222
Ind. Eng. Chem. Process Des. Dev. 1982, 27, 222-231
Qboja, 0. Ph.D. Thesis, University of Surrey. Guiiord, England, 1975. Sater, V. E.; Levensplel, 0. Ind. Eng. Chem. Fundem. 1988, 5 , 86. Thomas, W. J. A.B.C.M. Distllletkm Symposium, London, England, 1964. Thomas, W. J.; Campbell, M. Trans. Inst. U".Eng. 1987, 45, 53. Thomas, W. J.; Haq,M. A. I n d . Eng. Chem. Process Des. 'Dev. 1978, 75.
509.
Thomas, W. J.; OgboJa,0. Ind. Eng. Chem. Process D e s . D e v . 1970. 77, 429.
Received for review January 5 , 1981 Accepted September 18, 1981
Fischer-Tropsch Synthesis in the Slurry Phase on Mn/Fe Catalysts Wolf-Dieter Deckwer" and Yalcln Serpemen Institut fur Technische Chemie, Universitat (TH) Hsnnover, Calllnstrasse 3, D-3000 Hannover 1, Federal Republic of Germany
Mllos Ralek and Bruno Schmldt Institut fur Technische Chemie, Technische Universitat Berlin, D-1000 Berlin 72, Federal Republic of Germany
Fischer-Tropsch (FT) synthesis was studied on a specially prepared Mn/Fe catalyst in the slurry phase. Compared to classical K promoted Fe precipitation catalysts, the Mn/Fe catalyst gives higher yields of C, to C, olefins, i.e., about 60 g/Nm3 synthesis gas converted. The product slate follows the Schulz-Flory distribution and is little influenced by operational conditions. From the experimental conversion data, rate constants for overall synthesis gas conversion were calculated.
Introduction In the second wave of Fischer-Tropsch (FT) research starting after the oil embargo in 1973, activities were mainly concentrated on the development of new and more selective catalysts, as one of the major handicaps of all classical FT processes is their poor selectivity. Although the mechanism of the FT synthesis is still subject to many interpretations, there seems to be no doubt now that the synthesis is similar to a polymerizationtype process giving a broad product slate which follows a Schulz-Flory distribution (Anderson et al., 1951, Anderson, 1956, Henrici-Oliv6 and Oliv6,1976). In Germany, special efforts were made to improve the selectivity with regard to lower olefins which can be used as chemical feedstock. Due to high gasoline prices it is, however, also economically promising to produce C5 to Cll hydrocarbons on catalysts with improved selectivities. The various activities can roughly be divided into the following groups: (1)Fe catalysts modified with oxides of Ti, U, Mo, Mn, and Co (Biissemeier et al., 1976; Kitzelmann et al., 1977; Frohning, 1978); (2) partially poisoned Fe whisker promoted with K, Au, and Co (Kitzelmann and Vielstich, 1978); (3) Mn catalysts containing 10 to 20 wt % Fe (Kolbel and Tillmetz, 1976; Kolbel et al., 1978); (4) zeolites with encaged metals, for instance, Fe&O),,-NaY adduct (Ballivet-Tkatchenko et al., 1980) and Ru-Y zeolites (Jacobs, 1980); (5) two-component catalytic systems, Le., a transition metal FT catalyst with a CO hydrogenation function in combination with shape selective catalysts of high acidity (ZSM-5, silicalites) (Caesar et al., 1979) or ZSM-5 and silicalites impregnated with Fe or Co and promoted with K (Rao and Gormley, 1980). *Address all correspondence to this author at Fachbereich Chemie, Univeraitiit Oldenburg, Postfach 2509, Oldenburg, Federal Republic of Germany, D-2900. 0196-4305/82/1121-0222$01.25/0
Only the last two cases present pertinent approaches to avoid the broad product spectrum commonly encountered in FT synthesis, but only in case (4) is the non-SchulzFlory distribution obtained from primary reactions (Ballivet-Tkatchenko et al., 1980; Jacobs, 1980). In case (5) the non-Schulz-Flory distribution is a result of secondary reactions, namely, on the one hand, conversion of primary FT olefins and oxygenates to C5to Cl1 hydrocarbons and, on the other hand, possibly cracking and isomerization of higher hydrocarbons. A new approach for the production of low molecular weight olefins from syngas via methanol has been reported by Kaeding and Butter (1980). Applying a ZSM-5 class zeolite catalyst modified with phosphorus compounds, these authors obtained 70% selectivity to Cz-C4 olefins at 100% methanol conversion. The FT synthesis can be carried out in various reactors, for instance, fixed bed (ARGE Lurgi-Ruhrchemie), fluidized bed (Hydrocol), entrained bed (Kellogg-Sasol), and slurry phase reactors (Rheinpreussen-Koppers). The performance data of the different FT processes in industrial plants and demonstration units were compared and evaluated by Deckwer (1980a,b). Though each of these processes has some favorable features, it is particularly the reaction in the slurry phase which reveals salient advantages. These can be summarized as follows (Kolbel und Ralek, 1977, 1980; Deckwer, 1980a,b): high single-pass conversion, low methane formation, high yield of C3+ products, large content of transportation fuels in CB+ products, high catalyst and reactor performance, and the possibility of using synthesis gas of high CO content. The last point might become important because CO-rich synthesis gases (with a CO to H2 ratio of 1.5) are produced by second generation gasifiers which could be converted directly to FT products in catalytic slurry phase reactors. This would save an additional shift reaction and increase 0 1982 American Chemical Society
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 2, 1982 223
