Catalyst and reverse disproportionation process

A catalyst containing tungsten, an alkali or alkaline earth component, and iron on a support, preferably silica gel, is disclosed which is useful in reverse disproportionation of stilbene and ethylene to produce styrene.

BACKGROUND OF THE INVENTION 
1. Field of the Invention 
The present invention relates to a catalyst and a process for reverse 
disproportionation of ethylene and stilbene to produce styrene. The 
catalyst comprises tungsten, potassium, and iron on a silica support. 
2. Description of the Prior Art 
The production of styrene from stilbene and ethylene is disclosed in U.S. 
Pat. No. 3,965,206, the teachings of which are incorporated by reference. 
Use of conventional disproportionation catalysts such as cobalt molybdate 
on alumina, or tungsten oxide or silica, alumina or silica-alumina, for 
reverse disproportionation is taught. 
U.S. Pat. No. 3,764,635, Fattore, et al, the teachings of which are 
incorporated by reference, teaches a process for disproportionating 
olefins using a catalyst of tungsten and bismuth on a support, preferably 
silica. The catalyst is active for disproportionation without any 
activation step. 
U.S. Pat. No. 3,792,107, Fattore, et al, the teachings of which are 
incorporated by reference, discloses use of a catalyst of tungsten and 
copper or tungsten and Group VIII metals, preferably Fe, Co, or Ni, or 
silica or other support. It is claimed that this catalyst requires no 
activation before use in disproportionation. 
U.S. Pat. No. 3,728,414, Helden, et al, the teachings of which are 
incorporated by reference, teaches a conventional olefin 
disproportionation catalyst with a promoter, a Group IIIa metal on an 
alumina carrier. Conventional olefin disproportionation catalysts are said 
to contain titanium, vanadium, chromium, manganese, zirconium, niobium, 
molybdenum, technetium, ruthenium, rhodium, palladium, tin, hafnium, 
tantalum, tungsten, rhenium, osmium, and iridium. This reference teaches 
that additional components, e.g., coactivators, hydrogenating components, 
components for isomerization of the double bond, and the like may also be 
added. Coactivators listed include cobalt oxide, and compounds of iron, 
nickel, and bismuth. 
U.S. Pat. No. 4,192,961 teaches conversion of a mixture of dibenzyl and 
stilbene with ethylene in the presence of a catalyst of chromium oxide, 
tungsten oxide, an oxide of an alkali metal and silica or alumi-silicate. 
Styrene yields of 78 to 80 wt %, based upon conversion of ethylbenzene, 
dibenzyl and stilbene, are claimed. 
U.S. Pat. No. 3,658,930, Kenton, et al, the teachings of which are 
incorporated by reference, teaches disproportionation of olefins using a 
rhodium oxide promoter on conventional olefin disproportionation catalyst, 
e.g., tungsten, molybdenum, rhenium, or tellurium on silica. 
U.K. Patent specification No. 1,205,677 teaches disproportionation of 
olefins using a conventional catalyst, such as molybdenum trioxide, 
tungsten trioxide or rhenium heptoxide on alumina, silica, or 
alumina-silica, and incorporating into this conventional catalyst a second 
component to effect double bond isomerization of olefins. Group VIII noble 
metals are suggested as being suitble, with preferred isomerization 
catalysts containing platinum and especially palladium. An alkali or 
alkaline earth metal ions are added to the catalyst to serve as a base to 
inhibit the oligomerization of branched chain olefins. 
None of these prior art catalysts are believed to possess sufficient 
activity and stability to permit their use in a commercial reverse 
disproportionation process. 
Although it is desirable to have a catalyst which can be successfully 
reactivated at approximately the same temperature used for the reverse 
disproportionation reaction, the choice of a commercial catalyst must 
depend on other factors as well. A very significant factor is the cost of 
the catalyst. Same promoters, such as noble metals, are expensive. 
Sometimes the increased productivity of a catalyst, which requires a high 
temperature activation more than justifies the increased costs associated 
with wear and tear on equipment and catalyst due to swings in temperature. 
Similarly, a longer cycle length, between regenerations is important in a 
commercial unit. 
