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Catalytic conversion of fast pyrolysis oil to hydrocarbon fuels over HZSM-5 in a dual reactor system
Biomass and Bioenergy Vol. 5, No. 6, pp. 445-455, 1993 Copyright 0 1994 Ekvicr Science Ltd
Printed in Great Britain. All rights reserved 0961-9534/93$7.00+ 0.00
CATALYTIC CONVERSION OF FAST PYROLYSIS TO HYDROCARBON FUELS OVER HZSM-5 IN A DUAL REACTOR SYSTEM R.
OIL
K. SHARMAand N. N. BAKHSHI*
Department of Chemical Engineering, University of Saskatchewan, Saskatoon, S7N OWO, Canada (Received
20 February
1993; revised received
27 July
1993; accepted
I5 September
1993)
Ah&r&-The catalytic conversion of fast pyrolysis bio-oil to hydrocarbon fuels was studied over HZSM-5 at atmospheric pressure. Experiments were conducted in a dual reactor system having two reactors in series. The temperatures in these reactors were in the range 340~400°C (first reactor) and 350-450°C (second reactor). The bio-oil was co-processed with tetralin in all the runs. The objective was to maximize the organic distillate product with a high concentration of aromatic hydrocarbons. The maximum amount of organic distillate in the effluent from the second reactor was 21 wt% of the bio-oil feed and the highest concentration of aromatic hydrocarbons was 76 wt% of the distillate. The dual reactor system was particularly beneficial when the temperature in the first reactor was low. Thus, with the first reactor at 340°C. the yields of organic distillate and aromatic hydrocarbons were 15-16 wt% and 8-11 wt% of wood, respectively, which are nearly two-fold compared to those from a single reactor system operated at 340°C (7.8 wt% and 4.8 wt%). Under the above conditions, the coke plus char yields were 25-26 wt% of wood which are up to 10 wt% lower than from the single reactor system at 340°C (29 wt%). Keywords-Pyrolysis
oil, upgrading, HZSM-5, dual reactor system.
hydrotreating catalysts are only active at high pressures in the presence of hydrogen, whereas Biomass is a renewable source of fuels and the operation with HZSM-5 is at atmospheric chemicals which has negligible sulfur, nitrogen, pressure. Also, coking with HZSM-5 is low due and metal contents. It is also CO2 neutral. to its shape-selective properties and intraThermochemical conversion of biomass to crystalline network.’ The catalyst has mediumfuels and chemicals is a two-step process. The size pores which restrict the formation of large, first step is the fast pyrolysis or high pressure highly branched molecules which lead to coke liquefaction of biomass into a bio-oil which is formation. However, unlike in small-pore zeothen converted to hydrocarbon fuels in the lites, the HZSM-5 pores permit the entry of second step of this route. With woody biomass, simple branched molecules necessary for their the yields of bio-oils from fast pyrolysis are cracking. 75-80 wt% of wood compared to 30-35 wt% Due to the poor stability of the bio-oils, low from the high pressure liquefaction. The pyroconcentrations of bio-oil vapors in the feed have lytic oils contain up to 50 wt% oxygen and are been used in most studies in the literature.4,8,9 usually unstable resulting in their polymerization In some cases, only the low boiling fraction of at relatively low temperatures. These characterthe bio-oil was used as feed.” Churin et al.’ and istics make the bio-oils unsuitable as fuels. In Sheu er ~1.~ hydrotreated the bio-oils using order to stabilize the bio-oils and to lower their Co-Mo/Al,O, catalyst by co-processing the oxygen contents, it is necessary to convert them bio-oils with tetralin. Also, in our earlier work into hydrocarbon fuels. with a high pressure bio-oil, it was observed that The conversion of bio-oils to hydrocarbon the handling and processing of the bio-oil befuels has been studied in literature using both the came relatively easy after tetralin addition. The hydrotreating catalysts such as Co-Mo/Al,O, ,I_3 use of tetralin in coal liquefaction is already well and zeolite catalysts such as HZSM-5.4-6 The known. Therefore, it was considered desirable __to study the conversion of bio-oil over HZSM-5 *To whom correspondence should be addressed. with tetralin co-processing. 1. INTRODUCTION
445
446
R. K. SHAKMA
and N. N. BAKHSHI
Sharma and Bakhshi” studied the conversion of fast pyrolysis oil to hydrocarbon fuels over HZSM-5 in a fixed-bed reactor at 340-410°C both with and without tetralin co-processing. The co-processing runs were of special interest since the concentration of aromatic hydrocarbons in organic distillate in these runs was usually 15-20 wt% higher compared to that in the absence of tetralin. The highest amount of char in the co-processing runs was 27 wt% of bio-oil at 340°C which decreased at high temperatures. Interestingly. the char fraction also became increasingly carbonaceous in nature above 370°C. This was also indicated by the amount of water in the product, which increased from 23 wt% of bio-oil at 340°C to 42 wt% at 410°C. The maximum amount of organic distillate was 19 wt% of the bio-oil at 370°C while the maximum concentration of aromatic hydrocarbons in the distillate was 83 wt% at 390°C. The amount of coke increased with temperature to about 20 wt% of the bio-oil at 390 and 410°C. The conversion of the nonvolatile fraction of the bio-oil was highest at 410°C. The above study showed that the high temperatures were useful for cracking of the non-volatile fraction of the bio-oil. However, the coking increased with temperature. At high temperatures, the amount of water in the product was high, which was detrimental to hydrocarbon yields since the hydrogen content of the bio-oil was low. It was found that 26-39 wt% of the bio-oil was lost in coke and char fractions due to its poor stability which led to polymerization of the bio-oil components. Baker and Elliott’ observed that the coke and char formation during the hydrotreatment of bio-oil over Co-MO/A&O, catalysts decreased when the upgrading was done in a dual reactor system. They also found that some of the reactive groups in the bio-oil were converted to more stable compounds by the catalytic treatment at low temperatures (250-270°C) in the first reactor. Hence, it was of interest to investigate the performance of the dual reactor system for the conversion of bio-oil over HZSM-5 catalyst. In this work, the conversion of fast pyrolysis bio-oil was studied at atmospheric pressure in a dual reactor system using HZSM-5 catalyst. The dual reactor system consisted of two catalytic reactors in series. The temperatures in these reactors were in the range 340-4OO“C (first reactor) and 350-450°C (second reactor). The bio-oil was co-processed with tetralin in
all the runs. The objective was to maximize the amount of organic distillate product with a high concentration of aromatic hydrocarbons. Also, the results were compared with those from the single reactor system which are given elsewhere. 2. EXPERIMENTAL
SECTION
2.1. Catalyst preparation and characterization The catalyst HZSM-5 was prepared according to the method described by Chen et al.” The details of the preparation and characterization of the catalyst are given elsewhere.” The particle size of the catalyst in the reactor was 5001410 pm and its surface area was 365 m’g ‘. The sodium and aluminum contents were found to be 0.075 wt% and 0.76 wt%, respectively, and the Si0,/Al,03 weight ratio was 59. The unit cell composition was reported as Na,,, H,,,Al,,,Si,,,,O,,, 14H,0.‘3 The X-ray diffraction pattern of the finished catalyst was found to be similar to that reported by Argauer and Landolt.14 Prasad et aLI3 observed that the acid site density and stability of HZSM-5 were not altered by heating the catalyst at temperatures up to 500°C. 2.2. Bio-oil The bio-oil was obtained from the Ensyn Technologies, Inc., Gloucester, Ontario, where it was produced by Rapid Thermal Processing (RTP) of maple wood.i5 The yield of the bio-oil was 74 wt% of wood. The byproducts from pyrolysis were: char, 15 wt%; CO and COz, 10 wt%; and hydrocarbon gases, 1 wt%. The bio-oil was in the form of an emulsion containing 21 wt% water including 8 wt% moisture which was present in the wood. The elemental analysis of the bio-oil including the water was: 43.6 wt% C, 8 wt% H, and 0.5 wt% N. The balance 47.9 wt% was 0. Vacuum distillation of the bio-oil at 172 Pa showed that the nonvolatile fraction (pitch) amounted to 34 wt% of the bio-oil at 200°C whereas 60 wt% of the bio-oil was low boiling (below 60°C at 172 Pa). 2.3. Equipment, analysis
run procedures
and product
The experiments were carried out at atmospheric pressure in a continuous, downflow, dual reactor system. The details of the experimental equipment, run procedures and analytical methods are described elsewhere” and are given here only briefly. The dual reactor system
447
Catalytic conversion of fast pyrolysis oil Table I. Scope of experiments for bio-oil upgrading in the dual reactor system Temperature in Reactor 1 (“C) Temperature in Reactor 2 (“C) Bio-oil/tetralin (wt) Catalyst in each reactor (g)
340-400 350-450 1.0 2.0
consisted of two catalytic reactors in series. The temperatures in these reactors were in the range 340-400°C (first reactor) and 350-450°C (second reactor). The operating conditions are summarized in Table 1. The feed consisting of bio-oil and tetralin mixture, in 1: 1 weight ratio, was kept vigorously stirred and was introduced to the first reactor by a micro-metering pump (Eldex, model A-60-S). The effluent leaving this reactor directly entered the second reactor. The products from the second reactor were separated into gas and liquid fractions. The liquid product condensed in a trap whereas the gaseous product was collected over a saturated brine solution. The rate of gas formation was monitored by the frequency of gas bubbles appearing in the brine solution. The rate decreased slowly during the run, implying a continuous deactivation of the catalyst. The products were collected over the entire run period and an average analysis was obtained. The char particles were mainly formed above the catalyst bed and were removed from the reactors at the end of the run. The quartz wool plugs placed above each catalyst bed prevented the char particles from mixing with the catalyst during their removal. Finally, the spent catalyst was recovered and regenerated with air at 600°C for 1 h in order to determine the amount of coke. No attempt was made to recover the tetralin from the products. The gaseous product was analyzed by gas chromatography using a combination of packed columns and flame ionization (FID) and thermal conductivity detectors (TCD). The liquid product was in two phases: an organic phase and water. The organic phase was separated and was found to contain 334 wt% water. It was distilled at 200°C under a vacuum of 172 Pa to separate the desired “organic distillate” product from “residue”. The organic distillate was later analyzed by gas chromatography using a capillary column and FID. The water phase was also analyzed. The identity of the peaks was confirmed by GC/MS analysis and by the use of pure compounds. Among the other products, the residue, coke, and char fractions were not analyzed.
