U.S. patent number 3,781,195 [Application Number 05/214,468] was granted by the patent office on 1973-12-25 for process for the production of gaseous olefins from petroleum distillate feedstocks.
This patent grant is currently assigned to BP Chemicals International Limited. Invention is credited to Paul Trevor Davis, Terence George Glover, John Robert Jones.
United States Patent |
3,781,195 |
Davis , et al. |
December 25, 1973 |
PROCESS FOR THE PRODUCTION OF GASEOUS OLEFINS FROM PETROLEUM
DISTILLATE FEEDSTOCKS
Abstract
Olefins, e.g., ethylene and propylene, are produced by
hydrogenating a petroleum distillate feedstock, e.g., a petroleum
vacuum distillate feedstock (b. pt. 300.degree. - 650.degree.C),
over a catalyst containing combinations of Ni, Mo, Co and Won
SiO.sub.2 /Al.sub.2 O.sub.3 bases, preferably Ni/W/SiO.sub.2
/Al.sub.2 O.sub.3, under reaction conditions such that a
substantial quantity of aromatics in the feedstock are
hydrogenated. The whole product of hydrogenation is then thermally
cracked in the presence of steam, producing an increased yield of
gaseous olefin products.
Inventors: |
Davis; Paul Trevor (Feltham,
EN), Glover; Terence George (Isleworth,
EN), Jones; John Robert (Walton-on-Thames,
EN) |
Assignee: |
BP Chemicals International
Limited (London, EN)
|
Family
ID: |
26236060 |
Appl.
No.: |
05/214,468 |
Filed: |
December 30, 1971 |
Foreign Application Priority Data
|
|
|
|
|
Jan 6, 1971 [GB] |
|
|
605/71 |
Sep 14, 1971 [GB] |
|
|
42,765/71 |
|
Current U.S.
Class: |
585/251; 502/254;
502/315; 585/648; 208/57; 208/143; 502/314; 585/270 |
Current CPC
Class: |
C10G
69/06 (20130101); C10G 2400/20 (20130101) |
Current International
Class: |
C10G
69/06 (20060101); C10G 69/00 (20060101); C10g
037/00 () |
Field of
Search: |
;208/57,130
;260/683R |
References Cited
[Referenced By]
U.S. Patent Documents
Primary Examiner: Levine; Herbert
Claims
We claim:
1. A process for the production of olefins which process comprises
hydrogenating a petroleum vacuum distillate feedstock boiling in
the range of 300.degree. to 650.degree.C. in the presence of a
nickel/tungsten/silica/alumina hydrogenation catalyst and hydrogen
at a temperature in the range 50.degree. to 500.degree.C and a
pressure in the range 50 to 5,000 psig and thermally cracking the
resulting hydrogenated product in the presence of steam.
2. A process according to claim 1 wherein the hydrogenation
temperature is in the range 300.degree. to 400.degree.C.
3. A process according to claim 1 wherein the hydrogenation
pressure is in the range 200 to 3,000 psig.
4. A process according to claim 1 wherein the hydrocarbon liquid
Hourly Space Velocity (L.H.S.V.) is in the range 0.1 to 5.0.
5. A process according to claim 1 wherein the hydrogen is used on a
recycle basis at 5 to 10 times the molar rate of the hydrocarbon
feedstock.
6. A process according to claim 1 wherein the feedstock from the
hydrogenation reaction is varpourised in the presence of steam, at
a steam to hydrocarbon weight ratio of 0.5:1 to 2.0:1 and is passed
through a heated zone at a maximum temperature in the range
700.degree. to 1,000.degree. with a residence time in this
temperature range between 0.01 and 5 seconds.
Description
The present invention relates to a process for the production of
gaseous olefins from petroleum distillate feedstocks.
