Note: Descriptions are shown in the official language in which they were submitted.
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TS 8521
PROCESS FOR THE HYDROCRACKING OF A
HYDROCARBONACEOUS FEEDSTOCK
The present invention relates to a process for the hydro-
cracking of a hydrocarbonaceous feedstock containing a relatively
low amount of nitrogen.
There exists a large number of processes for hydrocracking
hydrocarbonaceous feedstocks and numerous catalysts that are used in
these processes. Many of these processes comprise two stages, a
hydrotreating stage and a hydrocracking stage, the two stages
operating with different types of catalysts.
Product from the first stage may be treated to remove ammonia,
hydrogen sulphide and other light gases prior to being passed to the
second stage, or product may be passed directly to the second stage.
In this two stage or series-flow mode of operation the hydrocracking
stage is frequently referred to as a second stage hydrocracker.
Hydrocracking is as such a well-established process in which
heavy hydrocarbons are contacted in the presence of hydrogen with a
hydrocracking catalyst. The temperature and pressure applied are
relatively high, so that the heavy hydrocarbons are cracked to
products of a lower average molecular weight and lower boiling
point.
These products include gaseous material, i.e. C1-Cq hydro-
carbons, naphtha and a middle distillate fraction, i.e. a kerosine
fraction and a gas oil fraction.
Since the gaseous products are not very much wanted and since
there is an increasing demand for middle distillates, it would be
advantageous to have a hydrocracking process available for
converting hydrocarbonaceous feedstocks that shows a considerable
selectivity towards middle distillates and a low gas make.
It has now been found that a surprisingly low gas make and a
high yield of middle distillates can be obtained if a hydro-
carbonaceous feedstock containing a relatively low amount of
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nitrogen is passed over a catalyst system comprising a specific
sequence of hydrocracking catalysts.
Accordingly, the present invention relates to a process for
hydrocracking a hydrocarbonaceous feedstock containing less than
500 ppmw N by contacting the feedstock at elevated temperature and
pressure in the presence of hydrogen with a hydrocracking catalyst A
and wherein at least a portion of the product so obtained is
subsequently contacted with a hydrocracking catalyst B at elevated
temperature and pressure and in the presence of hydrogen, whereby
the hydrocracking catalysts A and B have a RS > 1 and a RA.RV < 5.
In the context of the present invention the RS (ratio of
selectivities) of two respective hydrocracking catalysts is defined
as follows:
selectivity hydrocracking catalyst A
RS =
selectivity hydrocracking catalyst B
whereby the selectivities are expressed as selectivities towards the
C1-Cq hydrocarbons fraction (C1-C4), for conversion into 370 °C-
products, when the hydrocracking catalysts are applied under
particular (standard) process conditions. That is to say applying a
particular catalyst volume, temperature, pressure, feed, space
velocity, gas/feed ratio and a reactor loading-method. The
selectivity towards C1-Cq hydrocarbons (SC1-C4) is determined from
the formula:
(C1-C4)product - (C1-C4)feed
SCl-C4 = . 100$
(370 °C+)feed - (370 °C+)product
In the context of the present invention it is further observed
that the RA (ratio of activities) of two respective hydrocracking
catalysts is defined as follows:
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k hydrocracking catalyst A
RA -
k hydrocracking catalyst B
whereby k is defined as the reaction rate constant of the respective
hydrocracking catalysts, for conversion into 370 °C- products, when
applied under particular (standard) process conditions. This means a
particular catalyst volume, temperature, pressure, feed, space
velocity, gas/feed ratio and reactor loading-method.
In the context of the present invention it is further observed
that RV (ratio of volumes) is defined as follows:
reactor volume loaded with hydrocracking catalyst A
RV =
reactor volume loaded with hydrocracking catalyst B
whereby the same particular reactor loading-method is applied for
both hydrocracking catalysts.
Both the RS and the RA are to be determined after the hydro-
cracking catalysts have been allowed to stabilize under the
particular (standard) process conditions.
A set of particular (standard) process conditions as mentioned
hereinabove may include a temperature of 390 °C, an average hydrogen
partial pressure of 12.4 MPa (124 bar), a space velocity of 0.6
kg/1/hr, a hydrotreated flashed distillate containing 23 ppmw N, a
gas/feed ratio of 2000 N1/kg and a dense-bed loading method.
It is expected for a catalyst system of a hydrocracking
catalyst A and a hydrocracking catalyst B that the gas make and the
middle distillate yield will be the average of the individual
contributions of the catalysts weighed by their respective
activities and volumes. However, in accordance with the present
invention surprisingly a lower gas make and a higher middle
distillate yield, than expected, can be obtained by selecting a
sequence of hydrocracking catalysts wherein the first hydrocracking
catalyst has a higher gas make and a RA.RV < 5.
