U.S. patent application number 13/262364 was filed with the patent office on 2012-04-19 for process for producing middle distillates by hydroisomerization and hydrocracking of a heavy fraction derived from a fischer-tropsch effluent.
This patent application is currently assigned to IFP Energies nouvelles. Invention is credited to Vincent Coupard, Aurelie Dandeu.
Application Number | 20120091034 13/262364 |
Document ID | / |
Family ID | 41254657 |
Filed Date | 2012-04-19 |
United States Patent
Application |
20120091034 |
Kind Code |
A1 |
Dandeu; Aurelie ; et
al. |
April 19, 2012 |
PROCESS FOR PRODUCING MIDDLE DISTILLATES BY HYDROISOMERIZATION AND
HYDROCRACKING OF A HEAVY FRACTION DERIVED FROM A FISCHER-TROPSCH
EFFLUENT
Abstract
The present invention describes a process for producing middle
distillates from a C5+ liquid paraffinic fraction, termed a heavy
fraction, with an initial boiling point in the range 15.degree. C.
to 40.degree. C. produced by Fischer-Tropsch synthesis, comprising
the following steps in succession: passing said C5+ liquid
paraffinic fraction, termed a heavy fraction, over at least one ion
exchange resin at a temperature in the range 80.degree. C. to
150.degree. C., at a total pressure in the range 0.7 to 2.5 MPa, at
an hourly space velocity in the range 0.2 to 2.5 h.sup.-1;
eliminating at least a portion of the water formed in step a);
hydrogenating the unsaturated olefinic type compounds of at least a
portion of the effluent derived from step b) in the presence of
hydrogen and a hydrogenation catalyst; and
hydroisomerization/hydrocracking of at least a portion of the
hydrotreated effluent derived from step c) in the presence of
hydrogen and a hydroisomerization/hydrocracking catalyst.
Inventors: |
Dandeu; Aurelie; (Saint-Just
Chaleyssin, FR) ; Coupard; Vincent; (Villeurbanne,
FR) |
Assignee: |
IFP Energies nouvelles
Rueil-Malmaison Cedex
FR
ENI S.p.A
Rome
IT
|
Family ID: |
41254657 |
Appl. No.: |
13/262364 |
Filed: |
March 24, 2010 |
PCT Filed: |
March 24, 2010 |
PCT NO: |
PCT/FR2010/000251 |
371 Date: |
December 12, 2011 |
Current U.S.
Class: |
208/57 |
Current CPC
Class: |
B01J 23/42 20130101;
B01J 21/12 20130101; B01J 35/1042 20130101; B01J 23/44 20130101;
C10G 45/34 20130101; C10G 2300/4012 20130101; C10G 47/00 20130101;
B01J 39/04 20130101; B01J 35/1019 20130101; C10G 65/12 20130101;
C10G 2300/301 20130101; C10G 2300/205 20130101; C10G 2/30 20130101;
C10G 2300/4006 20130101; C10G 7/00 20130101; Y02P 30/20 20151101;
C10G 45/58 20130101; B01J 23/88 20130101; C10G 2300/4018 20130101;
B01J 35/10 20130101; C10G 25/02 20130101; C10G 2300/1022 20130101;
C10G 47/16 20130101; C10G 2300/1025 20130101; B01J 35/108 20130101;
B01J 31/08 20130101; C10G 2300/1011 20130101; B01J 35/1061
20130101; C10G 45/60 20130101; C10G 67/06 20130101; C10G 45/32
20130101 |
Class at
Publication: |
208/57 |
International
Class: |
C10G 69/02 20060101
C10G069/02 |
Foreign Application Data
Date |
Code |
Application Number |
Apr 3, 2009 |
FR |
09/01653 |
Claims
1. A process for producing middle distillates from a C5+ liquid
paraffinic fraction, termed a heavy fraction, with an initial
boiling point in the range 15.degree. C. to 40.degree. C., produced
by Fischer-Tropsch synthesis, comprising the following steps in
succession: a) passing said C5+ liquid paraffinic fraction, termed
a heavy fraction, over at least one ion exchange resin to allow
esterification of alcohols and carboxylic acids into esters and/or
to retain metals dissolved in the feed, at a temperature in the
range 80.degree. C. to 150.degree. C., at a total pressure in the
range 0.7 to 2.5 MPa, at an hourly space velocity in the range 0.2
to 2.5 h.sup.-1; b) eliminating at least a portion of the water
formed in step a); c) hydrogenating the unsaturated olefinic type
compounds of at least a portion of the effluent derived from step
b) in the presence of hydrogen and a hydrogenation catalyst; d)
hydroisomerization/hydrocracking of at least a portion of the
hydrotreated effluent derived from step c) in the presence of
hydrogen and a hydroisomerization/hydrocracking catalyst; e)
separating and recycling unreacted hydrogen and light gases to the
hydroisomerization/hydrocracking step d); f) distilling the
effluent derived from step e).