Ethylene, propylene and butadiene which are basic intermediates for
a large proportion of the petrochemical industry are obtained in
the main by thermal cracking of petroleum gases and distillates
such as naphtha and gas oil. There is a world wide increase in
demand on the use of these lighter components of petroleum and it
is desirable that heavier feedstocks be utilised for olefin
production. In the past a number of problems arose in the cracking
of heavier feedstocks which have so far prevented their use in the
economic production of light olefins. The principle problems
were:
1. Excessive coke deposition in the cracking tubes which reduces
heat transfer, thereby necessitating higher tube skin temperatures.
Excessive coke deposition also restricts flow in the cracking
tubes, and may ultimately lead to blockage. The coke must be
removed at frequent intervals by burning out, involving plant
shut-down for an excessive proportion of time on-stream.
2. Tar deposition in the transfer lines and heat exchangers; this
reduces the efficiency of heat recovery and requires plant
shut-down for cleaning, again impairing the overall efficiency of
operation.
3. Yields of olefin products are low compared with those from
lighter feedstocks necessitating increased feedstock and fuel
requirements with extra furnaces heat exchangers and other
equipment involving much higher initial capital investment.
4. Feedstocks from many sources contain high levels of sulphur;
this is not necessarily detrimental to the operation of the
cracking process but may increase the cost of plant construction.
Further the bulk of the sulphur is concentrated in the liquid
products boiling above 200.degree.C which are therefore less
valuable as fuel oil.
We have now found that by pretreating the petroleum distillate
feedstock with hydrogen in the presence of a catalyst under
conditions which lead to substantial reduction in the content of
aromatic compounds (especially of polycyclic aromatic compounds)
and of sulphur compounds, but without substantial re-organisation
or breakdown of the carbon structures of the various component
compounds, the disadvantages of the prior art referred to above are
substantially overcome.
Thus according to the present invention there is provided a process
for the production of olefins which process comprises hydrogenating
a petroleum distillate feedstock in the presence of a hydrogenation
catalyst and hydrogen and thermally cracking the resulting
hydrogenated product in the presence of steam.
Thermal cracking within the context of this application is intended
to include steam cracking but not catalytic cracking.
The preferred petroleum distillate feedstock is a vacuum distillate
boiling within the range (at atmospheric pressure) 300.degree. to
650.degree. C., though lighter distillate feedstocks such as gas
oil boiling within the range 200.degree. - 350.degree.C may be
used.
It is important to avoid excessive breakdown of the feedstock in a
hydrocracking type of reaction. A limited amount of breakdown can
be tolerated and may even give the benefit of producing a more
mobile product but excessive hydrocracking leads to the use of
larger quantities of hydrogen with increased manufacturing costs
and to the formation of products which do not give corresponding
benefits in further increases in the yield of olefins. Any catalyst
which is capable of catalysing the hydrogenation of compounds
containing aromatic rings without substantial structural alteration
or breakdwon may be used. Since most feedstocks contain sulphur and
nitrogen compounds it is desirable that the catalyst should also
possess some tolerance to these materials and their hydrogenation
products. Hydrogenation catalysts embodying these requisites
include for example nickel/molybdenum/alumina,
cobalt/tungsten/alumina, nickel/tungsten/alumina,
cobalt/molybdenum/alumina, nickel/cobalt/molybdenum/alumina,
cobalt/molybdenum/silica/alumina, nickel/molybdenum/silica/alumina,
cobalt/tungsten/silica/alumina. A particularly active hydrogenation
catalyst is nickel/tungsten/silica/alumina.
Although it will usually be convenient to employ the hydrogenation
catalyst without prior exposure to materials containing sulfur at
least initially, the catalyst may also be used in the sulfided
form.