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Suitably, the hydrocracking catalysts A and B are selected in
such a way that the RS > 1.1. Suitably, the hydrocracking catalysts
A and B are selected in such a way that 1.1 < RS < 4. Preferably,
the hydrocracking catalysts A and B are selected in such a way that
0.1 < RA. RV < 3.5, more preferably they are selected so that
0.2 < RA.RV < 2.
Suitably, the hydrocracking catalysts A and B are selected so
that RS/RA > 1. Suitably, the hydrocracking catalysts A and B are
selected so that 1 < RS/RA < 3.
At least part of the product obtained over hydrocracking
catalyst A is contacted with hydrocracking catalyst B. Suitably, at
least 50$ by volume of the product obtained over hydrocracking
catalyst A is contacted with hydrocracking catalyst B.
Suitably, at least part of the product obtained over
hydrocracking catalyst A can be recycled to hydrocracking
catalyst A.
Suitably, at least part of the product obtained over
hydrocracking catalyst B can be recycled to hydrocracking catalyst A
and/or hydrocracking catalyst B.
Suitably, the complete product obtained over hydrocracking
catalyst A is contacted with hydrocracking catalyst B.
The hydrocracking catalysts A and B, respectively, can be
arranged in one or more beds with hydrocracking catalyst A and one
or more beds with hydrocracking catalyst B. The bed or beds with
hydrocracking catalyst A and the bed or beds with hydrocracking
catalyst B can be arranged in one or more reactors. Suitably, the
hydrocracking catalysts A and B are applied in stacked-bed
configuration.
Suitably, the RV < 5, preferably in the range from 0.1 to 2.
Suitably, the hydrocracking catalysts A and B are selected in
such a way that RA < 4. Suitably, they are selected in such a way
that 1 < RA < 4.
Preferably, the process is carried out in such a way that the
product obtained over hydrocracking catalyst A comprises at least
50$ by weight of 370 °C- products.
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Suitably, the present process is carried out in such a way that
more than 30$ by weight conversion of 370 °C + boiling point
material is established. Preferably, more than 40$ by weight
conversion is established. Suitably, the amount of C1-C4
hydrocarbons produced is less than 20$ by weight on feed.
The hydrocracking catalysts A and B may be any conventional
hydrocracking catalyst provided that both the RS and RA. RV fulfil
the requirements as set out hereinabove. For example, the
hydrocracking catalysts A and B may be fresh and regenerated forms
respectively of the same catalyst.
Suitably, the hydrocracking catalysts A and B comprise a
support comprising a large pore molecular sieve and a binder.
The molecular sieves have pores larger than 6 ~, preferably
between 6 and 12 ~. Suitable wide pore molecular sieves are
described in the book Zeolite Molecular Sieves by Donald W. Breck,
Robert E. Krieger Publishing Co., Malabar, Fla., 1984. Suitable wide
pore molecular sieves comprise the crystalline aluminosilicates, the
crystalline aluminophosphates, the crystalline
silicaaluminophosphates and the crystalline borosilicates. Preferred
are the crystalline aluminosilicates or zeolites. The zeolites are
preferably selected from the group consisting of faujasite-type and
mordenite-type zeolites. Suitable examples of the faujasite-type
zeolites include zeolite Y and zeolite X. Other wide pore zeolites
such as zeolite L, beta and omega can also be used alone or in
combination with the more preferred zeolites.
The most preferred wide pore zeolite comprises a zeolite Y,
preferably an ultrastable zeolite Y (zeolite USY). The ultrastable
zeolites used herein are well known to those skilled in the art.
They are for instance exemplified in US 3,293,192 and US 3,499,070.
They are generally prepared from sodium zeolite Y using one or more
ammonium ion exchanges followed by steam calcination.
Suitably, hydrocracking catalysts A and B each comprise a wide
pore zeolite. Preferably, both hydrocracking catalysts A and B
comprise a zeolite Y, particularly a modified zeolite Y having a
unit cell size below 2.445 nm (24.95 ~). In the latter case,
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hydrocracking catalyst A has preferably a content of zeolite Y which
is at least equal to the zeolite Y content of hydrocracking catalyst
B. In other words, the ratio of the zeolite Y contents of
hydrocracking catalysts A and B (Y1/Y2) is at least 1. More
preferably this ratio is in the range from 2 to 12.
Hydrocracking catalysts A and B may each comprise an amount of
zeolite Y in the range from 1$ to 95~ by weight, based on total
support.
Hydrocracking catalysts A and B will each further comprise at
least one hydrogenation component of a Group VI metal and/or at
least one hydrogenation component of a Group VIII metal. Suitably,
the catalyst composition according to the present invention
comprises one or more components of nickel and/or cobalt and one or
more components of molybdenum and/or tungsten or one or more
components of platinum and/or palladium.
