.8-2.5 1-8 g/mn3

to to to to to to to to

18-28 18-26 9-14 2-5 4-11 6.5-8 2.1-2.6 2-300

vol %, dry basis vol % dry basis vol % dry basis vol % dry basis vol % wet basis MJ/mn3 mn,dry gas3/kg daf mg/mn3

Tar conversion increases with space time in the catalytic


reactor and H2O/C* in the flue gas and decreases upon increasing ER in the upstream gasifier. Tar* conversions close to 100% are easily reached at residence times g0.20 s, which is equivalent to space times g0.040 kg‚h/mTb,wet3. On the other hand, with a gas residence time of 0.07 s (equivalent to 0.014 kg‚h/ mTb,wet3), conversions between 70 and 90% are obtained. CH4 conversions g80% are simultaneously obtained with space times g0.043 kg‚h/mTb,wet3. Apparent kinetic constants for an overall first-order tar* elimination are 80-130 mTb,wet3/kg‚h for ER between 0.19 and 0.24 and 50-80 mTb,wet3/kg‚h for ER between 0.26 and 0.35 (for Tb ) 840 °C and τ ) 0.0180.056 kg‚h/mTb,wet3). Under the experimental conditions indicated in this work, catalyst activity for tar* elimination does not decrease in a “long-term” test of 50 h. This time-onstream is not enough for a commercial application, of course, and further testing of these catalysts in a very long-term test is required. Despite this limitation in this research (tests of relatively short times-on-stream), a clear step forward has been given in catalytic hot gas cleanup (in biomass gasification). It has been demonstrated how, for instance, the commercial ICI 46-1 catalyst in its full (commercial) size is highly active and useful for such gas cleanup. Nickel-based commercial steam-reforming catalysts (for naphthas) destroy tars present in the gas generated in fluidized-bed biomass gasifiers in a very efficient way and are a good solution for the technical feasibility of the overall gasification process. Acknowledgment This work has been carried out under the JOULE III Program of the EU DG-XII, Project JOR3-CT95-0053. The authors thank the European Commission for its financial support. The work also received small financial support too from the Spanish DGES Financed Project PB96-0743. Discussions on the use of commercial steamreforming catalysts for this process with Mr. Harald Sha¨fer and Roland Spahl of BASF AG, Ludwigshafen (Germany), with P. E. Hojlund Nielsen of Haldor Topsøe A/S of Lyngby (Denmark), with Dr. Pekka Simell from VTT (Finland), and with Dr. Twigg from Jonhson Mathey (U.K.) were fruitful and are recognized. The authors are also grateful for the invoice of samples of catalysts by BASF AG (Ludwigshafen), ICI- Katalko (Billingham, Cleveland, U.K.), and Haldor Topsøe A/S (Lyngby, DK). Nomenclature Ctar* ) tar* content in the flue gas (mg/mn3) daf ) dry, ash free dp ) particle size of the catalyst (mm) H2O/C* ) steam-to-(carbon-to-be-reformed) (CH4 + C2H2 + C2H4 + C2H6 + tars*) ratio in the flue gas at the inlet of the catalytic reactor (mol of H2O/(atom‚g of C*)) ER ) overall equivalence ratio (air-to-fuel weight ratio used in the experiment divided by the air-to-fuel weight ratio for stoichiometric combustion) (dimensionless) Hbed ) height of the catalytic bed (cm) kapp, tar* ) apparent kinetic constant for tar* removal (mTb,wet3/kg‚h) LHV ) low heating value of the gas (MJ/mn3) P ) pressure in the catalytic reactor (over atmospheric pressure) (mbar)

Q ) gas flow rate at the inlet of the catalytic reactor (mTb,wet3/h] tar* ) tar sampled and measured according to the method used by UZ and UCM (Spain), and described in ref 5 SV ) gas hourly space velocity in the catalytic reactor (mn,wet3/h‚m3) t ) time on stream (h) Tb or Tbed ) temperature measured in the center of the catalytic bed (°C) u0 ) superficial gas velocity of the gas at the inlet of the reactor (cm/s) umf ) minimum fluidization velocity (cm/s) W ) weight of catalyst (kg) XCH4 ) methane conversion (dimensionless) Xtar* ) tar* conversion (dimensionless) Ygas ) gas yield (mn,dry gas3/kg daf) Greek Symbols τ ) space time, defined as W/Q (kg‚h/mTb,wet3) τ0 ) gas residence time (s) Fbed ) catalytic bed density (kg/L)

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Received for review October 7, 1999 Revised manuscript received January 28, 2000 Accepted February 9, 2000 IE990738T

Biomass Gasification with Air in Fluidized Bed. Hot Gas Cleanup with

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