3. RESULTS
AND DISCUSSION
3.1. Overall material balance The overall material balance from various runs is presented in Table 2. A mass balance for tetralin over the two reactors showed that the conversion of tetralin was small in all the runs below 420°C. In runs where the temperature in the second reactor was 450°C the conversion of tetralin was 8-10 wt%, which affected both the organic distillate yields and the product compositions. The effect on organic distillate yields was taken into account by subtracting the amount of tetralin which was present in the product. On the other hand, the effect on product compositions could not be accounted for, since the exact nature of products from tetralin conversion in the presence of bio-oil was not known. However, an estimate of the contribution of tetralin conversion to the product composition was made. Based on our earlier work, in which tetralin alone was passed over HZSM-5,16 the main products of tetralin conversion were C,-C, aliphatic and C,-C,, aromatic hydrocarbons. Therefore, it is estimated that. in this study, 5-6 wt% of tetralin converted to aromatic hydrocarbons in the organic distillate fraction at 450°C while 3-4 wt% was converted to the gaseous product. The product analysis also indicated that the degree of tetralin dehydrogenation to naphthalene in the presence of bio-oil was small. Due to the low reactivity of tetralin, the results were calculated on a tetralin-free basis. The weight hourly space velocity (WHSV) in Table 2 is based only on the flow rate of bio-oil entering the first reactor. The run time was set at 45 min even though the catalyst was still active at the end of the run, as indicated by a significant rate of gas production. The material balance was better than 95 wt% in most runs. It is seen in Table 2 that the amount of organic distillate, i.e. the desired product, was in the narrow range of 18-21 wt% of the biooil. These amounts are slightly higher than those from the single reactor system, which were 19 wt% or lower in the co-processing runs at 340-41O‘C.” The composition of the organic distillate, discussed later, showed that it mostly consisted of hydrocarbons and should be suitable as a fuel or as a source of hydrocarbons. The amount of gaseous product was between 10 and 14 wt% of the bio-oil, as shown in Table 2. These amounts are lower than those with the hydrotreating catalysts which were up
448
R. K. SHARMA
and N. N. BAKHSHI to 35 wt% of the bio-oil.‘.” This suggests that
the secondary cracking of the organic distillate components was relatively slow over HZSM-5. Interestingly, the amount of residue in the product was below 5 wt% indicating that a substantial cracking of the non-volatile fraction of the bio-oil occurred during the upgrading. The product also contained 22-25 wt% water including water which was present in the bio-oil. Diebold et ~1.~ pyrolyzed wood in a vortex reactor and passed the pyrolytic vapors directly over HZSM-5 at 400°C using steam as carrier gas. The product was found to contain 32 wt% water. This shows that the fraction of oxygen which was lost in water by dehydration was higher than that in this study. The dehydration reactions were undesirable since the hydrogen content of the bio-oil was low. Another advantage of using the dual reactor system was a lower char formation than in the single reactor system. As seen in Table 2, the amount of char decreased slightly with the temperatures of the two reactors and was 13-l 6 wt % of the bio-oil. The maximum char amount from the single reactor system was 27 wt% of bio-oil at 340°C. It should be noted that the feed rates in the dual reactor system were higher even though the residence times in both the systems were similar. This indicates that the amount of char was dependent on the bio-oil feed rate. The char was a solid carbonaceous material, completely insoluble in acetone, suggesting that it was characteristically different from the bio-oil which was soluble. The char was formed by polymerization of the bio-oil components. A typical elemental analysis of the char fraction at 340°C was: 60 wt% C, 5 wt% H, and 0.3 wt% N. The balance 34.7 wt% was oxygen. Diebold et al4 and Milne et aL6 reported up to 16 wt% char when the pyrolytic vapors from the vortex reactor were passed over HZSM-5 at 400°C using steam as carrier gas. Considering that the feed in our study was the “whole” bio-oil, i.e. including the non-volatile fraction, the observed char amounts were rather low, which may be attributed to improvement in the bio-oil stability in the presence of tetralin. The fraction of bio-oil which converted to coke increased with the temperature of the first reactor from 20 wt% of bio-oil at 34O’C (Table 2, Run 13) to 27 wt% at 400°C (Run 3). Interestingly, these coke amounts are much higher than those from heavy oils and tar sands bitumens which are not known to have any coking problems below 400°C. This may be due
Catalytic conversion of fast pyrolysis oil
to the higher oxygen content of the bio-oil compared to the petroleum oils. High coking may also be due to high reactivity and instability of the bio-oil. The coke amounts, in this study, were higher than those which were obtained when only the non-phenolic fraction (NPF) of the bio-oil was treated with HZSM-5.‘(’ The NPF fraction was obtained after separating the phenolic/neutral (P/N) compounds from the bio-oil. The higher coking with the bio-oil (compared to NPF) indicates that, in this study, the phenolic compounds contributed substantially to the coke formation. Elliott” also suggested that the phenolic compounds were the main coke precursors in the hydrotreatment of bio-oil over the CO-MO catalyst. The other major difference between bio-oil and its NPF fraction was in their water contents, which were 21 wt% and 50 wt%, respectively. Ison and Gorte’* reported that the presence of excess water in the feed reduced the formation of coke over HZSM-5 due to partial gasification of coke by steam. However, since the temperatures during the NPF upgrading were lower than those at which significant gasification of coke by steam occurs (> 5Oo”C), the lower coke amounts with the NPF fraction were not due to its higher water content compared to the biooil. Thus, the increased coking with the bio-oil was mainly due to the presence of phenolic compounds. 3.2. Orgunic distillate composition The composition of the organic distillate is given in Table 3. The distillate mainly consisted of hydrocarbons, phenols, acids, ketones and alcohols. The aromatic hydrocarbons were of special interest since the aromatization reactions are particularly catalyzed by HZSM-5.7,‘8 The aromatic concentration was in the range of 53-76 wt% of the organic distillate. The highest concentration was in Run 2, where the temperatures in the first and second reactors were 400 and 420°C respectively. As already mentioned, in runs where the temperature in the second reactor was 450°C 5-6 wt% of tetralin also converted to aromatic hydrocarbons in the organic distillate fraction. The maximum aromatic concentration from the single reactor system was 83 wt%,” whereas Chantal et al.* reported up to 76 wt% concentration of CrC10 hydrocarbons using HZSM-5 catalyst. It was interesting to note that the temperatures at which the maximum aromatic concentration from the upgrading of NPF fraction was obtained were also the
450
R. K. SHARMA and N. N. BAKHSHI
same as those in this study. This indicates that the acidity, acid site density, and selectivity of HZSM-5 were not affected by the difference in the water contents of the two feeds since the water content of NPF fraction was much higher compared to that of the bio-oil. Similar observations were made by Chang et al.’ where no deactivation due to the presence of water was reported in the conversion of methanol over HZSM-5 at 400-520°C even though a significant amount of water was produced during the conversion. The aliphatic hydrocarbons were the other major components of the organic distillation fraction (Table 3). Their concentration varied with the temperature of the second reactor from 16 wt% of distillate at 400°C (Run 1) to 3 wt% at 450°C (Run 3). The organic distillate also contained up to 17 wt% aldehydes and 9 wt%, each, of phenols and ketones. Although the concentration of phenols in the organic distillate was low, these compounds may be utilized effectively by reacting with methanol to form methyl aryl ethers which are known octane enhancers.4 The organic distillate was found to be mostly low boiling (below 60°C at 172 Pa). It is seen in Table 3 that the concentration of low boiling fraction increased with the temperature of the first reactor and was 30-93 wt% of the organic distillate. The increase in the low boiling fraction was due to increased cracking of the heavy components of the bio-oil at high temperatures. The composition of the organic distillate suggests that it is an excellent source of aromatic hydrocarbons. Due to its high hydrocarbon content, it is likely that the combustion characteristics of the organic distillate would be superior to those of the bio-oil. The organic distillate may also be suitable as a component of the reformulated gasoline due to the presence of oxygenated compounds such as alcohols and ethers. However, these aspects were not studied in this work. 3.3. Composition
of the gaseous product
The composition of the gaseous product is presented in Table 4. The major components were CO and CO, with concentrations of 35-42 and 29-39 wt%, respectively. The concentration of CO appeared to increase with an increase in the temperatures of the two reactors whereas the CO, concentration decreased. The above CO, CO2 concentrations are 5-10 wt% higher compared to those from the single reactor system.”
