Ind. Eng. Chem. Process Des. Dev. 1082, 21, 217-222
217
Mass Transfer Studies on Sieve Trays with I-In. Diameter Perforations Walter J. Thomas Chemical Engineering Department, University of Surrey, GuMford, England GU2 5XH
Oluwasola OgboJa' Chemical Englneering Department, University of Lagos, Lagos, Nigeria
The mass transfer characteristics of sleve trays having 1-in. dlameter perforations have been studied employing two different designs of columns-a small 3 X 1 ft rectangular section metal column and a 32-in. internal diameter round glass column. The fluid systems employed were air-CO,/water and air-CO,/NaOH solution. Graphical correlations of the results have been presented. The results are similar in trend to and compare well with those in the literature. Some differences are apparent between the data from the two columns which are attributable to the effect of column size and geometry.
Introduction Researchers have not given much attention to sieve trays with perforation diameters larger than 3/a in. because it was erroneously held that such trays would be susceptible to excessive weeping and entrainment. Yet, they are preferable in operations where dirt, rust, and solid deposition can block the holes. Lemieux and Scotti (1969) and, more recently, Thomas and Ogboja (1978) have shown that large-hole trays have weeping and entrainment characteristics which compare very well with those from small-hole trays. On the basis of this finding, the advantage of large-hole trays may be exploited in design. However, the nonabundance of research data constitutes a hindrance. Researchers have not been considering tray size in particular as an important variable in the studies of tray characteristics. Ogboja (1975) has reported significant influence of tray size on mass transfer parameters such as interfacial area and mass transfer coefficient. Since experimental data obtained on small trays are used in designing industrial trays, appropriate criteria are needed for scaling up the data obtained on the former. The problem of scaling up can be avoided if experimental studies are carried out with tray size as one of the variables. The present work was carried out to establish the efficiency and mass transfer characteristics of sieve trays with 1-in. diameter perforations. Two columns were used in order to test the effect of shape and size, a small rectangular column and an industrial-size round column. Murphree tray efficiencies and mass transfer coefficients were determined using the system air-C02/water, and gaslliquid interfacial areas using air-C02/NaOH solution. These parameters were investigated as functions of liquid flow rate per foot of weir length, L,, and gas kinetic energy factor, FA. The values of the variables were chosen to span the range of operability of each column. Theory (a) Tray Efficiency. The Murphree tray efficiency in liquid terms is defined as
where xin and xOutare the mole fractions of the transferring component at the tray inlet, and exit respectively and xeqm will be the mole fraction at the exit if the liquid on the tray were in equilibrium with the vapor leaving it. The definition is based on the assumption that there is complete mixing on the tray. The complex gaslliquid interaction on the tray results in a high degree of back-mixing and therefore eq 1 is appropriate for the sieve tray. (b) Liquid Residence Time. The liquid mean residence time is obtained from the distribution of a residence time function
Usually, f(t) is the concentration of a tracer injected into the liquid at the tray inlet measured at the exit against time. The dimensionless variance which is a measure of the degree of back-mixing on the tray is given by (3)
where Jm(t - tnJ2f(t) dt ut =
J'f(t)
0196-4305/82/1121-0217$01.25/0
dt
.
(c) Mass Transfer Coefficient, kLa Murphree tray efficiency has been related to the dimensionless distribution function f ( 0 ) by Foss (1957) as 1EML
=
1*Direct correspondence to Dr. Olu. Ogboja, 2631 Univ. Blvd., North, Apt. G13, Jacksonville, FL 32211.
