Note: Descriptions are shown in the official language in which they were submitted.
A PROCESS FOR THE DISPROPORTIONATIO~ OF ALKE~ES
The present invention relates to the catalytic
di~proportionation of alkenes. Speci~ically, this
f invention relates to an improvement in the catalytic
5 disproportionation of alkenes by use of a distillation
column reactor.
According to this invention, the term
disproportionation refer~ to the conversion of a
10 hydrocarbon into similar hydrocarbons of both higher
and lower numbers of carbon atoms per molecule. In the
case of alkenes, a mixture of new products is obtained
comprising alkenes of both higher and lower molecular
weights. Such an operation is useful in many
15 instances. For example, a more plentiful hydrocarbon
can be converted to a less plentiful and, therefore,
more valuable hydrocarbon. One instance of such a
conversion occurs when the process of this invention is
used to balance the alkene production of a naphtha
20 cracking plant by disproportionating the large
30,442B-F -1-
--2--
quantities of butenes into ethylene and hexene, or
propylene and pentene. The disproportionation of
butenes is a particularly valuable disproportionation
reaction for the use of excess butenes. Approximately
5 equimolar quantities of the higher and lower molecular
weight alkenes may be produced by such
disproportionation reactions. The higher molecular
weight alkenes produced by disproportionation may be
oracked to yield additional ethylene or propylene.
Much of the prior art describes conven~ional
processes for the disproportionation of alkenes.
Typically, the conventional process is carried out
batchwise or in a continuous manner~ using the catalyst
in the ~orm o~ a fixed bed, a fluidized bed or a moving
bed. At the end of the reaction period, the
hydrocarbon phase is separated from the solid catalyst
phase and the hydrocarbon products are recovered.
Well-known techniques, such as fractional distillation,
solvent extraction and adsorption are employed for the
separation of the hydrocarbon products.
Conventional processes for the
disproportionation of alkenes, such as those described
hereinabove, may be operated at temperatures as high as
500C, but are unable to achieve simultaneous high
conversion and high selectivity to desired products.
See, e.g., Banks, R. L., J. Molecular CatalYsis~ V. 8,
pp. 269-276 (1980). Alkene disproportionation
reactions are reversible; therefore, theoretically the
maximum conversion which can be achieved is limited by
the thermodynamic equilibrium. In the dispropor-
tionation of butene-l to ethylene and trans-hexene-3,
for example, the conversion at equilibrium is
approximately 50 percent. Selectivity is normally
30 9 442B-F -2-
--3--
controlled by the catalyst and by process conditions,
such as temperature, pressure and residence ti~e. In a
fixed bed reactor, the residence time is determined by
the ~eed rateO In the disproportionation o~ butene-1,
for example, a slow feed rate will result in a longer
residence time for the propylene product, eventually
leading to propylene disproportionation and the
production of ethylene. On the other hand, a high ~eed
rate will result in a short residence time for
propylene and hence less ethylene in the product mix.
Unfortunately, high feed rates reduce conversion; thus
it is difficult to produce high conversion and high
selectivity to propylene simultaneously. In addition,
' 15 despite the numerous methods of controlling process
conditions, many conventional disproportionation
processes give broad product distributions, including
the by-products of isomerization and secondary
disproportionation reactions Several illustrations of
the prior art and its inherent limitations are
presented hereinbelow.
U.S. Patent 3,261,879 (1966) discloses a
disproportionation of olefin hydrocarbons by contact
with a catalyst containing molybdenum oxide or tungsten
oxide in a conventional reactor. Conversions are
taught to vary over a wide range; however, at high
conversion a broad distribution of C2_12 olefinic
products is shown. Isomerization yields are taught to -
3 be high.
U.S. Patent 3,463,827 (1969) discloses a
process for the disproportionation of ole~ins by
contact with a Group VIB metal carbonyl associated with
alumina, silica or silica-alumina. The process is
conducted in a fixed bed reactor, and the products are
30,442B-F -3-
_4_
separated in a fractionation col~mn. Conversions are
taught to be low, less than 20 percent for bu~ene-1,
and isomerization tends to be high.
U.S. Patent 3,448,163 (1969) describes a
conventional process for the disproportionation of
olefin~ employing a catalyst comprising rhenium
heptoxide impregnated on alumina. In the
disproportionation of butene-1, the combined
selectivity to ethylene and hexenes is taught to be 93
percent at a converxion of only 29 percent.
