U.S. patent application number 10/517100 was filed with the patent office on 2006-03-23 for process for the hydroprocessing of heavy hydrocarbon feeds using at least two reactors.
Invention is credited to Toshiyuki Ado, Katsushisa Fujita, Takeshi Hashiguchi, Naoto Kimbara, Yoshimitsu Miyauchi.
Application Number | 20060060509 10/517100 |
Document ID | / |
Family ID | 29727749 |
Filed Date | 2006-03-23 |
United States Patent
Application |
20060060509 |
Kind Code |
A1 |
Miyauchi; Yoshimitsu ; et
al. |
March 23, 2006 |
Process for the hydroprocessing of heavy hydrocarbon feeds using at
least two reactors
Abstract
The invention pertains to a process for hydroprocessing a heavy
hydrocarbon feed using at least two reactors in which a heavy
hydrocarbon feed is subjected sequentially to the steps of
hydroprocessing in a first hydroprocessing reactor, in which it is
subjected sequentially to a hydrodemetallization step, a
hydrodesulfurization step carried out at a temperature higher than
that of said hydrodemetallization step, and an asphaltene removal
step carried out at a temperature higher than that of said
hydrodesulfurization step, hydroprocessing in a second
hydroprocessing reactor, in which it is subjected sequentially to a
hydrodesulfurization step and an asphaltene removal step, which
latter is carried out at a temperature higher than that of said
hydrodesulfurization step.
Inventors: |
Miyauchi; Yoshimitsu;
(Minato-ku, JP) ; Ado; Toshiyuki; (Kanagawa,
JP) ; Hashiguchi; Takeshi; (Singapore, JP) ;
Kimbara; Naoto; (Minato-ku, JP) ; Fujita;
Katsushisa; (Ehime, JP) |
Correspondence
Address: |
Akzo Nobel;Intellectual Property Department
7 Livingstone Avenue
Dobbs Ferry
NY
10522-3408
US
|
Family ID: |
29727749 |
Appl. No.: |
10/517100 |
Filed: |
June 5, 2003 |
PCT Filed: |
June 5, 2003 |
PCT NO: |
PCT/EP03/06033 |
371 Date: |
August 18, 2005 |
Current U.S.
Class: |
208/210 ;
208/211; 208/212; 208/251H |
Current CPC
Class: |
B01J 35/10 20130101;
B01J 35/1061 20130101; C10G 65/04 20130101; B01J 35/1019
20130101 |
Class at
Publication: |
208/210 ;
208/211; 208/212; 208/251.00H |
International
Class: |
C10G 65/02 20060101
C10G065/02; C10G 65/04 20060101 C10G065/04 |
Foreign Application Data
Date |
Code |
Application Number |
Jun 11, 2002 |
JP |
2002-169868 |
Claims
1. A process for hydroprocessing a hydrocarbon feed containing
Conradson carbon using at least two reactors in which the
hydrocarbon feed is subjected sequentially to the steps of
hydroprocessing in a first hydroprocessing reactor, in which it is
subjected sequentially to a hydrodemetallization step, a
hydrodesulfurization step carried out at a temperature higher than
that of said hydrometallization step, and an asphaltene removal
step carried out at a temperature higher than that of said
hydrodesulfurization step, hydroprocessing in a second
hydroprocessing reactor, in which it is subjected sequentially to a
hydrodesulfurization step and an asphaltene removal step, which
latter is carried out at a temperature higher than that of said
hydrodesulfurization step.
2. The process of claim 1 wherein the hydrodemetallization step is
carried out using a hydrodemetallization catalyst, the
hydrodesulfurization step is carried out using a
hydrodesulfurization catalyst, and asphaltene removal is carried
out using an asphaltene removal catalyst.
3. The process of claim 2 wherein the hydrodemetallization catalyst
comprises a Group VIB metal component on a porous oxide carrier,
the catalyst having a surface area of 50-200 m.sup.2/g and an
average pore diameter of 10-35 nm, wherein the hydrodesulfurization
catalyst comprises a Group VIB metal component and a Group VIII
metal component on a porous oxide carrier, the catalyst having a
surface area of 50-400 m.sup.2/g and an average pore diameter of
5-20 nm, and wherein the asphaltene removal catalyst comprises a
Group VIB metal component on a porous oxide carrier, the catalyst
having a surface area of 50-200 m.sup.2/g and an average pore
diameter of 10-35 nm.
4. The process of claim 2 wherein the hydrodesulfurization catalyst
has a Group VIB metal content which is at least 2 wt. % higher than
the Group VIB metal content of the hydrodemetallization catalyst
and the Group VIB metal content of the asphaltene removal
catalyst.
5. The process of claim 2 wherein the hydrodesulfurization catalyst
has an average pore diameter which is at least 1 nm below the
average pore diameter of the hydrodemetallization catalyst and the
average pore diameter of the asphaltene removal catalyst.
6. The process of claim 1 wherein a third hydroprocessing reactor
is applied downstream of the second hydroprocessing reactor, in
which third hydroprocessing reactor at least part of the effluent
of the second hydroprocessing reactor is subjected sequentially to
a hydrodesulfurization step and an asphaltene removal step, which
latter is carried out at a temperature higher than that of said
hydrodesulfurization step.
7. The process of claim 1 wherein a third hydroprocessing reactor
is applied downstream of the second hydroprocessing reactor, in
which third hydroprocessing reactor at least part of the effluent
of the second hydroprocessing reactor is subjected to a
hydrodesulfurization step.