The catalysts may conveniently be prepared by impregnating the
support with an aqueous solution of a salt of each of the metals,
either consecutively or simultaneously. Thus nickel may be added in
the form of nickel nitrate, tungsten as ammonium metatungstate,
cobalt as cobalt nitrate, acetate, etc. and molybdenum as ammonium
molybdate. It will usually be found convenient to impregnate the
support first with the salt of the metal which is to be present in
the highest concentration in the finished catalyst, though this is
not essential. Other methods of preparing the catalyst include
precipitating the metals on the support from a solution of their
salts and coprecipitation of the metals with the hydrated support
material.
It is preferred that the catalysts be activated before use in the
reaction by contact with a stream of hydrogen at a temperature in
the range 100.degree. to 800.degree.C, preferably 300.degree. to
600.degree.C, for a period of 1 minute to 24 hours. The sulfided
form of the catalyst may conveniently be prepared by passing
hydrogen through liquid tetrahydrothiophene and then over the
catalyst maintained at a temperature in the range 100.degree.C to
800.degree.C, preferably 300.degree.C to 600.degree.C, for a period
of 1 minute to 24 hours.
Whilst the precise nature of the active species in the above
hydrogenation catalysts is not known it is possible that the
catalyst contains, in addition to the support, elemental metal,
metal oxides, metal sulfides and complex aluminium or silicon/metal
compounds.
Using nickel and cobalt catalysts the hydrogenation temperature may
be in the range 50.degree. to 500.degree.C, preferably 300.degree.
to 400.degree.C, and the pressure may be in the range 50 to 5000
p.s.i.g., preferably 200 to 3000 p.s.i.g.
The hydrocarbon Liquid Hourly Space Velocity (LHSV) may be in the
range of 0.1 to 5.0, preferably 0.25 to 2.0.
Hydrogen is preferably used on a recycle basis, preferably at about
5 to 10 times the molar rate of the hydrocarbon feedstock, and may
be passed through scrubbers to remove hydrogen sulfide and ammonia
before recycle. However other methods of operation may also be used
such as batch operation in an autoclave. For catalysts other than
those containing cobalt or nickel the reaction conditions may be
different.
Hydrogenation may be carried out in a single stage or in a series
of two or more operations using the same or different
catalysts.
The feedstock from the hydrogenation reaction is vaporised in the
presence of steam at a steam to hydrocarbon weight ratio of about
0.5:1 to 2.0:1 and passed through a heated zone, preferably a tube,
at a maximum temperature in the range 700.degree. to 1,000.degree.C
with a residence time in this temperature range between 0.01 and 5
seconds, preferably 0.1 to 2 seconds. The products are rapidly
cooled in a heat exchange system and separated and purified by
conventional means.
The invention is illustrated by the following examples:
COMPARISON TEST 1
A full range Agha Jari vacuum distillate with a hydrogen to carbon
atomic ratio of 1.73 and a sulfur content of 1.72 per cent weight
was steam cracked in a 26 ml quartz reactor at a maximum
temperature of 830.degree.C. Analysis by physical separation and
spectroscopic methods (including U.V. Absorbance) indicated that
the aromatic compound content was 49% weight.
The steam to hydrocarbon feed weight ratio was 1 to 1 with an
average total molar flow of 3.3 moles per hour. The ethylene and
propylene yields were 23 and 10 per cent weight respectively with a
total conversion to cracked gas of 53 per cent weight on feed. Coke
deposited in the cracking zone corresponded to 1,200 ppm of the
hydrocarbon feed.
This example is provided for purposes of comparison and is not an
example according to the invention.
EXAMPLE 1
The catalyst was prepared by calcining alumina at a temperature of
550.degree.C. The calcined alumina was then impregnated with an
aqueous solution of ammonium molybdate, evaporated to dryness and
further calcined at 550.degree.C. This procedure was then repeated
using an aqueous solution of cobalt nitrate. The catalyst was then
activated in a stream of hydrogen at 400.degree.C for 16 hours.