The amounts) of hydrogenation components) in the
hydrocracking catalysts is preferably in the range from 0.05 to 10$
by weight of Group VIII metal components) and from 2 to 40$ by
weight of Group VI metal component(s), calculated as metals) per
~0 100 parts by weight of total catalyst.
More preferably, the amounts) of hydrogenation components) in
the hydrocracking catalysts is in the range from 0.5 to 8$ by weight
of Group VIII metal components) and from 10 to 25$ by weight of
Group VI metal component(s), calculated as metals) per 100 parts by
weight of total catalyst.
Preferably, the total amount of hydrogenation components) in
hydrocracking catalyst A is equal to or less than the total amount
of hydrogenation components) in hydrocracking catalyst B. In other
words, the ratio of the amounts of hydrogenation components) in
hydrocracking catalysts A and B (hA/hB) is at most 1. More
preferably, hA/hB is in the range from 0.5 to 1. In a very
attractive embodiment of the present invention hA/hB is less than 1
whereas Y1/Y2 is more than 1. In a preferred embodiment of the
present invention hA/hB is in the range from 0.5 to 1 whereas Y1/Y2
is in the range from 2 to 12.
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The hydrogenation components in the hydrocracking catalysts may
be in the oxidic and/or sulphidic form, in particular in the
sulphidic form. If a combination of at least a Group VI and a Group
VIII metal component is present as (mixed) oxides, it will normally
be subjected to a sulphiding treatment prior to proper use in
hydrocracking.
Suitably, the supports of hydrocracking catalysts A and B may
comprise a zeolite Y, a binder and/or a dispersion of silica-alumina
in an alumina matrix.
Preferably, the support of hydrocracking catalyst B comprises
less than 25 by weight of the zeolite Y, more than 25$ by weight of
binder and at least 30$ by weight of the dispersion.
Preference is given to supports of hydrocracking catalyst B
comprising less than 15$ by weight of the zeolite Y.
Preferably, the support of hydrocracking catalyst B has a
binder/zeolite Y weight ratio in the range from 2 to 40.
Suitably, the support of hydrocracking catalyst B comprises 90
to 70$ by weight of the dispersion.
Suitably, the alumina matrix comprises a transitional alumina
matrix, preferably a gamma-alumina matrix.
The binders) present in the supports of hydrocracking
catalysts A and B as described hereinabove suitably comprise
inorganic oxides or mixtures of inorganic oxides. Both amorphous and
crystalline binders can be applied.
Examples of suitable binders comprise alumina, magnesia,
titania and clays. If desired, small amounts of other inorganic
oxides such as zirconia, titania, magnesia and silica may be
present. Alumina is a preferred binder.
Suitably, the crystalline alumonosilicate of the zeolite Y type
to be applied in hydrocracking catalysts A and B comprises a
modified zeolite Y having a unit cell size below 2.435 nm (24.35 ~.),
a degree of crystallinity which is at least retained at increasing
Si02/A1203 molar ratios, a water adsorption capacity (at 25 °C and
p/p0 value of 0.2) of at least 8$ by weight of modified zeolite and
a pore volume of at least 0.25 ml/g wherein between 10$ and 60$ of
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the total pore volume is made up of pores having a diameter of at
least 8 nm. This type of modified zeolite Y has been described in
detail in EP-B-247679, which is herein incorporated by reference.
Preferably, between 10$ and 40~ of the total pore volume of the
modified zeolite Y is made up of pores having a diameter of at least
8 nm.
Suitably, the modified zeolite Y has a water adsorption
capacity of 8 to 10$ by weight of modified zeolite.
Preferably, the modified zeolite Y has a Si02/A1203 molar ratio
in the range from 4 to 25, more preferably in the range from 8 to
15.
Suitable process conditions for the present hydrocracking
process comprise temperatures in the range from 250 to 500 °C,
hydrogen partial pressures of up to 30 MPa (300 bar) and space
velocities in the range from 0.1 to 10 kg feed per litre catalyst
per hour (kg/1/hr). Gas/feed ratios in the range from 100 to
5000 N1/kg can suitably be applied. Preferably, the present
hydrocracking process is carried out at a temperature in the range
from 300 to 450 °C, a hydrogen partial pressure in the range from
2.5 to 20 MPa (25 to 200 bar) and a space velocity in the range from
0.2 to 5 kg feed per litre catalyst per hour. Preferably, gas/feed
ratios in the range from 250 to 2500 N1/kg are applied.