2. A process according to claim 1, in which said C5+ liquid
fraction undergoes a step for decontamination by passage over a
guard bed containing at least one guard bed catalyst, before
passage over an ion exchange resin in accordance with step a).
3. A process according to claim 2, in which said guard bed catalyst
comprises a macroporous mercury volume for a mean diameter at 50 nm
of more than 0.1 cm.sup.3/g, and a total volume of more than 0.60
cm.sup.3/g.
4. A process according to claim 1, in which said C5+ liquid
paraffinic fraction passes over a single ion exchange resin in
order to carry out the simultaneous esterification of alcohols and
carboxylic acids into esters and capture of metals dissolved in the
feed.
5. A process according to claim 4, in which said resin is used at a
temperature in the range 100.degree. C. to 150.degree. C., at a
pressure in the range 1 to 2 MPa and at an hourly space velocity in
the range 0.5 to 1.5 h.sup.-1.
6. A process according to claim 4, in which said resin is
constituted by copolymers of divinyl benzene and polystyrene with a
degree of cross-linking in the range 20% to 35%, and an acid
strength, assayed by potentiometry during neutralization with a KOH
solution, in the range 0.2 to 6 mmol H+ equivalent/g.
7. A process according to claim 4, in which said resin is a
polysiloxane grafted with alkylsulphonic type acid groups (of the
--CH.sub.2--CH.sub.2--CH.sub.2--SO.sub.3H type), with a size in the
range 0.5 to 1 2 mm and with an acid strength, assayed by
potentiometry during neutralization with a KOH solution, of 0.4 to
1.5 mmol H+ equivalent/g.
8. A process according to claim 1, in which said C5+ liquid
paraffinic fraction passes over two distinct ion exchange resins
with different natures, in two different reactors.
9. A process according to claim 8, in which the reactor containing
the ion exchange resin allowing the capture of metals is used
upstream of the reactor containing the ion exchange resin allowing
the esterification of alcohols and carboxylic acids.
10. A process according to claim 8, in which said first resin is a
resin constituted by copolymers of divinyl benzene and polystyrene
with a degree of cross-linking in the range 1% to 20% and an acid
strength, assayed by potentiometry during neutralization with a KOH
solution, in the range 1 to 15 mmol H+ equivalent/g.
11. A process according to claim 8, in which said first resin is
used at a temperature in the range 80.degree. C. to 110.degree. C.,
at a pressure in the range 1 to 2 MPa and at an hourly space
velocity in the range 0.2 to 1.5 .sup.-1.
12. A process according to claim 8, in which the
hydroisomerization/hydrocracking catalyst contains at least one
hydrodehydrogenating element selected from noble metals from group
VIII, preferably platinum and/or palladium, and at least one
amorphous refractory support, preferably silica-alumina.
13. A process according to claim 1, in which the paraffinic feed
produced by Fischer-Tropsch synthesis is produced from a synthesis
gas produced from a natural gas using the gas-to-liquid, GTL,
route.
14. A process according to claim 1, in which the paraffinic feed
produced by Fischer-Tropsch synthesis is produced from a synthesis
gas produced from coal using the coal-to-liquid, CTL, route.
15. A process according to claim 1, in which the paraffinic feed
produced by Fischer-Tropsch synthesis is produced from a synthesis
gas produced from biomass using the biomass-to-liquid route.
Description
[0001] In the Fischer-Tropsch process, synthesis gas (CO+H.sub.2)
is catalytically transformed into oxygen-containing products and
into essentially straight-chain hydrocarbons in the gas, liquid or
solid form.
[0002] The paraffinic feed produced by Fischer-Tropsch synthesis
used in the process of the invention is produced from a synthesis
gas in the Fischer-Tropsch process; the synthesis gas (CO+H.sub.2)
is advantageously produced employing three routes.
[0003] In one preferred implementation, synthesis gas (CO+H.sub.2)
is produced from natural gas using the gas-to-liquid, GTL,
route.