U.SD Patent 3,641,189 (1972) discloses a
conventional olefin disproportionation process ~~ 15 utilizing a rhenium heptoxide catalyst supported on
alumina. High selectivities are accompanied by low
conversion of the feedstock.
U.S. Patent 3,642,931 (1972) teaches a
conventional olefin disproportionation process
employing a catalyst comprising rhenium heptoxide
supported on a refractory oxide of zirconium, thorium,
tin, or mixtures thereo~. The degree of
~isproportionation is taught to be less than 15
percent.
U.S. Patent 3~676,520 (1972) discloses a method
of disproportionating olefins by contacting the olefin
with a catalyst comprising rhenium oxide and a support,
3 such as alumina. The process is carried out in any of
the aforementioned ~tandard reactors.~ In the
disproportionation of propylene, the conversion is
taught to be less than 5 percent at high contact
temperature. There are no teachings on how to control
the selectivities of higher olefins, such as butenes.
30,442B-F _4_
- 5 - 64693-4101
In view of the deficiencies of the prior art methods, it
would be desirable to provide a process for the disproportionation
of alkenes which would be capable of achieving simultaneous high
conversion and high selectivity at moderate temperatures. It
would also be desirable to provide a process for the dispropor-
tionation of butenes, such that a high yield o~ ethylene or
propylene, and the corresponding hexenes or pentenes, could be
achieved by a simple adjustment of the operating conditions. Such
a process would easily meet the demands for varying olefin feed-
stocks. It would also be desirable to obtain the disproportion-
ation and the separation of products simultaneously, since any
reduction in the number of process steps offers cansiderable
economic advantages.
The present invention i5 such a process for the
disproportionation of alkenes, involving contacting at least one
alkene with a catalyst con~aining at least one rhenium compound or
; complex in a distillation column reactor under such reaction
fractionation conditions that there is formed at least one product
of the disproportionation of at least one of the alkenes. By use
of the distillation column reac~or, this disproportionation and
the separation of products are advantageously conducted simultane-
ously. Surprisingly, the process of the present invention
proceeds with high selectivlty at high conversion at moderate
temperatures. More surprisingly, the process of the present
invention, as applied to butenes, can be controlled to give high
yields of ethylene or propylene, whlchever is desired, and *he
corresponding hexenes and pentenes by simple adjustments in the
operating conditions. By all of the aforementioned accomplish-
~3
6 - 6~693-4101
ments, the present invention satisfies a long-felt need for
improvement in ~he art of disproportiona~ing alkenes.
Alkenes which are subject to disproportionation
according to the process of the present invention include acyclic
alkenes having at least 3 carbon atoms, and their aryl derivatives
and mixtures thereof. Preferred are alkenes having from 3 to 30
carbon atoms and mixtures thereof. More preferred are mono-l- and
- 2-alkenes, such as, ~or example, butene-l, and mixtures of these
alkenes, such as, for example, a mixture o:~ butene-l and butene-2.
Most preferably, ~he process of the present invention is applied
to butene-l or a mixture of butene-1 and butene-2. Optionally,
an inert material may be included in the alkene fed to the
distillation column reactor. Examples of said inert materials
include nitrogen, the inert gases of Group VIIIA, such as helium
and argon, and alkanes, such as methane, propane and hutane.
Catalysts suitable for use in ~he process of the present
invention are those materials which catalyze the disproportion-
ation reaction when used in the process of the present invention,
and include conventional catalysts used for the disproportionation
of alkenes. Examples of said conventional catalysts include
supported materials which contain rhenium, and which optio~ally
include a promoter r such as tetramethyltin or tetrabutyltin.
The catalysts comprise rhenium or a rhenium compound or complex
having an alumina support. Rhenium oxides are preferred for use
in the catalyst of the process of the present invention. The
support materials can be in a variety of forms, and may contain
other materials
' ~;B
which do not substantially promote undesirable side
reactions. Pre~erably, any conventional catalytic
grade o~ alumina or silica-alumina may be used as the
support. Gamma-alumina is the most preferred support
material.