the overall thermal efficiency of the indirect coal liquefaction route (Fischer and Hildebrand, 1979; Poutsma, 1980). In addition to the advantages pointed out above, the F'T synthesis in the slurry phase is also attractive from an engineering point of view as heat removal and reactor design is simple and there are no serious attrition and erosion problems either. The FT slurry process was studied intensively by Kolbel and co-workers at Rheinpreussen (Kolbel und Ackermann, 1956; Kolbel and Ralek, 1977,1980),investigators from the Bureau of Mines (Schlesinger et al., 1954), the UK Fuel Research Station (Hall et al., 1952, Calderbank et al., 1963; Farley and Ray, 19641, and others (Mitra and Roy, 1963; Kunugi et al., 1968; Sakai and Kunigi, 1974). All this early work carried out with promoted Fe catalysts has been reviewed by Kolbel and Ralek (1977, 1980). Assuming a simplified reaction rate law, i.e., first order in hydrogen and zero order in CO, Satterfield and Huff (1980) analyzed FT studies in the slurry phase with regard to mass transfer limitations. The analysis of these authors is based on literature data and estimations for the gas/ liquid mass transfer coefficient, the gas holdup, and the bubble diameter. The conclusion of Satterfield and Huff is that significant mass transfer resistances at the gas/ liquid interface may occur in the FT slurry process. In addition, Satterfield et al. (1980) observed experimentally that the selectivity of the FT synthesis carried out in a stirred vessel is modified by mass transfer effects. Zaidi et al. (1979) and Deckwer et al. (1980) measured gas holdup, heat and mass transfer coefficients, and bubble diameters under those conditions which prevail in the FT slurry process. On the basis of these results, Deckwer et al. (1981b) analyzed literature studies on the FT slurry process and concluded that the process is predominantly reaction controlled. Deckwer et al. (1981b) point out that the low mass transfer resistance has to be attributed mainly to the fact that the use of molten paraffin as a liquid medium to suspend the catalyst leads to high values of the gas holdup and the interfacial area. Quicker and Deckwer (1981) measured diameters of bubbles generated by different spargers and confirmed that in molten wax, bubble diameters are less than 1mm, which leads to high interfacial areas. Mohammed (1977) studied the FT synthesis in the slurry phase using a 3.8 cm diameter bubble column. The catalysts used were untreated red mud and an Mn/Fe catalyst which was precipitated in the usual batchwise manner as described by Kolbel and Tillmetz (1976). The precipitated catalyst precursor was activated in a fluidized bed and then suspended in molten paraffin. The measurements of Mohammed (1977) were based on statistical designs, and regression models were derived therefrom. The aim was to optimize the selectivity to short-chain (C, to C,) olefins. It was found (Mohammed, 1977; Mohammed et al., 1979) that the formation of C, to C4 olefins increased with increasing CO to H2 feed ratio and space velocity and decreasing synthesis temperature. Compared to red mud, the Mn/Fe catalyst used by Mohammed yielded higher selectivities of Cz to C4 olefins which were in the range of 40 to 60 g/Nm3 synthesis gas converted. However, the conversions attainable with the Mn/Fe catalyst were low.
Objective In the present paper, new experimental results on the FT synthesis in the slurry phase will be reported. The measurements were carried out in a bubble column (3.8 cm in diameter, suspension height 60 to 100 cm). An Mn/Fe catalyst precipitated by means of a special method
9 Separators tor liquid hydrocarbons
'
P
meter
We'
Tail gas
analysis
Figure 1. Experimental setup.
and activated in the slurry phase was used. From the experimental conversions overall rate constants will be developed by adopting a procedure similar to that of Satterfield and Huff (1980) and by using the hydrodynamic properties reported by Deckwer et al. (1980).
Experimental Section Bubble Column Slurry Reactor. The measurements were carried out in a stainless steel bubble column of 3.8 cm i.d. and a height of 2 m. An Fe/Mn catalyst of special preparation was used. A schematic picture of the experimental setup is given in Figure 1. The feed gas (CO, H,, N,) controlled by valves and mass flow meters passes an activated carbon filter and the preheater and then enters the reactor. The gas was sparged by a metallic porous plate (75 to 100 pm mean pore diameter). The liquid phase was molten wax (melting point about 90 "C) and contained the catalyst. The tail gas passes a baffled disengagement volume to prevent entrainment of the liquid phase. After pressure reduction, the liquid hydrocarbons are separated by cooling. The CO and C02 content in the gas flow is monitored by IR gas analyzers. The gas composition is analyzed by 3 gas chromatographsworking in parallel. On a Carbosieve column CHI, CO, CO,, and N, are analyzed using He as carrier gas and a heat conductivity detector. H2 is determined on an MS 13 X column with Ar as carrier gas. The hydrocarbons (C, to C4) are determined by means of a Poropak column using a temperature program, H2 as carrier gas, and FID. Although this study is mainly concerned with improving the selectivity with regard to C2to C4 olefins, the hydrocarbon liquid phase obtained in the cooled separators was analyzed for a few runs. The analysis was done by GC using a column with silicon oil, FID, and temperature program (60 to 250 "C, 5 K/min). Only overall hydrocarbon fractions up to CI3were determined. Catalyst Precipitation. When precipitating the twocomponent catalyst in the usual batch-wise manner by adding ammonia to the heated solution of the metal nitrates, it turned out that the ratio of the metals applied in the solution differed greatly from that found in the washed and dried precipitate. It is believed that this re-
224
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 2, 1982 Impeller speed
w I
*0° r p m
YJl
PH
\ Slurry exit
I L
I I;,
+
( t o filtration) Water out
Table I. Operating Conditions of Synthesis Runs (P = 1.2 MPa) run no.
T, "C
(CO/H,)I
SV,h-l
1-7 8-1 3 14-24 25-32 33-41 42-49 50-59 60-69
282 288 293 298 303 303 303 303
1.63-1.69 1.85 1.61 1.68 1.73 1.35 2.25-2.29 0.73
430-875 363-710 258-1 150 258-1 180 34 0- 1 650 237-1620 340-1010 335-1240
I 12"
sauare blades
'C
(CO/H,I,
298
161
293
168
T.
0.6 -
0
0
-
500
c
e- Hot water ( 9 5 ° C ) ( f r o m Thermostat ) *-- NH3 feed
0 200
LOO
600
800
1000
1200
Space velocity. h-' Figure 3. Conversion vs. space velocity a t 293 and 298 "C.
