Ind. Eng. Chem. Process
Des. Dev. 1085, 24,
and furnace oil is a t least as high in the two-stage cases (with metals-contaminated pretreatment substrate) as it is for the direct cracking cases. In addition, the two-stage process provides a significant reduction in coke yield and a corresponding increase in total liquid recovery. Conclusions
Petroleum residua from three different crude sources (West Texas, North Slope, and Kuwait crudes) have been upgraded by using a two-stage approach. The first stage involved pretreatment of the residua over a low surface area alumina substrate. At the conditions tested, 88-9670 metals removal, 27-33% sulfur removal, and 61-78% carbon residue removal were achieved. The effect of reaction temperature on pretreatment yields and contaminant removal effectiveness is consistent with typical thermal cracking operations. The second-stage experiments involved conventional catalytic cracking of the pretreated oils. Among the three residua tested, the North Slope resid proved to be the least susceptible to pretreatment and cracking. The effects of nickel and vanadium (deposited on substrate or catalyst) on yields during pretreatment are different from those observed during conventional cracking. The MAT unit is a useful tool for determining yield distributions from residual oil pretreatment and subsequent catalytic cracking of pretreated oils. The two-stage process is an excellent way to maximize the yield of middle distillate boiling range (450-650 O F )
1275-1281
1275
material from the processing of residual oils. Registry No. Ni, 7440-02-0; V, 7440-62-2. L i t e r a t u r e Cited Barthollc, D. 6.; Haseltine, R. P. Oil Gas J. 1981, 79 (22), 89-92. Busch, L. E.; Hetlinger, W. P.; Krock, R. P. OilGas J. 1984, 82(51), 54-56. Campagna, R. J.; Krishna, A. S.; Yanik, S. J. Oil Gas J. 1983, 81 (44), 128-134. Dale, G. H.; McKay, D. L. Hydrocarbon Process, 1977, 9, 97-102. Dean, R. R.; Mauleon, J. L.; Letzsch, W. S.; Legendre, M. Paper presented at the 1983 NPRA Annual Meeting, San Francisco, CA, March 20-22, 1983, AM-83-42. Henderson, D. S.; Ciapetta, F. G. Oil Gas J. 1967, 65 (42), 88-93. Itoh, M.; Suzuki, T.; Tsujlmoto, Y.; Takegami, Y.; Watanabe. Y. Ind. Eng. Chem. Process Des. Dev. 1984, 23, 622-625. Krishna, A. S.;Campagna, R . J.; Engllsh, A. R.; Kowalczyk, D. C. Paper presented at the 1984 NPRA Annual Meeting, San Antonio, TX, March 25-27, 1984, AM-84-51, Mills, G. A.; Boedeker, E. R.; Oblad, A. G. J. Am. Chem. SOC. 1950, 72, 1554. Mitchell, B. R. Ind. Eng. Chem. Prod. Res. Dev. 1980, 19, 209-213. Murphy, J. R.; Treese, S.A. OilGas J. 1979, 77(26), 135-142. Murphy, J. R. Oil Gas J. 1980, 78 (354, 108-110. Rosynek, M. P.; Shipman, G. F.; Yan, T. Y. US. Patent 3 983030, Sept. 28, 1976. Rush, J. 6. Chem. Eng. Prog. 1981, 77(12), 29-32. Teichman, D. P.; Bridge, A. G.; Reed, E. M. Hydrocarbon Process. 1982. 5 , 105-109. Vermillion, W. L.; Gearhart, J. A. Hydrocarbon Process. 1983, 9 , 89-91. Yan, T. Y. Ind. Eng. Chem. Process Des. Dev. 1984, 23, 415-419. Yanik, S. J.; Frayer, T.; Huling, G. P.; Somers, A. E. OilGas J . 1977, 75(20), 139-1 45. Zandona, 0. J.; Busch, L. E.; Hettiger, W. P. Paper presented at the 1982 NPRA Annual Meeting, San Antonio, TX, March 21-23, 1982, AM-82-61.