8. The process of claim 1 wherein a further hydroprocessing reactor
is applied between the first hydroprocessing reactor and the second
hydroprocessing reactor, in which further hydroprocessing reactor
at least part of the effluent of the first hydroprocessing reactor
is subjected to a hydrodesulfurization step, with at least part of
the effluent of the further hydroprocessing reactor being led to
the second hydroprocessing reactor.
9. The process of claim 1 wherein the feed is a heavy hydrocarbon
feed of which at least 40 wt. % boils above 538.degree. C. and
which comprises at least 2 wt. % of sulfur and at least 5 wt. % of
Conradson Carbon.
10. The process of claim 1 wherein in the first reactor the
temperature in the hydrodesulfurization zone is at least 2.degree.
C. above the temperature in the preceding hydrodemetallization
zone, while in the first, second, and optional further reactors the
temperature in the asphaltene removal zones is at least 2.degree.
C. above the temperature in the respective preceding
hydrodesulfurization zones.
Description
CROSS REFERENCE TO RELATED APPLICATIONS
[0001] This application is based on PCT Application No.
PCT/EP03/06033, filed on Jun. 5, 2003, and claims priority from
Japanese Patent Application No. 2002-169868, filed Jun. 11, 2002,
the content of both of said applications incorporated herein in
their entirety.
BACKGROUND OF THE INVENTION
[0002] 1. Field of the Invention
[0003] The present invention relates to a process for
hydroprocessing a heavy hydrocarbon feed using hydroprocessing
catalysts. More in particular, it relates to a hydroprocessing
process for efficiently decreasing the sulfur, metals, nitrogen,
and Conradson carbon residue (CCR) in a heavy hydrocarbon feed.
[0004] 2. Prior Art
[0005] The heavy hydrocarbon oils such as atmospheric residues and
vacuum residues that are produced when crude oils are refined
contain large amounts of various impurities such as sulfur, metals,
nitrogen, and Conradson carbon residue. In recent years there has
been increasing use of these heavy hydrocarbon oils as feedstocks
in hydroprocessing operations to remove impurities such as sulfur,
so as to obtain a product suitable for use as fuel oils and also
for conversion into more economically valuable light oils.
[0006] The impurities in the feedstock to be removed in the
hydroprocessing operation include sulfur, Conradson carbon residue
(CCR), various metals, more in particular nickel and vanadium,
nitrogen and asphaltenes. The exact amounts of the various
impurities to be removed differ from feedstock to feedstock.
[0007] Various processes for the hydroprocessing of heavy
hydrocarbon feeds have been described in the art.
[0008] U.S. Pat. No. 4,054,508 describes a three-zone process for
the demetallization and hydrodesulfurization of hydrocarbon oils,
preferably residual fractions, which can contain substantial
amounts of asphaltenes. The first reaction zone contains a
large-pore catalyst comprising a Group VIB metal, preferably Mo, on
an alumina support, and having at least 60% of its pore volume in
pores with a diameter of 100-200 .ANG., at least about 5% of its
pore volume in pores with a diameter greater than 500 .ANG., and
preferably a surface area of up to 110 m.sup.2/g. The second
reaction zone contains a small-pore catalyst containing a
hydrogenating component and having at least 50%, preferably at
least 60%, of its pore volume in pores with a diameter of 30-100
.ANG., and a surface area of at least 150 m.sup.2/g. The third
reaction zone contains the same catalyst as the first reaction zone
but in a smaller amount. The large-pore catalyst has higher
demetallization activity and lower hydrodesulfurization activity
compared to the small-pore catalyst.
[0009] U.S. Pat. No. 4,431,525 discloses a hydrotreating process of
heavy hydrocarbon feedstocks containing asphaltenes, metals,
nitrogen compounds, and sulfur compounds over a sequence of three
catalysts. The first catalyst is a demetallization catalyst and has
an average pore diameter of 12.5-35 nm, the second catalyst is a
hydrodesulfurization catalyst and has an average pore diameter of
8-13 nm. The third catalyst, for which no purpose is given,
contains Mo, Cr, and Co, and has a pore volume of 0.4-0.8 ml/g, a
surface area of 150-300 m.sup.2/g, and an average pore diameter of
10-20 nm.
[0010] U.S. Pat. No. 5,744,025 discloses a process for
hydrotreating a metal-containing heavy hydrocarbon feedstock over a
sequence of three catalysts. The first catalyst is a
demetallization catalyst and has at least 40% of its pore volume in
pores with a diameter of 17-25 nm and a surface area of 100-160
m.sup.2/g. The second catalyst, which can have demetallization
and/or hydrodesulfurization and/or hydrodenitrogenation and/or
Conradson carbon residue removal activity, has at least 40% of its
pore volume in pores with a diameter of 3-17 nm and a surface area
of 160-350 m.sup.2/g. The third catalyst is stated to be a
hydrodemetallization catalyst and comprises a Group VIB and a Group
VIII hydrogenation metal component, preferably on alumina, and has
at least 40%, preferably at least 60%, of its pore volume in pores
with a diameter of 17-25 nm and a surface area of 100-160
m.sup.2/g.
[0011] Japanese laid-open patent application No. 2001-003066
discloses a process for hydrotreating heavy gas oils over a
sequence of three catalysts with different pore size distributions,
followed by a hydrodesulfurization reactor.
[0012] Japanese laid-open patent application No. 2000-351978
describes a process for the hydroprocessing of heavy hydrocarbon
feeds wherein the feed is subjected to the sequential steps of
hydrodemetallization, hydrodesulfurization, and the combination of
isomerization and hydrodesulfurization.
