Note: Descriptions are shown in the official language in which they were submitted.
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TS 5510
LUBRICATING BASE OIL PREPARATION PROCESS
The present invention relates to a process for the preparation
of lubricating base oils, in particular the preparation of very high
viscosity index lubricating base oils from a mixture of carbon
monoxide and hydrogen.
The term "very high viscosity index" refers to a viscosity
index (VI) above 135, as determined by ASTM-D-2270.
The preparation of hydrocarbons from a mixture comprising
carbon monoxide and hydrogen by contacting the mixture with a
suitable synthesis catalyst at elevated temperatures and pressures
is known in the art as the Fischer-Tropsch synthesis. It is known
in the art to apply Fischer-Tropsch synthesis processes in the
preparation of a range of principally aliphatic hydrocarbons having
a wide range of molecular weights.
In order to improve the yield of valuable hydrocarbon products
from the Fischer-Tropsch synthesis process, a variety of process
schemes have been proposed for upgrading the Fischer-Tropsch
products. Thus, in US patent No. 4 125 566 a process scheme is
disclosed in which the highly olefinic effluent of a Fischer-Tropsch
synthesis is treated by one or more of distillation, polymerisation,
alkylation, hydrotreatment, cracking-decarboxylation, isomerisation
and hydroreforming. This process scheme yields products lying
mainly in the gasoline, kerosene and gasoil ranges.
From the variety of aforementioned processes which may be
applied in upgrading the products of a Fischer-Tropsch synthesis, a
number of process schemes have been proposed which rely upon the
application of hydrotreatment processes in the upgrading. Thus,
US patent specification No. 4 478 955 discloses a process scheme
comprising contacting the effluent of a Fischer-Tropsch synthesis
process with hydrogen in the presence of a suitable hydrogenation
catalyst. The effluent of the Fischer-Tropsch synthesis is
described in this US patent specification as comprising pre-
dominantly olefins and carboxylic acids. Under the action of the
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hydrogenation treatment, useful fuel components comprising alkanes,
alcohols and esters are produced.
In an alternative process scheme disclosed in US patent
specifications Nos. 4 059 648 and 4 080 397, the products of a
Fischer-Tropsch synthesis are upgraded by being subjected firstly to
a hydrotreatment and thereafter fractionated. Selected fractions of
the fractionated product are subsequently subjected to a selective
hydrocracking process in which the fractions are contacted with a
special zeolite catalyst capable of transforming the aliphatic
hydrocarbons present in the fractions into aromatic hydrocarbons.
The resulting aromatic-rich products are said to be useful as
gasoline and light and heavy fuel oils.
More recently, much interest has been paid to the application
of the Fischer-Tropsch synthesis in the preparation of substantially
paraffinic hydrocarbon products suitable for use as fuels. Whilst
it is possible to use the Fischer-Tropsch synthesis process to
directly prepare paraffinic hydrocarbons having boiling points in
the boiling point ranges of the fuel fractions, it has been found
advantageous to use the Fischer-Tropsch synthesis process to prepare
high molecular weight paraffinic hydrocarbons having a boiling point
above the upper limit of the boiling point range of the middle
distillates, that is a Fischer-Tropsch wax having a boiling point
above 370 °C, and subject the products so-obtained to a selective
hydrocracking process to yield the desired hydrocarbon fuels.
Thus, in British patent specification No. 2 077 289, a process
is disclosed comprising contacting a mixture of carbon monoxide and
hydrogen with a catalyst active in the Fischer-Tropsch synthesis and
thereafter cracking the resulting paraffinic hydrocarbons in the
presence of hydrogen to yield middle distillates. A similar process
scheme is disclosed in European patent application No. 0 147 873.
European patent application No. 0 583 836 discloses a process
for the preparation of hydrocarbon fuels from carbon monoxide and
hydrogen.
Of particular interest, however, is the use of the Fischer-
Tropsch synthesis to prepare hydrocarbons suitable for use as
lubricating base oil or lubricating base oil precursor, such as a
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Fischer-Tropsch wax.
Processes for the preparation of very high viscosity index
lubricating base oils from Fischer-Tropsch wax feed are known in the
art. United States patent specification No. 4 594 172 describes the
preparation of very high viscosity index lubricating base oils by
treating a C10-C19 fraction of a Fischer-Tropsch wax with an organic
peroxide to yield a C20+ oligomerised fraction and hydroisomerising
that fraction over a platinum-containing silica-alumina catalyst.
European patent application No. 0 515 256 describes a process
for the hydroisomerisation (hydroconversion) of Fischer-Tropsch
waxes, using a catalyst containing zeolite Y. Upon solvent dewaxing
of a fraction boiling above 380 °C a lubricating base oil is
obtained having a VI of at least 130 and a pour point of at least
-12 °C. As outlined in this patent application, the
hydroisomerisation treatment can be preceded by a hydrogenation
treatment to remove any unsaturated hydrocarbons and oxygenates from
the Fischer-Tropsch wax.