Use of iron, as a catalyst additive, significantly increases the ability of 
the catalyst to effect the reverse disproportionation reaction for 
extended periods of time, the iron slightly reduces productivity during 
the start of a run, but at the end of a 12-hour run the iron-promoted 
catalyst has a higher productivity than standard catalyst. 
SUMMARY OF THE INVENTION 
The present invention provides a catalyst comprising catalytically 
effective amounts of iron, tungsten, and an alkali or alkaline earth 
component on a carrier material. 
In another embodiment, the present invention provides a process for the 
reverse disproportionation of stilbene and ethylene which comprises 
contacting stilbene and ethylene at reverse disproportionation conditions 
with an activated catalyst containing iron, tungsten, and an alkali or 
alkaline earth component or compounds thereof supported on a carrier 
material, to produce styrene. 
In a more limited embodiment, the present invention provides a process for 
the reverse disproportionation of stilbene and ethylene into styrene 
comprising contacting the stilbene and ethylene at temperature of 300 to 
600 C. with an activated catalyst comprising tungsten, iron, and an alkali 
or alkaline earth metal or compounds thereof on silica gel carrier, and 
wherein the atomic ratio of iron to tungsten is from 1:20 to 2:1, to 
produce styrene, and continuing said contact until said catalyst has been 
at least partially deactivated by coke deposition, removing said 
deactivated catalyst from contact with reactants and regenerating said 
catalyst by oxidizing coke from said catalyst with an oxygen containing 
gas to produce an oxidized catalyst with reduced coke content and 
thereafter activating said catalyst by contacting said oxidized catalyst 
with activating gas at 400 to 600 C. for a time sufficient to activate 
said catalyst, and thereafter returning said catalyst to contact with 
stilbene and ethylene for further reverse disproportionation of stilbene 
and ethylene into styrene. 
DETAILED DESCRIPTION 
The Reverse Disproportionation Reaction 
The total reaction of this invention may be represented by the following 
equation: 
EQU C.sub.6 H.sub.5 CH.dbd.CHC.sub.6 H.sub.5 +CH.sub.2 .dbd.CH.sub.2 
.revreaction.2C.sub.6 H.sub.5 CH.dbd.CH.sub.2 
Catalyst 
The catalyst may contain from 0.1 to 10 wt % W, preferably 1 to 6 wt %, and 
0.005 to 5 wt % Fe, preferably 0.05 to 2 wt %. The catalyst also has 0.01 
to 2%, preferably 0.03 to 0.3 wt % alkali or alkaline earth metal ion, 
preferably potassium. Other promoters may be present. 
The support is preferably silica gel, but any other support used for 
conventional disproportionation catalysts may also be used, though the 
catalyst performance may change some. 
Reaction Conditions 
The reverse disproportionation reaction conditions are given in U.S. Pat. 
No. 3,965,206, this teachings of which are incorporated by reference. In 
general, temperatures of 300 to 600 C. are adequate. Pressures from 
subatmospheric to 1000 atm, absolute are suitable, but operation at 1 to 
10 atmospheres gives good results. 
Hydrogen or nitrogen or inerts may be present during the disproportionation 
reaction. Nitrogen is believed inert. Adding hydrogen may, or may not, cut 
down on catalyst coking, and may overreduce the catalyst. Nitrogen, and 
hydrogen and other inert gases will also cut down residence time of 
reactants in the reactors, if desired. I prefer to operate with the 
reactants as the sole feed to the reactor. The feed to the process of the 
present invention consists of relatively pure stilbene and ethylene. Other 
materials may be present, but polar materials act as catalyst poisons.

EXAMPLES 
Reactor 
The experimental apparatus used in all examples consisted of a 0.5-inch OD 
stainless steel tube, 18-31 cm long. The catalyst was maintained in the 
reactor as a fixed bed. Reactants flowed in a vapor phase, down flow, 
through the catalyst bed. The catalyst was supported on a quartz wool plug 
resting on an inert support. During the early phases of the study, 
1/8-inch alundum beads were used, but experiments showed that this 
material was not inert and caused some coking. The later studies were 
conducted using 1/8-inch long quartz billets cut from 2 mm rod as a 
support. 