[0004] In another preferred implementation, synthesis gas
(CO+H.sub.2) is produced from coal using the coal-to-liquid
process, CTL.
[0005] In another preferred implementation, synthesis gas
(CO+H.sub.2) is produced from biomass using the biomass-to-liquid
process, BTL.
[0006] However, such products, principally constituted by normal
paraffins, cannot be used as they are, in particular because their
cold properties are not very compatible with the normal uses of oil
cuts. As an example, the pour point of a straight-chain hydrocarbon
containing 20 carbon atoms per molecule (boiling point equal to
approximately 340.degree. C., i.e. usually included in the middle
distillates cut) is approximately 37.degree. C., which renders its
use impossible, since the specification for gas oil is -15.degree.
C. Thus, hydrocarbons derived from the Fischer-Tropsch process
comprising mainly n-paraffins have to be transformed into more
upgradeable products such as gas oil or kerosene, for example,
which are obtained, for example, after catalytic
hydrocracking/hydroisomerization reactions. In contrast, they may
have a non-negligible quantity of unsaturated compounds of the
olefinic type and oxygen-containing products (such as alcohols,
carboxylic acids, ketones, aldehydes and esters). Moreover, such
oxygen-containing and unsaturated compounds are concentrated in the
light fractions. Thus, in the C5+ fraction corresponding to
products boiling at an initial boiling point in the range
15.degree. C. to 40.degree. C., these compounds represent in the
range 10-20% by weight of unsaturated olefinic type compounds and
in the range 5-10% by weight of oxygen-containing compounds.
[0007] Such products are generally free of heteroatomic impurities
such as sulphur or nitrogen, but may contain small quantities of
Fe, Co, Zn, Ni or Mo originating from the dissolution of catalyst
fines by the carboxylic acids. Those metals may form complexes with
the oxygen-containing compounds. Said products contain no or
practically no aromatics, naphthenes and more generally cycles, in
particular in the case of cobalt catalysts.
[0008] The hydrogenation of unsaturated olefinic type compounds
present in hydrocarbons from the Fischer-Tropsch process is a
highly exothermic reaction. Thus, under the severe
hydrocracking/hydroisomerization operating conditions, the
transformation of said unsaturated compounds may have a negative
impact on the hydrocracking step, such as thermal runaway of the
reaction, substantial coking of the catalyst or the formation of
gum by oligomerization. In order to protect the hydrocracking step,
a hydrotreatment step is carried out under conditions which are
less severe than those of hydrocracking step. However, the
impurities in the feed, the oxygen-containing compounds and the
metals (Fe, Co, Zn, Ni, Mo) have a deleterious effect not only on
the activity of the hydrotreatment and hydrocracking catalysts, but
also on the stability of the hydrotreatment catalyst. In fact, in
the hydrotreatment reactor, the operating conditions as regards
temperature are such that the oxygen-containing compounds do not
decompose but are adsorbed onto the catalyst and form coke. In the
hydrocracking section, the severe operating conditions cause the
decomposition of oxygen-containing compounds into water, CO and
CO.sub.2 which are inhibitors of the acid functions (water) and the
hydrogenating function (CO, CO.sub.2) of the hydrocracking catalyst
and which thus modifies the activity and selectivity. As a
consequence, the presence of alcohol or acid type oxygen-containing
compounds present in the feeds necessitates an increase in the
temperature of the hydrotreatment and hydrocracking step in order
to compensate for the drop in activity and maintain the conversion.
Further, the carboxylic acids can extract active particles from the
hydrotreatment and hydrocracking catalysts, thus reducing the
service life of said catalysts. Similarly, the metals complexed by
the oxygen-containing compounds decompose on the active site of
said catalysts (HDT and HCK) in the presence of hydrogen and very
selectively poison the active sites of said catalysts.
[0009] One of the aims of the invention is thus to reduce the total
oxygen content of the feed and thus to limit the inhibiting effects
of the oxygen-containing compounds and thereby limit the increase
in the temperature in order to compensate for the drop in activity
and maintain the conversion on the two steps, hydrotreatment and
hydrocracking.
[0010] Thus, upstream of the hydrotreatment step and in order to
increase the service life of the hydrotreatment catalyst and the
hydrocracking catalyst, the Applicant has instigated a step
allowing transformation on an ion exchange resin, simultaneously or
otherwise, of the alcohols and carboxylic acids constituting the
oxygen-containing compounds into esters, and of capturing the
metals complexed by said oxygen-containing compounds.