European patent application No. 0 321 303 discloses the
preparation of middle distillate products from Fischer-Tropsch wax
by a hydroisomerisation (hydroconversion) treatment. At least a
portion of the bottoms fraction from the hydroisomerisation zone is
either (a) further processed in a second hydroisomerisation zone or
(b) fractionated and/or dewaxed for the production of a lubricating
oil fraction boiling in the range of 343.3 °C to 510 °C. If
desired,
oxygenates may be removed from the Fischer-Tropsch wax by
distillation. The hydroisomerisation catalyst to be used in the
first hydroisomerisation (hydroconversion) treatment is a platinum
on fluorided alumina catalyst which is particularly effective
(selective) at converting paraffinic Fischer-Tropsch wax to middle
distillate material. In particular, this catalyst is reported to be
more effective than a catalyst containing a zeolite, in particular
zeolite beta, as described in Example 3 of this patent application.
In contrast to this, the afore-mentioned European patent
application No. 0 515 256 describes in Example 3 that a zeolite Y
catalyst can be as effective (selective) in the preparation of
middle distillates from Fischer-Tropsch wax as a platinum on silica-
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alumina catalyst, but that a zeolite Y catalyst is much more active.
For a conversion of about 40$ by weight of Fischer-Tropsch wax to
products boiling below 400 °C, the zeolite Y catalyst required a
reactor temperature of 260 °C, whilst the platinum on silica-alumina
catalyst required a reactor temperature of 340 °C.
A hydroisomerisation (or hydroconversion} process involves both
hydrocracking of paraffinic hydrocarbons and isomerisation of linear
paraffinic hydrocarbons to branched paraffinic hydrocarbons. If is
is desired to prepare lubricating base oils, it is advantageous to
minimise the hydrocracking activity and to maximise the
hydroisomerisation activity. Nevertheless, some hydrocracking
activity is still required to crack the heaviest wax molecules to
lower boiling products. A disadvantage of a highly active catalyst,
such as a zeolitic catalyst like zeolite Y, is that normally the
hydrocracking activity is still too high and the hydroisomerisation
activity too low. As will be discussed hereinafter, other molecular
sieve catalysts are known such as silicoaluminophosphates and other
zeolitic catalysts wherein the activity (expressed in terms of
acidity} has been reduced to an alpha-value below 20 or even below
10 or 5. However, these catalysts normally do not have sufficient
hydrocracking activity.
Thus, European patent application No. 0 469 547 describes a
process for the preparation of high viscosity index lubricants from
slack wax by a two-step process wherein the first step the slack wax
is hydrocracked under mild conditions using an amorphous catalyst,
in particular a catalyst comprising Ni and W on a fluorided alumina
carrier, and in the second step hydroisomerised using a low acidity
zeolite beta catalyst, preferably having an alpha value of not
greater than 5.
The alpha value is an approximate indication of the catalytic
cracking activity of the catalyst compared to a standard catalyst.
The alpha test gives the relative rate constant (rate of normal
hexane conversion per volume of catalyst per unit time) of the test
catalyst relative to the standard catalyst which is taken as an
alpha of 1 (Rate Constant = 0.016 sec-1). The alpha test is
described in United States patent specification No. 3 354 078 and in
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J. Catalysis, 4, 527 (1965) 6, 278 (1966) and 61, 395 (1980) and
discussed in European patent application No. 0 464 547. The alpha
value is determined on the catalyst carrier not containing any
catalytically active metals.
For the purposes of this specification the term hydroconversion
process as used hereinafter refers to a process in which
hydrocracking reactions and hydroisomerisation reactions occur and
which is carried out in the presence of a catalyst comprising a
refractory oxide carrier.
The term hydroisomerisation process as used hereinafter refers
to a process in which hydroisomerisation reactions and hydrocracking
reactions occur, but which is carried out after a hydroconversion
treatment and in which process generally less hydrocracking occurs
than in the hydroconversion process.
A disadvantage of hydroconversion catalysts comprising a
refractory oxide carrier is the high operating temperature which is
required, in particular if the catalyst has been used for a
prolonged time such as more than 2 years. In order to compensate for
any catalyst deactivation generally the reaction temperature is
increased. Above 350 °C and in particular above 400 °C at least
part
of the Fischer-Tropsch wax is converted into undesired aromatic
compounds. Accordingly, it would be desirable to be able to provide
a process which allows a hydroconversion process to be carried out
at an operating temperature well below 400 °C and preferably below
350 °C, whilst using a catalyst containing a refractory oxide
carrier and a catalytically active metal having hydrogenation/-
dehydrogenation activity.
It has now surprisingly been found that if the Fischer-Tropsch
wax (hydrocarbon wax) is first contacted with hydrogen in the
presence of a hydrogenation catalyst under conditions such that
substantially no hydroisomerisation or hydrocracking of the
hydrocarbon wax occurs, the hydroconversion process can be operated
at a reaction temperature below 400 °C, and even below 350 °C,
even
after the hydroconversion catalyst has been used for a prolonged
time, that is for example more than 2 years. As compared with a
process in which the Fischer-Tropsch wax is contacted with a
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hydroconversion catalyst without a preceding hydrogenation step, the
reaction temperature can be at least 5 °C, preferably at least 10
°C
lower.