Catalytic
conversion
It was found that up to 22 wt% of the oxygen in the bio-oil was lost in CO and CO2 by decarboxylation in this study. The decarboxylation route was especially desirable for the rejection of oxygen since the hydrogen content of the biooil was low. In this way, the bio-oil hydrogen was mainly used in the formation of hydrocarbons. Mime et ~1.~ also observed that in the presence of suitable co-feeds, a significant fraction of the bio-oil oxygen was rejected in CO and CO2 using HZSM-5 catalyst. In addition to CO and CO*, the gaseous product contained 19-29 wt% C-C, hydrocarbons (Table 4). Among these, ethylene and propylene were the major ones, with concentrations of 5-8 wt% each: As previously mentioned, in runs where the temperature in the second reactor was 450°C 3-4 wt% of tetralin also converted to the gaseous product consisting of hydrocarbons. 3.4. Extent of upgrading The extent of upgrading in terms of pitch conversion and coke level was calculated as a measure of the degree of catalytic cracking of the heavy components into the desired product (see Appendix). The results are given in Table 5. It is seen that the amount of non-volatile pitch was 14-19 wt% of the total product. As already mentioned, the bio-oil contained 34 wt% pitch. This indicates that a maximum of 58 wt% of the pitch in the bio-oil was converted during upgrading. The fraction of pitch which converted to coke was high, as indicated by the coke level which was at its maximum 80 wt% of the pitch. Thus, although the non-volatile components
Table 5. Extent
of upgrading
of fast pyrolysis
oil
of the bio-oil were the major coke precursors, the volatile fraction also contributed to coking in most cases. 3.5. Comparison of the results with literature A further comparison of the results for the main product fractions was made on uniform wood basis (see Appendix). The fractions of interest were organic distillate, coke plus char, aromatic hydrocarbons, and CO and CO,. The effect of first reactor temperature on the yields of these fractions at temperatures of 400 and 450°C in the second reactor is plotted in Figs l-3. The trends at the other temperatures of the second reactor were similar. Figure 1 shows the effect of temperature on the yields of organic distillate, and coke plus char. It should be noted that the results in Fig. 1 are on wood basis whereas those in Table 2 are on bio-oil basis. As seen in the figure, the yields of organic distillate from the dual reactor system were 15-16 wt% of wood which are higher than those from the single reactor system, except when the latter is operated at 370°C.” The yield of coke plus char fraction, in this study, was at its minimum at about 25 wt% of wood when the temperatures in the first and second reactors were 340 and 400°C. As also seen in the figure, the coke plus char yields from the dual reactor system were mostly higher than those from the single reactor system, except when the latter was at 340°C. The low coke plus char yields at high temperatures in the single reactor system were due to the low char amounts.” As already mentioned, the amount of coke in the single
in the dual reactor
T, (“C)
I
400
2 3 4 5 6 7 8 9 10 II 12 13 14 15 16
400 400 380 3x0 380 370 370 360 360 360 340 340 340 340 340 *Tetrahn-free
400 420 450 400 420 450 370 ,450 450 420 400 450 420 350 370 400 basis.
Low boil.
High boil.
43.2 50.4 46.1 56.X 52.7 52.9 39.1 54.9 49.4 52.2 52.0 51.X 52.4 52.9 52.X 52.X
12.1 5.2 10.1 0.9 3.1 5.0 10.7 1.0 5.2 3.X 4.2 4.6 4.1 4.0 4.2 4.2
Pitch 14.x 16.5 14.3 15.9 16.9 16.3 19.5 15.2 16.1 16.6 15.9 16.3 17.X 17.3 16.5 17.2
system* Yield (wt% wood)
Amount (wt% of bio-oil) Run no.
451
Coke 26.5 26.X 27.2 23.4 24.7 24.5 24.5 23.5 23.7 22.4 22.7 21.x 20.0 20. I 20.2 20.4
Unidentified 3.4 1.1 1.7 3.0 2.0 1.3 6.2 5.4 5.6 5.0 5.2 5.5 5.7 5.7 6.3 5.4
Upgrading (wt% pitch)
Aromatics
CO +co,
Pitch conv.
1.9 11.5 11.2 12.3 10.1 9.0 8.8 10.3 7.5 9.6 10.2 X.4 9.1 9.1 10.6 10.8
7.6 7.0 7.1 7.6 6.6 7.9 5.7 11.2 7.1 6.5 7.2 6.9 7.1 7.0 6.7 6.9
56.5 51.5 57.9 53.2 50.3 52.0 42.6 55.3 52.6 51.2 53.2 52.0 47.6 49.1 51.5 49.4
Coke level 71.9 78.8 80.0 68.8 72.6 72.0 72.0 69. I 69.7 65.9 66.8 64.1 5X.X 59. I 59.4 60.0
452
R. K.
SHARMAand
N. N.
BAKHSHI
Legend ??Dbtilktte, .........
T =400% ..a.........................
0 Dbtiti,
T.=4!50%N-s-
. Diatilialu
Sindo moctw
s~alem
o Coke+Char T =4OO*C ,.,,,1,,,1,.,111.1,,Illldlll+llllll)ll,~~*~~,,,~~,.~.~~,,~~,,~~~~~,,,~~1 ACoke+CharT=45OX -,-,-LmL.-s-m-,X CokuXhar. Sir% moctw svstem
01 so
f
SW
420
T-7
first mac$C
Fig. 1. Relationship of temperature with organic distillate and coke plus char
reactor system above 370°C was substantially higher. Since coking reduces the run time before regeneration of the catalyst becomes necessary, the operation of the single reactor system at high temperatures was not entirely beneficial. Figure 2 shows the effect of temperature on the yield of aromatic hydrocarbons. It is seen that the aromatic yield was not strongly dependent on the temperature of the first reactor. On the other hand, with an increase in the temperature of the second reactor, the yield usually increased up to 400°C before decreasing at higher temperatures due to secondary crack-
ing of the aromatics. The runs in which the temperature of the second reactor was 400°C were of special significance. In these runs, the aromatic yield increased from 10.8 wt% of wood to a maximum of 12.3 wt% as the temperature in the first reactor increased from 340 to 380°C. Above 370°C (first reactor), the aromatic yields were similar to those from the single reactor system under similar operating conditions. When the first reactor temperature was 340°C the aromatic yields from the dual reactor system were 8-l 1 wt% of wood. These values are nearly double that of the aromatic yield (4.8 wt% of
Legend T =4OO*C .a......................... H’C
yields.
I___
Sinsle reactorsystem
Fig. 2. Relationship of temperature with aromatic hydrocarbon yield.
Catalyticconversionof fast
453
pyrolysis oil
lZm
0
?TG!MC__ ?
f
II i
T=4O@C .a.........*...............
rn~&sy&m
?? -. 0-
5
/+++:
/
\
..“..............
:..-...+z
I”
??
8 e 4.-: r!