Special precautions were taken to exclude oxygen from the apparatus and to 
keep the stilbene feed in the vapor phase. Special steam tracing, heating, 
and nitrogen purging of lines contacting stilbene are essential in a pilot 
plant, but may not be as critical in a large scale commercial plant. 
Catalyst Preparation 
A series of catalysts was prepared. The basic catalyst contained 0.56 wt % 
WO.sub.3 and 0.038 wt % K.sub.2 O. The catalyst was prepared by adding 20 
g of 14-35 mesh Davison Grade 59 silica gel, which had been freshly 
calcined, to 28 ml of a solution containing 5 ml of 0.0236N KOAc solution, 
10 ml of H.sub.2 O and 13 ml of concentrated NH.sub.4 OH. The silica gel 
was only minimally wetted by the 28 ml of liquid. The mixture was shaken 
for 30 minutes, then dried overnight in a stream of air on a filter, and 
finally calcined for 2 hours at 600 C. Various additives, those which were 
soluble in the alkaline solution described above, were simply added to the 
alkaline solution along with the potassium and tungsten components. In 
some cases, because of solubility limitations, ammonium hydroxide would 
not dissolve the additive, so in these cases a few drops of concentrated 
HNO.sub.3 was added to obtain a clear solution. In all cases the total 
liquid volume of impregnating solution was 28 ml, the exact volume was 
obtained by adjusting the amount of water added. In all cases, except 
where noted, additives were added sufficient to give an atomic ratio of 
tungsten: additive of 5:1. I believe the additives, the added metallic 
components, were present as oxides on the catalysts, because of the 
calcination in air for two hours at 600 C. 
When rhodium was added, a different procedure was used as no water soluble 
rhodium compound was readily available. A large batch of base catalyst 
(containing 0.56 wt % WO.sub.3 and 0.038 wt % K.sub.2 O) was made up as 
described above. A 20.12 g portion of this catalyst was then impregnated 
with 25 ml of a methanol solution containing 0.0375 g of Rh (acac). This 
alcoholic impregnating solution was sufficient to just impart wetness to 
the catalyst. After shaking for 30 minutes, drying in air, and calcining 
for 2 hours at 600 C. the catalysts were ready for use. 
Table I shows a listing of catalysts prepared. 
TABLE I 
__________________________________________________________________________ 
PREATION OF 0.56% WO.sub.3, 0.038% K.sub.2 O CATALYSTS WITH 
ADDED METALLIC COMPONENTS 
Compound Wt. of 
Additive 
Used Compound.sup.a, g. 
Comments 
__________________________________________________________________________ 
Pt Pt(NH.sub.3).sub.2 (ONO).sub.2 
0.0330.sup.b 
Pd Pd(NH.sub.3).sub.2 (ONO).sub.2 
0.0232 5 drops HNO.sub.3, boiled to dissolve salts 
Ni Ni(NO.sub.3).sub.3.6H.sub.2 O 
0.0281 No NH.sub.4 OH 
Zn Zn(OAc).sub.2.2H.sub.2 O 
0.0212 No NH.sub.4 OH; 5 drops 30% H.sub.2 O.sub.2 
Cr Cr(NO.sub.3).sub.3.9H.sub.2 O 
0.0387 No NH.sub.4 OH; 5 drops 30% H.sub.2 O.sub.2 
Fe Fe(NO.sub.3).sub.3.9H.sub.2 O 
0.0390 No NH.sub.4 OH; 5 drops 30% H.sub.2 O.sub.2 ; 10 
drops HNO.sub.3 
Ru RuNO(NO.sub.3).sub.3 
0.0306.sup.c 
5 drops 30% H.sub.2 O.sub.2 
Mo (NH.sub.4).sub.6 MO.sub.7 O.sub.2 4.4H.sub.2 O 
0.0171 
V NH.sub.4 VO.sub.3 
0.0114.sup.d 
Sn SnSO.sub.4 0.0218 No NH.sub.4 OH; 10 drops conc. H.sub.2 SO.sub.4, 
4 drops HNO.sub.3 
Re Re.sub.2 O.sub.7.3 (Dioxane) 
0.0362 NO NH.sub.4 OH; 5 drops 30% H.sub.2 O.sub.2 
Ag AgNO.sub.3 0.0164 
Ce Ce(NO.sub.3).sub.3.6H.sub.2 O 
0.0419 No NH.sub.4 OH 
Eu Eu(NO.sub.3).sub.3.6H.sub.2 O 
0.0431 No NH.sub.4 OH 
As As.sub.2 O.sub.5.NH.sub.2 O 
0.0127.sup.e 
No NH.sub.4 OH 
U UO.sub.2 (C.sub.2 H.sub.3 O.sub.2).sub.2.2H.sub.2 O 
0.0410 No NH.sub.4 OH 
Mn Mn(C.sub.2 H.sub.3 O.sub.2).sub.2.4H.sub.2 O 
0.0237 No NH.sub.4 OH 
Rh Rh(acac) 0.0375 Alcoholic impregnation. 