Received for review August 24, 1984 Revised manuscript receicied January 22, 1985
Coke Formation in the Pyrolysis of n-Hexane Mrltinjoy Pramanlk and Deepak Kunzru’ Department of Chemical Engineering, Indian Institute of Technology, Kanpur, India
Pyrolysis of n-hexane was studied in a jet-stirred mixed reactor in the temperature range of 993-1083 The kinetics of the overall pyrolysis, as well as the coke formation, was investigated. The overall pyrolysis was essentially first order, and the major products were ethylene, methane, propylene, and hydrogen. The effect of temperature, conversion, and molar ratio of steam to hexane on the rates of coke deposition was measured by periodically weighing a small cylinder suspended into the center of the reactor. Coking rates increased with increasing temperature, conversion, and initial partial pressure of n-hexane. The experimental coking rates could be adequately fitted by a model in which ethylene was the coke-forming species.
Pyrolysis of hydrocarbons such as ethane, propane, and naphtha is an important process for the production of olefins. Pyrolysis of any hydrocarbon is always accompanied by coke formation which deposits on the inner walls of the cracking coil. This accumulation of coke with time gradually increases the pressure drop across the tubular reactor and reduces the overall heat-transfer coefficient across the tube wall due to the extra thermal resistance of the deposited coke layer. The normal industrial practice is to periodically shut down the reactor and burn the deposited coke with a mixture of steam and air. Depending upon the feedstock, decoking is done anywhere after 30-60 days. Coke is a complex of carbon and hydrogen which can be formed either from the reactants or products or from both. Coke deposition depends upon several factors, such as the nature of the feedstock, hydrocarbon partial pressure, temperature of the reactor, conversion, presence of 0198-4305/85/1124-1275$01.50/0
sulfur compounds, and the nature of the reactor surface. Studies of the coking phenomena during pyrolysis of pure hydrocarbons should be helpful in understanding the more complex interactions occurring during the pyrolysis of natural feedstocks. However, very meager information is available on the coking kinetics of pure hydrocarbons. Hirt and Palmer (1963) studied the coke formation from methane pyrolysis between 1160 and 1370 K and obtained an activation energy of 432 kJ/mol. Johnson and Anderson (1962) investigated the coke formation from acetylene at temperatures ranging from 770 to 1270 K and observed both the coke and polymer as the products. They suggested that the coke and polymer were formed by two competing parallel reactions in the gas phase. Albright and McConnell(l978) reported that the rate of coke formation in ethane pyrolysis was significantly affected by the material of construction of the pyrolysis tube and also on the pretreatment given to the reactor. Albright and Yu (1978) 0 1985 American
Chemical Society
1276
Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 4, 1985 Closed f o r main pyrolysis r u n ,Closed
for coking runs
I c e cooled water
n-HEX4NE
S t e a m vaporiser
’I
Cylinder
TC2
Figure 1. Schematic diagram of the experimental setup.