(4)
1
Jm
(1 -
exp(-XEoG-e).f(e)de
1-exp(-XEoG.e).f(e) del
(5)
0
where 0 = t/tm and X = mG,/L,. 0 1982 American Chemical Society
G, and L, represent
218
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 2, 1982
the molar gas and liquid flow rates, respectively, and m is the slope of the equilibrium curve y = n x + c (where x and y are the mole fractions of the transferring component in the liquid and gas phases, respectively, and c is a constant). If the system is liquid film controlled, X >> 1 and
The interfacial area per unit volume of froth is
Hence It can be shown that NL
=
mOG
(7)
hence EML
= 1-
Amexp(-NL8)f(8) dtJ
(8)
Foss et al. (1958) have proposed the following relationship for 0 and f(8) f ( 8 ) = a@ exp(8)
(9)
for f(0) to be a normalized distribution it must satisfy the following conditions
&f(6) d8 = 1
Jm
(1 - 8)' f ( 6 ) d8 = u2
(12)
Based on these guidelines, Foss et al. (1958) and Thomas and Campbell (1967),using different approaches, obtained the following form for f ( 8 ) f(8) =
e(l-u2)/a2
exP (-6 / a2)
(u2)l/a2
r(i/
(13)
on substituting this expression for f ( 8 ) in eq 8, we have EML
+
= 1 - (1 NL.~2]-1/02
(14)
hence
It has been shown in the A.1.Ch.E. Bubble Tray Design Manual (1958) that NL = kLa*t,
(16)
hence
Thus, if E m , u' and tmare known from experiments, kLa can be calculated from eq 15 and 16. (d) Gas/Liquid Interfacial Area. For a pseudofirst-order reaction between NaOH and COz, the rate of mass transfer per unit interfacial area has been given by Danckwerts (1970) as
R =
C i d m i
(18)
If A, is the total interfacial area, then
R, = AtR Therefore
(19)
a=
A Z f C , d m (e) Mass Transfer Coefficient, kL. If kLa is derived from results of physical absorption experiments, and a from chemical absorption experiments, then the condition to be satisfied before kL can be deduced from eq 17 and 22 is kLa (chemical absorption) = kLa (physical absorption) (23) Eben and Pigford (1965) and Harris and Roper (1962, 1963) have shown that this condition can be satisfied if the chemical absorption reaction is pseudo first order. Experimental Section The Main Apparatus. The main apparatus consisted of a 3 X 1f t rectangular stainless steel column and a 32-in. internal diameter round glass column mounted side by side on a wooden platform and steel framework. The test column consisted of three boxes, the test tray being located between the middle and top boxes. The glass column also consisted of three boxes with the test tray being similarly located; however, the bottom tray whose function was to ensure even distribution of the inlet air was made of stainless steel. Detailed description of the auxiliary equipment and the details of the experimental trays have been presented in a previous paper by Thomas and Ogboja (1978). Liquid Residence Time Equipment. The liquid residence time distribution was measured by measuring the degree of light obscuration caused by dye (nigrosine) dissolved in water. A special dye detector was designed for the purpose. Essentially, the detector had three compartments-two end compartments and a middle compartment. One end compartment housed a powerful lamp, the other housed a selenium cell, while the middle compartment served as the conduit of liquid through the detector. The end compartments were water-tight and light rays from the lamp could reach the cell only via convex mirror windows. The locations of both lamp and cell in relation to the windows were such that the rays from the lamp were parallel in the middle compartment and were converged on the cell. The detector was designed to suspend on the exit weir and receive liquid from the central 50% of the weir length. A selenium cell can generate voltage across a resistor connected in parallel with it if light rays impinge on it. This characteristic was exploited in measuring the dye concentration continuously by means of the arrangement in Figure 1. By means of the balancing unit the voltage generated by the cell across a selected resistor was balanced against a standard voltage when there was no dye in the liquid stream. On introducing dye, an out-of-balancewhich was a measure of the concentration of the dye was obtained. A continuous record of this was made on paper tape by means of a punch drive, an Addo punch, and a scanner unit. Dye injection was effected for a desired period by incorporating a solenoid valve driven by a
Ind. Eng. Chem. Process Des. Dev., Vol. 21,No. 2, 1982 219 Pressur i r e d
..
Solenoid volve
Troy
i
\
D y e i n j e c t i o n point
Balancing unit
Timer
G
Punch drive A d d o punch S c o n n e r unit
Voltmeter
I U
DVM
J
Figure 1. Residence time measuring equipment.