The composite catalyst is prepared by suitable
methods such as dry mixing, impregnation or
coprecipitation. Catalytic metal oxides or compounds
convertible to catalytic metal oxides by calcination
are suitably employed in the catalyst preparation. A
convenient method for the preparation of the catalyst
is to dry blend the catalytic metal oxide, such as
rhenium oxide, and the support in a ball mill where
intimate contact between the finely divided particles
is achieved. The milled composite can be pressed into
pellets or tablets of various sizes and shapes.
Additionally, the ~inished catalyst may be in the form
of granules as well as in other shapes, such as, for
example J agglomerates, spheres and extrudates, or it
may be employed in such conventional distillation
packing shapes as, for example9 Raschig rings and
saddles. If desired 7 pelleted catalysts can be crushed
to obtain particles having specific mesh size.
After the catalytic metal oxide or compound
which may be converted to a catalytic metal oxide by
calcination is associated with the support, the
composite is subjected to a calcination or activation
step before being utilized in the olefin conversion
process. The activation technique comprises heating at
elevated temperatures in the presence of a suitable
flowing gas. Air is a preferred activation gas,
although other gases, for example, inert gases SUC}l as
nitrogen or the noble gases, may be used, provided that
30,442B-F -7-
--8--
at least part of the catalytic metal present in the
catalyst composition is in the oxide form at the
completion of the activation. In some instances, the
catalyst may be heated serially in more than one gasO
The catalysts are subjected to a temperature which is
generally in the range of from 300C to 700C for 0.5 to
20 hours or longer. Generally, longer activation
periods are used with lower temperatures, and shorter
activation periods are used with higher temperatures.
Either way, the selectivities to the disproportionation
products are the same.
The activated catalyst may be used, without
regeneration, for runs of up to several days or more,
r 15 and may be regenerated. The regeneration is
accomplished by suitable methods ~or regenerating oxide
catalysts and may comprise the same steps used in the
activation procedure.
The distillation column reactor employed in the
present invention is suitably a distillation column
having therein a catalyst for the disproportionation o~
alkenes. Any type of distillation tower may be
employed in the process of the present invention,
provided that a fixed bed of catalyst may be created
therein to fill the reaction-distillation zone or
portions thereof. The catalyst packing is of such a
nature as to allow vapor flow through the catalyst,
while also providing sufficient surface area to
catalyze the disproportionation reaction.
The distillation column reactor is operated so
as to disproportionate the alkenes in the feed stream
and separate the products therefrom simultaneously.
Accordingly, the reactor may be operated under any
30,442B-F -8-
~onditions, e.g., temperature and pressure, at which
disproportionation is achieved. Thus, the temperature
within the column will be related to the boiling point
of the alkene starting material, and to the pressure in
the column. Typically, the operating temperature in
the distillation column reactor may range from -50C to
300C, and preferably will be from 0C to 150C. The
pressure in the distillation column reactor typically
will be from zero to lO00 psig (6900 kPa gauge) and
1~ preferably will be from 50 to 30Q psig (350 to 2100 kPa
gauge). Higher or lower temperatures and pressures may
be employed; however, beyond the lower end of the
range~ the reaction will proceed slowly, if at all, and
,- 15 beyond the higher end of the range, undesirable side
reactions and coke formation may occur. Additionally,
it will probably be more expensive to operate outside
the ranges given.
The process of the present invention is a
method for the simultaneous disproportionation of
alkenes and d;stillation of the product mixture as it
forms. For example, when butene-1 is fed to the
distillation column reactor, it contacts the catalyst
and is selectively disproportionated to ethylene and 3-
hexene. The ethylene immediately upon formation
ascends through the distillation column reactor, and
the 3-hexene descends in accordance with conventional
principles of distillation. Consequently, the reverse
3 reaction, whereby 3-hexene and ethylene combine to form
1-butene, does not occur. Thus, in contrast to prior
art disproportionation methods, which were capable of
achieving high conversion with low selectivity or high
selectivity at low conversion, the process of the
30,442B-F -9-
--10--
present invention surprisingly is capable o~ obtaining
simultaneous high conversion and high selectivity.
For the purposes of this invention, the term
conversion refers to the elimination of the alkenes in
the feed stream from the reaction mixture. For
example, in the practice of this invention, butene-1
may be converted substantially to ethylene and hexene-3
under the proper conditions using a dlstillation column
reactor wherein butene-l is held in the central portion
of the column and is not allowed to escape. For the
purposes of the present invention, the term selectivity
refers to the percentage of the converted feed which
goes to the desired major products.