studied the effect of the material of construction of the reactor on the rates of coke formation in the pyrolysis of acetylene, butadiene, ethylene, and propylene. Kinney and Del Bel (1954) studied the pyrolytic behavior of unsubstituted aromatic hydrocarbons and suggested that coke is formed from a stepwise condensation of the aromatic rings. Shah et al. (1976) investigated the coke formation in the pyrolysis of n-octane between 1023 and 1073 K. The coking rate was found to be high initially and then reduced to an asymptotic value. All the above-mentioned studies were conducted in tubular flow reactors, and the total coke formed was measured by either weighing the reactor before and after the experiment or by burning the coke and measuring the effluent gaseous amounts. Since coke is not deposited uniformly along the reactor, this results in only average coking rates and precludes a kinetic analysis. To overcome this problem, Sundaram and Froment (1979) developed a jet-stirred completely mixed reactor which allowed the kinetics of the main pyrolysis reaction and the kinetics of coking to be determined simultaneously. For propane pyrolysis, the coking rates were best represented by a model which assumed that coke was formed from propylene by a first-order reaction, whereas the coke formation in ethane pyrolysis was modeled as a first-order reaction from the C4+hydrocarbons (Sundaram et al., 1981). The objective of this study was to investigate and model the effect of temperature, conversion, and inlet hydrocarbon partial pressure on the rates of coke deposition during the pyrolysis of n-hexane. n-Hexane was chosen as the reactant because, compared to natural liquid feedstocks such as naphtha and gas oil, the product distribution is much simpler and well established. Moreover, no quantitative treatment of coke deposition in n-hexane pyrolysis has been reported. Only a few investigations have been published on the homogeneous gas-phase pyrolysis of n-hexane. It is generally accepted that the overall decomposition of n-hexane is first order (Frey and Hepp, 1933; Illes et al., 1965; Ebert et al., 1983). Zdonik et al. (1967) calculated the distribution of primary products from pyrolysis of n-hexane in accordance with the Rice mechanism. Murata et al. (1973) reported a simulation model for the prediction of the initial product distribution from pyrolysis of normal paraffinic hydrocarbons and found it to be in good agreement with the experimental results for a series of normal paraffins from n-butane to n-hexadecane. Their model was extended to a higher conversion range by Murata and Saito (1975). With increasing conversion, the selectivities (moles
of productf100 moles of feed decomposed) of methane, ethylene, and aromatics increased while that of propylene and butadiene went through a maximum. None of the above investigations measured the amount or rate of coke formed during pyrolysis.
Experimental Section The experiments were conducted in a jet-stirred reactor which had provision for measurement of kinetics of the main reaction as well as the kinetics of coking. A schematic diagram of the experimental setup is shown in Figure 1. Steam, which was used as an inert, was generated in a vaporizer and mixed with liquid n-hexane before the preheater. Continuous charging of hexane and water to the reactor was done by a duplex reciprocating pump which could be adjusted for different flow rates. The preheated and mixed vapors of hexane and steam entered the preheater-reactor assembly at approximately 773 K. The preheater-reactor assembly was constructed from a 50 mm i.d. 304 SS tubing and had provisions for suspending a small cylinder through a central tube which also served as an outlet for the reactor effluents. The reactor volume was approximately 103 cm3. The plate separating the preheater and reactor contained 44 1.5-mm holes distributed uniformly. The preheated reactants were issued as jets from these holes and provided the mixing in the reactor. The preheater-reactor was heated in a manually controlled electric furnace, and the temperature in the reactor was measured by a chromel-alumel thermocouple placed in a thermowell. The temperature variation along the length of the preheater-reactor assembly was also measured by inserting a thermocouple from the outlet tube. In all the runs, the axial temperature gradient in the main reactor was negligible. Due to nonavailability of a continuous recording electrobalance, the apparatus had to be modified depending on whether the product distribution or coke deposition was .being measured. For measuring the coke deposition, a small cylinder of the same material as the reactor was suspended by means of a fine platinum wire into the central portion of the reactor. The amount of coke formed on the cylinder was measured by removing the cylinder through the central tube. The cylinder was weighed by suspending on an electrobalance and resuspended at the center of the reactor. This was done at periodic intervals during the course of a run. To prevent oxidation of the deposited coke, prior to each weighing, the cylinder was cooled by raising it from the center of the reactor into the relatively cooler portion of the outlet tube. Attempts to measure the
Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 4, 1985 1277 1.01
I
I
1
Table I. Reaction Order for Hexane Cracking (Temperature = 1023 K)
I
6=16
(cAo)Z, 7,
s
1.08 0.68 0.51 0.35
(C,,),,
g-mol L-'
1.703 1.703 1.703 1.703
X X X X
g-mol L-' 7.012 X IO-' 7.012 X 7.012 X 7.012 X
(XA)l
0.792 0.740 0.684 0.590
(XA)z 0.810 0.737 0.678 0.595
n 1.02 1.09
1.10 1.05
Temp I K ) 0
1083 1053
A
1023
0
993
0
0
LOO
800 1200 V/FAO,(lit. mole-' s)
1600
Figure 2. Variation of n-hexane conversion with v/FA,.
coke continuously by suspending the cylinder inside the reactor from a single-pan balance were unsuccessful due to vibrational problems. In these experiments, no condensers were used, and the hot reador effluent was directly vented. For measuring the product distribution, an identical run was repeated without the suspended cylinder but with the reactor effluent quenched in two icewatercooled condensers placed in series. The noncondensables were passed through a sampling valve and a gas flow meter and then vented. The gas and liquid products formed in the pyrolysis were analyzed on a gas chromatograph. At the completion of the run, the reactor was flushed with steam, the coke deposited on the cylinder, and reactor burned by passing heated air.