SUMMARY OF THE INVENTION
[0013] In one embodiment, the present invention is a process for
hydroprocessing a hydrocarbon feed containing Conradson carbon
using at least two reactors in which the hydrocarbon feed is
subjected sequentially to the steps of [0014] hydroprocessing in a
first hydroprocessing reactor, in which it is subjected
sequentially to a hydrodemetallisation step, a hydrodesulfurisation
step carried out at a temperature higher than that of said
hydrodemetallisation step, and an asphaltene removal step carried
out at a temperature higher than that of said hydrodesulfurisation
step, [0015] hydroprocessing in a second hydroprocessing reactor,
in which it is subjected sequentially to a hydrodesulfurisation
step and an asphaltene removal step, which latter is carried out at
a temperature higher than that of said hydrodesulfurisation
step.
[0016] Other embodiments of the invention relate to details
concerning catalyst composition and physical properties, process
flow details and feed composition, all of which are described
hereinbelow.
DETAILED DESCRIPTION OF THE INVENTION
[0017] A problem encountered in heavy feed hydroprocessing
operations in which the removal of Conradson carbon residue is an
important goal is the formation of sludge. The mechanism of sludge
formation is believed to be as follows. Conradson carbon residue is
considered to contain asphaltenes as its main component.
Asphaltenes are complicated high-molecular materials made up of
nuclei of condensed polycyclic aromatic hydrocarbons with saturated
hydrocarbon chains and naphthenic rings bound to them. They also
contain sulfur, nitrogen, and oxygen. They are insoluble in
n-hexane and soluble in carbon disulfide and benzene. In the
feedstock, asphaltenes are present in a dispersed form due to the
presence of solvent components. However, if during the
hydroprocessing operation the asphaltenes and the solvent
components are hydrogenated to a too great extent, the asphaltenes
cohere to form sludge (insoluble granular materials) and
sediment.
[0018] If sludge is formed during a hydroprocessing operation, it
settles and is deposited in the various parts of the refinery such
as heat exchangers, reactors, etc. In so doing it can seriously
disturb the hydroprocessing operation, even to the extent that the
operation may have to be stopped prematurely.
[0019] Furthermore, the presence of sludge will lead to catalyst
deactivation due to coke formation. As a result thereof, the
operating temperature of the process needs to be increased to
obtain the same product properties. This leads not only to a
shortened cycle length, but also to a waste of energy.
[0020] Although the processes described in the above references
work well in effecting hydrodemetallization and
hydrodesulfurization, there is still room for improvement. More in
particular, there is need for a hydroprocessing process that can
achieve efficient contaminant removal (hydrodemetallization,
hydrodesulfurization, hydrodenitrogenation, Conrandson carbon
residue removal), wherein the removal of Conradson carbon residue
and other contaminants is not accompanied by excessive sludge
formation. The present invention solves this problem by providing
the process described above.
[0021] The crux of the process according to the invention is that
the feed is treated in two hydroprocessing reactors, with an
asphaltene removal catalyst being present at the high-temperature
end of both of the reactors. It is this particular feature of the
present invention, which in combination with the specific reaction
sequence, ensures that sludge formation is kept limited while
efficient contaminant removal is obtained. This combination of
features makes for an efficient and highly stable process.
[0022] The entire effluent can be led from the first reactor to the
second reactor, but it is also possible to lead only the liquid
fraction from the first reactor to the second reactor, or to lead
only part of the liquid fraction from the first reactor to the
second. It is preferred to lead the entire effluent or the entire
liquid effluent to the further reactor. In the latter case, a
gas/liquid stripper will be present between the reactors.
[0023] The option wherein the entire liquid effluent is led to the
further reactor is considered most preferred at this point in time.
The above also holds for streams from the second to any further
reactor(s). It is possible to apply intermediate reactors between
the two reactors having an asphaltene removal catalyst at the
high-temperature end of both of the reactors. In that case, the
above also holds for streams to any intermediate reactor.
[0024] The feed can be led through the reactors from top to bottom
or from bottom to top. Operation from top to bottom is most common.
The catalyst sequences in the present application are given in the
direction of the feed.
[0025] The reactors used in the process according to the invention
are the ones conventionally used in hydroprocessing and oil
refining. These reactors are generally adiabatic. It is preferred
that the reactor is provided with a means for supplying hydrogen
through a quench line, etc. for controlling the inside reaction
temperature. Stripping equipment for removing hydrogen sulfide,
ammonia, etc. can be installed between the reactors.
[0026] The different steps of the process according to the
invention are carried out at different temperatures, with the
temperature of the hydrodesulfurization step in the first reactor
being higher than that of the hydrodemetallization step, while the
temperature of the asphaltene removal step in its turn is higher
than the temperature in the hydrodesulfurization step. In the
second reactor, the temperature of the asphaltene removal step is
higher again than the temperature in the hydrodesulfurization step
in the second reactor.
[0027] The hydrodemetallization step in the first reactor is
preferably carried out at a temperature of 300-400.degree. C., more
preferably 350-380.degree. C.
[0028] The subsequent hydrodesulfurization step is carried out at a
temperature above that of the hydrodemetallization step. The
hydrodesulfurization temperature preferably is 320 to 420.degree.
C., more preferably 360-410.degree. C.
[0029] The asphaltene removal step in the first reactor is carried
out at a temperature which is higher than that of the
hydrodesulfurization step. It is preferably carried out at a
temperature of 350-450.degree. C., more preferably 370-420.degree.
