Note: Descriptions are shown in the official language in which they were submitted.
s `
PROCESS FOR SELECTIVE HYDROGENATION OF DIENES IN PYROLYSIS
GASOLINE
This invention relates to a process for the selective
hydrogenation of dienes in pyrolysis gasoline.
As will be known, pyrolysis gasoline is obtained as a by-
product in the preparation of ethene and/or propene by means of
high-temperature pyrolysis (e.g. cracking in the presence of
steam) of gaseous or liquid hydrocarbons, such as naphtha or
gas oil.
Pyrolysis gasolines on the one hand are extremely unstable
owing to the presence of a relatively high proportion of highly
olefinically unsaturated hydrocarbons, and on the other hand
contain aromatic compounds and alkenes having a high octane
number which are particularly valuable and are in themselves
useful as stable motor gasoline components.
In order to obtain a stable gasoline with high octane
number from a pyrolysis gasoline the highly olefinically un-
saturated compounds, which mainly consist of dienes~ for example
those of the cyclopentadiene type, have to be removed there-
from. This remo~al may be achieved by partial hydrogenation of
the dienes to mono-olefins. Because the hydrogenation of mono-
olefins in general leads to a reduction in octane num~er, sucha hydrogenation is to be avoided as much as possible. Moreover,
in doing so the hydrogen consumption is kept to a desired low
level.
:,
Qfi5
It is known to hydrogenate dienes present in pyrolysis
gasolines selectively with the aid of catalysts with hydrogen-
ating activity such as supported catalysts comprising a metal
of Group VI and/or Group VIII o~ the Periodic System of
Elements.
In order to convert a high proportion of the dienes
originally present without incurring unacceptably high process
temperatures, it has been proposed to recycle part o~ the product
obtained after the selective hydrogenation to the reactor in
which the selective hydrogenation is being carried out. How-
ever, it is felt as a drawback that during such a process an
appreciable amount of mono-olefins is also hydrogenated, which
leads to an unattractively high hydrogen consum~tion.
The invention provides a solution for this problem by using
at least three consecutive reactors and a specific recirct~ation
pattern.
Accordingly, the invention provides a process for the
selective hydrogenation of dienes in pyrolysis gasoline which
comprises catalytic hydrogenation of the pyrolysis gasoline in
at least three consecutive reactors, in at least two of the said
consecutive reactors recirculating part of the hydrocarbon mixture
emerging from a reactor over that reætor~no recirculation of the
hydrocarbon mixture emerging from the last o~ the consecutive
reactors being carried out over that reactor.
It is to be understood that recirculating part of the hydro-
carbon mixture emerging from a reactor over that reactor, stands
for the direct recirculation of the said hydrocarbon mixture,
no further hydrogenation thereof being carried out in any
subse~uent reactor before the recirculation.
It is essential ~at in at least two of the consecutive
reactors recirculation of part of the hydrocarbon mixture
emerging from each of these reactors over the relevant reactor
is carried out, because in case two consecutive reactors are
- ~P4~J~fi~
used and the recirculation is carried out over one reactor only
the amount of dienes still present in the ultimate product of
the process cannot be brought to an acceptable low level without
ex-tensive hydrogenation of the mono-olefins, the latter gi~ing
rise to an lmdesired high level of hydrogen consumption.
Preferably, the number of eonsecutive reactors is three.
More reaetors may be used, and the reeireulation as described
may be carried out over more than two reactors, but in general
the advantage to be achieved (less hydrogenation of mono-olefins
at a preset amount of dienes in the ultimate product) does not
compensate for the drawbacks which consist of building and
handling of an extra reactor.
The catalys-t with hydrogenating activity to be used in the
reactors very suitably consists of a support comprising one or
more metals of Group VIB and/or Group VIII of the Periodic System
of Elements and/or compounds of these metals. The support very
suitably consists of alumina, silica or silica alumina. Catalysts
comprising platinum or palladium are very suitable. The most
preferred catalyst comprises partially sulphided nickel on
alumina as a support.
The catalyst ve~J suitably is in the form of one or more
fixed beds in the reaetors~ and the catalytic hydrogenations are
preferably carried out by passing a mixture of liquid and
hydrogen-containing gas in downflow over the catalyst according
to the trickle flow technique. In this technique, the starting
hydrocarbon oil which is present partly in the liquid phase
and partly in the vapour phase is allowed to flow downward in
the presence of hydrogen or of a hydrogen-containing gas over a
catalyst in the form of a fixed bed, the unvaporized part of
the starting material flowing over the catalyst particles in
the form of a thin liquid layer.