Results and Discussion Kinetics of n -Hexane To determine the kinetics of hexane cracking, experiments were conducted a t atmospheric pressure in the temperature range of 993-1083 K. Steam was used as an inert diluent and the molar ratio of steam to hexane, 6, was maintained a t either 16 or 6. The reactor space time was varied between 0.25 and 1.2 s. Typical conversion vs. v/FA, curves are shown in Figure 2. As expected, conversion increases with increasing V/FA, and temperature. These data were used to obtain the kinetics of the overall pyrolysis reaction. The material balance of n-hexane in a completely mixed reactor may be written as
FA^ =
FA,(^ - XA)+ (-rA)V
Space time, s
Figure 3. Determination of reaction rate constant for n-hexane pyrolysis at different temperatures.
the pyrolysis a t different initial hexane partial pressures. From eq 2 and 3, we obtain
or
If data are taken at the same space time and temperature but different initial reactant concentration, eq 6 yields
(1)
or
The rate equation for a nth order reaction can be expressed as
where cA, the coefficient of volume expansion, can be written as (4)
The reaction order, n, can be determined by conducting
The conversion of hexane was measured a t two different initial partial pressures and eq 7 used to evaluate n a t different conversion levels. Values of n thus obtained are shown in Table I. As can be seen from this table, the pyrolysis of n-hexane can be assumed to be first order. For a first-order reaction, eq 6 simplifies to
The plot of the right-hand side of eq 8 vs. space time is shown in Figure 3 for the different temperatures. The plots are linear, and the slope of these lines yields the rate constants. For'a first-order reaction, for the same temperature, the plots at two different partial pressures should be identical; however, for n-hexane pyrolysis, some devi-
1278
Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 4, 1985
i
81
-
0
0
30
60
90
120
1 9
Time, min
Figure 6. Variation of rate of coke formation with run time. 88
I
I
I
9.L
9.6
9.8
1
1
9.0
9.2
I
10.0
1 x 104, K-' 1 Figure 4. Arrhenius plot for the pyrolysis of n-hexane.
$140-
6 =6
5
Temp = 993 H
50
60
70
80
90
Convarsion, par cent
Figure 5. Product distribution vs. conversion for the pyrolysis of n-hexane at 993 K.
ation was shown, indicating that the order is not exactly unity, but for our calculations the pyrolysis has been assumed to be first order. The Arrhenius plot for the first-order rate constant is shown in Figure 4. The frequency factor and activation energy were calculated from a least-squares analysis, and values of klo and E,, together with the 95% confidence limits, were (1.56 f 0.02) X 1O1O s-l and 187.5 f 0.8 kJ/mol, respectively. Product Distribution. The major products obtained during n-hexane pyrolysis were ethylene, methane, hydrogen, propylene, and 1-butene, together with small amounts of ethane. Products higher than C4 were not analyzed. The variation of the product yield with conversion at 993 K and a molar steam-to-hexane ratio of 6 is shown in Figure 5. A t a fixed partial pressure, with increasing conversion yields of ethylene, methane, and ethane increased, hydrogen and 1-butene yields decreased, whereas propylene yield passed through a maximum. For a fixed conversion and inlet hexane concentration, the product distribution was nearly independent of tempera-