__________________________________________________________________________ 
.sup.a 20 g of silica gel base 
.sup.b 61.00% Pt 
.sup.c 35.87% Ru 
.sup.d 76.90% V.sub.2 O.sub.5 
.sup.e 87.65% As.sub.2 O.sub.5 
Catalyst Activation 
Catalysts were activated, in situ, by passing 200 scc/min of CO over the 
catalyst at a specified temperature for specified time. It is possible to 
use other activating gases, or no gas at all, but a CO activation 
procedure was chosen as a standard one to permit screening of the effects 
of various additives on catalyst activation. 
Test Procedure 
The activated catalyst was then tested for its activity on a standard feed 
consisting of 200 scc/min ethylene and 40 scc/min of stilbene. The 
residence time in the catalyst bed was 0.3 seconds. The products were 
analyzed by gas chromatography. 
After the catalyst lost activity, it was regenerated by contacting it with 
48 scc/min of air for 45 minutes at 575 C. 
A typical operating sequence is presented below: 
A. Activation Cycle 
1. 25 min. Nitrogen purge of lines and reactor system (200 cc/min.). 
Ethylene purge of line up to oxygen trap. Reactor temperature equilibrated 
to activation temperature. 
2. 5 min. Nitrogen purge continuing. Ethylene purge of lines through oxygen 
trap to vent, located at ethylene-to-saturator feed valve. CO flow to vent 
to purge CO line in panel control board. 
3. 55 min., typical. Nitrogen off. CO feed to reactor for activation, feed 
rate typically 200 cc/min. Ethylene purge to vent continuing, with oxygen 
meter (Teledyne Trace Oxygen Analyzer Model 311-1) connected to vent to 
monitor ethylene quality. 
4. 5 min. Nitrogen purge to vent to clear lines in control panel. CO feed 
to reactor and ethylene feed to vent continuing. 
5. 15 min. Nitrogen purge of reactor and lines. Ethylene feed to vent 
continuing. Temperature changed to disproportionation run temperature. 
6. 5 min. Ethylene feed to reactor, bypassing saturator. Saturator feed 
valve open to reactor to equalize pressure. 
B. Disproportionation Cycle 
1. 30 min. Ethylene feed through stilbene saturator and thence to reactor; 
GC sampling progam called during last 60 sec. of cycle. Step is repeated 
as desired. 
C. Burn-off Cycle 
1. 5 min. Nitrogen purged to vent to clear lines in control panel. Ethylene 
feed to saturator off, but saturator feed valve to reactor open to 
equalize pressure. 
2. 60 min. Nitrogen purge to reactor (48 cc/min). Temperature changed to 
burn-off temperature, usually 575 C. Air purged to vent to equilibrate 
pressure in line. 
3. 45 min. Air feed to reactor, 48 cc/min. GC analysis for CO.sub.2 called. 
4. 15 min. Nitrogen purge, 200 cc/min., through reactor system. 
5. Shut down or recycle. 
This procedure was used to test the different catalyst formulations. 
Experimental results are shown as productivity, measured as moles of 
styrene per liter of catalyst per hour. 34 moles per liter per hour 
represents about 83% conversion of stilbene to styrene. Productivity is 
reported both for the start of run conditions (initial) and at the end of 
the run, i.e., after 4.5 hours of operation (final). The data are 
presented below in Table IIA. 