Further, it has been found that the hydroconversion catalyst
deactivates much slower in a process which involves a preceding
hydrogenation step. Thus, the rate of reaction temperature increase
required to compensate for any loss of catalyst activity can be much
slower.
Further, most surprisingly, it has been found that the process
involving a hydrogenation step and a hydroconversion step exhibits a
higher selectivity to valuable hydrocarbons boiling in the
lubricating base oil range, compared with a prior art process
comprising only a hydroconversion step.
Therefore, the present invention relates to a process for the
preparation of lubricating base oils comprising subjecting a waxy
raffinate to a pour point reducing treatment, and recovering a
lubricating base oil therefrom, which waxy raffinate has been
prepared by contacting a hydrocarbon product with hydrogen in the
presence of a hydroconversion catalyst, comprising a catalytically
active metal having hydrogenation/dehydrogenation activity supported
on a refractory oxide carrier, under conditions such that
hydrocracking and hydroisomerisation of the hydrocarbon product
occur to yield the waxy raffinate, wherein the hydrocarbon product
has been prepared by:
(a) contacting a mixture of carbon monoxide and hydrogen with a
hydrocarbon synthesis catalyst at elevated temperature and pressure
to prepare a substantially paraffinic hydrocarbon wax; and
(b) contacting the hydrocarbon wax so-obtained with hydrogen in the
presence of a hydrogenation catalyst under conditions such that
substantially no hydroisomerisation or hydrocracking of the
hydrocarbon wax occurs to yield the hydrocarbon product.
For the purposes of this specification, conditions such that
substantially no hydrocracking or hydroisomerisation occurs in step
(b) of the process of the present invention are defined as such
conditions that the conversion in step (b) of the fraction of the
feed boiling above 370 °C in $ by weight, to a fraction boiling
CA 02141925 2005-02-11
63293-3650
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below 370°C, is less than 10%.
In the hydroconversion step conditions such that
hydrocracking and hydroisomerisation of the hydrocarbon
product occur are defined as such conditions that the
conversion as defined hereinabove is at least 15%.
In a preferred embodiment, the invention relates to
a process for the preparation of lubricating base oils
comprising subjecting a waxy raffinate to a hydroisomerisation
process involving contacting the waxy raffinate with hydrogen
in the presence of a hydroisomerisation catalyst, wherein the
hydroisomerisation catalyst comprises a molecular sieve and a
Group VIII noble metal, at a temperature of from 175 to 380°C
and a pressure from 10 to 250 bars, such that conversion of
products boiling above 370°C to products boiling below 370°C
is less than 25%, and recovering a lubricating base oil
therefrom having a pour point of less than -20°C, which waxy
raffinate has been prepared by contacting a hydrocarbon
product with hydrogen in the presence of a hydroconversion
catalyst, comprising a catalytically active metal having
hydrogenation/dehydrogenation activity supported on a
refractory oxide carrier under conditions such that
hydrocracking and hydroisomerisation of the hydrocarbon
product occur at a conversion of at least 25% to yield the
waxy raffinate, wherein the hydrocarbon product has been
prepared by: (a) contacting a mixture of carbon monoxide and
hydrogen with a hydrocarbon synthesis catalyst at elevated
temperature and pressure to prepare a substantially paraffinic
hydrocarbon wax; and (b) contacting the hydrocarbon wax so-
obtained with hydrogen in the presence of a hydrogenation
catalyst at a temperature of between 100 and 300°C and under
conditions such that a percent weight fraction of a feed
boiling above 370°C which is converted to a fraction boiling
below 370°C is below 10% to yield the hydrocarbon product.
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For the purposes of this specification, the term "substantially
paraffinic" when used in connection with hydrocarbon wax ,refers to a
hydrocarbon mixture comprising at least 70 %wt (% by weight)
paraffins, preferably at least 80 %wt paraffins. Hydrocarbon wax
produced by the process of this invention typically comprises at
least 90 %wt paraffins, more typically at least 95 %wt paraffins.
In step (a) of the process of the present invention, a feed
comprising a mixture of carbon monoxide and hydrogen is contacted at
elevated temperature and pressure with a catalyst active in the
synthesis of paraffinic hydrocarbons. Suitable processes for the
preparation of the mixture of carbon monoxide and hydrogen are well
known in the art and include such processes as the partial oxidation
of methane, typically in the form of natural gas, and the steam
reforming of methane. The relative amounts of carbon monoxide and
1$ hydrogen present in the feed may vary over a wide range and may be
selected according to the precise catalyst and process operating
conditions being employed. Typically, the feed contacting the
catalyst comprises carbon monoxide and hydrogen in a hydrogen/carbon
monoxide molar ratio of below 2.5, preferably below 1.75. More
preferably, the hydrogen/carbon monoxide ratio is in the range of
from 0.9 to 1.5, especially from 0.9 to 1.3. Unconverted carbon
monoxide and/or hydrogen may be separated from the synthesis product
and recycled to the inlet of the synthesis reactor.