The recirculation ratios over the respective reactors are
to be chosen such that in each of these reactors the ratio of
the dienes hydrogenated to mono-olefins on the one hand and
Q6~ ,
the mono-olefins hydrogenated to paraffins on the other hand,
is high because in this way the overall hydrogen consumption
is kept low. It has been found that the weigkt ratio of the
hydrocarbon mixture recirculated to the first reactor and the
pyrolysis gasoline fed thereto very suitably is from 5 to 15,
in particular from 9 to 11. The preferred weight ratio of the
hydrocarbon mixture recirculated to the second reactor and
the hydrocarbon mixture emerging from the first reactor which
is fed to the second reactor, has been found to be from 2 to 4.
No recirculation is to be used in the final reactor of
the consecutive reactors, because the amount of dienes still
present in the hydrocarbon mixture fed to that reactor is so
low that at relatively high space velocity hydrogenation thereof
to mono-olefins can lead to the desired concentration of dienes
in the effluent without undue hydrogenation of mono-olefins to
paraffins.
The hydrogen to be employed in the catalytic hydrogenation
may be pure or in the form of a hydrogen-containing gas. The
gases employed should preferably contain more than 50% by volume
of hydrogen. Very suitable are, for example, the hydrogen-con-
taining gases obtained in the catalytic reforming or steam-
reforming of gasoline fractions, and mixtures of hydrogen and
light hydrocarbons. Any excess of hydrogen-containing gas is
advantageously recycled, possibly after the previous removal of
undesired components there~rom.
The catalytic hydrogenations are very suitably carried out
at the following conditions in the reactors: a temperature in
the range from 50-250 C, preferably 50-150 ~, a total pressure
in the range from 10-80 bar a anda hydrogen partial pressure in
the range from 5-60 bar a.
The liquid hourly space velocity of the hydrocarbon mixture
which is fed to a reactor for the first -time (e.g., the ~resh
feed fed to the first reactor or that part o~ the effluent of a
l~OQ~S
,
reactor which is fed to the next reactor) may vary for any reactor.
In order to calculate the space velocity of the mixture obtained
after combination of the recycle liquid and the hydrocarbon mixture
which is fed to a particular reactor for the first time, the space
velocity of the last-mentioned hydrocarbon mixture is to be
multiplied by the recirculation ratio used for that reactor.
Very suitably the liquid hourly space velocities of the mixtures
obtained after combination of the recycle liquid and the hydro-
carbon mixture which is fed to a particular reactor for the first
time, are between 5 and 50 and preferably between 10 and 20 kg
mixture per litre catalyst per hour. At lower space velocities
the extent of heat release of the hydrogenation reaction may be
such that temperature control becomes difficult if not impossible.
The space velocities of the hydrocarbon mixtures fed for the
first time to a particular reactor to which recycling is taking
place, will in general be less than 5 kg per litre catalyst per
hour.
In the last reactor, the feed of which only contains
relatively low amounts of dienes, the space velocity very suit-
ably ranges from 2-20 and preferably from 5-10 kg feed per litre
catalyst per hour.
The ratio of fresh gas to fresh feed very suitably is from
50-500 Nl gas per kg feed, and the ratio of recycle gas to fresh
feed from 200-500 Nl gas per kg feed.
The accompanying figure I represents a simplified flow
diagram of a suitable embodiment of the process according to
the invention. In this figure various au,ciliary devices, such
as pumps, cocks, valves, control valves, etc. have been
omitted.
A pyrolysis gasoline is fed in via line 1, and after mixing
with a hydrogen-containing gas supplied via line 2, introduced
into reactor Rl via line 3. Reactor R1 contains one or more
fixed beds of catalyst. The effluent of R1 is led via line 4
to separation vessel V1 in which gas and liquid (the latter
.
~4~Q~S
consisting substantially of h~drocarbons) are separated. Part
of the liquid is transported via line 5, mixed with pyrolysis
gasoline from line 1 and recirculated to R1 via line 3. The
remainder of the liquid in V1 and the gas are ~orwarded via
line 6~ mixed with a liquid stream emerging from separating
vessel V2 via line 7, and fed to reactor R2 via line 8.
Reactor R2 contains one or more fixed beds of catalyst. The
effluent from reactor R2 is forwarded via line 9 to separation
vessel V2, in which vessel liquid and gas are separated. Part
of the liquid is recycled via line 7 as described, and the
remainder of the liquid and the gas from V2 are fed to reactor
R3 via line 10. The effluent from reactor R3, which reactor
contains one or more fi~ed beds of catalyst, is led to
separation vessel V3 via line 11. In V3 gas and liquid are
separated. The liquid is removed via line 12 as the final
product of the process, the gas from V3 is (if desired after
purification) forwarded via line 13 to line 2, in which line
fresh hydrogen-containing gas is fed via line 14.
EXAMPLE
The process was carried out according to the flow scheme
given in Figure I.