O! s40
\.A
I
SW
I
400
4?D
TutphZnoetor:C
Fig. 3. Relationship of temperature with total yield of CO and CO,
wood) which was observed in the single reactor system operated at 340°C (Fig. 2). Another important aspect of upgrading was the CO + CO, yield which governed the ultimate yield of hydrocarbons. As mentioned before, the high yields of CO + CO2 were desirable in order to utilize the bio-oil hydrogen mainly in the formation of hydrocarbons. The effect of temperature on CO + COz yield is shown in Fig. 3. It is seen that the yield slightly increased with the temperature of the first reactor and was 6-l 1 wt% of wood. The maximum yield was obtained when the temperatures in the first and second reactors were 370” and 450°C respectively (Run 8). The above CO + COz yields are higher than those from the single reactor system, which were below 6 wt% of wood (Fig. 3). Thus, the dual reactor system was particularly advantageous when the temperature in the first reactor was low (<36O”C). In runs where the first reactor was at 34O”C, the organic distillate yields were 15-16 wt% of wood, which are
nearly two-fold compared to those from the single reactor system operated at 340°C (7.8 wt%). It is remarkable that the coke plus char yields in these runs were 25-26 wt%, which are up to 10 wt% lower than in the single reactor system at 340°C (29 wt%). At high temperature; (first reactor), the performance of the dual reactor system was similar to that of the single reactor system. However, the operation of the single reactor system at high temperatures was not particularly desirable since it resulted in high coke amounts and in a decrease in the degree of decarboxylation to CO and CO*, which were both detrimental to the formation of hydrocarbons. 3.6. Optimum yield of hydrocarbon fuels The optimum product yields from this are summarized in Table 6. The maximum of hydrocarbon fuels was 15.9 wt% of where the yield of the light hydrocarbon (C,-C,) was 3.5 wt% of wood. Under
study yield wood gases these
Table 6. Comparison of optimum product yields* Reactor system
Hydrocarbon fuels Char and coke Hydrocarbon gases CO and CO, Water Total *On wt% wood basis.
Dual (this work)
Single (Sharma and Bakhshi)”
Single (Diebold er af.)4
Single (Evans and Milne)’
15.9 44.1 3.5 17.8 18.1
14.8 38.7 6.2 15.9 24.4
12.0 20.0 4.6 26.4 37.0
9.8 25.4 8.7 37.6 24.9
100.0
100.0
100.0
100.0
454
R.K. SHARMA and N.N.
conditions, the coke plus char yield was 45 wt% of wood. This fraction may be used as a process fuel, Alternatively, the char fraction may be liquefied indirectly via gasification and methanol synthesis. The above yield of hydrocarbon fuels was somewhat higher compared to the yield from the single reactor system which was 14.6 wt%” whereas the yield of light hydrocarbon gases was lower in this study. Another significant advantage of using the dual reactor system was that the CO + CO? yield was higher and the yield of water was lower compared to those from the single reactor system (Table 6). The results were also compared with the other studies. The optimum yield of hydrocarbon fuels, in this work, was also found to be higher compared to the yields reported by Diebold et ~1.~ and Evans and Milne,’ which were 12 wt% and 10 wt% of wood respectively (Table 6). Besides, the latter yields were obtained at low concentrations of the bio-oil vapors in the feed. The fraction of oxygen which was rejected in water was relatively low in this study, which was especially desirable. Although the yields of coke plus char, reported by Diebold et al4 and Evans and Milne,’ were low, they were associated with increased decarboxylation and dehydration of the coke and char fraction to CO, CO, and water instead of its conversion to hydrocarbon fuels (Table 6). This study showed that the dual reactor system was useful for converting the bio-oil into hydrocarbon fuels using HZSM-5 as catalyst. The operation was at atmospheric pressure without the cost of producing high pressure hydrogen as in conventional hydrotreatment. Also, the formation of gasoline range hydrocarbons was favored by HZSM-5 since its medium-size pores exclude the formation of bulky branched molecules with over 9 or 10 carbon atoms. Although the catalyst deactivated rapidly, it was found that it was easily regenerated by oxidation with air at 600°C. The results indicated that a significant fraction of the bio-oil oxygen was lost by decarboxylation into CO and CO,. Milne et a1.6 reported that the decarboxylation reactions over HZSM-5 were enhanced in the presence of suitable co-feeds. Thus, it seems that in this study the decarboxylation reactions were enhanced by the tetralin co-processing. In the literature, tetralin has normally been used as a hydrogen donor solvent at high hydrogen pressures so that any naphthalene which was formed by tetralin dehydrogenation
BAKHSHI
was hydrogenated back to tetralin by the hydrogen. 2,3However, in this study, the amount of tetralin dehydrogenated was small compared to the amount of tetralin which was present in the feed, so that the role of tetralin appeared to be mainly as a diluent. In runs where the temperature in the second reactor was 45OC, the contribution of tetralin conversion to the product composition was estimated based on our earlier work in which tetralin alone was passed over HZSM-5. 4. CONCLUSIONS
The dual reactor system was particularly beneficial when the temperature in the first reactor was low, as indicated by the following conclusions: (1) When the temperature in the first reactor was 340°C the yields of organic distillate from the dual reactor system were 15-16 wt% of wood, which are nearly two-fold compared to that from the single reactor system operated at 340°C (7.8 wt%). (2) In runs where the first reactor temperature was 340°C the aromatic hydrocarbon yields from the dual reactor system were 8-l 1 wt% of wood, which are also nearly two-fold compared to the yield from the single reactor system at the same temperature (4.8 wt%). (3) At low temperatures in the first reactor (34O”Q the overall coke plus char amounts were 25-26 wt% of wood, which are up to 10 wt% lower than in the single reactor system at 340°C (29 wt%). (4) The extent of decarboxylation was generally higher in the dual reactor system compared to the single reactor system. This was especially desirable since the hydrogen content of the bio-oil was low. APPENDIX
(A) Pitch = amount residue.
of char + amount
of
% Coke level = amount of coke in the product x 100 amount of pitch in the bio-oil % Yield of a fraction on wood basis = % yield of that fraction on bio-oil basis x yield of bio-oil on wood basis. (B) Average molecular weight of the gas product, M 9’ = X W,M; ‘, where M, and W, are the molecular weight and weight fraction,
Catalytic conversion of fast pyrolysis oil
respectively, of component i in the gaseous product. As an example, MBfor Run 1 was 35.1. Acknowledgements-The authors are grateful to the Bioenergy Development Program of Energy, Mines and Resources (EMR), Canada, for the financial support of this research.
REFERENCES 1. E. G. Baker and D. C. Elliott, Catalytic upgrading of biomass pyrolytic oils. In Research in Thermochemical Biomass Conversion (A. V. Bridgwater and J. L. Kuester, Eds), pp. 883895. Elsevier Applied Science, London (1988). 2. E. Churin, P. Grange and B. Delmon, Catalytic upgrading of pyrolysis oils. In Biomass for Energy and Industry (G. Grassi, G. Gosse and G. Santos, Eds), __ pp. 2.6162.620. Elsevier Applied Science, London (1990). 3. Y. H. E. Sheu. R. G. Anthonv and E. J. Soltes. Kinetic studies of upgrading pine pyrolytic oils by hydrotreatment. Fuel Process. Technol. 19, 31-50 (1988). 4. J. Diebold, J. Scahill, R. Bain, H. Chum, S. Black, T. Milne, R. Evans and B. Rejai, Biomass liquefaction at SERI. In Biomass Thermal Processing (E. Hogan, J. Robert, G. Grassi and A. V. Bridgwater, Eds), pp. 101-108. CPL Press, London (1992). 5. R. J. Evans and T. A. Milne, Molecular-beam mass spectrometric studies of wood vapor and model compounds over HZSM-5 catalyst. In Pyrolysis Oils from pp. 311Biomass (E. J. Soltes and T. A. Milne, Eds), __ 327. ACS Symposium Series 376, American Chemical Societv. Washinaton. DC (1988). 6. T. A. Milne, R. j. Evans and J. Filley, Molecular beam mass spectrometric studies of HZSM-5 activity during wood pyrolysis product conversion. In Research in Thermochemical Biomass Conversion (A. V. Bridgwater and J. L. Kuester, Eds), pp. 910-926. Elsevier Applied Science, London (1988). 7. C. D. Chang, W. H. Lange and R. L. Smith, The conversion of methanol and other O-containing compounds to hydrocarbons over zeolite catalyst. J. Catal. 56, 169-173 (1979).