Suitable catalysts for use in the synthesis of paraffinic
hydrocarbons are known in the art. Typically, the catalyst
comprises, as the catalytically active component, a metal from Group
VIII of the Periodic Table of the Elements. Particular
catalytically active metals from Group VIII include ruthenium, iron,
cobalt and nickel. For the process of the present invention, a
catalyst comprising cobalt as the catalytically active ~tal is
preferred.
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The catalytically active metal is preferably supported on a
porous carrier. The porous carrier may be selected from any
suitable refractory metal oxide or silicate or mixture thereof.
Particular examples of preferred carriers include silica, alumina,
titania, zirconia and mixtures thereof. Carriers comprising silica
and/or alumina are especially preferred.
The catalytically active metal may be applied to the carrier by
any of the techniques known in the art, for example comulling,
impregnation or precipitation. Impregnation is a particularly
preferred technique, in which the carrier is contacted with a
compound of the catalytically active metal in the presence of a
liquid, most conveniently in the form of a solution of the metal
compound. The compound of the active metal may be inorganic or
organic, with inorganic compounds being preferred, in particular
nitrates. The liquid employed may also be either organic or
inorganic. Water is a most convenient liquid.
The amount of catalytically active metal present on the carrier
is typically in the range of from 1 to 100 parts by weight,
preferably 10 to 50 parts by weight, per 100 parts by weight of
carrier material.
The catalytically active metal may be present in the catalyst
together with one or more metal promoters or co-catalysts. The
promoters may be present as metals or as the metal oxide, depending
upon the particular promoter concerned. Suitable metal oxide
promoters include oxides of metals from Groups IIA, IIIB, IVB, VB or
VIB of the Periodic Table, oxides of the lanthanides and/or the
actinides. Preferably, the catalyst comprises an oxide of an
element in Group IVB of the Periodic Table, in particular titanium
or zirconium. Catalysts comprising zirconium are especially
preferred. As an alternative or in addition to the metal oxide
promoter, the catalyst may comprise a metal promoter selected from
Groups VIIB and/or VIII of the Periodic Table. Preferred metal
promoters include platinum and palladium. A most suitable catalyst
comprises cobalt as the catalytically active metal and zirconium as
a promoter. The promoter may be incorporated in the catalyst using
any of the methods discussed hereinbefore with respect to the
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catalytically active component.
The promoter, if present in the catalyst, is typically present
in an amount of from 1 to 60 parts by weight, preferably from 2 to
40 parts by weight, per 100 parts by weight of carrier material.
The hydrocarbon synthesis is conducted under conditions of
elevated temperature and pressure. Typically, the synthesis is
effected at a temperature in the range of from 125 to 300 °C,
preferably from 175 to 250 °C. The reaction pressure is typically
in the range of from 5 to 100 bar, preferably from 12 to 50 bar.
The synthesis may be conducted using a variety of reactor types and
reaction regimes, for example in a fixed bed regime, a slurry phase
regime or an ebullating bed regime.
The hydrocarbon wax of the synthesis step (a) is subjected to a
hydrogenation treatment in step (b) of the process of the present
invention. The entire effluent of the synthesis stage may be led
directly to the hydrogenation step. However, it is preferred to
separate from the hydrocarbon wax of the synthesis stage unconverted
carbon monoxide and hydrogen and water formed during the synthesis.
If desired, the low molecular weight products of the synthesis
stage, in particular the Cq- fraction, for example methane, ethane
and propane, may also be removed prior to the hydrogenation
treatment. The separation is conveniently effected using
distillation techniques well known in the art. Alternatively, the
hydrocarbon wax can be separated into a low boiling fraction,
boiling for example below 330 °C or below 370 °C and at least
one
high boiling fraction, boiling above 330 °C or above 370 °C and
treat
the high boiling fraction in the process of the present invention.
Separation may be effected using vacuum distillation or
alternatively short path distillation such as vacuum film
distillation (using wiped film evaporators).
In the hydrogenation stage, stage (b), the hydrocarbon product
is contacted with hydrogen in the presence of a hydrogenation
catalyst. Suitable catalysts for use in this stage are known in the
art. Typically, the catalyst comprises as catalytically active
component one or more metals selected from Groups VIB and VIII of
the Periodic Table of Elements, in particular one or more metals
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selected from molybdenum, tungsten, cobalt, nickel, ruthenium,
iridium, osmium, platinum and palladium. Preferably, the catalyst
comprises on or more metals selected from nickel, platinum and
palladium as the catalytically active component.
A particularly suitable catalyst comprises nickel as a
catalytically active component.
Catalysts for use in the hydrogenation stage typically comprise
a refractory metal oxide or silicate as a carrier. Suitable carrier
materials include silica, alumina, silica-alumina, zirconia, titania
and mixtures thereof. Preferred carrier materials for inclusion in
the catalyst for use in the process of this invention are silica,
alumina, silica-alumina, and diatomaceous earth (kieselguhr).