Fresh feed in line 1 consisted of a pyrolysis gasoline
which contained 60.2~ow dienes and 20.0~w mono-olefins. This
feed was added at a space velocity of 1.48 kg/l catalyst in R1/
hour. The catalyst in all reactors consisted of partially
sulphided nickel on alumina, the amount of nickel being 10.7%w
on carrier. The inlet temperature of R1 was 63C, the temper-
ature at the outlet of this reactor was 90C, the average
pressure in R1 was 62.5 bar a (H2 partial pressure 45.4 bar a).
30 The effluent of R1 was separated in V1 at 90C, and 10 times
the amount of liquid fed via line 1 was recycled from V1 via
line 5 (recycle ratio 10). This recycle liquid contained 18.4%w
dienes. The remainder of the liquid and the gas in V1 were led
. .
~4~6~ ,
via line 6 at a space velocity of 4.13 kg/1 catalyst in R2/hour
-to reactor R2. The inlet temperature of R2 was ô3 C, the outlet
temperature 104 C and the average pressure 61.5 bar a (H2 partial
pressure 40.6 bar a). The effluent from R2 was separated at
104 C in separator V2 in liquid and gas, and 3.2 times the amount
of liquid fed via line 6 was recycled from V2 via line 7 to R2
(recycle ratio 3.2). The recycle liquid con-tained 5.8%w dienes.
The remainder of the liquid and gas in V2 were forwarded via
line 10 to reactor R3 at a space velocity of 6.89 kg/l catalyst
in R3/hour. The inlet temperature of R3 was 90 C, the outlet
temperature 123 C and the average total pressure 59.5 bar a
(H2 partial pressure 34.4 bar a). The effluent of R3 was
separated in V3, and the liquid removed as final product of
the process via line 12. This product contained 0.5~ow diolefins
and 72.7%w mono-olefins. The gas from V3 was recycled via line 13
and mixed with fresh hydrogen-containing gas, the latter con-
taining 94.1 mol.% hydrogen, 4.6 mol.% methane, 1.0 mol.%
nitrogen and 0.3 mol.% water. Fresh gas was supplied via line
14 in an amount of 233 Nl H2/kg fresh feed. Recycle gas via
line 13 amounted to 300 ~l/kg fresh feed and contained 75.0 mol.%
hydrogen, 11.8 mol.% methane, 6.2 mol.% nitrogen, 2.0 mol.% water,
the remainder consisting of hydrocarbons with at most 6 carbon
atoms. The amount of hydrogen consumed per kg feed for undesired
mono-olefin-saturation was 0.002 kg.
Comparative Exam~le
The process was carried out according to the flow scheme
given in Figure 2.
Fresh feed in line 1 consisted of a pyrolysis gasoline which
contained 60.2%w dienes and 20.0%w mono-olefirs.This feed was
added at a space velocity of 0.42 kg/l catalyst in R1/hour. The
catalyst in both reactors consisted of partially sulphided nickel
on alumina, the amount of nickel being 10.7%w on carrier. The
inlet temperature of R1 was 63 C, the temperature at the outlet
of this reactor was 90 C, the average pressure in R1 was 62.5
bar a (H2 partial pressure of 4~ bar a at reac-tor outlet. The
effluent of R1 was separated in V1 at 90 C, and 13 times the
amount of liquid fed via line 1 was recycled from V1 via line 5
(recycle ratio 13). This recycle liquid contained 5~8~ow dienes.
The remainder of the liquid and the gas in V1 were forwarded
via line 6 to reactor R2 at a space velocity of 6 . 89 kg/l
catalyst in R2/hour. The inlet temperature of R2 was 90C,
the outlet temperature 123 C and the average total pressure
59.5 bar a (H2 partial pressure 32 bar a at reactor outlet).
The effluent of R2 was separated in V2 and the liquid removed
as final product of the process via line 8. This product
contained 0.5%w diolefins and 69.9%w mono-olefins. The gas
from V2 was recycled via line 9 and mixed with fresh
hydrogen-~ontaining gas, containing 94.1 mol.% hydrogen, I~.6
mol.% methane, I.0 mol.% nitrogen and 0.3 mol.% water. Fresh
gas was supplied via line 10 in an amount of 244 Nl/kgfresh feed.
Recycle gas via line 9 amounted to 300 ~l/kg fresh feed and
contained 73.8 mol.% hydrogen, 12.8 mol.% methane, 6.7 mol.%
nitrogen, 2.3 mol.% water, the remainder consisting of hydro-
carbons with at most 6 carbon atoms. The amount of hydrogen
consu~ed per kg of feed for undesired mono-olefin saturation
was 0.003 kg, which is 50% higher than in the Example according
to the invention.
;'