[0011] This step is followed by separation of water before the
hydrotreatment step, which can reduce the total oxygen content and
thus limit the inhibiting effects of the oxygen-containing
compounds, and thereby limit the increase in temperature in order
to compensate for the drop in activity and maintain the conversion
over the two steps, hydrotreatment and hydrocracking. The water
separation can also wash and capture CO and CO.sub.2, which are
inhibitors, dissolved in the feed.
PRIOR ART
[0012] Shell's patent application (EP-0 583 836) describes a
process for producing middle distillates from a feed obtained by
the Fischer-Tropsch process. In this process, the feed from the
Fischer-Tropsch synthesis may be treated in its entirety, but
preferably the C4- fraction is removed from the feed so that only
the C5+ fraction boiling at a temperature of over 15.degree. C. is
introduced into the subsequent step. Said feed undergoes
hydrotreatment in order to hydrogenate the olefins and alcohols in
the presence of a large excess of hydrogen, so that the conversion
of products boiling above 370.degree. C. into products with a lower
boiling point is less than 20%. The hydrotreated effluent
constituted by paraffinic hydrocarbons with a high molecular weight
is preferably separated from hydrocarbon compounds with a low
molecular weight, in particular the C4- fraction, before the second
hydroconversion step. At least a portion of the remaining C5+
fraction then undergoes a hydrocracking/hydroisomerization step
with at least 40% by weight conversion of products boiling above
370.degree. C. into products with a lower boiling point.
[0013] Neither the presence of impurities in the feed nor the
presence of steps for eliminating such impurities is mentioned in
that application. Thus, Shell's patent application (EP-0 583 836)
does not deal with the problem of eliminating the impurities
present in the feed from the Fischer-Tropsch process.
[0014] SASOL's patent applications (WO-06/005084) and WO-06/005085
concern the elimination of metals complexed by the
oxygen-containing compounds present in a paraffinic feed derived
from the Fischer-Tropsch process. Those patents describe
decomposition after adding water in a zone for hydrothermal
conversion of those compounds. That decomposition is followed by a
physical treatment which can remove the metals after decomposition.
Those applications require adding water to the system, the
existence of three steps (reaction, water separation, filtration)
and do not affect the oxygen-containing compounds present in the
feed. The present invention can reduce the number of steps
required, dispense with the addition of water and simultaneously
carry out transformation of the oxygen-containing compounds present
in the feed.
[0015] More precisely, the present invention concerns a process for
producing middle distillates from a C5+ liquid paraffinic fraction,
termed a heavy fraction, with an initial boiling point in the range
15.degree. C. to 40.degree. C., produced by Fischer-Tropsch
synthesis, comprising the following steps in succession: [0016] a)
passing said C5+ liquid paraffinic fraction, termed a heavy
fraction, over at least one ion exchange resin at a temperature in
the range 80.degree. C. to 150.degree. C., at a total pressure in
the range 0.7 to 2.5 MPa, at an hourly space velocity in the range
0.2 to 2.5 h.sup.-1; [0017] b) eliminating at least a portion of
the water formed in step a); [0018] c) hydrogenating the
unsaturated olefinic type compounds of at least a portion of the
effluent derived from step b) in the presence of hydrogen and a
hydrogenation catalyst; [0019] d) hydroisomerization/hydrocracking
of at least a portion of the hydrotreated effluent derived from
step c) in the presence of hydrogen and a
hydroisomerization/hydrocracking catalyst; [0020] e) separating and
recycling unreacted hydrogen and light gases to the
hydroisomerization/hydrocracking step d); [0021] f) distilling the
effluent derived from step e).
BRIEF DESCRIPTION OF THE DRAWINGS
[0022] FIGS. 1 and 2 show preferred implementations of the process
of the invention without limiting its scope.
DETAILED DESCRIPTION OF THE INVENTION
[0023] Throughout the remainder of the description, we shall detail
the various steps in the process of the invention by referring to
FIGS. 1 and 2.
[0024] At the outlet from the Fischer-Tropsch synthesis unit, the
effluent derived from the Fischer-Tropsch synthesis unit is
advantageously divided into two fractions, a light fraction termed
the cold condensate, and a heavy fraction, termed waxes.