TABLE IIA 
______________________________________ 
EFFECT OF ADDITIVES ON CATALYST ACTIVATION 
Activation 
Catalyst.sup.a 
Conditions Productivity.sup.b 
Additive Time Temp. Initial 
Final 
______________________________________ 
None (STD) None 4.19 6.74 
1 hr 450 8.05 12.65 
8 hr 450 27.65 19.43 
Ce 1 hr 450 8.13 12.30 
Eu None 4.01 6.27 
1 hr 450 6.91 12.57 
As None 2.16 3.99 
1 hr 450 8.62 12.91 
1hr 450 8.92 13.31 
Fe 1 hr 450 9.07 17.79 
Cr 1 hr 450 7.10 12.20 
Ni 1 hr 450 7.51 13.82 
Ru 1 hr 450 4.68 11.19 
Pt 1 hr 450 9.38.sup.c 
7.21.sup.c 
Pd 1 hr 450 7.29.sup.c 
6.35.sup.c 
______________________________________ 
.sup.a All additives at 5:1 W:additive mole ratio unless otherwise noted. 
.sup.b All runs were for 4.5 hr. Run temp., 425 C. 
.sup.c Average for 2 runs. 
The test apparatus was then partially dismantled and rebuilt. A number of 
additional tests were then run. The main difference between operations 
reported in Table IIA and Table IIB, presented hereafter, is the amount of 
oxygen contamination. I believe that the data presented in IIA reflect 
less oxygen contamination than those in Table IIB. Since the testing 
occurred under superatmospheric pressure, it was thought that there would 
be no air contamination due to leaks in the piping. Reactants might leak 
out, but air would not get in. Several ppm oxygen diffused into the test 
apparatus through a leak to increase the oxygen level, and decrease the 
catalyst activity. Oxygen is a catalyst poison. The amount of O.sub.2 
contamination was relatively constant during the IIA testing period, I 
estimate about 0.2 ppm O.sub.2 by volume. For the IIB testing period about 
0.3 ppm O.sub.2 by volume was present. I checked the activity of my 
standard, or reference, catalyst periodically during the IIA and IIB 
testing periods. The standard, or reference, catalyst consistently gave 
lower productivity during the IIB tests. The results of the more O.sub.2 
-contaminated runs are reported in Table IIB. All tests were conducted at 
about 3 psig, or 1.2 atm, absolute. 
TABLE IIB 
______________________________________ 
EFFECT OF ADDITIVES ON CATALYST ACTIVATION 
Activation 
Catalyst.sup.a 
Conditions Productivity 
Additive Time Temp. Initial 
Final 
______________________________________ 
None (STD) 1 hr 450 6.09 11.04 
1 hr 450 5.87 11.49 
1 hr 450 4.53 9.63 
Rh 1 hr 450 9.82 10.60 
1 hr 450 2.98 8.16 
1 hr 450 2.41 7.45 
8 hr 450 2.76 6.53 
Mo 1 hr 450 3.92 7.75 
1 hr 450 4.75 7.56 
V 1 hr 450 1.05 2.86 
Sn 1 hr 450 0.04 (0.18) 
1 hr 450 1.34 (1.61) 
Re 1 hr 450 4.49 4.26 
Zn 1 hr 450 2.82 4.32 
Ag 1 hr 450 3.42 8.02 
1 hr 450 4.32 8.48 
U 1 hr 450 4.98 5.47 
Mn 1 hr 450 3.76 4.04 
______________________________________ 
.sup.a All additives at 5:1 W:additive mole ratio unless otherwise noted. 
It is believed that results can be compared very well within Table IIA, or 
within Table IIB. Direct comparison of an additive listed in Table IIB 
with an additive from the Table IIA is harder to make, because of the 
increased oxygen contamination in those runs presented in Table IIB. It is 
believed that the relative activities, i.e., activity of a catalyst in 
Table IIB with an additive compared to activity of a catalyst with no 
additives from Table IIB can be compared. These data, relative activation, 
for initial activity, are reported in Table III. The relative activities 
are probably more significant than relative end of run activities, so 
comparisons were made based on initial activity. A single number means the 
catalyst was stable after repeated regenerations, multiple numbers means 
activity declined. 