C.
[0030] The preferred temperature ranges for the
hydrodesulfurization and the asphaltene removal step carried out in
the second reactor and optional further reactors are the same as
those given above for the corresponding processes in the first
reactor.
[0031] The reaction temperatures given above for the various
reaction zones are the average feed temperatures over the entire
reaction zone filled with a certain type of catalyst. The
temperature in the hydrodesulfurization zone generally is at least
2.degree. C. above the reaction temperature in the preceding
demetallizationhydrodemetallization zone, preferably at least
4.degree. C. The reaction temperature in the asphaltene removal
zone generally is at least 2.degree. C. above the reaction
temperature in the preceding hydrodesulfurization zone, preferably
at least 4.degree. C. This applies to all reactors which contain
the specified reaction zones in the specified order.
[0032] Other than the reaction temperature, the hydroprocessing
conditions are not especially limited. Conventional hydroprocessing
conditions may be applied. These include a hydrogen partial
pressure of generally 2 to 22 MPa, preferably 10-20 MPa, a hydrogen
to feedstock ratio of generally 300-1500 Nl/l, preferably 600-1000
Nl/l, and a liquid hourly space velocity (LHSV) of generally 0.1-10
h.sup.-1, preferably 0.2-2.0 h.sup.-1.
[0033] FIG. 1 is a process chart of the process according to the
invention. As is illustrated in this figure, a hydrocarbon feed is
first fed to a reactor 1, and passes through a low-temperature zone
11 containing a hydrodemetallization catalyst, an
intermediate-temperature zone 12 containing a hydrodesulfurization
catalyst, and a high-temperature zone 13 containing an asphaltene
removal catalyst. The effluent from the first reactor is led to a
reactor 2, where it is passed through a lower-temperature zone 21
containing a hydrodesulfurization catalyst and a higher-temperature
zone 22 comprising an asphaltene removal catalyst, resulting in the
formation of a product that is discharged from the bottom of the
reactor 2.
[0034] FIG. 2 shows a variation on the process of the invention
wherein a third reactor is present downstream of the reactors 1 and
2 described above for FIG. 1. In this embodiment the effluent from
reactor 2 is led to a third reactor 3, where it is passed through a
lower-temperature zone 31 containing a hydrodesulfurization
catalyst and a higher-temperature zone 32 comprising an asphaltene
removal catalyst, resulting in the formation of a product that is
discharged from the bottom of the reactor 3.
[0035] FIG. 3 shows a variation on the process according to the
invention, wherein a reactor 2 containing only a
hydrodesulfurization catalyst is applied between a reactor 1 which
comprises, from inlet to outlet, a low-temperature zone 11
containing a hydrodemetallization catalyst, an
intermediate-temperature zone 12 containing a hydrodesulfurization
catalyst, and a high-temperature zone 13 containing an asphaltene
removal catalyst, and a reactor 3, which contains, from inlet to
outlet, a lower-temperature zone 31 containing a
hydrodesulfurization catalyst and a higher-temperature zone 32
comprising an asphaltene removal catalyst, resulting in the
formation of a product that is discharged from the bottom of the
reactor 3.
[0036] FIG. 4 shows a further variation of the process according to
the invention, wherein a third reactor is present downstream of the
reactors 1 and 2 described above for FIG. 1. In this embodiment the
effluent from reactor 2 is led to a third reactor 3, where it is
passed through a hydrodesulfurization catalyst, resulting in the
formation of a product that is discharged from the bottom of the
reactor 3.
[0037] In all these embodiments it is preferred that the entire
effluent, or the entire liquid effluent if a stripper is present
between the reactors, is led to the further reactor.
[0038] It is preferred for the volume ratio in the first reactor of
the hydrodemetallization catalyst layer, the hydrodesulfurization
catalyst layer, and the asphaltene removal catalyst layer to be
3-50:30-95:2-30, more preferably 5-45:35-95:2-25.
[0039] It is preferred for the volume ratio in the second reactor
between the hydrodesulfurization catalyst layer and the asphaltene
removal catalyst layer to be 60-97:3-40, more preferably
65-97:3-35.
[0040] If so desired, a limited amount of scale removal catalyst
may be present in the top of the second and optionally further
reactors. Scale removal catalysts are known to the skilled person.
If present, the scale removal catalyst will generally be applied in
an amount of 1-3 wt. %.
[0041] As indicated above, it is possible to apply a third or even
further reactors downstream of the second reactor, the third or
further reactors containing, from inlet to outlet, a
hydrodesulfurization catalyst and an asphaltene removal catalyst.
The preferred volume range between the hydrodesulfurization
catalyst and the asphaltene removal catalyst in the third and
further reactors is within the same ranges as indicated above for
the second reactor.
[0042] The process of the present invention is particularly
suitable for the hydroprocessing of heavy hydrocarbon feeds. It is
particularly suitable for hydroprocessing heavy feedstocks of which
at least 40 wt. %, preferably at least 60 wt. %, boils above
538.degree. C. (1000.degree. F.) and which comprise at least 0.1
wt. %, preferably at least 1 wt. %, of sulfur and at least 5 wt. %
of Conradson carbon. The sulfur content of the feedstock may be
above 2 wt. %. Its Conradson carbon content may be above 8 wt. %,
preferably above 10 wt. %. The feedstock will contain contaminant
metals, such as nickel and vanadium. Typically, these metals are
present in an amount of at least 15 wtppm, calculated on the total
of Ni and V, more particularly in an amount of at least 30
wtppm.