[0025] The two fractions defined thereby comprise water, carbon
dioxide (CO.sub.2), carbon monoxide (CO) and unreacted hydrogen
(H.sub.2). Further, the light fraction, the cold condensate,
contains light C1 to C4 hydrocarbons, termed the C4- fraction, in
the gaseous form.
[0026] In accordance with a preferred implementation which is not
shown in the figures, the light fraction, termed the cold
condensate, and the heavy fraction, termed waxes, are treated
separately then recombined in the line in order to obtain a single
C5+ fraction 1, termed the heavy fraction, with an initial boiling
point in the range 15.degree. C. to 40.degree. C., preferably with
a boiling point of 20.degree. C. or more.
[0027] The light fraction, termed the cold condensate, enters a
fractionation means which is not shown in the figures. The
fractionation means may, for example, be constituted by processes
which are well known to the skilled person such as a flash drum, a
distillation or a stripper. Advantageously, the fractionation means
is a distillation column which can eliminate light and gaseous C1
to C4 hydrocarbons, termed the gaseous fraction C4-, corresponding
to products boiling at a temperature of less than 20.degree. C.,
preferably less than 10.degree. C. and more preferably less than
0.degree. C., and recovery of a C5+ fraction, termed the heavy
fraction, with an initial boiling point in the range 15.degree. C.
to 40.degree. C., preferably with a boiling point of 20.degree. C.
or more.
[0028] The stabilized effluent derived from the fractionation means
and said heavy fraction, termed waxes, are then recombined in order
to obtain a stabilized C5+ liquid fraction 1, corresponding to
products boiling at an initial boiling point in the range
15.degree. C. to 40.degree. C., preferably with a boiling point or
20.degree. C. or more. Said C5+ fraction 1 constitutes the feed
used in the process of the invention.
[0029] Before passage over an ion exchange resin in accordance with
step a) of the process of the invention, said liquid C5+ fraction
may optionally initially undergo a step for decontamination in a
reactor 2 by passage over a guard bed containing at least one guard
bed catalyst.
[0030] The treated heavy fractions may possibly contain solid
particles such as mineral solids. They may possibly contain metals
contained in hydrocarbon structures such as organometallic
compounds of greater or lesser solubility. The term "fines" means
fines resulting from physical or chemical attrition of the
catalyst. They may be on the micron or sub-micron scale. These
mineral particles thus contain the active components of these
catalysts; a non limiting list is as follows: alumina, silica,
titanium, zirconia, cobalt oxide, iron oxide, tungsten, ruthenium
oxide, etc. These solid minerals may be in the form of a calcined
mixed oxide: examples are alumina-cobalt, alumina-iron,
alumina-silica, alumina-zirconia, alumina-titanium,
alumina-silica-cobalt, alumina-zirconia-cobalt, etc.
[0031] Said heavy fractions may also contain metals within
hydrocarbon structures, which may possibly contain oxygen or
organometallic compounds of greater or lesser solubility. More
particularly, said compounds may be silicon-based. As an example,
they may be anti-foaming agents used in the synthesis process. As
an example, the solutions of a silicone type silicon compound or
silicone oil emulsion are more particularly contained in the heavy
fraction.
[0032] Further, the catalyst fines described above may have a
silica content which is greater than the formulation for the
catalyst, resulting from intimate interaction between the catalyst
fines and the anti-foaming agents described above.
[0033] The problem which thus arises is to reduce the quantity of
solid mineral particles and possibly to reduce the quantity of
metallic compounds which are deleterious to the
hydroisomerization-hydrocracking catalyst.
Characteristics of Catalysts used in the Guard Beds
[0034] The guard beds advantageously contain at least one
catalyst.
Shape of Catalysts
[0035] The catalysts in the guard beds used in the optional
decontamination step of the process of the invention may
advantageously have the shape of spheres or extrudates. However, it
is advantageous for the catalyst to be in the shape of extrudates
with a diameter in the range 0.5 to 5 mm, more particularly in the
range 0.7 to 2.5 mm. The shapes are cylinders (which may or may not
be hollow), twisted cylinders, multilobes (2, 3, 4 or 5 lobes, for
example), or rings. The cylindrical shape is preferred, but any
other shape may be used.
[0036] To accommodate the presence of contaminants and/or poisons
in the feed, in a further preferred implementation, the guard bed
catalysts may have more particular geometrical forms to increase
their void fraction. The void fraction of said catalysts is in the
range 0.2 to 0.75. Their external diameter may be between 1 and 35
mm Possible particular non-limiting forms are: hollow cylinders,
hollow rings, Raschig rings, toothed hollow cylinders, crenellated
hollow cylinders, pentaring cartwheels, multiple holed cylinders,
etc.