ture. At a fixed conversion, on reducing the inlet partial pressure of hexane, the ethylene yield increased, whereas the propylene yield decreased. There was no significant effect of partial pressure on the yields of the other products. These product yields compare favorably with the experimental results reported in the literature. For example, a t 700 "C, 6 = 7, and a conversion of 80%, Murata and Saito (1975) reported the ethylene, methane, propylene, hydrogen, and 1-C4yields as 112,85,55.5, 35.5, and 10 mol/100 mol of hexane decomposed, respectively. In this study a t 720 "C, 6 = 6, and a conversion of 80%, the yields of the above products are 115, 88.5, 55, 52, and 8 mol, respectively. Kinetics of Coke Formation. The effect of temperature, conversion, and inlet partial pressure of hexane on the rates of coke formation was investigated. For most of the runs, the molar ratio of steam to hexane was maintained at 6:1, but to study the effect of inlet partial pressure of hexane on the coking rate, some runs were taken at a molar ratio of 4:l. Each run was continued until two successive coking rates were identical which was achieved in approximately 1'/2-2 h. Typical plots of coking rate vs. time at two different temperatures but identical conversion and partial pressure are shown in Figure 6. Similar plots were obtained for all the runs. As can be seen from this figure, the coking rate is initially high and then tends to an asymptotic value within 90-120 min. The high coking rate in the beginning of the experiment is presumably due to the catalytic effect of the metal surface. This catalytic effect progressively reduces as coke deposits on the surface and the rate of coke formation decreases to a constant value corresponding to the metal surface being completely covered with a coke layer. A similar phenomenon has been observed in the pyrolysis of ethane (Sundaram et al., 1981), propane (Sundaram and Froment, 1979), n-octane (Shah et al., 1976), and naphtha (Kumar and Kunzru, 1985). To use the coking rates obtained in this study to predict the rates in another reactor, the effect of the surface area on the coking rate has to be known. The surface area of the suspended cylinders was changed by a factor of 2, and in all cases, the rate of coking was directly proportional to the surface area; i.e., the specific rate of coking, expressed as g coke/[h (m2of metal area)], was the same at otherwise identical conditions. The rates of coke formation varied with reactor temperature, conversion, and initial partial pressure of hexane, and the variation is shown in Figure 7. Since coke is formed from secondary reactions and once formed does not react further, rc increases with conversion and tem-
Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 4, 1985
Table 111. Parameter Estimates for Linear freq fact, g coke/ [mol i/(L m2 rate coeff h)l (1.56 f 0.02) X kl k2 (1.69 f 0.34) X 1O'O k3 (4.17 f 1.37) X 10" k4 (9.91 f 2.79) X 10" k5 (1.23 f 1.28) X 10" k6 -(1.59 f 9.84) X 10" k7 (4.88 f 1.46) X 10" k8 -(1.23 f 0.61) X 10" k9 -(3.84 f 132.7) X lo9 k1o (2.85 f 6.12) X 10"
L
I L
c h I
E cn
Y
1279
Coking Models activat energy, kcal/mol 44.6 f 0.2 30.3 f 6.0 35.0 f 9.8 33.3 f 8.3 34.1 f 11.6 31.3 f 384.7 30.3 f 2.0 32.4 f 28.1 30.3 f 17.23 35.4 f 20.6
Homogeneous; units, s-*.