455
8. P. Chantal, S. Kaliaguine, J. L. Grandmaison and A. Mahay, Production of hydrocarbons from aspen poplar pyrolytic oils over HZSM-5. Appl. . . Catal. 10, 317-332 (1984). 9. N. Y. Chen. D. E. Walsh and L. R. Koenine. Fluidized bed upgrading of wood pyrolytic liquids &d related compounds. In Pyrolysis Oilsfrom Biomass (E. J. Soltes and T. A. Mime, Eds), pp. 277-289. ACS Symposium Series 376. American Chemical Society, Washington, DC (1988). 10. J. F. Mathews, E. G. Tepylo, R. L. Eager and J. M. Pepper, Upgrading of aspen poplar wood oil over HZSM-5 zeolite catalyst. Can. J. Chem. Eng. 63. 686-692
(1985).
Il. R. K. Sharma and N. N. Bakhshi, Catalytic upgrading of fast pyrolysis oil over HZSM-5. Can. J. Chem. Eng. 71, 383-391 (1993). 12. N. Y. Chen, J. N. Maile and W. J. Reagan, Preparation of zeolites. U.S. Pat. 4,112,056 (1978). 13. Y. S. Prasad, H. Yaoliang and N. N. Bakhshi, Effect of hydrothermal treatment of HZSM-5 catalyst on its performance in the conversion of canola and mustard oils to hydrocarbons. Ind. Eng. Chem. Prod. Res. Dezt. 25, 251-257 (1986).
14. R. J. Argauer and G. R. Landolt, Crystalline zeolite HZSM-5 and method of preparing the same. U.S. Pat. 3,702,886 (1972).
15. B. A. Free], R. G. Graham, D. R. H&man and A. J. Vogiatzis, Rapid thermal processing of biomass (RTP): Development, demonstration and commercialization. Paper Presented at Energy from Biomass and Wastes XVZ, Inst. Gas Technol., Chicago, March (1992). 16. R. K. Sharma and N. N. Bakhshi, Conversion of nonphenolic fraction of biomass-derived pyrolytic oil to hydrocarbon fuels over HZSM-5 using a dual reactor system. Biores. Technol. 45, 195-203 (1993). 17. D. C. Elliott, Relation of reaction time and temperature to chemical composition of pyrolysis oils. In P_vrolysis Oils from Biomass (E. J. Soltes and T. A. Milne, Eds). pp. 55-65. ACS Symposium Series 376, American Chemical Society, Washington, DC (1988). 18. A. Ison and R. J. Gorte, The adsorption of methanol and water on HZSM-5. J. Carol. 89, 150-158 (1984).
Printed in Great Britain. All rights reserved 0961-9534/93$7.00+ 0.00
CATALYTIC CONVERSION OF FAST PYROLYSIS TO HYDROCARBON FUELS OVER HZSM-5 IN A DUAL REACTOR SYSTEM R.
OIL
K. SHARMAand N. N. BAKHSHI*
Department of Chemical Engineering, University of Saskatchewan, Saskatoon, S7N OWO, Canada (Received
20 February
1993; revised received
27 July
1993; accepted
I5 September
1993)
Ah&r&-The catalytic conversion of fast pyrolysis bio-oil to hydrocarbon fuels was studied over HZSM-5 at atmospheric pressure. Experiments were conducted in a dual reactor system having two reactors in series. The temperatures in these reactors were in the range 340~400°C (first reactor) and 350-450°C (second reactor). The bio-oil was co-processed with tetralin in all the runs. The objective was to maximize the organic distillate product with a high concentration of aromatic hydrocarbons. The maximum amount of organic distillate in the effluent from the second reactor was 21 wt% of the bio-oil feed and the highest concentration of aromatic hydrocarbons was 76 wt% of the distillate. The dual reactor system was particularly beneficial when the temperature in the first reactor was low. Thus, with the first reactor at 340°C. the yields of organic distillate and aromatic hydrocarbons were 15-16 wt% and 8-11 wt% of wood, respectively, which are nearly two-fold compared to those from a single reactor system operated at 340°C (7.8 wt% and 4.8 wt%). Under the above conditions, the coke plus char yields were 25-26 wt% of wood which are up to 10 wt% lower than from the single reactor system at 340°C (29 wt%). Keywords-Pyrolysis
oil, upgrading, HZSM-5, dual reactor system.
hydrotreating catalysts are only active at high pressures in the presence of hydrogen, whereas Biomass is a renewable source of fuels and the operation with HZSM-5 is at atmospheric chemicals which has negligible sulfur, nitrogen, pressure. Also, coking with HZSM-5 is low due and metal contents. It is also CO2 neutral. to its shape-selective properties and intraThermochemical conversion of biomass to crystalline network.’ The catalyst has mediumfuels and chemicals is a two-step process. The size pores which restrict the formation of large, first step is the fast pyrolysis or high pressure highly branched molecules which lead to coke liquefaction of biomass into a bio-oil which is formation. However, unlike in small-pore zeothen converted to hydrocarbon fuels in the lites, the HZSM-5 pores permit the entry of second step of this route. With woody biomass, simple branched molecules necessary for their the yields of bio-oils from fast pyrolysis are cracking. 75-80 wt% of wood compared to 30-35 wt% Due to the poor stability of the bio-oils, low from the high pressure liquefaction. The pyroconcentrations of bio-oil vapors in the feed have lytic oils contain up to 50 wt% oxygen and are been used in most studies in the literature.4,8,9 usually unstable resulting in their polymerization In some cases, only the low boiling fraction of at relatively low temperatures. These characterthe bio-oil was used as feed.” Churin et al.’ and istics make the bio-oils unsuitable as fuels. In Sheu er ~1.~ hydrotreated the bio-oils using order to stabilize the bio-oils and to lower their Co-Mo/Al,O, catalyst by co-processing the oxygen contents, it is necessary to convert them bio-oils with tetralin. Also, in our earlier work into hydrocarbon fuels. with a high pressure bio-oil, it was observed that The conversion of bio-oils to hydrocarbon the handling and processing of the bio-oil befuels has been studied in literature using both the came relatively easy after tetralin addition. The hydrotreating catalysts such as Co-Mo/Al,O, ,I_3 use of tetralin in coal liquefaction is already well and zeolite catalysts such as HZSM-5.4-6 The known. Therefore, it was considered desirable __to study the conversion of bio-oil over HZSM-5 *To whom correspondence should be addressed. with tetralin co-processing. 1. INTRODUCTION
445
446
R. K. SHAKMA
and N. N. BAKHSHI
Sharma and Bakhshi” studied the conversion of fast pyrolysis oil to hydrocarbon fuels over HZSM-5 in a fixed-bed reactor at 340-410°C both with and without tetralin co-processing. The co-processing runs were of special interest since the concentration of aromatic hydrocarbons in organic distillate in these runs was usually 15-20 wt% higher compared to that in the absence of tetralin. The highest amount of char in the co-processing runs was 27 wt% of bio-oil at 340°C which decreased at high temperatures. Interestingly. the char fraction also became increasingly carbonaceous in nature above 370°C. This was also indicated by the amount of water in the product, which increased from 23 wt% of bio-oil at 340°C to 42 wt% at 410°C. The maximum amount of organic distillate was 19 wt% of the bio-oil at 370°C while the maximum concentration of aromatic hydrocarbons in the distillate was 83 wt% at 390°C. The amount of coke increased with temperature to about 20 wt% of the bio-oil at 390 and 410°C. The conversion of the nonvolatile fraction of the bio-oil was highest at 410°C. The above study showed that the high temperatures were useful for cracking of the non-volatile fraction of the bio-oil. However, the coking increased with temperature. At high temperatures, the amount of water in the product was high, which was detrimental to hydrocarbon yields since the hydrogen content of the bio-oil was low. It was found that 26-39 wt% of the bio-oil was lost in coke and char fractions due to its poor stability which led to polymerization of the bio-oil components. Baker and Elliott’ observed that the coke and char formation during the hydrotreatment of bio-oil over Co-MO/A&O, catalysts decreased when the upgrading was done in a dual reactor system. They also found that some of the reactive groups in the bio-oil were converted to more stable compounds by the catalytic treatment at low temperatures (250-270°C) in the first reactor. Hence, it was of interest to investigate the performance of the dual reactor system for the conversion of bio-oil over HZSM-5 catalyst. In this work, the conversion of fast pyrolysis bio-oil was studied at atmospheric pressure in a dual reactor system using HZSM-5 catalyst. The dual reactor system consisted of two catalytic reactors in series. The temperatures in these reactors were in the range 340-4OO“C (first reactor) and 350-450°C (second reactor). The bio-oil was co-processed with tetralin in
all the runs. The objective was to maximize the amount of organic distillate product with a high concentration of aromatic hydrocarbons. Also, the results were compared with those from the single reactor system which are given elsewhere. 2. EXPERIMENTAL
SECTION