The catalyst may comprise the catalytically active component in
an amount of from 0.05 to 80 parts by weight, preferably from 0.1 to
70 parts by weight, per 100 parts by weight of carrier material.
The amount of catalytically active metal present in the catalyst
will vary according to the specific metal concerned. One
particularly suitable catalyst for use in the hydrogenation stage
comprises nickel in an amount in the range of from 30 to 70 parts by
weight per 100 parts by weight of carrier material. A second
particularly suitable catalyst comprises platinum in an amount in
the range of from 0.05 to 2.0 parts by weight per 100 parts by
weight of carrier material.
Suitable catalysts for use in the hydrogenation stage of the
process of this invention are available commercially, or may be
prepared by methods well known in the art, for example the methods
discussed hereinbefore with reference to the preparation of the
hydrocarbon synthesis catalyst.
In the hydrogenation stage, the hydrocarbon wax is contacted
with hydrogen at elevated temperature and pressure. The operating
temperature may typically range from 100 to 300 °C, more preferably
from 150 to 275 °C, in particular from 175 to 250 °C. Typically,
the
operating pressure ranges from 5 to 150 bars, preferably from 10 to
50 bars. Hydrogen may be supplied to the hydrogenation stage at a
gas hourly space velocity in the range of from 100 to 10000 N1/1/hr,
more preferably from 250 to 5000 N1/1/hr. The hydrocarbon wax being
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treated is typically supplied to the hydrogenation stage at a weight
hourly space velocity in the range of from 0.1 to 5 kg/1/hr, more
preferably from 0.25 to 2.5 kg/1/hr. The ratio of hydrogen to
hydrocarbon wax may range from 100 to 5000 Nl/kg and is preferably
from 250 to 3000 N1/kg.
The hydrogenation stage is operated under conditions such that
substantially no isomerisation or hydrocracking of the feed occurs.
The precise operating conditions required to achieve the desired
degree of hydrogenation without substantial hydrocracking or
hydroisomerisation occurring will vary according to the composition
of the hydrocarbon wax being fed to the hydrogenation stage and the
particular catalyst being employed. As outlined hereinbefore, as a
measure of the severity of the conditions prevailing in the
hydrogenation stage and, hence, the degree of hydrocracking and
isomerisation occurring, the degree of conversion of the feed
hydrocarbon may be determined. In this respect, conversion, in
percent, is defined as the percent weight of the fraction of the
feed boiling above 370 °C which is converted during the
hydrogenation to a fraction boiling below 370 °C. The conversion of
the hydrogenation stage is below 10$, preferably below 8$, more
preferably below 5$.
In the process of the present invention, the hydrocarbon
product leaving the hydrogenation stage substantially consists of
high molecular weight, paraffinic hydrocarbons having a boiling
point range within and above that of lubricating base oils.
Lubricating base oils typically have a 5$ by weight boiling point of
at least 330 °C, preferably at least 370 °C. The boiling point
range
of lubricating base oils may range up to 650 °C, preferably up to
600 °C. It will be appreciated that the above boiling points and
boiling point ranges refer to boiling point (ranges) at atmospheric
pressure. At least a part of this hydrocarbon product is subjected
to the hydroconversion step of the process of this invention, to
yield the waxy raffinate. If desired, the entire effluent of the
hydrogenation stage may be led directly to the hydroconversion
stage. However, it is preferred to separate the low molecular
weight hydrocarbons, especially the Cq- fraction, from the higher
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molecular weight hydrocarbons prior to the hydroconversion stage.
The separation may be conveniently achieved using distillation
techniques well known in the art. At least a part of the remaining
C5+ fraction of the hydrocarbon product is then used as feed for the
hydroconversion stage.
In the hydroconversion stage, waxy raffinate is prepared from
the hydrocarbon product of the hydrogenation stage by hydrocracking
and hydroisomerising the product with hydrogen in the presence of a
suitable catalyst. Typically, the catalyst comprises as
catalytically active component one or more metals selected from
Groups VIB and VIII of the Periodic Table of Elements, in particular
one or more metals selected from molybdenum, tungsten, cobalt,
nickel, ruthenium, iridium, osmium, platinum and palladium.
Preferably, the catalyst comprises one or more metals selected from
nickel, platinum and palladium as the catalytically active
component. Catalysts comprising platinum as the catalytically
active component have been found to be particularly suitable for use
in the hydroconversion stage.
Catalysts for use in the hydroconversion stage typically
comprise a refractory metal oxide as a carrier. Suitable carrier
materials include silica, alumina, silica-alumina, zirconia, titania
and mixtures thereof. Preferred carrier materials for inclusion in
the catalyst for use in the process of this invention are silica,
alumina and silica-alumina. A particularly preferred catalyst
comprises platinum supported on a silica-alumina carrier. If
desired, the acidity of the catalyst carrier may be enhanced by
applying a halogen moiety, in particular fluorine, or a phosphorous
moiety to the carrier. This may be especially preferred if the
catalyst carrier itself is not acidic, for example if the catalyst
carrier contains alumina or silica.