Feedstocks which can suitably be applied in the process
according to the present invention comprise all sorts of
hydrocarbonaceous feedstocks as long as they fulfil the requirement
to contain less than 500 ppmw N. Suitably, the feedstock contains
less than 200 ppmw N. Suitably, the feedstocks comprise gas oils,
deasphalted oils, coker gas oils and other thermally or
catalytically cracked gas oils and syncrudes, optionally originating
from tar sands, shale oils, residue upgrading processes or biomass
or combinations thereof, which may have been hydrotreated before
being contacted with hydrocracking catalyst A. The feedstocks can
for instance suitably be contacted with an alumina containing
hydrotreating catalyst prior to contact with hydrotreating catalyst
A. The feedstock will generally be such that a major part, say over
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50$ by weight, has a boiling point above 370 °C.
Suitably, the complete product obtained from such a hydro-
treating step is applied as feedstock in accordance with the present
invention.
In an attractive embodiment of the present invention use is
made of three reaction zones arranged in series whereby the complete
effluent from a first reaction zone is passed to a second reaction
zone, and the complete effluent from the second reaction zone is
passed to a third reaction zone. The first reaction zone comprises
IO an amorphous hydrotreating catalyst as described hereinbefore, the
second reaction zone comprises a first zeolitic hydrocracking
catalyst which contains at least one metal of Group VIB and/or at
least one metal of Group VIII, and the third reaction zone comprises
a second zeolitic hydrocracking catalyst which contains at least one
metal of Group VIB and/or at least one metal of Group VIII. The
zeolitic hydrocracking catalysts fulfil the RS and RA. RV
requirements in accordance with the invention.
A hydrotreatment as described hereinabove and the process
according to the present invention can suitably be carried out in
reactors in series or in a stacked-bed configuration.
In another embodiment of the present invention the process is
carried out in a two-stage mode of hydrocracking operation. In this
type of operation the effluent obtained from the first reaction zone
comprising an amorphous hydrotreating catalyst is subjected to a
separation treatment to remove from the effluent a gaseous phase and
a liquid phase including a naphtha and a middle distillate fraction.
The remaining effluent is subsequently subjected to the process
according to the present invention, whereby at least a part of the
residual fraction obtained is recycled to hydrocracking catalyst A.
The hydrocarbonaceous feedstock may be the effluent obtained
from one or more hydrocracking stages arranged upstream in respect
of hydrocracking catalyst A. The product obtained from hydrocracking
catalyst B may subsequently be contacted with a further catalyst,
for instance, an amorphous silica-alumina containing catalyst.
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The present invention will be further understood from the
following illustrative example.
Example I
A heavy vacuum gas oil feed having a sulphur content of 2.0 $w
(as determined according to standard test method ASTM D2622), a
nitrogen content of 1000 ppmw (as determined according to standard
test method ASTM D4629), an initial boiling point of 340 °C, a 50?s
boiling point of 470 °C and a final boiling point in excess of 540
°C
is first hydrotreated in the presence of C-424 catalyst (ex
Criterion) to reduce the nitrogen content to less than 200 ppmw.
The hydrotreated feed so obtained is then contacted, in a once-
through operation, in a stacked bed reactor at a temperature of
400 °C, a hydrogen partial pressure of 10.9 x 105 Pa (109 bar) and a
space velocity of one (1) kg of feed per litre of catalyst per hour
(kg.l-l.h-1) with a first bed of hydrocracking catalyst Z-713 (ex
Zeolyst International) (Catalyst A) and then with a second bed of
hydrocracking catalyst Z-603 (ex Zeolyst International) (Catalyst
B). The ratio of the selectivity of Catalyst A to the selectivity
of Catalyst B (RS) is 5.2/3.1, i.e. 1.7. The ratio of the activity
of Catalyst A to the activity of Catalyst B (RA) is 0.95/0.19, i.e.
5.0, and the ratio of the Catalyst A reactor volume to the Catalyst
B reactor volume (RV) is 125/175, i.e. 0.7. Thus RA.RV is 5.0 x 0.7
- 3.5.
The distribution of 300 °C- product (in $w based on total feed)
at 96$ conversion is as follows:
C1 - C4 . 6.4
C5 - 150 °C . 37.6
140 °C - 300 °C . 53.8
Comparison Example
If Example I is repeated but using hydrocracking catalyst Z-603
in the first catalyst bed (Catalyst A') and hydrocracking catalyst
Z-713 in the second catalyst bed (Catalyst B'), then RS is 3.1/5.2 =
0.6, RA is 0.19/0.95 = 0.2, RV is 175/125, i.e. 1.4, and RA.RV is
0.2 x 1.4 = 0.3.
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In this case, the distribution of 300 °C- product (in $w based
on total feed) at 96$ conversion is as follows:
C1 - C4 . 7.0
C5 - 150 °C . 37.6
140 °C - 300 °C . 53.2
Thus, the above results clearly demonstrate that by using the
process of the present invention in which a particular sequence of
hydrocracking catalysts is employed such that RS > 1 and RA.RV < 5,
it is possible to achieve higher yields of middle distillates with
lower gas make compared to processes Where one or both of the RS and
RA. RV values do not fulfil the above requirements.