[0043] Suitable feedstocks include atmospheric residue, vacuum
residue, residues blended with gas oils, vacuum gas oils, coker gas
oils, crudes, shale oils, tar sand oils, solvent deasphalted oil,
coal liquefied oil, etc. Typically they are atmospheric residue
(AR), vacuum residue (VR), and mixtures thereof.
[0044] The hydrodemetallization catalyst used in the process
according to the invention generally comprises a Group VIB metal
component, preferably molybdenum, on a porous oxidic carrier. The
catalyst may or may not also contain a Group VIII metal component,
preferably nickel and/or cobalt, for reasons of economics and
performance more preferably nickel. The Group VIB metal component
is generally present in an amount of 1.5-20 wt. %, calculated as
trioxide, more preferably 4-16 wt. %. If a Group VIII metal
component is present in the hydrodemetallization catalyst, it is
generally present in an amount of 0.3-6 wt. %, preferably 1-5 wt.
%.
[0045] The hydrodemetallization catalyst is based on a porous
oxidic carrier. Preferably, the carrier comprises at least 50 wt. %
of alumina, the balance being made up of oxides of one or more of
silicon, titanium, or zirconium. Preferably, the carrier contains
at least 75 wt. % of alumina, more preferably at least 95 wt.
%.
[0046] The hydrodemetallization catalyst may contain limited
amounts of other ingredients known in the art as suitable for use
in hydrodemetallization catalysts, including phosphorus, boron,
alkaline metal components, and alkaline earth metal components.
[0047] The demetallizationhydrodemetallization catalyst has a total
pore volume which is generally between 0.2 and 1.4 ml/g, preferably
0.4-1.2 ml/g, more preferably 0.5-0.9 ml/g.
[0048] The hydrodemetallization catalyst has a surface area of
50-250 m.sup.2/g, preferably 80-200 m.sup.2/g, more preferably
100-180 m.sup.2/g. If the surface area of the catalyst is too low,
the catalytic activity will be insufficient. If the surface area of
the catalyst is too high, the average pore diameter of the catalyst
may be too low. The average pore diameter of the catalyst is
defined in the context of the present specification as the pore
diameter at which half of the pore volume is present in pores with
a diameter above this value while the other half of the pore volume
is present in pores with a diameter below this value. The
hydrodemetallization catalyst used in the process according to the
invention generally has an average pore diameter of 10-35 nm,
preferably 15-30 nm, more preferably 20-30 nm, still more
preferably 24-30 nm.
[0049] It is preferred for the hydrodemetallization catalyst used
in the process according to the invention to have less than 15%,
more preferably less than 10%, of pore volume present in pores with
a diameter below 100 .ANG..
[0050] Hydrodemetallization catalysts suitable for use in the
present invention are known in the art. They are, e.g.,
commercially available from Nippon Ketjen Co. Ltd under the
designations KFR 10, KFR 20, KFR 22, and KG 5.
[0051] The hydrodesulfurization catalyst to be used in the process
according to the invention generally comprises a Group VIB metal
component, preferably molybdenum, and a Group VIII metal component,
preferably nickel and/or cobalt, more preferably nickel, on a
porous oxidic carrier. The Group VIB metal component is generally
present in an amount of 9-30 wt. %, calculated as trioxide, more
preferably 10-27 wt. %. The Group VIII metal component is generally
present in an amount of 2-12 wt. %, preferably 2-8 wt. %.
[0052] The hydrodesulfurization catalyst is based on a porous
oxidic carrier. Preferably, the carrier comprises at least 50 wt. %
of alumina, the balance being made up of oxides of one or more of
silicon, titanium, or zirconium. Preferably, the carrier contains
at least 75 wt. % of alumina, more preferably at least 95 wt.
%.
[0053] The hydrodesulfurization catalyst may contain limited
amounts of other ingredients known in the art as suitable for use
in hydrodesulfurization catalysts, including phosphorus, boron,
alkaline metal components, and alkaline earth metal components.
[0054] The hydrodesulfurization catalyst has a total pore volume
which is generally between 0.2 and 1.4 ml/g, preferably 0.4-1.2
ml/g, more preferably 0.5-0.9 ml/g. The hydrodesulfurization
catalyst generally has a surface area of 50-400 m.sup.2/g,
preferably 100-300 m.sup.2/g, more preferably 200-300 m.sup.2/g. If
the surface area of the catalyst is too low, the catalytic activity
will be insufficient. If the surface area of the catalyst is too
high, the average pore diameter of the catalyst may be too low. The
hydrodesulfurization catalyst used in the process according to the
invention generally has an average pore diameter of 5-20 nm,
preferably 7-15 nm, more preferably 7-12 nm.
[0055] Hydrodesulfurization catalysts suitable for use in the
present invention are known in the art. They are, e.g.,
commercially available from Nippon Ketjen Co. Ltd under the
designations KFR 70, KFR 72, and KFR 70B.
[0056] The hydrodesulfurization catalyst used in the first reactor,
the hydrodesulfurization catalyst used in the second reactor and
any hydrodesulfurization catalyst(s) used in any further reactor
may be the same or different.
[0057] The asphaltene removal catalyst used in the process
according to the invention generally comprises a Group VIB metal
component, preferably molybdenum, and a Group VIII metal component,
preferably nickel and/or cobalt, more preferably nickel, on a
porous oxidic carrier. The Group VIB metal component is generally
present in an amount of 2-20 wt. %, calculated as trioxide, more
preferably 4-16 wt. %. If the Group VIB metal content is too low,
the activity of the catalyst will be impaired. If the Group VIB
metal content is too high, the activity will not be improved
further.