The concept of simultaneous high selectivity
and high conversion may be expresscd conveniently in
terms of yield. For the purposes of the present
invention, the term "yield" refers to the numerical
product of conversion and selectivity (yield =
conversion x selectivity). For example, a process
according to the present invention operating at a
conversion of 0.75 and a selectivity of 0.90 would have
a yield of 0.675, which is the numerical product of
0.75 and 0.90. The process of the present invention
may be operated to give higher yields than prior art
disproportionation processes. Typical yields of the
process of the presen~t invention are at least 65
percent, based upon moles of alkene in the feed stream.
Pre~erably, the yield will be at least 75 percent.
Most preferably, the yield will be at least 85 percent.
A great deal of control over the rate of
reaction and the distribution of products can be
achieved simply by adjusting the operating conditions
30~442B-F -10-
of the reaction system. For example, the temperature
in the system may be increased by increasing the
pressure, and the feed rate may be adjusted to control
the percent conversion. The process of the present
invention is additionally advantageous in that a great
degree of control may be exercised over the percentage
of propylene and ethylene produced simply by adjusting
the residence time of the process. It is known that
the residence time may be altered by changing any of a
0 number o~ process variables, such as the feed rate, the
length or height of the distillation column reactor,
and the overhead temperature. When a low overhead
temperature is maintained, propylene remains in contact
1~ with the catalyst for a Ionger time 9 eventually
disproportionating to ethylene and butene-2.
Alternatively, when a high overhead temperature is
maintai~ed, propylene lea~es the column faster. Thus,
~or any fixed column length, the selectivity to
ethylene or propylene can be controlled to a large
extent by simply controlling the overhead temperature.
In either case, the conversion is high in the
distillation column reactor; thus, high yields o~
ethylene or propylene can be produced, as desired, to
meet the varying olefin feedstock requirements.
The number of theoretical trays, the
fractionation conditions, including temperature and
pressure, the reflux and reboil control system, the
3 number and location of side streams, the flow rate,
etc., are those which are used in conventional
engineering practice and can be determined by
conventional design calculations and procedures. The
fractionating apparatus employed in the process of the
30,442B-F -11-
q~
-12-
present invention may be operated using known process
control techniques.
When an alkene or a mixture of alkenes is fed
to a distillation column reactor under conditions
previously described herein, the alkene or alkenes will
be disproportionated to products having boiling points
lower and higher than the boiling point of the alkene
feed stream. Further, the process proceeds with high
conversion and simultaneous high selectivity.
The following examples and catalyst
preparations are given to illustrate the invention and
should not be construed as limiting its scope. All
percentages in the examples are mole percent unless
otherwise indicated. Two comparative examples are
given to illustrate the reaction in a conventional
fixed bed reactor.
Preparation of Catal~st A
A mass (25 g) of gamma-alumina in the form of
1/8" x 1/8" (3.2 mm x 3.2 mm) tablets was added to a
solution prepared by dissolving 3.35 g of Re207 in 250
ml of aqueous ethanol (5-10 percent H20). The
resulting mixture was stirred under vacuum for a few
minutes and was then evaporated to dryness on a steam
bath. The resulting light grey pellets were dried at
100C for 3 hours and were then calcined in a dry air
3 stream at 600C for 1 hour. Plasma emission analysis,
using the 346.05 nm rhenium line, indicated a 9.04
percent rhenium loading.
30,442B-F -12-
-13-
Preparation of Catal~st B
A catalyst was prepared according to the method
of preparation of Catalyst A except that sufficien~
NH4ReO4 was added to the aqueous ethanal to give a 2
percent rhenium loading. Analysis o~ the resulting
catalyst indicated a 1.55 percent rhenium loading.
Preparation of Catalyst C
A catalyst was prepared according to the method
of preparation of Catalyst A except that the catalyst
was subjected to an additional calcination treatment at
500C in a stream of water-saturated air (0.2 SCFH) for
2 hours. The catalyst had a 9.2 percent rhenium
loading, as determined by plasma emission analysis.