Venner timer. The timer and the data punch drive were operated by the same starter so as to commence the record of data from the instant of dye injection. A pulse of tracer injection requires the delivery of the desired quantity of tracer within a few seconds. Also, the tracer delivery must be such that the liquid concentration is the same at all points of the inlet at the moment of injection. In order to meet these requirements a dye tank was constructed and provided with an opening to a 5-atm pressure source. Dye from the bottom of the tank was led through a copper tubing to injection points at the tray inlet via the solenoid valve. The injection points were spaced evenly and fitted with pressure valves to prevent the issuing of more dye to the tray after each injection. Liquid Residence Time Measurement. In a preliminary experiment the detector was calibrated outside the column by measuring the cell voltages across a selected resistor corresponding to various dye concentrations. A linear relationship was found to exist between the cell voltage and dye concentration over a range of dye concentrations. The widest range was obtained with a resistor of 100 kQ and was therefore selected for the studies. After the calibration, the detector was securely installed in the column and the experiment started. The desired flow rates of liquid and gas were set and the balancing unit was adjusted to make the reading on the voltmeter zero. Dye solution was injected at the tray inlet and a continuous digital record of dye concentration was obtained on paper tape from the tray exit. Measurement of Tray Efficiency. Carbon dioxide was injected continuously into the air stream on the suction side of the fan to ensure complete mixing with the air before reaching the tray. Liquid samples were taken at the tray inlet and exit when the system reached equilibrium. Adequate care was taken to ensure constant COz concentration in the air and constant air and water temperatures. The solubility of C 0 2in water was measured for the tray conditions by passing a stream of gas bled from the gas line at a point before the tray into water in a bottle for 1 h. In order to prevent C 0 2 desorption, the samples were mixed with dilute NaOH solution. The COz concentration was later determined by means of an automatic titrator. Measurement of Interfacial Area. The procedure was the same as for tray efficiency except that the liquid was NaOH solution. Liquid samples were similarly taken and later analyzed for the rate of C02 absorption.
10
15
20
30
25 L ,
35
40
45
Qpmllt
b F,= 1 7 2 . Round Column
F,= 2 06, Rectangular C 0 1 u m n F ~ ‘ ~ ’ ~Work ’~ A FA 2 10. Reclanguior Column+ Thomas and Hag (1975)
Figure 2. Murphree tray efficiency.
Results The Murphree tray efficiency Em was calculated from eq 1 and the mean liquid residence time t, from eq 2. Usually, liquid residence time distribution is terminated by a long tail which tends to give a large value to the mean residence time. It has been known from experience that the amount of tracer represented by the tail is a negIigible fraction of the total amount injected. Sater and Levenspiel (1966) have suggested a technique for estimating the end of such distributions. The dimensionless variance of the distribution was obtained from eq 3 and 4 and hence the mass transfer coefficient kLa from eq 17. The diffusivity D,of C 0 2and the second-order reaction rate constant k2was obtained as outlined by Barrett (1966) and the interfacial concentration Ci of C02as outlined by Danckwerts (1970). The gas/liquid interfacial area a was then calculated from eq 22. The value of kL was then obtained from a and kLa. The graphical correlations of the results are shown in Figures 2-9. Discussion The parameters studied, (Em,t,, 2,a, kLa,kL),have been correlated mainly against liquid flow rate, (represented by Lw), because gas flow rate, (represented by FA), was found to have no significant effect on the parameters. For example, by increasing FA from 1.70 to 3.00, an increase of 2% was achieved in the efficiency of the smaller tray. These FA values correspond to 300 and 600 ft3/min of air, respectively. It may therefore be appreciated that a significant increase in the amount of gas flow hardly resulted in a change in the measured value of EML. The reason for this insensitivity to increased gas rate may not be immediately apparent since more gas is ex-
Ind. Eng. Chem. Process Des. Dev., Vol.
220
21, No. 2, 1982
-
.
0'6-
m
5
.
012.
I
.
014-
A
&
30
35
I
5;
L
0
5
F2
15
'0
25
L,
30
Roclangulor Column
ilaunc
35
45
40
L
gpmlfl
O6
FA
fa; 1 7 2
column
2 6
Figure 3. Liquid residence time.
:
:
2
2 ~
50
L
12
8
4
20
16 1 ,
m
24
32
1148
i
34
10
e Round Column
15
.
5
0
Column.
Rectangular
28
6
n
Fa
' 2 06
A
FA: 1 7 2
Figure 4. Murphree tray efficiency as a function of liquid residence time.
20
25
40
45
Lw g p m l l t
F A =1 7 2
Round
column
F~ i 2 06, Rectangular FA * 2
lo
Rectangular
columri CGlUmn>
Dresant
Work
ThO'POS
aPd Foq
35
40
1'976)
Figure 8. Gas-liquid interfacial area.