[0058] The Group VIII metal component is generally present in an
amount of 0.5-6 wt. %, preferably 1-5 wt. %. If the Group VIII
metal content is too low, the activity of the catalyst will be
impaired. If the Group VIII metal content is too high, the activity
will not be improved further.
[0059] The asphaltene removal catalyst is based on a porous oxidic
carrier. Preferably, the carrier comprises at least 50 wt. % of
alumina, the balance being made up of oxides of one or more of
silicon, titanium, or zirconium. Preferably, the carrier contains
at least 75 wt. % of alumina, more preferably at least 95 wt. %.
The asphaltene removal catalyst may contain limited amounts of
other ingredients known in the art, including phosphorus, boron,
alkaline metal components, and alkaline earth metal components.
[0060] The asphaltene removal catalyst has a total pore volume
which generally is at least 0.4 ml/g, preferably at least 0.55
ml/g. The pore volume generally is at most 1.4 ml/g, preferably at
most 1.2 ml/g, more preferably at most 0.9 ml/g. The asphaltene
removal catalyst generally has a surface area of at least 50
m.sup.2/g and at most 200 m.sup.2/g, preferably 100-180
m.sup.2/g.
[0061] If the surface area of the catalyst is too low, the
catalytic activity will be insufficient. If the surface area of the
catalyst is too high, the average pore diameter of the catalyst may
be too low. The asphaltene removal catalyst used in the process
according to the invention generally has an average pore diameter
of 10-35 nm, preferably 15-30 nm, more preferably 18-28 nm.
[0062] If the average pore diameter is too low, the decomposition
of asphaltenes will be insufficient and excessive hydrogenation may
take place to increase sludge formation. If the average pore
diameter is too high, the asphaltene decomposition rate may also
decrease, with inherent sludge formation.
[0063] It is preferred for the asphaltene removal catalyst used in
the process according to the invention to have less than 15%, more
preferably less than 10%, of pore volume present in pores with a
diameter below 100 .ANG..
[0064] The asphaltene removal catalyst generally has 0.3 ml/g or
less of pore volume in pores with a diameter of 100 nm or larger,
preferably 0.2 ml/g or less. If the amount of pore volume present
in this range is too high, the balance between the decomposition of
the asphaltenes and the resins which act as their solvent is lost,
and sludge is liable to be formed. Moreover, if the amount of pore
volume present in this range is too high, the mechanical strength
of the catalyst will decrease.
[0065] Asphaltene removal catalysts suitable for use in the present
invention are known in the art. They are, e.g., commercially
available from Nippon Ketjen Co. Ltd under the designations KFR 10,
KFR 20, KFR 22, and KG 5.
[0066] The asphaltene removal catalyst used in the first reactor,
the asphaltene removal catalyst used in the second reactor, and any
asphaltene removal catalyst(s) used in any further reactor may be
the same or different.
[0067] As can be seen from the above description, the asphaltene
removal catalyst and the hydrodemetallization removal catalyst used
in the process according to the invention are quite similar. This
means that it is within the scope of the present invention for the
demetallizationhydrodemetallization catalyst and the asphaltene
removal catalyst to be used in the process according to the
invention to be the same.
[0068] The hydrodesulfurization catalyst used in the process
according to the invention is always different from the
hydrodemetallization catalyst and the asphaltene removal catalyst.
More in particular, the hydrodesulfurization catalyst has a Group
VIB metal content which is at least 2 wt. %, preferably at least 3
wt. %, higher than the Group VIB metal content of the
hydrodemetallization catalyst and the Group VIB metal content of
the asphaltene removal catalyst, calculated as trioxide.
[0069] Additionally, the average pore diameter of the
hydrodesulfurization catalyst is at least 1 nm, preferably at least
2 nm, more preferably at least 3 nm, below the average pore
diameter of the demetallizationhydrodemetallization catalyst and
the average pore diameter of the asphaltene removal catalyst.
[0070] The catalyst particles may have the shapes and dimensions
common to the art. Thus, the particles may be spherical,
cylindrical, or polylobal and their diameter may range from 0.5 to
10 mm. Particles with a diameter of 0.5-3 mm, preferably 0.7-2 mm,
for example 1.2-1.5 mm, and a length of 1.5-10 mm, for example
2.5-4.5 mm, are preferred. Because they lead to a reduced pressure
drop in demetallizationhydrodemetallization operations, polylobal
particles may be preferred.
[0071] If so desired, a small amount of the catalyst used in the
top layer of one or more of the reactors may have a different
particle shape and size for the purpose of regularising the flow of
the feedstock passing through the reactor. As is conventional in
the art, the various types of catalysts can be present in a single
bed in the unit, or in a number of beds one above the other
containing the same type of catalyst.
[0072] The way in which the catalysts used in the process according
to the invention are prepared is not critical. By way of example, a
suitable preparation process is described below.
[0073] A typical production process for preparing a catalyst
carrier comprising alumina is coprecipitation of sodium aluminate
and aluminum sulfate. The resulting gel is dried, extruded, and
calcined, to obtain an alumina-containing carrier. Optionally,
other components such as silica may be added before, during, or
after precipitation. By way of example, a process for preparing an
alumina gel will be described below.