5
I
I
I
06
0.7
08
0.9
XA
Figure 7. Variation of the asymptotic coking rate with conversion. Table 11. Various Linear Coking Models I C6H14 % products C2H45 coke I1 C&l4 % products C3H65 coke I11 c~H14k products 1-C4Hs4 coke IV C6H14 k products C2H4h coke C3H6 b coke V C6H14 % products CzH4k coke 1-C4Hsh coke VI C6H14 % products C3Hs h coke l-C4H8% coke
perature. The monotonous increase of rc with conversion at constant 6 and temperature also shows that coking is not mass transfer controlled. The mass-transfer coefficient increases and conversion decreases with decreasing T ; therefore, if mass transfer was controlling the rate of coke formation, rc would either decrease or show a maximum with increasing conversion. Moreover, the activation energy for rc (discussed later) is much larger than would be obtained if the coking was mass transfer controlled. The rate of coke formation decreases with a decrease in the inlet hexane partial pressure. On decreasing the inlet hexane concentration, the concentration of coke precursors is decreased, thus reducing the rate of coke formation. At these temperatures, the carbon-steam reaction is not expected to be significant (Biba et al., 1978) and would not contribute to the reduction in rc with increasing partial pressure of steam. Modeling of Coke Kinetics. Coking is a complex phenomenon, and the mechanism probably involves the formation of coke precursors by a homogeneous gas-phase reaction involving the reactants and/or products, transfer of these coke precursors to the walls of the reactor, and a further heterogeneous reaction of these precursors with the walls to form coke. At high temperatures, the chemical reaction rates are large and the rate of coke formation is controlled by the mass transfer of the coke precursors to the reactor walls, whereas a t low temperatures the coke formation is kinetically controlled (Fernandez-Baujin and Solomon, 1975). Due to the above complexities, only
simplified models involving the various products were used to explain the data. The various models tried were that coke can be formed from either ethylene, propylene, 1-butene, ethylene and propylene, ethylene, and 1-butene or propylene and 1butene. For these models, it was assumed that each reaction was first order. Such first-order coking models have been found to satisfactorily represent the coking rates in ethane (Sundaram et al., 1981) and propane pyrolysis (Sundaram and Froment, 1979). The various first-order coking models investigated are shown in Table 11. For each model, an expression for rc can then be written. For instance, for model I,
or
where yC2& is the selectivity of CzH4(moles produced/mole of hexane cracked) and a is the total number of moles produced/mole of hexane cracked. cy was determined experimentally and generally varied from 3.3 to 3.6. Since the experiments were conducted in a mixed reactor, the experimental values of rc and the product concentration were available. The model parameters were estimated by minimizing the residual sum of squares (RSS) of the deviation between the calculated and experimental rates of coke formation by the appropriate choice of the frequency factor and the activation energy. The regression was performed by using combined Gauss-Newton and quasiNewton methods. Due to the strong correlation between the frequency factor and the activation energy, a reparameterization was necessary. For the reparameterization, the rate constants were expressed as
ki = ki, exp(-Ei/RT,)
[ :(f i ) ]
exp - -
--
-
(11)
where T, is the average temperature for all the runs. The estimated parameters together with their 95% confidence intervals are shown in Table 111. Models IV, V, and VI were discarded due to nonsignificantly determined parameters. The parameters of the other three models were positive and statistically significant. However, preference was given to model I because of the lower RSS. Similar results have been reported for coke formation during ethane and propane pyrolysis (Sundaram et al., 1981; Sundaram and Froment, 1979). For ethane pyrolysis, the model parameters were positive and significant for first-order models with ethylene, propylene, or C,+ as the
1280
Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 4, 1985
A
I
./I
1.6 .c
Table V. Parameter Estimates for Models freq fact, g coke/[(mol i/L)"' m2 rate coeff h1 kl (1.56 f 0.02) X 10'O" k,, (7.54 f 0.8) X 10" kn (7.81 f 0.74) X 10'O k13 (0.29 f 1.34) X loz1
Nonlinear Coking activat energy, kcalimol 44.6 f 0.2 23.5 f 3.5 33.2 f 11.7 26.6 f 7.6
react order 1.0 2.4 f 0.24 0.88 i 0.36 3.42 i 1.8
Homogeneous; units, s-l,
OLY0, 1 ~
0
0
0.4
08
12
Experimental rc , m:.cm-*.
,
16
20
hr-I
Figure 8. Comparison between calculated and experimental rates of coke formation (linear model).
i
O/
:
12t
Table IV. Various Nonlinear Coking Models VI1 CsH14 k products C2H, h . coke ~ VI11 C6HI4k products C3H, h coke IX
C,H14 4 products l-C4H8h+coke
coke precursors. However, the model involving C4+ as the coke precursor was preferred due to the lower RSS. The comparison between the experimental and calculated rc using model I is shown in Figure 8. The average deviation was 8%, and the maximum error between the predicted and experimental coking rate was 16%. To improve the model prediction, the restriction of first-order dependency was removed, and models postulating that coke formed from CzH4,C3H6,or 1-C4 were again tested. The nonlinear models tested are shown in Table IV. For the nonlinear model VII, r, can be written as
0 0
I
1.6
20
Experimentol rc , mg cm-2 hr-1
Figure 9. Comparison between calculated and experimental rates of coke formation (nonlinear model).