I
0
1
* *
b 02i
I
0
1/ 5
'0
'5
20
iw
3c
25
I
Rcclongulor Columr 0
R3und
45
LO
35
gpnllt
Columl
Fa = 1 7 7
'5
10
45
I
c@lumn 1 >resent W o r k Fa: 2 0 6 , R e c t a n g u l a r c @ l u m i r
A fa 3 361
210
:
Ractangular
-
co1umn-L-Thamos
Haq '19761
Figure 9. Mass transfer coefficient, k ~ .
0 351
c
30
gpmift
% = I 72, Round &
c
25 L,
Figure 5. Dimensionless variance.
d
20
FA:' 7 2
I
341
331 32!
Y
ci5 0 24 ' 2
14
' 6
18
20
2 2
24
2 5
2 8
3 0
FA- FACTOR
a e c t o n g u i a r c~~llmn, L y 0
Rarird
cotumlr
Lw
I
2 o g m ~ ) i t C I we r 2 5 9Fp,ii $1 ye,r
Figure 6. Dimensionless variance as a function of F A factor.
1
pected to cause more frothing and therefore better mixing and mass transfer. In the present work, the dimensionless variance of residence time distribution has been employed as a measure of the degree of liquid back-mixing. Figure 6 shows that an increase in F A produced an increase in back-mixing (2).However, the change is only an increase of 9% in uz for an almost 100% increase in F A and it may therefore be incorrect to expect higher mass transfer as a consequence. Mass transfer depends to a large extent on the contact time between the liquid and the gas phases. Obviously if achieving higher mixing will result in less gas/liquid contact time, then it cannot be immediately concluded that
Ind. Eng. Chem. Process Des. Dev., Vol. 21, No. 2, 1982 221
higher mass transfer will be achieved. The experience in this work was that while an increase in back-mixing was achieved with an increase in F A , the corresponding effect on the liquid residence time was found to be very insignificant. It therefore follows that the gasfliquid contact time was not affected by an increase in FA. Consequently, increasing FAdid not sufficiently affect the degree of liquid back-mixing or gas/liquid contact time as to produce a significant increase in mass transfer. Hence, some amount of energy may be saved by operating the tray at low values of FA compatible with operability. The remaining mass transfer parameters a,kLa,and k also depend on the degree of back-mixing and gasfliquid contact time; it is therefore not surprising that they are not significantly affected by an increase in FA. The variable L, in its own case produced an easily appreciable effect on the studied parameters. Increasing L, produced a decreasing effect on the tray efficiency, E M L (Figure 2). The reason for this result can be deduced from Figure 3 where it can be seen that t , decreases with increase in L,. Since decrease in residence time results in decrease in gasfliquid contact time, increase in liquid rate will therefore be accompanied by decrease in efficiency. Figure 4 has been deduced from Figure 2 and 3 to show that higher efficiency values can be obtained with increase in liquid residence time. Figure 5 shows that the degree of back-mixing, 2,was not affected by increase in liquid flow rate, L,. This suggests that higher values of L, did not increase the turbulence on the tray which may be a surprise on first thought. However, higher L , produced higher liquid holdup on the tray (Thomas and Ogboja 1978), thus reducing the turbulence induced by the gas. The constant values of u2 therefore imply that the turbulence induced by increasing the liquid flow rate was balanced by the suppression of the turbulence-inducing effect of the gas due to the consequent increase in liquid holdup. Higher gas/liquid interfacial area, a, was obtained with increase in liquid flow rate, L,. This is in order since the tray froth volume increased with increase in L, and a was a measure of the area of the active gas pockets in the froth. The values of kLa and kL have been deduced from the measured values of Em, a2,t,, and a. Therefore, their mode of variation with the experimental variables (L,, FA) must necessarily reflect their relationship to the measured quantities and the variation of the measured quantities with experimental variables. Correlations for both columns have been presented on the same figures to facilitate some comparison. Figure 2 shows that higher values of E M L have been obtained for the smaller column. This is to be expected since liquid residence time and therefore gasfliquid contact are higher for the smaller column, Figure 3. Also mixing is higher on the smaller column, Figure 5. Higher interfacial area values were also obtained for the smaller column, Figure 8. Since the smaller tray is more efficient than the larger tray and a is calculated from the rate of mass transfer on the trays, higher a values are therefore expected on the smaller tray. This result is in agreement with the observation of Thomas (19641, which is that, in general, small columns tend to support higher froth heights. In consequence, since higher froth heights (or volumes) have more mass-transferring gas pockets, larger a values are therefore expected on small trays. Literature data on sieve trays reflect very much the perculiarities of the experimental setups from which they have been obtained. This has made the comparison of the