Teflon
A Figure 2. Continuous precipitator for Mn/Fe catalyst.
sults from local pH spots and possibly temperature gradients which are difficult to avoid as the viscosity of the suspension increases considerably during precipitation. In general, it was not possible to prepare a catalyst with reproducible properties by applying this technique. Therefore, an experimental program was started which had the purpose of developing a method for the preparation of reproducible catalysts. In addition, the composition of the catalyst and the conditions of its precipitation and activation were optimized with regard to high selectivities of C2to C4olefm. The details of this catalyst development will be given elsewhere (Lehmann et al., 1980, 1981). As to the new catalyst preparation method, the precipitation is carried out continuously. The precipitator is shown in Figure 2, and its design is similar to that reported by Kolbel and Ralek (1980). It is a glass tube of 24 mm i.d. and a length of 500 mm. The tube contains a stirrer shaft to which 12-mm square blades are fiied, the distance between them being 12 mm. The residence time distribution corresponds to a cascade of about 10 vessels. The preheated solutions of the metal nitrates and the preheated NH, solution are introduced tangentially at the bottom. The precipitation is carried out at a temperature of about 95 "C and a stirrer speed of 800 rpm. The concentrations of the solutions should be chosen such that the Viscosity of the precipitate slurry will not be too high. The pH should be adjusted so that, on the one hand, precipitation of the metal hydroxides easily takes place while, on the other hand, formation of manganese-ammonia complexes is avoided. A pH of 7 was found to be the
optimum. The value was adjusted by controlling the feed pumps. The residence time in the precipitator can easily be varied from 20 to 60 s. Precipitation of 10 L of metal solution giving about 900 g of dried catalyst (nonactivated) can be carried out in less than 45 min. The precipitate (Fe203/Mn203)was dried at 130 "C under vacuum for 24 h. This was followed by grinding in a mortar to particles smaller than 50 pm. The majority of the particles had a diameter in the range of 10 pm. Activation of the catalyst was carried out in suspension; 64 g of the catalyst were suspended in wax at 130 O C . The suspension was introduced into the preheated reactor (150 "C and N2 flow). After heating to 270 "C the catalyst suspension was treated for 24 h with N2 (flow rate 50 NL/h) at 250 kPa. This was followed by CO treatment and reduction with H2 at the same temperature, pressure and flow rate and for 24 h in both cases. After that the catalyst was in active form and the synthesis runs were started. It should be pointed out that the catalyst denoted here throughout as Mn/Fe catalyst does not contain Fe and Mn in the metallic state. Actually, the catalytic phase is a mixture of Fe compounds (magnetite, carbides) and MnO (Podesta, 1981). Synthesis Measurements. The conversion measurements were carried out at 1.2 MPa under variation of space velocity, CO/H2 inlet ratio, and temperature in the order given in Table I. A new steady state was attained within 2 h, if the space velocity or the feed ratio was changed. With an increase in temperature (by 5 "C) it took about 5 h to establish the new steady state. Results The course of the conversion of synthesis gas as a function of the space velocity (volume of gas per catalyst volume and hour) is shown in Figure 3 for two temperatures. The usual behavior is observed; i.e., the conversion decreases with increasing space velocity. Figure 4 gives
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 2, 1982 225
1
303'C
Temperature
0.6
I C O / H 2 1,
0.5 -
0.4 -
0.30.2 -
0.1
i
1
1
6 00
200
1000
1400
Space velocity, h-' Figure 4. Conversion vs. space velocity at various CO to H2feed ratios.
lot ~
~~
800 1000 1200 Space velocity, h-I Figure 6. CHI and C6++ yields vs. space velocity, T = 283-298 "C. 200
All hydrocarbons
L
600
LOO
I
C2 - C,, 0
0
ICOIH,),
:165
T, ' C 0 0
A
282 293 298
A
r t
0
0
b
60E
2 ?
-
G
-
I
t 200
l
i LOO
I
I
600
I
I
,
800 Space velocity ,h-'
:/
A
I C O I t i 2 11 O n l y o l e f i n s C,-C,
lb -
_--__---
1200
Figure 5. C2 to C, yields a t various temperatures.
the measured conversion vs. the space velocity for different CO/H2 inlet ratios at 303 "C. High conversions were obtained for the CO/H2 ratios of 1.35 and 1.73, whereas higher and lower values give a considerably lower overall conversion. Particularly, if the synthesis gas is rich in Hz, the attainable conversion is remarkably small. At the present time, it is not possible to explain reasonably why the conversion in the slurry reactor decreases for H, rich inlet gas. Additional work is required to clarify this point. In this paper, the selectivities (yields) are expressed as grams of hydrocarbons obtained per Nm3 synthesis gas converted. As 1 Nm3 synthesis gas gives about 208 g hydrocarbons (CH,), if inlet and usage ratio are equal, seIectivities (in grams per Nm3 synthesis converted) can be converted to w t ?& by dividing by 208. The selectivities are plotted vs. space velocity in Figures 5 to 8. Figure 5 shows Cz-C4 yields for a CO to H2inlet ratio of 1.65 and different temperatures. An influence of the temperature cannot be recognized for the range studied. The overall
-.8--
/ _ - - -
1
,
1000
LO
8
20
A
.-r----*
,8*8'--
0 8
073
0 .
135
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173
o*
227
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 2, 1982
226
1
1
1
1
1
1
1
1
1
1
1
1
x , = 1PI-l ( 1 - P I 2 P = 0 675
8
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0
0
0
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-
1
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to2:
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c,.