[0074] First, a tank containing tap water or warm water is charged
with an alkali solution of sodium aluminate, aluminum hydroxide or
sodium hydroxide, etc., and an acidic aluminum solution of aluminum
sulfate or aluminum nitrate, etc. is added for mixing. The hydrogen
ion concentration (pH) of the mixed solution changes with the
progress of the reaction. It is preferable that when the addition
of the acidic aluminum solution is complete, the pH is 7 to 9, and
that during mixing the temperature is 60 to 75.degree. C. The
mixture is then kept at that temperature for in general 0.5-1.5
hours, preferably for 40-80 minutes.
[0075] In a following stage, the gel is separated from the
solution, and any commercially used washing treatment, for example
a washing treatment using tap water or hot water, is carried out to
remove impurities, mainly salts, from the gel. Then, the gel is
shaped into particles in a manner known in the art, e.g., by way of
extrusion, beading or pelletizing.
[0076] Finally, the shaped particles are dried and calcined. The
drying is generally carried out at a temperature between room
temperature and up to 200.degree. C., generally in the presence of
air. The calcining is generally carried out at a temperature of 300
to 950.degree. C., preferably 600 to 900.degree. C., generally in
the presence of air, for a period of 30 minutes to six hours. If so
desired, the calcination may be carried out in the presence of
steam to influence the crystal growth in the oxide.
[0077] By the above production process, it is possible to obtain a
carrier having properties which will give a catalyst with the
surface area, pore volume, and pore size distribution
characteristics specified above. The surface area, pore volume, and
pore size distribution characteristics can be adjusted in a manner
known to the skilled person, for example by the addition during the
mixing or shaping stage of an acid, such as nitric acid, acetic
acid or formic acid, or other compounds as moulding auxiliary, or
by regulating the water content of the gel by adding or removing
water.
[0078] The Group VIB metal components and, where appropriate, Group
VIII metal components or other components such as phosphorus, can
be incorporated into the catalyst carrier in a conventional manner,
e.g., by impregnation and/or by incorporation into the support
material before it is shaped into particles. At this point in time
it is considered preferred to first prepare the carrier and
incorporate the catalytic materials into the carrier after it has
been dried and calcined.
[0079] The metal components can be incorporated into the catalyst
composition in the form of suitable precursors, preferably by
impregnating the catalyst with an acidic or basic impregnation
solution comprising suitable metal precursors. For the Group VIB
metals, ammonium heptamolybdate, ammonium dimolybdate, and ammonium
tungstenate may be mentioned as suitable precursors. Other
compounds, such as oxides, hydroxides, carbonates, nitrates,
chlorides, and organic acid salts, may also be used.
[0080] For the Group VIII metals, suitable precursors include
oxides, hydroxides, carbonates, nitrates, chlorides, and organic
acid salts. Carbonates and nitrates are particularly suitable. The
impregnation solution, if applied, may contain other compounds the
use of which is known in the art, such as organic acids, e.g.,
citric acid, ammonia water, hydrogen peroxide water, gluconic acid,
tartaric acid, malic acid or EDTA (ethylenediamine tetraacetic
acid).
[0081] It will be clear to the skilled person that there is a wide
range of variations on this process. Thus, it is possible to apply
a plurality of impregnating stages, the impregnating solutions to
be used containing one or more of the component precursors that are
to be deposited, or a portion thereof. Instead of impregnating
techniques, dipping processes, spraying processes, etc. can be
used. In the case of multiple impregnation, dipping, etc., drying
and/or calcining may be carried out in between.
[0082] After the metals have been incorporated into the catalyst
composition, it is optionally dried, e.g., in air flow for about
0.5 to 16 hours at a temperature between room temperature and
200.degree. C., and subsequently calcined, generally in air, for
about 1 to 6 hours, preferably 1-3 hours, at 200-800.degree. C.,
preferably 450-700.degree. C. The drying is done to physically
remove the deposited water. The calcining is done to bring at least
part, preferably all, of the metal component precursors to the
oxide form.
[0083] It may be desirable to convert the catalysts, i.e., the
Group VIB and optional Group VIII metal components present therein,
into the sulfidic form prior to their use in the hydroprocessing of
hydrocarbon feedstocks. This can be done in an otherwise
conventional manner, e.g., by contacting the catalyst in the
reactor at increasing temperature with hydrogen and a
sulfur-containing feedstock, or with a mixture of hydrogen and
hydrogen sulfide. Ex situ presulfiding is also possible.
[0084] In the present specification, the indications Group VIB
metal and a Group VIII metal refer to the Periodic Table of
Elements applied by Chemical Abstract Services (CAS system). The
specific surface area is determined by nitrogen (N.sub.2)
adsorption using the BET method.
[0085] The determination of the total pore volume and the pore size
distribution is effected via mercury penetration at a contact angle
of 140.degree. with a surface tension of 480 dynes/cm, using, for
example, a mercury porosimeter Autopore II (trade name) produced by
Micrometrics.
EXAMPLES
[0086] Three catalysts were selected for use in the present
example. The hydrodemetallization (HDM) catalyst contained 9 wt. %
of molybdenum, calculated as oxide, and 2 wt. % of nickel,
calculated as oxide, on an alumina carrier. The catalyst had an
average pore diameter of 180 .ANG..
[0087] The hydrodesulfurization (HDS) catalyst contained 12 wt. %
of molybdenum, calculated as oxide, and 3 wt. % of nickel,
calculated as oxide, on an alumina carrier. The catalyst had an
average pore diameter of 120 .ANG..