A 250g sample of the Agha Jari vacuum distillate used in Comparison
Test 1 was hydrogenated in a 1 litre rocking autoclave at
350.degree.C under 1,500 p.s.i.g. of hydrogen during 8 h using 100g
of the cobalt/molybdenum/alumina catalyst prepared as above. The
recovered hydrogenated vacuum distillate, sample A, had a hydrogen
to carbon atomic ratio of 1.84 and a sulfur content of less than
0.05 per cent weight. Analysis by physical separation and
spectroscopic methods (including U.V. Absorbance) indicated that
the aromatic compound content was 19% weight. This material was
steam cracked under the same conditions as were used in Comparison
Test 1. The ethylene and propylene yields were 26 and 10 per cent
weight on feed respectively with a total conversion to cracked gas
of 58 per cent. There was also a substantial reduction in the
heavier products compared with those formed from untreated vacuum
distillate. In particular tarry material condensed in the transfer
line from the reactor was formed in only half the amount observed
in Comparison Test 1. The amount of coke deposited in the reactor
zone was only 250 ppm on feed.
EXAMPLE 2
A 100 g sample of the hydrogenated vacuum distillate sample A, was
further hydrogenated in a rocking autoclave at 350.degree.C and
1,500 psig of hydrogen during 18 hours using 40 g of a 5 per cent
nickel on silica catalyst prepared by impregnation as in Example 1.
This further hydrogenated vacuum distillate has a hydrogen to
carbon atomic ratio of 1.91 and a sulphur content of less than 0.02
per cent weight. Analysis by physical separation and spectroscopic
methods (including U.V. Absorbance) indicated that the aromatic
compound content was less than 2 per cent weight. On steam cracking
this sample under the conditions used in Comparison Test 1 the
ethylene and propylene yields were found to be 29 and 11 per cent
weight respectively, and the total cracked gas yield has increased
to 64 per cent. The yields of fuel oil and tarry material were
reduced to a quarter of the values obtained from the untreated
vacuum distillate while the amount of coke deposited in the reactor
was only 100 ppm on feed.
COMPARISON TEST 2
A full range Kuwait vacuum distillate with a hydrogen to carbon
atomic ratio of 1.74 and a sulfur content of 2.78 per cent weight
was steam cracked in a 20 ml quartz reactor at a temperature of
830.degree.C. Analysis by physical separation and spectroscopic
methods (including UV absorbance) indicated that the aromatic
compound content was 52 per cent weight.
The steam to hydrocarbon feed weight ratio was 1 to 1 with an
average hydrocarbon feed rate of 62 g per hour. The ethylene and
propylene yields were 23 and 10 per cent weight respectively with
52 per cent weight of the feed converted to cracked gas. Coke
deposited in the cracking zone corresponded to 1,050 ppm weight of
the hydrocarbon feed. The sulfur content of the fuel oil was 6.8
per cent weight.
This Example is provided for purposes of comparison and is not an
example according to the invention.
EXAMPLE 3
A 200 g sample of Kuwait vacuum distillate used in Comparison Test
2 was hydrogenated in a 1 litre rocking autoclave at 350.degree.C
under 2,200 psig of hydrogen during 8 hours using 50 g of
cobalt/molybdenum/alumina catalyst prepared by impregnation as in
Example 1.
The recovered hydrogenated vacuum distillate has a hydrogen to
carbon atomic ratio of 1.87 and a sulfur content of less than 400
ppm weight. Analysis by physical separation and spectroscopic
methods (including UV absorbance) indicated that the aromatic
compound content was 16 per cent weight.
This material was steam cracked under the same conditions as
Comparison Test 2. The ethylene and propylene yields were 27 and 12
per cent weight on feed respectively with 62 per cent weight of the
feed converted to cracked gas. The yields of fuel oil and tarry
material were reduced to one third of the values obtained from the
untreated vacuum distillate. Coke deposited in the cracking zone
corresponded to 150 ppm weight of the hydrocarbon feed.