Preparation of Catalyst D
A catalyst was prepared according to the method
of preparation of Catalyst A except that sufficient
NH4ReO4 was added to the aqueous ethanol to give a 2.5
percent rhenium loading. The catalyst was subjected to
a caloination treatment at 600C for 3 hours in a muffle
~urnace open to the air. This treatment allowed the
catalyst to collect moisture from the air upon cooling.
The catalyst had a 2.0 percent rhenium loading, as
determined by plasma emission analysis.
Example 1
A distillation column reaction vessel was
constructed from 316 stainless steel pipe. The main
body of the vessel had a 3/4-inch (19 mm) outside
diameter and the remainder of the ~essel had a 3/8-inch
(9.5 mm) outside diameter. The main body of the vessel
30,442B-F -13-
~q~J~
-14-
was filled to a depth of approximately l/2 inch (13 mm)
with glass Raschig rings. Catalyst A (25 g) was added
to the main body of the vessel to form a catalyst bed
having an approximate size of 9 inches (230 mm) by 5/8
inch (16 mm). The remaining space in the main body of
the vessel above the catal~st was filled with 1/8 inch
(3.2 mm) diameter 304 stainless steel beads. The void
volume of the reactor including reboiler was 70 ml,
measured with the catalyst and inert packing in place.
0 A thermosiphon reboiler was attached below the main
body of the vessel. The vessel was equipped with a
condensing means, means for controlling the temperature
in the reboiler, means for heating the main body of the
~~ 15 vessel, means for observing and recording the
temperature at various points in the vessel, means for
observing and controlling the pressure~ and means for
emergency relief o~ an overpressure condition~
Butene-l was fed at a rate of 9.0 g per hour
(60 ml/min 9 measured at Standard Temperature and
Pressure, STP) into the reboiler of the vessel. The
temperature in the reboiler was maintained at 115C, the
temperature in the overhead condenser was maintained at
-50C, and the pressure in the system was maintained
autogenously at approximately 80 psig (550 kPa gauge).
The butene-1 flashed from the reboiler into the
catalyst bed where the bulk o~ the butene reacted at a
temperature of less than 100C. The higher boiling
hexene-3 product fractionated immediately from the
butene. The lower boiling ethylene product separated
from the butene, passed through the condenser, and was
removed from the system. The average flow rate of
vapor from the top of the vessel was 30 cm3/min.,
measured at STP.
30,442B-F -14-
'7~2 ~
-15- .
The overhead stream was analyzed using vapor
phase chromatography and was found to be 90 mole
percent ethylene and lO mo.le percent propylene. The
bottoms stream was analyzed using vapor phase
chromatography and was found to be approximately 69.7
mole percent hexene-3, approximately 11.0 mole percent
pentene, approximately 14.7 mole percent butene-1,
approximately 3.7 mole percent heptene, and
approximately 0.9 mole percent octene. The overall
conversion was 92 mole percent, calculated according to
the formula:
% conversion = moles of butene-l converted x 100
moles of butene-1 fed
f 15
The selectivity to the various products and by-
products, based on the moles of butene-l converted, was
as ~ollows:
Species Mole % Seleativity
; Ethylene 45.1
Propylene 4.9
Pentenes 6.5
~5 Hexenes 41 . 3
Heptenes 2.2
Octenes o,5
wherein
% selectivity = moles of species formed x 100.
moles of butene-1 converted
Absolute mole values, rather than relative mole ratios,
were used in making these calculations, because the
overhead stream and the bottoms stream contained
30 , 4 4 2B-F -15-
-16-
different total moles of products~ Thus, the yield of
ethylene and hexene was approximately 79.5 percent 7
calculated as follows:
selectivity to (C2 + C6) - 0.451 ~ 0.413 - 0.864
conversion = 0.92
yield = o.g2 x 0.864 = o~795.
Exam~le 2
The method of Example 1 was repeated with the
following exceptions:
(a) Cat lyst B was employed (24 g);
(b~ the feed rate of 1~butene was 9.6 g/hour;
(c) the reboiler temperature was 130C; and
(d) the system pressure was maintained
autogeneously at 90 psig ~620 kPa gauge).
Analyses of the product streams and the selec-
tivities, calculated as in Example 1, are shown in
Table I.