2.1. Catalyst preparation and characterization The catalyst HZSM-5 was prepared according to the method described by Chen et al.” The details of the preparation and characterization of the catalyst are given elsewhere.” The particle size of the catalyst in the reactor was 5001410 pm and its surface area was 365 m’g ‘. The sodium and aluminum contents were found to be 0.075 wt% and 0.76 wt%, respectively, and the Si0,/Al,03 weight ratio was 59. The unit cell composition was reported as Na,,, H,,,Al,,,Si,,,,O,,, 14H,0.‘3 The X-ray diffraction pattern of the finished catalyst was found to be similar to that reported by Argauer and Landolt.14 Prasad et aLI3 observed that the acid site density and stability of HZSM-5 were not altered by heating the catalyst at temperatures up to 500°C. 2.2. Bio-oil The bio-oil was obtained from the Ensyn Technologies, Inc., Gloucester, Ontario, where it was produced by Rapid Thermal Processing (RTP) of maple wood.i5 The yield of the bio-oil was 74 wt% of wood. The byproducts from pyrolysis were: char, 15 wt%; CO and COz, 10 wt%; and hydrocarbon gases, 1 wt%. The bio-oil was in the form of an emulsion containing 21 wt% water including 8 wt% moisture which was present in the wood. The elemental analysis of the bio-oil including the water was: 43.6 wt% C, 8 wt% H, and 0.5 wt% N. The balance 47.9 wt% was 0. Vacuum distillation of the bio-oil at 172 Pa showed that the nonvolatile fraction (pitch) amounted to 34 wt% of the bio-oil at 200°C whereas 60 wt% of the bio-oil was low boiling (below 60°C at 172 Pa). 2.3. Equipment, analysis
run procedures
and product
The experiments were carried out at atmospheric pressure in a continuous, downflow, dual reactor system. The details of the experimental equipment, run procedures and analytical methods are described elsewhere” and are given here only briefly. The dual reactor system
447
Catalytic conversion of fast pyrolysis oil Table I. Scope of experiments for bio-oil upgrading in the dual reactor system Temperature in Reactor 1 (“C) Temperature in Reactor 2 (“C) Bio-oil/tetralin (wt) Catalyst in each reactor (g)
340-400 350-450 1.0 2.0
consisted of two catalytic reactors in series. The temperatures in these reactors were in the range 340-400°C (first reactor) and 350-450°C (second reactor). The operating conditions are summarized in Table 1. The feed consisting of bio-oil and tetralin mixture, in 1: 1 weight ratio, was kept vigorously stirred and was introduced to the first reactor by a micro-metering pump (Eldex, model A-60-S). The effluent leaving this reactor directly entered the second reactor. The products from the second reactor were separated into gas and liquid fractions. The liquid product condensed in a trap whereas the gaseous product was collected over a saturated brine solution. The rate of gas formation was monitored by the frequency of gas bubbles appearing in the brine solution. The rate decreased slowly during the run, implying a continuous deactivation of the catalyst. The products were collected over the entire run period and an average analysis was obtained. The char particles were mainly formed above the catalyst bed and were removed from the reactors at the end of the run. The quartz wool plugs placed above each catalyst bed prevented the char particles from mixing with the catalyst during their removal. Finally, the spent catalyst was recovered and regenerated with air at 600°C for 1 h in order to determine the amount of coke. No attempt was made to recover the tetralin from the products. The gaseous product was analyzed by gas chromatography using a combination of packed columns and flame ionization (FID) and thermal conductivity detectors (TCD). The liquid product was in two phases: an organic phase and water. The organic phase was separated and was found to contain 334 wt% water. It was distilled at 200°C under a vacuum of 172 Pa to separate the desired “organic distillate” product from “residue”. The organic distillate was later analyzed by gas chromatography using a capillary column and FID. The water phase was also analyzed. The identity of the peaks was confirmed by GC/MS analysis and by the use of pure compounds. Among the other products, the residue, coke, and char fractions were not analyzed.
3. RESULTS
AND DISCUSSION
3.1. Overall material balance The overall material balance from various runs is presented in Table 2. A mass balance for tetralin over the two reactors showed that the conversion of tetralin was small in all the runs below 420°C. In runs where the temperature in the second reactor was 450°C the conversion of tetralin was 8-10 wt%, which affected both the organic distillate yields and the product compositions. The effect on organic distillate yields was taken into account by subtracting the amount of tetralin which was present in the product. On the other hand, the effect on product compositions could not be accounted for, since the exact nature of products from tetralin conversion in the presence of bio-oil was not known. However, an estimate of the contribution of tetralin conversion to the product composition was made. Based on our earlier work, in which tetralin alone was passed over HZSM-5,16 the main products of tetralin conversion were C,-C, aliphatic and C,-C,, aromatic hydrocarbons. Therefore, it is estimated that. in this study, 5-6 wt% of tetralin converted to aromatic hydrocarbons in the organic distillate fraction at 450°C while 3-4 wt% was converted to the gaseous product. The product analysis also indicated that the degree of tetralin dehydrogenation to naphthalene in the presence of bio-oil was small. Due to the low reactivity of tetralin, the results were calculated on a tetralin-free basis. The weight hourly space velocity (WHSV) in Table 2 is based only on the flow rate of bio-oil entering the first reactor. The run time was set at 45 min even though the catalyst was still active at the end of the run, as indicated by a significant rate of gas production. The material balance was better than 95 wt% in most runs. It is seen in Table 2 that the amount of organic distillate, i.e. the desired product, was in the narrow range of 18-21 wt% of the biooil. These amounts are slightly higher than those from the single reactor system, which were 19 wt% or lower in the co-processing runs at 340-41O‘C.” The composition of the organic distillate, discussed later, showed that it mostly consisted of hydrocarbons and should be suitable as a fuel or as a source of hydrocarbons. The amount of gaseous product was between 10 and 14 wt% of the bio-oil, as shown in Table 2. These amounts are lower than those with the hydrotreating catalysts which were up
448
R. K. SHARMA
and N. N. BAKHSHI to 35 wt% of the bio-oil.‘.” This suggests that
the secondary cracking of the organic distillate components was relatively slow over HZSM-5. Interestingly, the amount of residue in the product was below 5 wt% indicating that a substantial cracking of the non-volatile fraction of the bio-oil occurred during the upgrading. The product also contained 22-25 wt% water including water which was present in the bio-oil. Diebold et ~1.~ pyrolyzed wood in a vortex reactor and passed the pyrolytic vapors directly over HZSM-5 at 400°C using steam as carrier gas. The product was found to contain 32 wt% water. This shows that the fraction of oxygen which was lost in water by dehydration was higher than that in this study. The dehydration reactions were undesirable since the hydrogen content of the bio-oil was low. Another advantage of using the dual reactor system was a lower char formation than in the single reactor system. As seen in Table 2, the amount of char decreased slightly with the temperatures of the two reactors and was 13-l 6 wt % of the bio-oil. The maximum char amount from the single reactor system was 27 wt% of bio-oil at 340°C. It should be noted that the feed rates in the dual reactor system were higher even though the residence times in both the systems were similar. This indicates that the amount of char was dependent on the bio-oil feed rate. The char was a solid carbonaceous material, completely insoluble in acetone, suggesting that it was characteristically different from the bio-oil which was soluble. The char was formed by polymerization of the bio-oil components. A typical elemental analysis of the char fraction at 340°C was: 60 wt% C, 5 wt% H, and 0.3 wt% N. The balance 34.7 wt% was oxygen. Diebold et al4 and Milne et aL6 reported up to 16 wt% char when the pyrolytic vapors from the vortex reactor were passed over HZSM-5 at 400°C using steam as carrier gas. Considering that the feed in our study was the “whole” bio-oil, i.e. including the non-volatile fraction, the observed char amounts were rather low, which may be attributed to improvement in the bio-oil stability in the presence of tetralin. The fraction of bio-oil which converted to coke increased with the temperature of the first reactor from 20 wt% of bio-oil at 34O’C (Table 2, Run 13) to 27 wt% at 400°C (Run 3). Interestingly, these coke amounts are much higher than those from heavy oils and tar sands bitumens which are not known to have any coking problems below 400°C. This may be due
Catalytic conversion of fast pyrolysis oil