600
800
1000
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Figure 8. CHI and C6+yields vs. space velocity, effect of CO to H, inlet ratio. ,
0
1.5 -
I 0
T "C
m mVI a -
I
303
0
298
A
293
o
288 282
x
0
u
05
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-
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I
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w
-
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?\
0
\ O
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-
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XI
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1
CO/H,
20
25
inlet r a t i o 1
Figure 9. Usage ratio va. feed ratio (CO to Hz).
is drastically reduced. The influence of I is also demonstrated by the results shown in Figure 8. The CHI and Cg+ yields depend only slightly on the space velocity and the inlet ratio provided I 1 1.35, and these results are comparable to those obtained for lower temperatures and shown in Figure 6. But if the CO to H, inlet ratio is reduced to 0.73, the CH, fraction amounts to about 100 g/Nm3 synthesis gas converted and the C6+fraction decreases to very low values. An important variable of the FT process, in general, is the CO to H2 usage ratio. The usage ratio is defined as moles of CO converted per mole of Hz converted. As stated by Kolbel and Ralek (1977), in practice, the feed ratio should be chosen to equal the usage ratio approximately. Figure 9 shows the usage ratio as a function of the inlet ratio for the present investigation. Obviously, for the applied catalyst, the usage ratio is about equal to the inlet ratio only if CO to Hz is close to 1.5. In CO rich gases, Le., (CO/H2)I> 2, the hydrogen conversion is considerably larger than the CO conversion while the opposite is found to be true at low CO to H2feed ratios, i.e., (CO/HJI = 0.73. If (CO/HJI is in the range of 1.5, the water gas shift reaction obviously yields just the required amount of H2 to make I = U. With regard to Figure 9, the usage ratio seems to depend only slightly on the temperature. The variation of U indicated in Figure 9 is due to changes in space velocities. Discussion of Product Distribution. The mechanism of the FT synthesis is still a field of active research and subject to many interpretations. However, Anderson et al. (1951) and Henrici-Olive and Olive (1976) proposed a polymerization type reaction mechanism, and it was shown
I
I
2
I
I
L
I
I
6
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I
8
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I
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12
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 2, 1982 227 Table 11. Product Distribution of Measurements in Fixed-Bed and Slurry Phasea slurry phase
T,"C (CO/H,)r
sv, S - l
XCO+H,
methane ethane ethylene prop a ne propylene butane butylenes1 butylene-2 2 (C,-C,)HC 2
(C,-C4)ol
c, + a
Table 111. Product Distribution in the Slurry Phase with Mn/Fe Catalyst and Shift on ZSM-5Type Catalysta
fixed bed
293 1.61 370 0.443 23.4 11.8 9.2 6.4 26.6 7.6 13.0 3.5
298 1.68 880 0.304
295 1.87 36 5 0.766
295 1.87 27 5 0.849
22.6 9.3 11.4 5.0 27.5 5.9 15.8 2.3
25.3 12.4 8.1 4.6 30.1 6.7 14.9 3.4
24.0 12.6 7.4 4.8 27.5 6.8 12.4 3.0
78.1 52.3 91.9
77.2 57.1 94.5
00.2 56.5 95.3
74.5 50.3 105.2
In g/Nm3 converted.
probability P as compared to the classical potassium promoted Fe catalyst. In fixed bed runs which were carried out for testing catalytic activity in the catalyst development program mentioned above, the C1 to C4distributions were measured (Lehmann et al., 1981). In Table I1 product distributions (Cl to C4)are given which were obtained from slurry phase and fixed bed runs carried out under approximately equal conditions. Although the conversions attainable in the two reactors differ greatly, the product slates differ only slightly. Actually, the scatter observed in the distributions for one reactor type is about the same as that observable for product distribution from the two different reactors. One may therefore conclude that both operation in the fixed bed and in the slurry phase yield the same product slate for the Mn/Fe catalysts used in this study. This result also implies that the chain growth probability does not depend on the kind of operation. This is not very surprising as it is well known that in fixed bed synthesis the internal and external surface of the catalyst is covered with higher product hydrocarbons. In addition, it can be estimated on the basis of the relevant hydrodynamic properties of the slurry reactor (Deckwer et al., 1980) that mass transfer resistances are neglibible. Hence, the FT synthesis in both fixed bed and slurry phase operations is mainly reaction controlled. It can therefore also be expected that the product distributions obtainable from the two operation modes are the same or similar, at least. Kolbel et al. (1978) reported C1 to C4 product distributions on a manganese precipitation catalyst which differ for fixed bed and slurry phase, the yields in the fixed bed being larger while the yields in the slurry phase are smaller than those found in this study. The reason for this discrepancy is not clear. Maybe it is caused by the greatly differing operational conditions applied by Kolbel et al. (1978). With the data of these authors it is not possible to make a Schulz-Flory plot. However, in general, the yields reported are in the same range as observed in this study. If the C2 to C4 fraction y is calculated from eq 1 using the experimentally observed value of P, Le., P = 0.675, one obtains y = 0.417 which corresponds to a yield of 87 g/Nm3 synthesis gas converted. The experimentally found C2 to C4fractions amount to about 80 g/Nm3. Therefore, only 92% of the theoretical value is obtained. As can be seen from the Schulz-Flory plot (Figure lo), this is due to the lower C2 fraction. According to Dry (1976) this low C2 value can be explained by assuming that the lowest olefin, i.e., ethylene, or its precursor, respectively, has a greater
Xl
2
= = a
x,-x, x,-x,, x11+
A
B
C
10.56 41.70 41.66 6.08
10.56 38.36 41.66 9.42
10.56 9.59 70.43 9.43
All figures in % wt. A: calculated theoretically with
P = 0.675; B: experimental, with consideration of (Ex,= 0.92 (cx2-x4);C: after shift o f C , t o C, olefins (75% of C, t o C, fraction) to C, t o C , , hydrocarbons o n ZSM-5.