and cyclization of the resulting radicals, followed by hydrogen abstraction from the cyclic species by metathetical reactions, such as
Herriott et al. (1972), in their pyrolysis model, assumed the coking reaction as CPH,
The model parameters were estimated as before with the only difference that the objective function was minimized by adjusting all three parameter, viz. kio, E , and m. The estimated parameters together with their 95% confidence intervals are shown in Table V. Model IX was rejected due to the nonsignificantly determined frequency factor. Preference was given to model VI1 due to the significantly lower RSS. Thus, the best model considering both the linear and the nonlinear model is model VII. The comparison between the experimental and calculated rCusing model VI1 is shown in Figure 9. Comparing Figures 8 and 9, we can see that the model predictions using model VI1 are much better than model I. The average percent deviation using model VI1 was 4 % and the maximum deviation 9%. This model is not a mechanistic picture of the complex coking reaction and should not be taken to suggest that coke is only formed from ethylene, although ethylene does take part in some of the coke forming reactions. As Ebert et al. (1983) have shown, most of the aromatization reaction occurs via addition of allyl-type radicals to alkenes
1.2
0.8
0.L
-
2C
+ 2Hrj
where it may involve acetylene as an intermediate. Snow and Schult (1957) in their design of an ethane pyrolysis reactor assumed the coke to be formed from CzH4by a second-order reaction with an activation energy of 24 500 cal/g-mol compared to the value of 23 500 cal/gmol obtained in this study. The experimental activation energy for coke formation during pyrolysis of n-paraffins varies over a wide range. For example, activation energies of 28.25, 73.58,50.78, and 51.0 kcal/g-mol have been reported for coke formation from pyrolysis of ethane (Sundaram et al., 1981), propane (Sundaram and Froment, 1979), naphtha (Kumar and Kunzru, 1985), and acetylene (Oxley et al., 1961), respectively.
Conclusions The results of this study show that the overall hexane pyrolysis is essentially first-order and the rate of coke formation during hexane pyrolysis can be satisfactorily represented by the following model C6H,, products
--
C2H4
coke
with the rate expression for coking
Ind. Eng. Chem. Process Des. Dev. 1985, 2 4 , 1281-1287
rc = 7.54 x
1012 exp(
-)cc2H:" -23 500 RT
It is expected that such simplified models might also be used to model the coking rates during the pyrolysis of other hydrocarbons. Nomenclature CA = concentration of n-hexane, mol/L CAo = inlet concentration of n-hexane, mol/L CclH4= concentration of ethylene, mol/L Ei = activation energy of the ith reaction, cal/g-mol or kJ/ g-mol F A o = inlet molar flow rate of hexane, mol/s ki = reaction rate coefficient for ith reaction kio = frequency factor for the ith reaction rate coefficient, s-l or g coke/[(mol i/L)" m2 h] m = reaction order for the coking reaction n = reaction order for hexane pyrolysis PT = total pressure, atm R, R, = gas constant, cal/(g-mol K) or L atm/(mol K) -rA = rate of reaction for hexane pyrol sis, mol/(L s) rc = rate of coke formation, g coke/(m 7h) T = temperature, K T,,, = average temperature for all the runs, K V = reactor volume, L XA = conversion of hexane yCzH4 = ethylene selectivity, mol produced/mol of hexane cracked
7
1281
= space time, s
Registry No. CBH14,110-54-3; C2H4,74-85-1; CHI, 74-82-8; C3H8, 115-07-1; Hz, 1333-74-0; C, 7440-44-0.