to the higher oxygen content of the bio-oil compared to the petroleum oils. High coking may also be due to high reactivity and instability of the bio-oil. The coke amounts, in this study, were higher than those which were obtained when only the non-phenolic fraction (NPF) of the bio-oil was treated with HZSM-5.‘(’ The NPF fraction was obtained after separating the phenolic/neutral (P/N) compounds from the bio-oil. The higher coking with the bio-oil (compared to NPF) indicates that, in this study, the phenolic compounds contributed substantially to the coke formation. Elliott” also suggested that the phenolic compounds were the main coke precursors in the hydrotreatment of bio-oil over the CO-MO catalyst. The other major difference between bio-oil and its NPF fraction was in their water contents, which were 21 wt% and 50 wt%, respectively. Ison and Gorte’* reported that the presence of excess water in the feed reduced the formation of coke over HZSM-5 due to partial gasification of coke by steam. However, since the temperatures during the NPF upgrading were lower than those at which significant gasification of coke by steam occurs (> 5Oo”C), the lower coke amounts with the NPF fraction were not due to its higher water content compared to the biooil. Thus, the increased coking with the bio-oil was mainly due to the presence of phenolic compounds. 3.2. Orgunic distillate composition The composition of the organic distillate is given in Table 3. The distillate mainly consisted of hydrocarbons, phenols, acids, ketones and alcohols. The aromatic hydrocarbons were of special interest since the aromatization reactions are particularly catalyzed by HZSM-5.7,‘8 The aromatic concentration was in the range of 53-76 wt% of the organic distillate. The highest concentration was in Run 2, where the temperatures in the first and second reactors were 400 and 420°C respectively. As already mentioned, in runs where the temperature in the second reactor was 450°C 5-6 wt% of tetralin also converted to aromatic hydrocarbons in the organic distillate fraction. The maximum aromatic concentration from the single reactor system was 83 wt%,” whereas Chantal et al.* reported up to 76 wt% concentration of CrC10 hydrocarbons using HZSM-5 catalyst. It was interesting to note that the temperatures at which the maximum aromatic concentration from the upgrading of NPF fraction was obtained were also the
450
R. K. SHARMA and N. N. BAKHSHI
same as those in this study. This indicates that the acidity, acid site density, and selectivity of HZSM-5 were not affected by the difference in the water contents of the two feeds since the water content of NPF fraction was much higher compared to that of the bio-oil. Similar observations were made by Chang et al.’ where no deactivation due to the presence of water was reported in the conversion of methanol over HZSM-5 at 400-520°C even though a significant amount of water was produced during the conversion. The aliphatic hydrocarbons were the other major components of the organic distillation fraction (Table 3). Their concentration varied with the temperature of the second reactor from 16 wt% of distillate at 400°C (Run 1) to 3 wt% at 450°C (Run 3). The organic distillate also contained up to 17 wt% aldehydes and 9 wt%, each, of phenols and ketones. Although the concentration of phenols in the organic distillate was low, these compounds may be utilized effectively by reacting with methanol to form methyl aryl ethers which are known octane enhancers.4 The organic distillate was found to be mostly low boiling (below 60°C at 172 Pa). It is seen in Table 3 that the concentration of low boiling fraction increased with the temperature of the first reactor and was 30-93 wt% of the organic distillate. The increase in the low boiling fraction was due to increased cracking of the heavy components of the bio-oil at high temperatures. The composition of the organic distillate suggests that it is an excellent source of aromatic hydrocarbons. Due to its high hydrocarbon content, it is likely that the combustion characteristics of the organic distillate would be superior to those of the bio-oil. The organic distillate may also be suitable as a component of the reformulated gasoline due to the presence of oxygenated compounds such as alcohols and ethers. However, these aspects were not studied in this work. 3.3. Composition
of the gaseous product
The composition of the gaseous product is presented in Table 4. The major components were CO and CO, with concentrations of 35-42 and 29-39 wt%, respectively. The concentration of CO appeared to increase with an increase in the temperatures of the two reactors whereas the CO, concentration decreased. The above CO, CO2 concentrations are 5-10 wt% higher compared to those from the single reactor system.”
Catalytic
conversion
It was found that up to 22 wt% of the oxygen in the bio-oil was lost in CO and CO2 by decarboxylation in this study. The decarboxylation route was especially desirable for the rejection of oxygen since the hydrogen content of the biooil was low. In this way, the bio-oil hydrogen was mainly used in the formation of hydrocarbons. Mime et ~1.~ also observed that in the presence of suitable co-feeds, a significant fraction of the bio-oil oxygen was rejected in CO and CO2 using HZSM-5 catalyst. In addition to CO and CO*, the gaseous product contained 19-29 wt% C-C, hydrocarbons (Table 4). Among these, ethylene and propylene were the major ones, with concentrations of 5-8 wt% each: As previously mentioned, in runs where the temperature in the second reactor was 450°C 3-4 wt% of tetralin also converted to the gaseous product consisting of hydrocarbons. 3.4. Extent of upgrading The extent of upgrading in terms of pitch conversion and coke level was calculated as a measure of the degree of catalytic cracking of the heavy components into the desired product (see Appendix). The results are given in Table 5. It is seen that the amount of non-volatile pitch was 14-19 wt% of the total product. As already mentioned, the bio-oil contained 34 wt% pitch. This indicates that a maximum of 58 wt% of the pitch in the bio-oil was converted during upgrading. The fraction of pitch which converted to coke was high, as indicated by the coke level which was at its maximum 80 wt% of the pitch. Thus, although the non-volatile components
Table 5. Extent
of upgrading
of fast pyrolysis
oil
of the bio-oil were the major coke precursors, the volatile fraction also contributed to coking in most cases. 3.5. Comparison of the results with literature A further comparison of the results for the main product fractions was made on uniform wood basis (see Appendix). The fractions of interest were organic distillate, coke plus char, aromatic hydrocarbons, and CO and CO,. The effect of first reactor temperature on the yields of these fractions at temperatures of 400 and 450°C in the second reactor is plotted in Figs l-3. The trends at the other temperatures of the second reactor were similar. Figure 1 shows the effect of temperature on the yields of organic distillate, and coke plus char. It should be noted that the results in Fig. 1 are on wood basis whereas those in Table 2 are on bio-oil basis. As seen in the figure, the yields of organic distillate from the dual reactor system were 15-16 wt% of wood which are higher than those from the single reactor system, except when the latter is operated at 370°C.” The yield of coke plus char fraction, in this study, was at its minimum at about 25 wt% of wood when the temperatures in the first and second reactors were 340 and 400°C. As also seen in the figure, the coke plus char yields from the dual reactor system were mostly higher than those from the single reactor system, except when the latter was at 340°C. The low coke plus char yields at high temperatures in the single reactor system were due to the low char amounts.” As already mentioned, the amount of coke in the single
in the dual reactor
T, (“C)
I
400
2 3 4 5 6 7 8 9 10 II 12 13 14 15 16
400 400 380 3x0 380 370 370 360 360 360 340 340 340 340 340 *Tetrahn-free
400 420 450 400 420 450 370 ,450 450 420 400 450 420 350 370 400 basis.
Low boil.
High boil.
43.2 50.4 46.1 56.X 52.7 52.9 39.1 54.9 49.4 52.2 52.0 51.X 52.4 52.9 52.X 52.X
12.1 5.2 10.1 0.9 3.1 5.0 10.7 1.0 5.2 3.X 4.2 4.6 4.1 4.0 4.2 4.2
Pitch 14.x 16.5 14.3 15.9 16.9 16.3 19.5 15.2 16.1 16.6 15.9 16.3 17.X 17.3 16.5 17.2
system* Yield (wt% wood)
Amount (wt% of bio-oil) Run no.
451
Coke 26.5 26.X 27.2 23.4 24.7 24.5 24.5 23.5 23.7 22.4 22.7 21.x 20.0 20. I 20.2 20.4
Unidentified 3.4 1.1 1.7 3.0 2.0 1.3 6.2 5.4 5.6 5.0 5.2 5.5 5.7 5.7 6.3 5.4
Upgrading (wt% pitch)
Aromatics
CO +co,
Pitch conv.
1.9 11.5 11.2 12.3 10.1 9.0 8.8 10.3 7.5 9.6 10.2 X.4 9.1 9.1 10.6 10.8
7.6 7.0 7.1 7.6 6.6 7.9 5.7 11.2 7.1 6.5 7.2 6.9 7.1 7.0 6.7 6.9
56.5 51.5 57.9 53.2 50.3 52.0 42.6 55.3 52.6 51.2 53.2 52.0 47.6 49.1 51.5 49.4
Coke level 71.9 78.8 80.0 68.8 72.6 72.0 72.0 69. I 69.7 65.9 66.8 64.1 5X.X 59. I 59.4 60.0
452
R. K.
SHARMAand
N. N.
BAKHSHI
Legend ??Dbtilktte, .........