reactivity which causes a higher value of P and hence a lower amount of C2 hydrocarbons. On the other hand, a more reasonable explanation of the lower value of the C2 fraction was given by Satterfield and Huff (1981). According to these authors the deviation from the SchulzFlory distribution may stem from various experimental artifacts. In particular, the oxygenates which collect in the condensed water phase are frequently not taken into account. Also in this study, the oxygenates were not analyzed. It is believed that this causes the deviation of the C2 fraction; see Figure 10. For a catalyst giving a product slate which follows the Schulz-Flory distribution the maximum amount of C2 to C4 hydrocarbons would be obtainable if P = 0.464 corresponding to 118 g/Nm3 of synthesis gas converted. However, at the same time the amount of CHI would increase to about 60 g/Nm3 which is about 50 % of the C2 to C4 yield. Such a high methane yield is usually not desirable. When starting the present investigations on the FT synthesis on Mn/Fe catalysts suspended in molten wax the main objective was to improve the C2 to C4 olefin selectivity as such short chain olefins are valuable as chemical feedstock. But in the meantime, gasoline prices increased considerably and production of C5 to Cll hydrocarbons is economically attractive too. A promising aspect seems particularly a combination of FT synthesis in the slurry phase with Mobil's process (Poutsma, 1980). The FT process could use synthesis gases of high CO to H2 ratios which are produced by second generation gasifiers. The short-chain olefins obtained therefrom can then be converted to C5to CI1hydrocarbons on shape selective catalysts (ZSM-5 zeolites) (Caesar et al., 1979; Dejaifve et al., 1980). Table I11 shows rough estimates on product distributions based on P = 0.675 obtained for Mn/Fe catalysts of this study (Column A). Column B gives the data actually found, namely, only 92 % of that value predicted for the C2 to C4 fraction from eq 2 with P = 0.675. In column C, the product distribution is given for the case that the C2 to C4 olefins are processed to C5 to Cll hydrocarbons. Then, about 70% of the entire hydrocarbons are in the gasoline range, and the fractions of methane, C2 to C4paraffins, and C12hydrocarbons amount to about 10 wt % each. Evaluation of Kinetic Constants. Satterfield and Huff (1980) proposed a model for FT synthesis in the slurry phase. This model is based on the assumption of a well-mixed slurry phase and the gas phase being in plug flow. The rate law used by Satterfield and Huff is first order in hydrogen and zero order in CO. Such a rate law can be regarded as a limiting case of the rate expressions proposed by Anderson (1956),Dry (1976),and Atwood and Bennett (1979). In addition to these assumptions, pore diffusion and external liquid/solid mass transfer resistances have been neglected in the model of Satterfield and Huff (1980). Deckwer et al. (1981b) slightly modified this model by accounting for the stoichiometry by introducing
228
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 2, 1982
the CO to H2usage ratio, and the variation of the gas flow rate was considered by using an overall contraction factor. This model, after integration, finally leads to the equation St = - a* XH - (1 + a*) In (1 - X,) (3) In this equation, X H is the hydrogen conversion. a* is given by
1+u
QI*=a-
(4) 1+I Here U and I are the CO to H2usage and inlet ratio, respectively. a is defined by V G ( ~ C O +=H ,1)- VGWCO+H,= 0 ) cy= (5) V G ( ~ C O +=H 0,)
VGis the volume flowrate and XCO+H,is the overall synthesis gas conversion. The Stanton number in eq 3 is defined by
200
6 00
1000
lL00
Space v e l o c i t y h - '
Figure 11. Experimental contraction factors. . . . , I
.
E, : l o 9
kJ/mol
in k ,
In this equation, K A presents the reciprocal overall resistance which contains both the mass transfer resistance a t the gas/liquid interface and the reaction resistance _1 -- -1 + --1 (7) KA h a ~ H E L If the quantities involved in eq 3 to 7 can be estimated, the kinetic constants kH can be evaluated. Deckwer et al. (1980) measured the hydrodynamic properties of the wax-catalyst slurry in the same apparatus as that used in this study. These authors report for the gas holdup tG
= O.O53ii~'.'
(8)
and the interfacial area
-L
295
300
290 I ,
1
295'C I
,
1,8
1.75
io3/
T, K - ~
Figure 12. Arrhenius plot of overall rate constants.
a = 4.5&'"
The liquid side mass transfer coefficient kL was found to follow closely the correlation of Calderbank and MooYoung (1961). For small bubble sizes this correlation is
If the relation for the hydrogen diffusivity in molten wax reported by Satterfield and Huff (1980) is introduced the mass transfer coefficient of H2 (in cm/s) can be calculated from kL = 0.1165
-3
(i)'3
exp(-1523/7?
(11)
The mean densities and viscosities can be obtained from data given by Hammer (1968) and should be corrected for the solid contents (Deckwer et al., 1980). The Henry constant in eq 6 can be calculated from hydrogen solubility in molten wax measured by Peter and Weinert (1955). The contraction factor a defined by eq 5 can be obtained from the experimental data. An overall nitrogen balance leads to
where xN,, and x N ~are , ~ the nitrogen mole fractions at the reactor injet and outlet, respectively. The contraction factors calculated from the experimental data of this study (at 303 "C) are shown in Figure 11 for various CO to Hz inlet ratios. The mean value of N for all the measurements
with a CO to Hz feed ratio of about 1.65 is -0.55. From the rate constants, ItH, which refer to hydrogen conversion, the overall rate constants for synthesis gas conversion are obtained by ko = kH (1 + (13) In this study, the majority of the runs was carried out with a CO/H2 inlet ratio of about 1.7, and for this ratio the highest overall synthesis gas conversions could be achieved; see Figure 4. Only for these runs consistent rate constants ko could be evaluated by means of eq 3 to 13. It is suspected that for H2 rich synthesis gas the rate low which assumes first order for Hz and zero order for CO does not apply. The overall rate constants evaluated from the runs with (CO/HJI of about 1.7 are plotted vs. the inverse temperature in Figure 12. The least-square fit of all 41 data gives an acitivation energy of 109 kJ/mol, which is in the reasonable range (Satterfield and Huff, 1980). Though the description of the data shown in Figure 1 2 is not very good, one is inclined to conclude that, in general, the first-order rate law presupposed in the data analysis is obviously applicable. An Arrhenius plot of the reciprocal overall resistance KA leads to an activation energy of 81 kJ/mol. This indicates that there is some mass transfer resistance, but this is moderate. Indeed, the evaluation of the measured data of this study which was carried out a very low gas velocities shows that the relative reaction resistance
p=-
~ / ~ H E L
1/ K A
(14)
is usually larger than 70% except for the lowest space velocities, and the highest temperature /3 decreases to 0.4.