TABLE I
Overhead Bottoms Selectivity
(mole %) (mole %) (mole %)
Ethylene 98 - 49
Propylene 2
Butenes - 31
3 Pentenes - 1 007
Hexenes - 68 49.3
Heptenes - - -
Octenes _ _ _
30,442B-F -16-
-17-
The overall conversion was 85 mole percent and the
combined yield of C2 + C6 hydrocarbons was
approximately 84 percent.
Example 3
The method of Example 1 was repeated, with the
following exceptions:
(a) Catalyst C was employed (30 g);
(b) the feed rate of 1-butene was 9.6 g/hour;
(c7 the reboiler te~perature was 120C;
f 15 ~d) the overhead temperature was approximately
28C; and
(e) the system pressure was maintained
autogeneously at 100 psig (690 kPa gauge).
Analy~is o~ the product streams and the selec-
tivities, calculated as in Example 1, are shown in
Table II.
TABLE II
Overhead Bottoms Selectivity
(mole %) (mole %) (mole %)
Ethylene 26 - 17
30 Propylene 49 - 32
Butenes 24 40
Pentenes 1 12 11
Hexenes - 41 36
Heptenes - 6 4
35 Octenes - <1 <1
30,442B~F -17-
-18-
The overall conversion was 68 mole percent.
Higher conversions may be readily attained by
maintaining the overhead temperature below 0C to
constrain the 1-butene to the reactive zone. This
example demonstrates the disproportionation of butene-1
into desired C2, C3, C5 and C6 products in a yield of
approximately 65 percent. Thus, as can be seen by
comparing the results of Examples 1 and 3, the product
mix may be varied by changing 9 in a simple manner, the
catalyst preparation and the process operating
conditions, most notably the overhead temperature.
Example 4
f 15 The method of Example 1 was repeated, with the
following exceptions:
- (a) Catalyst D was employed (23 g);
(b) the feed rate of 1-butene was 9.6
g/hour;
(c) the reboiler temperature was 130C;
(d) the overhead temperature was 0C; and
(e) the system pressure was maintained
autogeneously at 90 psig (620 kPa
gauge~.
Analyses of the product streams and the selec-
tivities, calculated as in Example 1, are shown in
Table III.
30,442B-F -18-
-19-
TABLE III
Overhead Bottoms Selectivity
(mole %) (mole %) tmole %)
Ethylene 2 - 1
Propylene 96 - 49
Butenes 2 14
Pentenes - 35 20
10 Hexenes _ 47 27
Heptenes - 3.5 2
Octenes - 0,5 0.3
The overall conversion was 92 mole percent. The
; 15 combined yield of C3 and C5 hydrocarbons was
approximately 64 percent, while the co~bined yield of
C2, C3, C5 and C6 hydrocarbons was approximately 89
percent. Thus, as can be seen by comparing the results
of Examples 2 and 4, the product mix may be varied by
changing the catalyst preparation in a simple manner
and by changing the overhead temperature.
Comparative Experiment 1 (C.E. 1)
A catalyst comprising gamma-alumina in the form
of 1/8" x 1/8" (3.2 mm x 3.2 mm) tablets and ~e207 was
prepared following the procedure described in the
preparation of Catalyst A except the catalyst was
calcined overnight in dry air at 530C. Plasma emission
30 analysis, using the 346.05 nm rhenium line indicated an
8.38 percent rhenium loading. The catalyst (25 g) was
added to a vessel to form a fixed bed reactor having an
approximate size of 9 inches by 5/8 inch (320 mm x 16
mm). Butene-1 was fed at a rate of 50, 150 or 200
ml/min into the reactor. The temperature of the bed
was maintained at 50C or 100C. The pressure was
30,442B-F -19-
-20-
maintained at 1 atmosphere in order to keep the butene
in the gaseous state. The product stream was analyzed
using vapor phase chromatography, and the product
distribution is given in Table IV.
TABLE IV
Selectivity (mole %)
Flow Temp Conv.
C.E.1ml/min Cmole % C2H4_3~6C5_10 C__12
(a) 50 10044.8 35.032.6 18.0 14.4
(b) 150 50 40.5 40.4 5.2 4.0 50.4
(c) 200 50 34.8 39.1 3.8 3.6 53.5
(d) 200 10040.8 36.519.9 15.3 28.3
Comparison o~ Example 1 with Comparative
Experiment 1~a), shows that the conversion in the fixed
bed reaotor i5 less than one-half the conversion in the
distillation column reactor. Furthermor-e, the combined
yield of ethylene and hexene, calculated as in Example
1, was only 22.1 percent in the fixed bed reactor, or
less than one-third of the analogous yield in the
distillation column reactor. At higher flow rates
illustrated by Comparative Experiments 1(b), (c) and
(d), the conversion in the fixed bed reactor decreased.