T =400% ..a.........................
0 Dbtiti,
T.=4!50%N-s-
. Diatilialu
Sindo moctw
s~alem
o Coke+Char T =4OO*C ,.,,,1,,,1,.,111.1,,Illldlll+llllll)ll,~~*~~,,,~~,.~.~~,,~~,,~~~~~,,,~~1 ACoke+CharT=45OX -,-,-LmL.-s-m-,X CokuXhar. Sir% moctw svstem
01 so
f
SW
420
T-7
first mac$C
Fig. 1. Relationship of temperature with organic distillate and coke plus char
reactor system above 370°C was substantially higher. Since coking reduces the run time before regeneration of the catalyst becomes necessary, the operation of the single reactor system at high temperatures was not entirely beneficial. Figure 2 shows the effect of temperature on the yield of aromatic hydrocarbons. It is seen that the aromatic yield was not strongly dependent on the temperature of the first reactor. On the other hand, with an increase in the temperature of the second reactor, the yield usually increased up to 400°C before decreasing at higher temperatures due to secondary crack-
ing of the aromatics. The runs in which the temperature of the second reactor was 400°C were of special significance. In these runs, the aromatic yield increased from 10.8 wt% of wood to a maximum of 12.3 wt% as the temperature in the first reactor increased from 340 to 380°C. Above 370°C (first reactor), the aromatic yields were similar to those from the single reactor system under similar operating conditions. When the first reactor temperature was 340°C the aromatic yields from the dual reactor system were 8-l 1 wt% of wood. These values are nearly double that of the aromatic yield (4.8 wt% of
Legend T =4OO*C .a......................... H’C
yields.
I___
Sinsle reactorsystem
Fig. 2. Relationship of temperature with aromatic hydrocarbon yield.
Catalyticconversionof fast
453
pyrolysis oil
lZm
0
?TG!MC__ ?
f
II i
T=4O@C .a.........*...............
rn~&sy&m
?? -. 0-
5
/+++:
/
\
..“..............
:..-...+z
I”
??
8 e 4.-: r!
O! s40
\.A
I
SW
I
400
4?D
TutphZnoetor:C
Fig. 3. Relationship of temperature with total yield of CO and CO,
wood) which was observed in the single reactor system operated at 340°C (Fig. 2). Another important aspect of upgrading was the CO + CO, yield which governed the ultimate yield of hydrocarbons. As mentioned before, the high yields of CO + CO2 were desirable in order to utilize the bio-oil hydrogen mainly in the formation of hydrocarbons. The effect of temperature on CO + COz yield is shown in Fig. 3. It is seen that the yield slightly increased with the temperature of the first reactor and was 6-l 1 wt% of wood. The maximum yield was obtained when the temperatures in the first and second reactors were 370” and 450°C respectively (Run 8). The above CO + COz yields are higher than those from the single reactor system, which were below 6 wt% of wood (Fig. 3). Thus, the dual reactor system was particularly advantageous when the temperature in the first reactor was low (<36O”C). In runs where the first reactor was at 34O”C, the organic distillate yields were 15-16 wt% of wood, which are
nearly two-fold compared to those from the single reactor system operated at 340°C (7.8 wt%). It is remarkable that the coke plus char yields in these runs were 25-26 wt%, which are up to 10 wt% lower than in the single reactor system at 340°C (29 wt%). At high temperature; (first reactor), the performance of the dual reactor system was similar to that of the single reactor system. However, the operation of the single reactor system at high temperatures was not particularly desirable since it resulted in high coke amounts and in a decrease in the degree of decarboxylation to CO and CO*, which were both detrimental to the formation of hydrocarbons. 3.6. Optimum yield of hydrocarbon fuels The optimum product yields from this are summarized in Table 6. The maximum of hydrocarbon fuels was 15.9 wt% of where the yield of the light hydrocarbon (C,-C,) was 3.5 wt% of wood. Under
study yield wood gases these
Table 6. Comparison of optimum product yields* Reactor system
Hydrocarbon fuels Char and coke Hydrocarbon gases CO and CO, Water Total *On wt% wood basis.
Dual (this work)
Single (Sharma and Bakhshi)”
Single (Diebold er af.)4
Single (Evans and Milne)’
15.9 44.1 3.5 17.8 18.1
14.8 38.7 6.2 15.9 24.4
12.0 20.0 4.6 26.4 37.0
9.8 25.4 8.7 37.6 24.9
100.0
100.0
100.0
100.0
454
R.K. SHARMA and N.N.
conditions, the coke plus char yield was 45 wt% of wood. This fraction may be used as a process fuel, Alternatively, the char fraction may be liquefied indirectly via gasification and methanol synthesis. The above yield of hydrocarbon fuels was somewhat higher compared to the yield from the single reactor system which was 14.6 wt%” whereas the yield of light hydrocarbon gases was lower in this study. Another significant advantage of using the dual reactor system was that the CO + CO? yield was higher and the yield of water was lower compared to those from the single reactor system (Table 6). The results were also compared with the other studies. The optimum yield of hydrocarbon fuels, in this work, was also found to be higher compared to the yields reported by Diebold et ~1.~ and Evans and Milne,’ which were 12 wt% and 10 wt% of wood respectively (Table 6). Besides, the latter yields were obtained at low concentrations of the bio-oil vapors in the feed. The fraction of oxygen which was rejected in water was relatively low in this study, which was especially desirable. Although the yields of coke plus char, reported by Diebold et al4 and Evans and Milne,’ were low, they were associated with increased decarboxylation and dehydration of the coke and char fraction to CO, CO, and water instead of its conversion to hydrocarbon fuels (Table 6). This study showed that the dual reactor system was useful for converting the bio-oil into hydrocarbon fuels using HZSM-5 as catalyst. The operation was at atmospheric pressure without the cost of producing high pressure hydrogen as in conventional hydrotreatment. Also, the formation of gasoline range hydrocarbons was favored by HZSM-5 since its medium-size pores exclude the formation of bulky branched molecules with over 9 or 10 carbon atoms. Although the catalyst deactivated rapidly, it was found that it was easily regenerated by oxidation with air at 600°C. The results indicated that a significant fraction of the bio-oil oxygen was lost by decarboxylation into CO and CO,. Milne et a1.6 reported that the decarboxylation reactions over HZSM-5 were enhanced in the presence of suitable co-feeds. Thus, it seems that in this study the decarboxylation reactions were enhanced by the tetralin co-processing. In the literature, tetralin has normally been used as a hydrogen donor solvent at high hydrogen pressures so that any naphthalene which was formed by tetralin dehydrogenation
BAKHSHI
was hydrogenated back to tetralin by the hydrogen. 2,3However, in this study, the amount of tetralin dehydrogenated was small compared to the amount of tetralin which was present in the feed, so that the role of tetralin appeared to be mainly as a diluent. In runs where the temperature in the second reactor was 45OC, the contribution of tetralin conversion to the product composition was estimated based on our earlier work in which tetralin alone was passed over HZSM-5. 4. CONCLUSIONS
The dual reactor system was particularly beneficial when the temperature in the first reactor was low, as indicated by the following conclusions: (1) When the temperature in the first reactor was 340°C the yields of organic distillate from the dual reactor system were 15-16 wt% of wood, which are nearly two-fold compared to that from the single reactor system operated at 340°C (7.8 wt%). (2) In runs where the first reactor temperature was 340°C the aromatic hydrocarbon yields from the dual reactor system were 8-l 1 wt% of wood, which are also nearly two-fold compared to the yield from the single reactor system at the same temperature (4.8 wt%). (3) At low temperatures in the first reactor (34O”Q the overall coke plus char amounts were 25-26 wt% of wood, which are up to 10 wt% lower than in the single reactor system at 340°C (29 wt%). (4) The extent of decarboxylation was generally higher in the dual reactor system compared to the single reactor system. This was especially desirable since the hydrogen content of the bio-oil was low. APPENDIX
(A) Pitch = amount residue.
of char + amount
of
% Coke level = amount of coke in the product x 100 amount of pitch in the bio-oil % Yield of a fraction on wood basis = % yield of that fraction on bio-oil basis x yield of bio-oil on wood basis. (B) Average molecular weight of the gas product, M 9’ = X W,M; ‘, where M, and W, are the molecular weight and weight fraction,
Catalytic conversion of fast pyrolysis oil
respectively, of component i in the gaseous product. As an example, MBfor Run 1 was 35.1. Acknowledgements-The authors are grateful to the Bioenergy Development Program of Energy, Mines and Resources (EMR), Canada, for the financial support of this research.
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