Ind. Eng. Chem. Process Des. Dev., Vol. 2 1 , No. 2, 1982 229 Table IV. Applied Catalysts in Liquid Phase FT Synthesis catalyst
activation at T = 270 "C in compn (nonactivated)
preparation
( A ) red mud (B) Mn/Fe
batch precipitation
fixed bed fluidized bed
(C) Fe/Cu
continuous precipitation
slurry phase
( D ) Mn/Fe
continuous precipitation
slurry phase
content in slurry, % wt (nonactivated)
48.1% wt Fe Mn/Fe = 3.1:1 14.8% wt Fe Fe:Cu= 1 9 : 1 55.3% wt Fe Mn/Fe = 6 : l 9.2% wt Fe
20.8 14.3 15.0
13.8
1 = 3l.L " C ICO'H21j='85
0030Red m u d
T
k,
-o---o-
-0
5-1
0025-
05
I5
I
2
PMPa
-
Figure 14. First-order rate constants as a function of pressure evaluated from data of Mohammed (1977).
M n I F e pptd batchwise
E A :7 2 k J / m o l
Kunugi et a I . / & h : \ \
103
, 16
3fO
3f0,
17
300
280
18
260
2LO
19
2201,'C
\
20
Mitra.Roy
\
IO~/T,K-' Figure 13. Rate constants for synthesis gas conversion from studies with various catalysts (see Table IV) in same experimental setup.
The procedure for evaluating rate constants from the conversion measurements of this study with a specially prepared Mn/Fe catalyst was also applied to some previous measurements with other catalysts but carried out in the same experimental setup; see Figure 1. The catalysts used are listed in Table IV. Mohammed (1977) used a Mn/Fe catalyst (B) which was precipitated batchwise and activated in a fluidized bed. For reasons of comparison Mohammed also did some measurements with untreated red mud activated at 270 "C in a fixed bed. As shown in Figure 13, the ko values evaluated from Mohammed's results reveal a wide scatter. The red mud is very active compared to the Mn/Fe catalyst but its selectivity to C2 to C4olefins is poor. Compared to the Mn/Fe catalyst (B) of Mohammed the catalyst employed in this study is more active by a factor of 6 to 10. This result presents a considerable improvement and underlines the importance of catalyst preparation. In addition, the present high activity catalyst precipitated continuously and reduced in the slurry phase is as selective with regard to C2to C4 olefins as the low activity catalyst of Mohammed (1977). The Fe/Cu catalyst (C) precipitated continuously is again considerably more active than the Mn/Fe catalyst (D) used in this study. The last one reaches the same activity only at a temperature of 300 "C which the Fe/Cu catalyst already has at 220 "C, but for this catalyst too, the C2 to C4
t11 17
300
280 I 1
18
260 I
I
19
240 T . " C '\ 20 IO3/ T. K'
1
Figure 15. Rate constants from literature studies.
olefin selectivity is small (Deckwer et al., 1981a). Mohammed (1977) also carried out some measurements under variation of synthesis pressure. Figure 14 shows overall rate constants evaluated from these runs at 344 "C. The CO to H, inlet ratio is 1.85 and the data points of Figure 14 are mean values for various space velocities. It can be seen that the overall rate constants are independent of the entire synthesis pressure. This confirms a posteriori the applicability of the simplified kinetic expression introduced by Satterfield and Huff (1980) and employed in this study too. The model described by Deckwer et al. (1981) which eventually leads to eq 3 can also be applied to estimate approximate rate constants for other FT studies carried out in the liquid phase. The results are shown in Figure 15 and compared with the findings with catalysts (C) and
230
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 2, 1982
Table V. Rate Constants for Overall Synthesis Gas Conversion %
authr
k,' ( s % wt
wt Fe
(in slurry)
T, "C
k,, s-'
Fe).'
23.3 11.0 14.0 10.1 3.2
268 266 250 2 50 260
0.4420 0.2450 0.0560 0.1017 0.1470
0.0190 0.0223 0.0040 0.0100 0.0459
10.0 2.12
280 300
0.2179 0.0068
0.0218 0.0032
8.3 1.26
250 303
0.2306 0.0690
0.0278 0.0544
Kolbel, Ackermann dR = 1 2 9 cm d~ = 4.7 cm
Schlesinger et al. Mitra and Roy Kunigi et al. Mohammed ( A ) red mud (B) Mn/Fe
this study
( C ) Fe/Cu ( D ) Mn/Fe
(D). Schlesinger et al. (1954) used a fused Fe catalyst which had a high selectivity with regard to oxygenates. The selectivecatalyst of Schlesinger et al. is about a~ active as the Mn/Fe catalyst (D) though the temperature ranges of their applicability differ by 50 "C. The Fe catalyst used by Kolbel and Ackermann (1956), Kunigi et al. (1968), Mitra and Roy (1963), and the Fe/Cu catalyst (C) give first-order rate constants which vary by a factor of about 4. Table V lists the numerical ko values. The data refer to different temperatures, different concentrations of the catalyst in the slurry, and to various catalyst preparations, of course. Nevertheless, one can suspect that the iron present in the slurry forms mainly the catalytically active component. It is therefore challenging to refer the rate constant ko to the Fe content in the slurry. It should be stressed that there is no scientific basis for doing this. However, it is surprising to note that if the rate constants, ko, are simply divided by the Fe content of the slurry, the k,,' values (in s-l (5% wt Fe)-') obtained from the data reported by Kolbel and Ackermann (1956) for the two slurry reactors with largely differing diameters are practically the same, and these data are also in reasonable agreement with the k,' values evaluated for red mud and the Fe/Cu catalyst (C). The k,' values of all the precipitated Fe based catalysts are in the range from 0.01 to 0.04 (s % wt Fe)-'. As a rule of thumb, a value of about 0.02 can be recommended for this type of catalyst and a temperature range of 260-270 "C. Some of the rate constants, kd, given in Table V differ slightly from those reported by Deckwer et al. (1981b),which is caused by an incorrect value of the Fe content assumed by Deckwer et al. (1981b). Summary and Conclusions The Mn/Fe catalyst used in this study gives a C2to