[0088] The asphaltene removal (HDAsp) catalyst contained 8.0 wt. %
of molybdenum, calculated as oxide, and 2.2 wt. % of nickel,
calculated as oxide, on an alumina carrier. The catalyst had an
average pore diameter of 260 .ANG., a surface area of 130
m.sup.2/g, a total pore volume of 0.8 ml/g, and a pore volume in
pores with a diameter of 1000 .ANG. and larger of 0.25 ml/g.
[0089] The feedstock used in the present example was a South Sea
atmospheric residue with the following properties: TABLE-US-00001
Sulphur 0.2 wt. % nitrogen 3000 wt. ppm Metals (nickel + vanadium)
70 wt. ppm Conradson Carbon residue 7 wt. % Asphaltenes.sup.1 2.5
wt. % Density (15.degree. C.) 0.93 g/ml Distillation properties
ASTM-D 5307 Initial boiling point 301.degree. C. 50 wt. %
563.degree. C. 70 wt. % 624.degree. C. .sup.1matter insoluble in
n-heptane
[0090] A small reactor system consisting of three reactors
connected in series was packed with the selected catalysts. The
packing of the reactors applied in Example 1 according to the
invention and in the Comparative Example is given below:
TABLE-US-00002 Example 1 Comparative Example 1.sup.st reactor HDM
catalyst - 20 vol. % HDM catalyst - 20 vol. % HDS catalyst - 55
vol. % HDS catalyst - 80 vol. % HDAsp catalyst - 25 vol. % 2.sup.nd
reactor HDS catalyst - 75 vol. % HDS catalyst - 100 vol. % HDAsp
catalyst - 25 vol. % 3.sup.rd reactor HDS catalyst - 95 vol. % HDS
catalyst - 100 vol. % HDAsp catalyst - 5 vol. %
[0091] After packing the reactor system with the respective
catalysts, the catalysts were sulfided in a conventional manner by
contacting them with a light gas oil (LGO) containing 2.5 wt. % of
dimethyl disulfide (DMDS), and the feedstock described above was
led to the reactor system at a pressure of 16.5 MPa, a liquid
hourly space velocity (LHSV) of 0.3 h.sup.-1, and a hydrogen to
feed ratio (H.sub.2/oil ratio) of 850 Nl/l. The reaction
temperature was adjusted to achieve a Conradson Carbon residue
content of 3.0 wt. % in the product oil, with the temperature in
the second bed being higher than the temperature in the first bed
in each reactor, while, if applicable, the temperature in the third
bed was higher than the temperature in the second bed. The test was
carried out for a period of 200 days.
[0092] In Example 1, where an asphaltene removal catalyst was
present in the high-temperature bottom region of each of the three
reactors, the process still operated well after 200 days, and could
be continued after that date. By comparison, the process of the
Comparative Example, where no asphaltene removal catalyst was
present in the high-temperature bottom region of each of the three
reactors, had to be discontinued on day 135, because the reactors
were clogged with sludge.
[0093] The following table gives the temperatures of the various
catalyst layers in Example 1 and the comparative Example.
TABLE-US-00003 Comparative Example Example 1 (day 200) (day 135)
1.sup.st reactor HDM catalyst 390.degree. C. HDM catalyst
390.degree. C. HDS catalyst 398.degree. C. HDS catalyst 400.degree.
C. HDAsp catalyst 405.degree. C. 2.sup.nd reactor HDS catalyst
413.degree. C. HDS catalyst 413.degree. C. HDAsp catalyst
422.degree. C. 3.sup.rd reactor HDS catalyst 421.degree. C. HDS
catalyst 422.degree. C. HDAsp catalyst 423.degree. C.
[0094] The processes of Example 1 and the Comparative Example were
operated under such a regimen that the amount of Conradson carbon
residue in the product was kept at 3.0 wt. %. This means that, as
is conventional in the art of hydroprocessing, the reaction
temperature is slowly increased to compensate for catalyst
deactivation to keep the amount of Conradson Carbon residue in the
product at the desired value.
[0095] FIG. 4 shows the change of the average reaction temperature
during the process at a constant Conradson Carbon content in the
product for the process according to the invention and for the
Comparative Example. It can be seen that for the process according
to the invention of Example 1, the temperature increase stabilised
to a value of less than 0.8.degree. C. per month. In contrast, for
the Comparative Example, by the end of the run a temperature
increase of 2.degree. C. per month was obtained.
[0096] Apparently, even if the Conradson Carbon Content of the
product is the same, the presence of an asphaltene removal catalyst
in the high-temperature bottom end of the hydroprocessing reactors
ensures that less sludge is formed, which leads to less clogging of
the unit and a longer and more stable operation. The presence of
the asphaltene removal catalyst also appears to result in a more
stable process, as can be seen from a lower increase in reaction
temperature being necessary to keep the Conradson carbon content of
the product at a value of 3.0 wt. %.
[0097] Although not wishing to be bound by theory, the inventors
believe that the presence of an asphaltene removal catalyst in the
high-temperature lower end of the hydroprocessing reactors ensures
that the decomposition of the asphaltenes and the resins which act
as their solvents is well balanced, resulting in low sludge
formation. On the other hand, it is believed that in the
comparative example, where no asphaltene removal catalyst was
present, the resins that act as solvent for the asphaltenes were
hydrogenated. As a result, the asphaltenes became insoluble in the
hydrocarbon feed, causing the formation of sludge. In turn, the
presence of sludge caused deactivation of the catalyst due to coke
being formed thereon.
* * * * *