EXAMPLE 4
A 200 g sample of Kuwait vacuum distillate used in Comparison Test
2 was hydrogenated in a 1 litre rocking autoclave at 350.degree.C
under 2200 psig of hydrogen during 8 hours using 50 g of
nickel/tungsten/silica/alumina catalyst prepared by impregnation as
in Example 1. The recovered hydrogenated vacuum distillate had a
hydrogen to carbon atomic ratio of 1.95 and a sulfur content of 100
ppm weight. Analysis by physical separation and spectroscopic
methods (including UV absorbance) indicated that the aromatic
compound content was less than 2 per cent weight.
This material was steam cracked under the same conditions as
Comparison Test 2. The ethylene and propylene yields were 28 and 14
per cent weight on feed respectively with 68 per cent of the feed
converted to cracked gas. The yields of fuel oil and tarry material
were reduced to one quarter of the values obtained from the
untreated vacuum distillate while the amount of coke deposited in
the reactor was only 100 ppm. The sulfur content of the fuel oil
was only 300 ppm.
EXAMPLE 5
A 200 g sample of the Kuwait vacuum distillate used in Comparison
Test 2 was hydrogenated in a 3 litre rocking autoclave at
350.degree.C under 2200 psig of hydrogen during 75 hours using 60.0
g of a cobalt/molybdenum/alumina catalyst prepared by impregnation
as in Example 1. The recovered hydrogenated vacuum distillate had a
hydrogen to carbon atomic ratio of 1.95 and a sulphur content of
125 ppm weight. Analysis by physical separation and spectroscopic
methods (including UV absorbance) indicated that the aromatic
compound content was less than 2 per cent weight.
This material was steam cracked under the same conditions as
Comparison Test 2. The ethylene and propylene yields were 28 and 14
per cent weight on feed respectively, with 67 per cent weight of
the feed converted to cracked gas. The yields of fuel oil and tarry
material were reduced to a quarter of the values obtained from the
untreated vacuum distillate while the amount of coke deposited in
the reactor was only 100 ppm.
EXAMPLE 6
A 200 g sample of Kuwait vacuum distillate used in Comparison Test
2 was hydrogenated in a 1 litre rocking autoclave at 350.degree.C
under 2,200 psig of hydrogen during 8 hours using 50 g of
nickel/tungsten/alumina catalyst prepared by impregnation as in
Example 1. The recovered hydrogenated vaccum distillate had a
hydrogen to carbon atomic ratio of 1.88 and a sulphur content of
less than 300 ppm weight. Analysis by physical separation and
spectrosopic methods (including UV absorbance) indicated that the
aromatic compound content was 15 per cent weight.
This material was steam cracked under the same conditions as
Comparison Test 2. The ethylene and propylene yields were 27 and 13
per cent weight on feed respectively with 62 per cent weight of the
feed converted to cracked gas. The yields of fuel oil and tarry
material were reduced to one third of the values obtained from
untreated vacuum distillate, while the amount of coke deposited in
the reactor was only 120 ppm.
Comparison Tests 1 and 2 show the ethylene and propylene yields,
cracked gas yield and coke lay-down resulting from the cracking of
straight run wax distillate without pretreatment. The Examples
demonstrate that hydrogenation of wax distillates prior to cracking
leads to an improvement in all the above parameters and, in
addition, not only is the sulfur content reduced but the yields of
fuel oil and tarry material are reduced also. The increased degree
of hydrogenation achieved by using a nickel/tungsten/silica alumina
catalyst is demonstrated in Example 4. The degree of hydrogenation
is similar to that of Example 5 but is achieved in a much shorter
time period.
In all examples according to the invention it was further observed
that the fuel oil product had a much lower viscosity and a reduced
tendency for emulsification with water after condensation of the
product and diluent steam by comparison with the fuel oil produced
in the Comparison Tests, thus facilitating both pumping and
handling at lower temperatures and separation and fractionation of
the liquid products respectively.
* * * * *