C4 olefin yield of about 60 g/Nm3 synthesis gas converted. The overall amount of C2 to C4 hydrocarbons attainable is about 80 g/Nm3. For the present catalyst used in slurry phase operation the product distribution is little influenced by the temperature, the space velocity, and the CO to H2 feed ratio provided (CO/H,), 1 1.35. If (CO/H2)I is reduced to 0.73, the product distribution changes drastically and methane is the main product (95 g/Nm3 synthesis gas converted). The C number weight fractions (up to CI3) follow the Schulz-Flory distribution, and the product slates (up to C,) obtained from fixed bed and slurry phase operations are the same for the catalyst studied though the conversions in the two reactors differ greatly. On the basis of simplifying assumptions, kinetic constants for overall synthesis gas conversion have been evaluated. The Mn/Fe catalyst used in this study is considerably more active than the one applied by Mohammed. This underlines the importance of proper cat-
alyst preparation. However, from an analysis of literature studies it follows that the Mn/Fe catalysts are still less active than conventional Fe-based catalysts. The optimum temperature range of the Mn/Fe catalyst is about 300 "C while other Fe based catalysts are used in the range of 230 to 270 "C. In slurry phase operation, the optimum CO to H2 feed ratio is in the range of 1.5 because in this range high conversion can be obtained. This result makes the Mn/Fe catalyst and operation in the slurry phase attractive for direct processing of synthesis gases from so-called second generation gasifiers and for a combination with Mobil's route for producing Cs to CI1hydrocarbons. An estimate indicates that such an FT-Mobil combination would give about 70% of the synthesized hydrocarbons in the gasoline range. Acknowledgment The authors gratefully acknowledge the support from the German Federal Ministry of Research and Technology (BMFT) and the Stiftung Volkswagenwerk. Nomenclature a = specific gas-liquid interfacial area referred to gassed susupension volume, cm-' D = diffusivity in the liquid phase dR = reactor diameter g = gravitational acceleration, cm/s2 H = Henry's constant, kPa cm3/mol i = carbon number I = CO to H2 inlet ratio, N c o , o / N ~ ~ , ~ k , = first-order rate constant for overall synthesis gas conversion, defined by eq 13, s-l kH = first-order rate constant for hydrogen conversion, s-l kL = liquid side mass transfer coefficient, cm/s KA = reciprocal overall resistance, defined by eq 7, s-l L = height of gassed suspension, cm N = mol flow rate, mol/s P = chain growth probability R = gas constant, kPa cm3/mol K S c = Schmidt number, p L / ( p L D ) St = Stanton number, defined by eq 6 S V = space velocity, NL synthesis gas/L of catalyst and h, h-l
T = temperature, OC or K uG, = superficial gas velocity at reactor inlet, cm/s LzG = mean superficial gas velocity, UG (1 + O&&O+H2), cm/s U = CO to H2 usage ratio, .4Nc /&H, VG = gas volume flow rate, Nm /h xi = mass fraction of hydrocarbons with i carbon atoms xN2 = nitrogen mole fraction in inlet (0) or outlet (1) gas XH = hydrogen conversion = synthesis gas conversion XCO+H, Y = yield, g of hydrocarbons/Nm3 synthesis gas converted
B
Greek Letters a = volume contraction factor, defined by eq 5 and calculated
from eq 1 2 a* = parameter, defined by eq 4 p = relative reaction resistance, eq 14 y = C2 to C4 hydrocarbon mass fraction 6 = relative volume fraction pL = liquid viscosity, P p = suspension viscosity, P
liquid density, g/cm3 p = suspension density (gas-free),g/cm3 pL =
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Bailiiet-Tkatchenko. D.; Condurier, G.; Mozzanego, H. "Proceedings of the International Symposium on Catalysis by Zeolites", Ecully: Lyon, France, 1980.
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Birssemeler, 8.; Frohning, C. D.; Cornlls, B. Hydrocarbon Process 1078, 7 7 , 105. Caesar, P. D.; Brennan, J. A.; Garwood, W. E.; Cirlc, J. J. Catal. 1070, 56, 274. Calderbank, P. H.; Evans, F.; Farley, R.; Jepson, 0.; Poll, A,, "Catalysis in Practice"; Symposium Proceedings. Institution of Chemical Engineers: London, 1963 p 68. Calderbank, P. H.; Moo-Young, M. Chem. Eng. Scl. 1081, 76, 39. Deckwer. W.-D. "Proceedings, 7th Annual International Conference on Coal Gaslflcation, Liquefaction, and Conversion to Electricity", Pittsburgh, Pa., Aug 5-7 196Oa. Deckwer, W.-D. 011 Oes J. 1080b, 78(45). 198. Deckwer. W.D.; Lehmann, H.J.; Ralek, M.; Schmidt, B. Chem. Ing. Tech. to be published, 198la. Deckwer, W.D.; Louisi. Y.; ZaMI, A.; Ralek. M. Ind. €no. Chem. Process Des. Dev. 1080, 79, 699. Deckwer, W.D.; Serpemen, Y.; Ralek, M.; Schmidt, B. Chem. Eng. Scl. 1081b, 36, 785. 791. DeJelfve, P.; Vdrlne, J. C.; Bolls, V.; Derouane, E. G. J . Catal. 1080. 63, 331. DryYM: E. Ind. Eng. Chem. Prod. Res. Dev. 1078, 75, 282. Farley, R.; Ray, D. J. Inst. Pet. 1084, 50, 27. Fischer, R. H.; Hildebrand, R. E. "Transportation Fuels from Synthesis Gas", paper presented at Methanol Symposium, 13th Middle Atlantic Regional Meeting of the American Chemical Society, Atlanta, Mar 23, 1979. Frohnlng, C. D. 2nd Kolloquium Fischer-Tropsch-Synthese, Jirllch, West Germany, 1978. Hall, C. C.; Gall, D.; Smith, S. L. J. Inst. Pet. 1052, 38, 845. Hammer, H. Habllitatlonsschrlft, TU Berlin, Berlin, West Germany, 1968. Henrlcl-Olhr6, G.; Oliv6, S. Angew. Chem. Int. Ed. Engl. 1076, 75, 136. Jacobs, P. A. "Proceedings of the International Symposium on Catalysis by Zeolites"; Ecully: Lyon, France, 1980. Kaeding, W. W.; Butter, S. A. J. Catal. 1080, 67,155. Kitrelmann, D.; Vlelstich, W. 2nd Kolloquium Flscher-Tropsch-Synthese, Jirllch. West Germany, 1978. Kitzelmann. D.; Vielstlch, W.; Ditrich, T. Chem. Ing. Tech. 1077, 4 9 , 463. KGlbel, H.; Ackermann, P. Chem. Ing. Tech. 1058, 28, 381.