Thus, in contrast to Example 1, a high conversi3n and
high selectivity cannot be obtained simultaneously in
the fixed bed reactor. Accordingly, a high yield of C2
+ C6 hydrocarbons cannot be achieved under mild
conditions.
30,442B-F -20-
-21-
Exa~le 5
The method of Example 1 was repeated~ with the
following exceptions:
(a) the ~eed was an equimolar mixture of
butene-l and butene-2, with the butene-2 being
50 percent cis and 50 percent trans;
(b) the feed rate was 20 g/hour;
(c) the reboiler temperature was 130C;
(d) the overheads temperature was
approximately 5C;
f 15
(e~ the system pressure was maintained
autogenously at 100 p9ig ( 690 kPa gauge); and
(~) the average vapor flow rate from the top
o~ the vessel was 120 cm3/min.
Analyses of the product streams and selectivities,
calculated as in Example 1, are shown in Table V.
TABLE V
Overhead Bottoms Selectlvity
(mole ~) (moles %) I mole %)
Ethylene 7 - 3
30 Propylene 92 - 47
Butenes 1 23
Pentenes - 49 30O7
Hexenes - 27 19O3
Heptenes
35 Other _ <1
30,442~-F --21-
3~3~
-22-
The overall conversion was 88 mole percent and
the yield to propylene and pentenes was approximately
68.4 percent. The combined yiald of C2, C3, C5 and C6
hydrocarbons was 88 mole percent. Thus, it is seen
that by making simple adjustments in the feedst~ck and
the operating conditions of Example 1, the selectivity
can be controlled to provide a high combined yield of
C3 and C5 hydrocarbons.
Comparative Experiment 2 (C.E. 2)
A rhenium catalyst (8.38 percent Re) was
prepared according to the procedure of Example 1,
except that the catalyst was calcined at 530C overnight
f 15 in air. The fixed bed reactor of Comparative
Experiment 1 was employed with 25 g of the rhenium
catalyst. An equimolar mixture o~ butene-1 and butene-
2~ with butene-2 being 50 percent cis and 50 percent
trans, was fed to the reactor at a rate of 20 g/hour.
The temperature of the bed was maintained at 50C, 75C
or 100C. The pressure was maintained at 1 atmosphere
50 as to keep the butenes in the gaseous state. The
product stream was analyzed using vapor phase
chromatography and the product distribution is given in
Table VI.
TABLE VI
Selectivity (mole ~)
3 Temp Conv.
C.E~2 C mole ~ _2H4 _3H6_5H10 -6H12
(a) 50 48.6 36.6 7.67 3.71 34.5
(b) 75 40.3 36.5 20.611.7 34.5
35 (c) 100 61.7 5.11 44.434.8 13.5
30,442B-F -22-
-23-
Comparison of Example 5 with Comparative
Experiment 2(b) 5 shows that the conversion in the fixed
bed reactor ;~ less than 50 percent of the conversion
in the distillation column reactor. Furthermore, the
eombined yield of C3 and C5 hydrocarbons in the fixed
bed reactor was only 13 percent, compared with 68
percent in the distillation column reactor. Even at
100C, the combined yield of C3 and C5 hydrocarbons in
the fixed bed reactor was only 49 percent~ as shown in
Comparative Experiment 2(c).
It may be noted that the conversions shown in
Examples 1-5 are for a column fed at the reboiler.
Conversions will be higher for a properly designed and
operated distillation column reactor having the feed
point located at a point in the column such that the
feed alkene(s) may be retained in the catalyst bed
until said alkene(s) are disproportionated.
AY previously mentioned, the preceding examples
serve only to illustrate the invention and its
advantages, and they should not be interpreted as
limiting since further modifications of the disclosed
invention will be apparent to those skilled in the art.
All such modifications are deemed to be within the
scope of the invention as defined by the following
claims.
30, 442B F -23-