The catalyst may comprise the catalytically active component in
an amount of from 0.05 to 80 parts by weight, preferably from 0.1 to
70 parts by weight, per 100 parts by weight of carrier material.
The amount of catalytically active metal present in the catalyst
will vary according to the specific metal concerned. A particularly
preferred catalyst for use in the hydroconversion stage comprises
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platinum in an amount in the range of from 0.05 to 2 parts by
weight, more preferably from 0.1 to 1 parts by weight, per 100 parts
by weight of carrier material.
Suitable catalysts for use in the hydroconversion stage of the
process of this invention are available commercially, or may be
prepared by methods well known in the art, for example the methods
discussed hereinbefore with reference to the preparation of the
hydrocarbon synthesis catalyst.
In the hydroconversion stage of this process, the hydrocarbon
product of the hydrogenation stage is contacted with hydrogen in the
presence of the catalyst at elevated temperature and pressure.
Typically, the temperatures necessary to yield the waxy raffinate
will lie in the range of from 175 to 380 °C, preferably from 250 to
350 °C, more preferably from 250 to 330 °C. The pressure
typically
applied ranges from 10 to 250 bars, more preferably from 25 to
250 bars. Hydrogen may be supplied at a gas hourly space velocity
of from 100 to 10000 N1/1/hr, preferably from 500 to 5000 N1/1/hr.
The hydrocarbon feed may be provided at a weight hourly space
velocity of from 0.1 to 5 kg/1/hr, preferably from 0.25 to
2 kg/1/hr. The ratio of hydrogen to hydrocarbon feed may range from
100 to 5000 N1/kg and is preferably from 250 to 2500 N1/kg.
As discussed hereinbefore in connection with the hydrogenation
stage, the degree of hydrocracking and isomerisation occurring in
the hydroconversion stage may be measured by determining the degree
of conversion of the fraction boiling above 370 °C, as hereinbefore
defined. Typically, the hydroconversion stage is operated at a
conversion of at least 20$, preferably at least 25$, but preferably
not more than 50~, more preferably not more than 45$.
The hydrogen required for the operation of both the
hydrogenation and the hydroconversion stages may be generated by
processes well known in the art, for example by the steam reforming
of a refinery fuel gas.
The waxy raffinate is then subjected to a pour point reducing
treatment to reduce the pour point to at least -12 °C, preferably at
least -18 °C, more preferably at least -24 °C.
Pour point reducing treatments are well known to those skilled
2m~9'z~
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in the art and include solvent dewaxing, catalytic dewaxing, (hydro)
isomerisation (dewaxing) and/or addition of pour point depressing
agents. The latter treatment is generally not preferred in the
preparation of lubricating base oils as additives contained in a
base oil may deteriorate rather quickly and blending of different
base oils and additive packages to prepare a finished base oil may
become a problem.
Catalytic dewaxing is well known to those skilled in the art.
In a catalytic dewaxing process straight chain paraffins and
slightly branched paraffins are cracked to products boiling below
the lubricating base oil boiling point range. However, the catalysts
Which are employed are not fully selective to wax molecules only. In
fact also branched paraffins having a very high VI and a
sufficiently low pour point are cracked to lower boiling products.
Thus, cracking of those compounds to products boiling below the
lubricating base oil range results in lubricating base oils having a
lower VI than lubricating base oils which have been prepared by a
solvent dewaxing process. Further, as compared with solvent
dewaxing, lubricating base oils are prepared in lower yield.
Nevertheless, catalytic dewaxing is being applied commercially as
generally the operation of the process is cheaper than operation of
a solvent dewaxing process.
Catalysts which can be used in a catalytic dewaxing process
include zeolites having a constraint index from 1 to 12, in
particular of the MFI structure type, such as ZSM-5, -11, -22, -23,
-35 as well as Ferrierite and composite crystalline silicates
described in European patent application Nos. 0 100 115, 0 178 699
and 0 380 180. Another suitable catalytic dewaxing catalyst
comprises mordenite. If desired, waxy raffinate feed can be
separated into various fractions and the various fraction can then
be treated separately using different dewaxing catalysts, as
disclosed in European patent application No. 0 237 655 and also
European patent application No. 0 161 833. The catalysts typically
contain at least one catalytically active metal chosen from Groups
VIb, VIIb and VIII of the Periodic Table of the Elements.
The catalytic dewaxing process is typically carried out at a
2~41~2~
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temperature from 200 °C to 500 °C, a hydrogen pressure from 5 to
100
bar, a space velocity from 0.1 to 5 kg/1/h and a hydrogen/oil ratio
from 100 to 2500 N1/kg.
Solvent dewaxing is well known to those skilled in the art and
involves admixture of one or more solvents and/or wax precipitating
agents with waxy raffinate and cooling the mixture to a temperature
in the range of from -10 °C to -40 °C, preferably in the range
of
from -20 °C to -35 °C, to separate the wax from the oil. The oil
containing the wax is usually filtered through a filter cloth which
can be made of textile fibres, such as cottony porous metal cloth:
or cloth made of synthetic materials.