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Kolbel, H.; Ralek, M. in "Chemlerohstoffe aus Kohle"; Falbe, J., Ed.; G. Thieme Verlag: Stuttgart. 1977; p 257. KGlbel, H.; Ralek, M. Catal. Rev. Sci. Eng. 1080, 27,225. KGlbel, H.; R a w , M.; Tlllmetz, K. D. Proc. 13th Intersoc. Energy Conv. Eng. Conf., SOC.of Automot. Engrs., 1978; p 482. Kolbel, H.; Tlllmetz, K.-D. Belgian Patent 837 628, 1976. Kunugi, T.; Sakal. T.; Negishl, N. Seklyu Gakkai Shl 1088, 7 7 , 636. Lehmann, H.J.; Ralek, M.; Deckwer, W.4. Paper No. 103d presented at 73rd Annual Meeting of AIChE, Chicago. Nov 16-20, 1980. Lehmann, H.J.; Schmidt, 6.; Ralek, M.; Deckwer, W.-D. Appl. Catal. submitted, 1981. Mltra, A.; Roy, A. Ind. Chem. Eng. July 1083, 127. Mohammed, M. Dr.-Ing. Thesis, TU Berlin, Berlin, West Germany, 1977. Mohammed, M.; Schmldt. B.; Schneidt, D.; Ralek, M. Chem. Ing. Tech. 1070, 57, 739. Peter, S.; Weinert, M. Z.f h y s . Chem. (Frankfurt am M a h ) 1955, 5 , 114. Podesta, W. Diploma work, Institut far Physikallsche Chemie, Universlt Hamburg, Hamburg, Federal Republic of Germany, 1981. Poutsma, M. L. Oak Rage National Laboratory Report 5635, 1980. Quicker, G.; Deckwer, W.-D. Chem. Eng. Sci. 1081, 36, 1577. Rao, V. U. S.: Gormley, R. J. Hydrocarbon Process. Nov 1080, 139. Sakai, T.; Kunugi, T. Seklyu Gakkal Shi 1074, 77, 863. Satterfield, C. N.; Huff, G. A. Chem. Eng. Scl. 1080, 35, 195. Satterfield, C. N.; Huff, G. A. I n d . Eng. Chem. Process D e s . Dev. to be published, 1981. Satterfield, C. N.; Longwell, J. P.; Huff, 0.A. Paper No. 102a presented at 73rd Annual Meeting, Amerlcan Institute of Chemical Enpineers, Chicam, NOV18-20, 1980. Schlesinger, M.; Benson, H.; Murphy, E.; Storch, H. H. Ind. Eng. Chem. 1054, 46, 1322. Zaidl, A.; Loulsi, Y.; Ralek, M.; Deckwer, W.-D. Ger. Chem. Eng. 1970, 2 , 94.
Received for review February 20, 1981 Revised manuscript received September 15, 1981 Accepted September 30, 1981
Modeling the Fischer-Tropsch Synthesis in the Slurry Phase Wolf-Dleter Deckwer' and Yalcln Serpemen Institut fur Technlsche Chemie, Universitat (TH) Hannover, Callinstrasse 3, 0 3 0 0 0 Hannover 7, Federal Republic of Germany
Milos Ralek and Bruno Schmldt Institut fur Technlsche Chemie, Technlsche Universltat Berlln, 0 7000 Berlin 72,Federal Republic of Germany
A model of the FT synthesis in the slurry phase is being developed. I t is based on a first-order rate expression for hydrogen consumption but considers all relevant hydrodynamic phenomena. The parameters involved in the model equations were estimated from literature correlations and recent experimental results. The basic set of operational conditions was that applied in the Rheinpreussen-Koppers demonstration plant. The computations show that liquid solid mass transfer resistances and catalyst settling can be neglected in larger diameter columns. Gas-liquid mass transfer limitations are also small as long as conventional catalysts and favorable gas velocities are applied. The simulated results are in general agreement with practical experience.
Introduction Since the oil embargo in 1973, the Fischel-Tropsch (FT) process, i.e., the synthesis of hydrocarbons from carbon monoxide and hydrogen, has regained great actuality. The FT synthesis can be carried out in fixed, fluidized, and entrained fluidized beds and in the slurry phase as well. The various FT processes were described in detail by Frohning et al. (1977), and a brief comparison was given
* Address all correspondence to this author at Fachbereich Chemie, Univemitiit Oldenburg, Postfach 2509, Oldenburg, Federal Republic of Germany, D-2900. 0196-4305/82/1121-0231$01.25/0
by Deckwer (1980). The literature on the FT synthesis in the slurry phase has been reviewed by Kolbel and Ralek (1977, 1980) and Poutsma (1980). Due to its salient advantages the FT process in the slurry phase appears to be particularly attractive (Fischer and Hildebrand, 1979; Poutsma, 1980). Purpose In previous papers, we have analyzed FT slurry reactor performance data from lab-scale reactors on the basis of a simple model (Deckwer et al., 1981a,b). Rough estimates of first-order rate constants for synthesis gas conversion 0 1982 American
Chemical Society