Literature Cited Aibright, L. F.; McConneil, C. F. "Abstracts of Papers", 175th National Meeting of the American Chemical Society, Annaheim, CA, March 1978;American Chemical Society: Washington, DC, 1978. Albright, L. F.; Yu, Y. C. "Abstracts of Papers", 175th National Meeting of the American Chemical Society, Annaheim, CA, March 1978; American Chemical Society: Washington, DC, 1978. Biba, V.; Macak, J.; Klose, E.; Maiecha, J. Ind. Eng. Chem. Process Des. Dev. 1078, 17,92. Ebert, K. H.; Ederer, H. J.; Isbarn, G. Int. J . Chem. Kinet. 1083, 15, 475. Fernandez-Baujin, J. M.; Solomon, S. M. Paper presented at the First Chemical Congress of the North American Continent, Mexico City, 1975. Frey, F. E.; Hepp, H. J. Ind. Eng. Chem. 1033, 25, 441. Herriott, G. E., Eckert, K. E.; Aibright, L. F. AIChE J . 1072, 18, 81. Hirt, T. J.; Palmer, H. 8.Carbon 1063, 1 , 65. Iiles, V.; Piaszkats, I.; Szepesy, L. Conf. Chem. Proc, Petro/. Natural Gas 1985, 1. Johnson, G. L.; Anderson, R. C. Proc. Carbon Conf. 5th 1082, 1 , 395. Kinney, C. R.; Del Bel, E. Ind. Eng. Chem. 1054, 4 6 , 548. Kumar, P.; Kunzru, D., Can. J . Chem. Eng., in press. Murata, M.; Saito, S., Amano, A., Macde, S. J. Chem. Eng. Jpn. 1073, 6 ,
252. Murata, M.; Saito, S. J . Chem. Eng. Jpn. 1075, 8 , 39. Oxley, J. 0.; Secrest, A. C.; Veigel, N. D.; Biocher, J..M. AIChE J . 1061, 7 ,
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Greek Letters a = moles of product formed per mole of hexane decomposed 6 = inlet molar ratio of steam to hexane eA = relative volume change during pyrolysis
Received for review June 26, 1984 Revised manuscript received February 20, 1985 Accepted March 5, 1985
Gasification of Sawdust in an Air-Blown Cyclone Gasifier James W. Couslns' and Wllilam H. Robinson Physics and Engineering Laboratory, Private Bag, Lower Hutt, New Zealand
An air-blown cyclone gasifier was designed to gasify sawdust at rates of 30-180 kg/h, giving cold gas outputs of 0.1-0.6 MW. A number of experiments were conducted with the air/wood ratio being varied. As the ratio was increased from 1.2 to 2.5 scm/kg of ovendry wood, the calorific value of the gas decreased from 5 to 2.5 Ml/scm and the cold gas energy yleld from 60% to 45% of the input wood energy. The gas consisted of N,, CO, H,, CO,, CH, and small amounts of higher hydrocarbons. The lowest operating airlwood ratio of 1.3 scm/kg was set by the accumulation of charcoal in the gasifier. The recommended input of 1.7 scm/kg gave a cold gas output of 0.55 MW for a wood input of 180 kg/h. The cyclone gasifier was simple to operate, reliable, the responded rapidly to changes in demand for gas.
The most common way of converting wood to wood gas is by partial combustion in air. Although it is a century-old technique, it is not a well-developed one. For example, most designs of gasifiers contain a slowly moving bed of charcoal or fuel through which gases and vapors have to flow, and so the performance is highly dependent on the porosity of the bed (Palmer et al., 1982). If a suitable level of porosity and a steady flow of solids through the reactor are to be achieved simultaneously, much of the fuel has to be in the form of small blocks ranging in size from 20 to 200 mm. Not all species of wood are useable, and
neither are sawdust, hogfuel, small wood chips, and many pelleted products. Small particle materials like sawdust make up a large part of the residue from a modern sawmill. Wood-free bark and slabwood can usually be sold, the bark as a horticultural potting mix and the slabwood for pulp chips, leaving the sawdust as the waste most readily available for fuel. Sawdust can be burned for industrial heating, but direct combustion is not always the best approach. In the retrofitting of oil- or gas-fired boilers, for example, it is often
0 196-430518511124- 128 1$0 1.5010 Published 1985 by the American Chemical Society