Examples of solvents which may be employed in the solvent
dewaxing process are C3-C6 ketones (e. g. methyl ethyl ketone, methyl
isobutyl ketone and mixtures thereof), Cg-C10 aromatic hydrocarbons
(e. g. toluene), mixtures of ketones and aromatics (e. g. methyl ethyl
ketone and toluene), autorefrigerative solvents such as liquefied,
normally gaseous C2-Cq hydrocarbons such as propane, propylene,
butane, butylene and mixtures thereof. Mixtures of methyl ethyl
ketone and toluene or methyl ethyl ketone and methyl isobutyl ketone
are generally preferred.
The solvents may be recovered from the wax and the lubricating
base oil by filtration and recirculation of the solvents into the
process. It will be appreciated that although solvents are
recirculated the process is still rather expensive as a large amount
of solvents is required and cooling of the waxy raffinate/solvent
mixture requires much energy.
The wax that is separated in the solvent dewaxing process may
be recycled to the hydroconversion stage, or alternatively, may be
sent to a hydroisomerisation stage if for example, the pour point
reducing treatment involves both a solvent dewaxing stage and a
hydroisomerisation stage. The wax may be subjected to a deoiling
treatment prior to recycling. Another possibility is to fractionate
the wax and sell one or more of the fractions on the wax market.
Fractionation is typically effected using short path distillation.
A very suitable pour point reducing treatment comprises a
hydroisomerisation treatment, in the art sometimes referred to as
21~~~~z
- 16 -
isomerisation dewaxing or iso-dewaxing. The hydroisomerisation
treatment typically comprises contacting the waxy raffinate with
hydrogen in the presence of a hydroisomerisation catalyst. As
compared with a solvent dewaxing treatment the hydroisomerisation
treatment is cheaper to operate and the hydroisomerisation treatment
does substantially not suffer from the disadvantages of catalytic
dewaxing, that is a lower VI and a lower yield as compared with
solvent dewaxing. In the hydroisomerisation process straight chain
paraffins are isomerised to branched paraffins boiling within the
lubricating base oil boiling point range, still having a high VI but
also a low pour point. It will be appreciated that whilst
hydroisomerisation reactions are preferred, hydrocracking reaction
need to be avoided as much as possible. The conversion of products
boiling above 370 °C to products boiling below 370 °C typically
is
less than 25~, preferably less than 205. The conversion is usually
more than 10$.
Thus, preferably a hydroisomerisation catalyst is used having a
high activity for catalysing hydroisomerisation reactions, but a low
activity for catalysing hydrocracking reactions. It has been found
that in order to achieve this the acidity of the catalyst, as
expressed by the alpha value, should be below 20. Preferably the
catalyst comprises a molecular sieve. Accordingly, in a preferred
embodiment the hydroisomerisation catalyst comprises a molecular
sieve having an alpha value below 20, more preferably below 10, even
more preferably below 5. The experimental conditions of the alpha
test to be used to determine the alpha values referred to in this
specification include a constant temperature of 538 °C and a
variable flow rate as described in detail in J. Catalysis, 61, 395
(1980).
In one embodiment of the invention, the molecular sieve
operating in hydroisomerisation duty, is a zeolite, preferably
having a silica/alumina molar ratio of at least 10, more preferably
at least 30. A zeolite having a high silica/alumina ratio generally
has a lower acidity than a zeolite having a low silica/alumina
ratio. A high silica/alumina ratio may be obtained by synthesis of
the zeolite at a high silica/alumina ratio and/or by a dealumination
21~~.92~
- 17 -
treatment such as steaming. Both methods are well known to those
skilled in the art. Alternatively, framework aluminium may be
replaced by another trivalent element such as boron which results in
a lower level of acidity.
Preferably, the molecular sieve is chosen from the group of
ZSM-12, mordenite and zeolite beta, more preferably zeolite beta.
The low acidity forms of zeolite beta may be obtained by synthesis
of a highly siliceous form of the zeolite, e.g. with a
silica/alumina ratio above 50 or by steaming zeolites of lower
silica-alumina ratio to the requisite acidity level. Another method
is by replacement of a portion of the framework aluminium of the
zeolite with another trivalent element such as boron. Preferably,
the zeolite contains framework boron, typically at least 0.1$ by
weight, preferably, at least 0.5$ by weight. The zeolite also may
contain material in the pores of the structure which do not form
part of the framework constituting the characteristic structure of
the zeolite. The term "framework boron" as used herein refers to
boron that is actually present in the framework of the zeolite. As
opposed to material present in the pores of the zeolite, framework
boron contributes to the ion exchange capacity of the zeolite.
Methods for preparing zeolites having a high silica/trivalent
metal ratio and containing framework boron have been described in
United States patent Nos. 4 269 813 and 4 672 049. Typically,
zeolite beta to be used in the process of the present invention
contains at least 0.1$ by weight. The boron content will usually be
not more than 5$ by weight, preferably not more than 2$ by weight.