TABLE III 
______________________________________ 
Additive Relative Activities 
______________________________________ 
Fe 1.13 
As 1.10 
Pt 1.09 
Ce 1.01 
Rh 1.01, 0.56, 0.45 
U 1.03 
None (Standard) 1.00 
Re 0.93 
Ag 0.89, 0.71 
Mo 0.89, 0.74 
Ni 0.93, 0.77 
Cr 0.88 
Eu 0.86 
Pd 0.85 
Mn 0.78 
Zn 0.58 
Ru 0.58 
Sn 0.31 
V 0.16, 0.17, 0.20, 0.20 
______________________________________ 
From these data, it is apparent that a reverse disproportionation catalyst 
containing iron is more active than the reference catalyst. These results 
were obtained with a relatively mild activation, 450 C. This relatively 
low temperature was picked because the reverse disproportionation reaction 
runs at 450 C. I decided to conduct further tests to see what effect a 
higher temperature activation step had on catalyst activity and life. 
Catalyst life may be a very significant factor because the presence of 
polars, such as oxygen, hurts catalyst activity. 
Oxygen is a reversible but strong catalyst poison. I tried to eliminate as 
much oxygen as I could from my test apparatus; however, the lowest oxygen 
level consistently obtainable was 0.3 ppm at this point in the testing 
program. This O.sub.2 level was estimated from measurements made at the 
reactor outlet using the Teledyne Trace Oxygen Analyzer. 
Oxygen contamination is a very serious problem in a pilot plant because the 
sizes of the feed streams are relatively small compared to the leaks that 
may be present in the plant. This may not be as much of a problem in a 
commercial unit, however, it may be a significant problem commercially if 
it is not possible to remove sufficient oxygen or polars from the feed. 
To accomodate this eventuality, or to design a process for use in reverse 
disproportionation of feed streams which contain significant amounts of 
oxygen contamination, or other polar contaminants, I tried to find a 
catalyst which would withstand relatively more catalyst poisons than could 
the prior art systems. 
In this series of tests, a slightly different activation procedure was 
used, all catalysts were activated at 575 C. for one hour. This is a much 
more severe activation step than the 450 C., one hour activation used in 
most of my prior work. The reason for using the higher activation 
temperature was to give the catalyst maximum activity. In the prior tests, 
using the 450 C. activation, the purpose of the experiments was to find a 
catalyst which could be sufficiently activated at a temperature near that 
of the reverse disproportionation reactor. 
Under the conditions of the reactor flow described above, the feed to the 
reactor zone contained about 0.3 ppm volume of oxygen. Results of these 
experiments are reported to the Tables IVA, B, and C. 
TABLE IVA 
______________________________________ 
EFFECT OF ADDITIVES WITH 
HIGH TEMPERATURE ACTIVATION 
Catalyst.sup.a 
Productivity.sup.b 
Coke 
Additive Initial Final %, Cat 
______________________________________ 
Fe 32.40 29.42 0.25 
None (STD) 33.73 26.48 0.43 
Pt 32.90 22.47 0.37 
Ce 32.87 25.29 0.25 
As 31.52 25.14 0.24 
Ru 30.94 26.95 0.23 
Cr 28.86 22.61 0.35 
Ni 21.68 24.89 0.16 
21.77 24.01 0.25 
Eu 30.43 20.93 0.27 
______________________________________ 
TABLE IVB 
______________________________________ 
Catalyst Productivity Coke 
Additive Initial Final %, Cat 
______________________________________ 
None (STD) 33.53 25.86 0.06 
V 11.47 7.78 -- 
Mo 15.88 10.24 -- 
Sn 16.3 -- -- 
Re 30.85 9.46 0.19 
Zn 12.44 10.13 0.15 
______________________________________ 
TABLE IVC 
______________________________________ 
Catalyst Productivity Coke 
Additive Initial Final %, Cat 
______________________________________ 
None (STD) 30.91 13.20 -- 
28.94 12.49 -- 
Ag 28.62 13.11 0.13 
U 27.09 10.42 0.06 
Mn 20.02 7.84 0.09 
______________________________________ 
.sup.a All additives at 5:1 W:additive mole ratio unless otherwise noted. 
.sup.b Run for 12 hrs, Run Temp, 425 C; Feed Rate, 200 scc/min ethylene. 
40 scc/min SB. Res. Time, 0.3 sec. 
Activation, 1 hr at 575 C. 