data from the present work with those from the literature difficult. The smaller tray of the present work was similar in description to that employed by Thomas and Haq (1976) except that theirs had 3/s-in. diameter perforations. Similar fluid systems too have been employed. It is encouraging to see that the data from the two sources compare very well. The data on the larger tray, however, show some degree of difference. The difference can only be attributed to tray size and geometry. Conclusion Within the range of operability, gas flow rate does not significantly affect the mass transfer characteristics of sieve trays. Large-hole trays exhibit characteristics similar to small-hole trays; hence large-hole trays will be an advantage in operations where rust or solid deposition may block the perforations. Higher efficiency values have been reported for the smaller column. This is as a result of higher mixing which normally takes place on small trays. As a consequence, some scale-up criteria are called for before data obtained on experimental columns can be employed for designing industrial trays. Nomenclature A = active area of tray, cm2 A, = total interfacial area, cm2 a = interfacial area per unit volume of froth, cm-' Bo = concentration of NaOH, g-mol/L Ci = concentration of C02 at interface, g-mol/L D,= diffusivity of C02 in solution, cm2/s E M L = Murphree tray efficiency in liquid terms EOG = overall point efficiency in vapor terms FA = u, (P,,)'/~ = kinetic energy factor, (lb/ft s2)1/2 f ( t ) = tracer concentration distribution function f ( 6 ) = dimensionless distribution function G , = gas flow rate, g-mol/cm2s k L = liquid film mass transfer coefficient, cm/s k 2 = second-order reaction rate constant, cm3/mol s L , = liquid flow rate, g-mol/cm2s L, = liquid flow rate per foot of weir length, gal/min ft weir m = slope of equilibrium line N L = number of liquid film transfer units R = mass transfer rate per unit interfacial area, g-mol/cm2 S
R, = total mass transfer rate, g-mol/s
t = liquid residence time, s t, = mean liquid residence time, s u, = gas superficial velocity based on tray active area, ft/s x = mole fraction of transferring component in the liquid phase xi,, = mole fraction of C02 at tray inlet xOut= mole fraction of C 0 2 at tray exit xeqm= mole fraction of CO at equilibrium y = mole fraction of transferring component in the gas phase Zf = height of froth above tray, cm Greek Letters 6 = dimensionless time = time based variance of residence time distribution, s2 Q = dimensionless variance of residence time distribution pv = gas density, lb/ft3 = m(Gm/Lm)
.I:
Literature Cited A.1.Ch.E. "Bubble Tray Design Manual"; New York, 1958. Barrett, P. V. L. Ph.D. Thesis, University of Cambridge, Cambrklge, England, 1966. Danckwerts, P. V. "Gas-LiauM Reactions": McGraw-Hill: London, 1970; Chapter 5. Eben, C. D.; Pigford, R. L. Chem. Eng. Sci. 1965, 20, 803. Foss, A. S. Ph.D. Thesis, University of Delaware, Newark, DE, 1957. Foss. A. S.:Gerster. J. A.: Plaford. R. L. AIChE J. 1956. 4 . 231. Harris, I.J.1 Roper, 'I. Can. C&m. Eng. 1962, 40, 245. Harris, 1. J.; Roper, I.Can. J. Chem. Eng. 1863, 4 1 . 158. Lemieux, E. J.; Scotti, L. J. Chem. Eng. Frog. 1969, 65(3), 52.
222
Ind. Eng. Chem. Process Des. Dev. 1982, 27, 222-231 Thomas, W. J.; OgboJa,0. Ind. Eng. Chem. Process D e s . D e v . 1970. 77, 429.
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.
Received for review January 5 , 1981 Accepted September 18, 1981
509.
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 polymerization type 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).
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
*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
0
1982 American Chemical Society