The silica/alumina ratio of the as-synthesized zeolite is typically
below 30. Preferably, the boron containing zeolite is steamed to
reduce the alpha value to not more than 10, preferably not more than
5. Typical steaming conditions are known and have been described in
European patent application No. 0 464 547.
The zeolite will usually be composited with a matrix material
(binder) to form the finished catalyst. Non-acidic refractory oxide
binder materials like silica, titania or alumina are preferred.
Silica is especially preferred. The zeolite is usually composited
with the matrix in amounts from 20 to 80$ by weight, preferably from
21~~~~
- 18 -
50 to 80$ by weight. Methods for extruding zeolite with the binder
are known to those skilled in the art.
In another embodiment of the invention the molecular sieve is
an aluminophosphate. Aluminophosphates are well known in the art and
have been described in, for example, United States patent
Nos. 4 310 440, 4 440 871, 4 567 029, and 4 793 984. Alumino-
phosphates have the advantage of an intrinsically lower acidity as
compared with zeolites.
For the purposes of this specification reference to
aluminophosphates is to be understood as reference to the class of
aluminophosphates, that is including metallo-aluminophosphates,
silico-aluminophosphates, metallo-silico-aluminophosphates as well
as non-metal substituted aluminophosphates and silico-
aluminophosphates.
Preferably, the molecular sieve has at least some intrinsic
acidity and therefore the aluminophosphate is chosen from the group
of metallo-aluminophosphates, wherein the further metal present in
the framework of the aluminophosphate is not a trivalent metal,
silico-aluminophosphates or metallo-silico-aluminophosphates. In a
particularly preferred embodiment the aluminophosphate is a silico-
aluminophosphate.
In one embodiment, the process is preferably carried out with a
hydroisomerisation catalyst comprising an aluminophosphate, in
particular a silico-aluminophosphate, chosen from the group of
structure types 11, 31 and 41, more preferably structure type 11.
Silico-aluminophosphates of structure types 11, 31 and 41 have been
described in international patent application No. WO 90/09362. In a
particular preferred embodiment of the invention, the hydro-
isomerisation catalyst comprises a silico-aluminophosphate of
structure type 11, and having a special silica/alumina distribution
over the crystalline particle. In particular, the silico-
aluminophosphate molecular sieve is characterised by an X-ray
diffraction pattern according to Table I.
21419'
- 19 -
TABLE I
d I/I * 100
9.41-9.17 m
4.37-4.31 m
4.23-4.17 vs
4.02-3.99 m
3.95-3.92 m
~.84-3.81
m = 20-70
s = 70-90
vs = 90-100
wherein I/IO * 100 is the relative intensity, where IO is the
intensity of the strongest line, and d is the interplanar spacing in
angstroms corresponding to the recorded lines. X-ray powder
diffraction patterns can be determined by standard techniques, using
K-alpha/doublet copper radiation.
The silico-aluminophosphate is further characterised by a P205
to alumina molar ratio at the surface of 0.80 or less and a P205 to
alumina ratio of the bulk of the silico-aluminophosphate is 0.96 or
greater, and the silica to alumina mole ratio at the surface is
greater than in the bulk of the silico-aluminophosphate. This
silico-aluminophosphate is known in the art as SM-3. Preparation of
SM-3 has been disclosed in international patent application No.
WO 91/13132.
In yet another embodiment of the invention, the isomerisation
catalyst comprises an inorganic, non-layered, porous, crystalline
phase material as described in international patent application No.
WO 93/02161.
The hydroisomerisation catalyst typically comprises a
catalytically active metal having hydrogenation/dehydrogenation
activity, such as those of Groups VIb and VIII. Preferably, the
hydroisomerisation catalyst comprises a Group VIII metal, in
particular a Group VIII noble metal such as platinum and/or
- 20 -
palladium. Means for incorporating the metal into the catalyst
carrier, comprising a molecular sieve as described hereinbefore, are
well known to those skilled in the art and have been described
hereinbefore. The amount of noble metal is typically in the range of
from 0.5 to 5$ by weight of the total catalyst, preferably in the
range of from 0.5 to 2~ by weight.
In the hydroisomerisation stage of this process, the waxy
raffinate is contacted with hydrogen in the presence of a catalyst
as described hereinabove at elevated temperature and pressure.
Typically, the temperatures necessary to yield the lubricating base
oil will lie in the range of from 175 to 380 °C, preferably from 200
to 350 °C. The pressure typically applied ranges from 10 to 250
bars, more preferably from 25 to 250 bars. The waxy raffinate may
be provided at a weight hourly space velocity of from 0.1 to
20 kg/1/hr, preferably from 0.1 to 5 kg/1/hr.
The finished lubricating base oil preferably has a pour point
of less than -15 °C, more preferably less than -20 °C and a VI
of
more than 135, preferably more than 140. Preferably, it should not
be necessary to treat the hydrogenated hydroconverted hydro-
isomerised lubricating base oil with an additional pour point
reducing treatment. However, it may be desired to isomerise the waxy
raffinate to an intermediate pour point and solvent dewax the
isomerised waxy raffinate to the final pour point.