Productivities in moles of Sm/L/hr; a productivity of 34 m/1/hr. 
.apprxeq.83% conversion. 
As was previously discussed, data should only be compared within each 
table, as it is believed that oxygen levels increased somewhat between the 
data presented in Tables IVA, IVB, and IVC. Data in Table IVA and Table 
IVB represent operation with about 0.3 ppm volume oxygen. Data from Table 
IVC represent an oxygen level of about 0.5 ppm volume oxygen. Although the 
Fe-promoted catalyst had a slightly lower initial activity than the 
reference catalyst, the iron-promoted catalyst had a final productivity of 
29.42 after 12 hours, compared to the reference catalyst with a 
productivity of 26.48. 
The iron-promoted catalyst also had a relatively low coke on catalyst level 
at the end of the 12 hour run, 0.25 wt % coke on catalyst as compared to 
0.43 wt % coke on the reference catalyst. Because of small sample sizes 
and difficulties encountered in reproducing other coke numbers. I do not 
have as much confidence in the reported coke levels as I do on the 
reported productivity levels, however, the reduced coke level is 
consistent with increased activity. Somewhat contradictorily, a number of 
other catalysts had poor initial and final productivities and yet also had 
relatively low coke levels on the catalyst. 
Also of interest, the iron-promoted catalyst does not show its outstanding 
productivity when subjected to relatively mild activation. Comparing 
catalysts promoted with Fe with the standard catalyst using a 450 C. 
activation for 1 hour with CO gas, showed some improvement of initial 
activity, with significantly enhanced activity later on, although none of 
the activities of either the standard catalyst or the iron catalyst are 
considered sufficiently high to be viable commercial catalysts. The 
superiority of the iron-promoted catalyst over conventional catalysts 
becomes apparent with the more severe activation step. 
I am not sure what the optimum amount of iron is. I know a 1:5 Fe:W atomic 
ratio gives good results. 
I believe that good results can be obtained with Fe:W atomic ratios of 1:20 
to 2:1 and preferably 1:10 to 1:3. 
If I were designing a commerical plant today, I would conduct further 
experiments to see if the various catalytic components could be optimized 
further. 
I would probably use a catalyst containing 2 to 10 times as much metal 
content as those catalysts used in the experiments. Commercially you want 
more active catalysts, and smaller reaction vessels, and would use 
catalysts with a higher metal loading. I used very lightly loaded 
catalysts for my experiments because the catalyst was extremely active. 
"Full strength" catalyst established equilibrium conditions so rapidly 
that I could not discern relatively smaller differences caused by 
different additives. Based on other experimental work, metal loadings ten 
times as high as probably be achieved using similar impregnation 
procedures, with five or tenfold increase in activity. Phrased another 
way, the reactants see the active metals, not the support, and the amount 
of conversion per gram of catalytic components (excluding support) is 
roughly constant. More metal on the support should provide more resistance 
to poisoning by trace amounts of O.sub.2 and polars. 
I would like to learn more about the active form of the catalysts I tested. 
The active form may be a simple oxide or may be a mixed heteropolyacid of 
SiO.sub.2, WO.sub.3 and MO.sub.x, where M is the additive metal. It is 
possible that the oxides mentioned and claimed do not exist as discrete 
oxides, but instead form some complex polymeric structure. 
Commercially, I would operate the plant with whatever oxygen stripping 
columns or oxygen and water absorbers were necessary to ensure oxygen, and 
other contaminants, especially polar ones, were excluded from the plant. 
My catalyst can be disposed within the reactor as a fixed bed, fluidized 
bed, moving bed, ebullating bed, or any other reactor configuration. The 
advantage of the fluidized, moving and ebullated bed reactors is that 
catalyst addition and withdrawal can be performed continuously. Thus, 
coke, or carbon deposition on the catalyt can be burned off, the catalyst 
activated, and returned to the reactor without shutting down the reactor. 
The disadvantage of this mode of operation is that the reactor designs are 
fairly complicated, as compared to simple fixed-bed down-flow design. When 
fixed-bed reactors are used, preferably, two or three reactors are 
provided in parallel, permitting one or more reactors to be taken off 
stream for carbon burn-off and activation while the other reactor(s) 
remain on stream.