Active Phase
[0037] Said catalysts of the guard beds used in the optional
decontamination step of the process of the invention may
advantageously have been impregnated with a phase which may or may
not be active. Preferably, the catalysts are impregnated with a
hydrodehydrogenating phase. Highly preferably, the CoMo or NiMo
phase is used. Still more preferably, the NiMo phase is used.
[0038] Preferably, the supports for said guard bed catalysts are
porous refectory oxides, preferably selected from alumina and
silica-alumina.
[0039] Said guard bed catalysts may advantageously have
macroporosity.
[0040] Said catalysts advantageously have a macroporous mercury
volume for a mean diameter of 50 nm which is more than 0.1
cm.sup.3/g, preferably in the range 0.125 to 0.175 cm.sup.3/g, and
a total volume of more than 0.60 cm.sup.3/g, preferably in the
range 0.625 to 0.8 cm.sup.3/g, and is advantageously impregnated
with an active phase, preferably based on nickel and molybdenum,
such as ACT961, for example. In this preferred embodiment, the Ni
content as the weight of oxide is generally in the range 1% to 10%
and the Mo content as the weight of oxide is in the range 5% to
15%. The surface areas, expressed as the SBET, of the supports for
said catalysts are in the range 30 m.sup.2/g to 220 m.sup.2/g.
[0041] In a first embodiment, the guard bed advantageously also
comprises at least one other catalyst having a mercury volume for a
pore diameter of more than 1 micron of more than 0.2 cm.sup.3/g and
preferably more than 0.5 cm.sup.3/g, and a mercury volume for a
pore diameter of more than 10 microns of more than 0.25 cm.sup.3/g
and preferably less than 0.4 cm.sup.3/g, said catalyst
advantageously being placed upstream of the catalyst of the
invention.
[0042] In a second embodiment, the guard bed advantageously also
comprises at least one other catalyst with a mercury volume for a
pore diameter of more than 50 nm of more than 0.25 cm.sup.3/g, the
mercury volume for a pore diameter of more than 100 nm being more
than 0.15 cm.sup.3/g and a total pore volume of more than 0.80
cm.sup.3/g.
[0043] Said guard bed catalyst and the catalyst of the first
embodiment may advantageously be associated in a mixed bed or a
combined bed. In general, the impregnated catalyst of the active
phase constitutes the majority of the guard bed and the catalyst of
the first embodiment which is preferred is added as a complement of
0 to 50% by volume with respect to the first catalyst, preferably 0
to 30%, more preferably 1% to 20%.
[0044] The combination of the catalyst of the invention and the
catalyst of the first embodiment does not limit the scope of the
invention. The catalysts which can be used in the guard beds may
advantageously be used alone or as a mixture; in a non-exhaustive
manner, they may be selected from those sold by
Norton-Saint-Gobain, for example MacroTrap.RTM. guard beds, or
catalysts sold by Axens from the ACT family: ACT077, ACT935, ACT961
or HMC841, HMC845, HMC941 or HMC945.
[0045] Preferred guard beds for use in the invention are the HMCs
and ACT961.
[0046] It may be particularly advantageous to superimpose these
catalysts in at least two different beds of varying heights. The
catalysts with the highest void ratio are preferably used in the
first catalytic bed or beds at the inlet to the catalytic reactor.
It may also be advantageous to use at least two different reactors
for said catalysts.
[0047] Advantageously, an association of said guard bed catalyst
with the catalysts of the first and second implementation is also
possible in a mixed bed or a combined bed. In this case, the
catalysts are placed with the void capacity decreasing in the
direction of flow.
[0048] After passing over said guard bed, the quantity of solid
particles is less than 20 ppm, preferably less than 10 ppm and more
preferably less than 5 ppm. The soluble silicon content is less
than 5 ppm, preferably less than 2 ppm and more preferably less
than 1 ppm.
Step a)
[0049] In accordance with step a) of the process of the invention,
said C5+ liquid paraffinic fraction, termed the heavy fraction,
with an initial boiling point in the range 15.degree. C. to
40.degree. C. produced by Fischer-Tropsch synthesis passes over at
least one ion exchange resin which can esterify the alcohols and
carboxylic acids into esters and/or capture metals dissolved in the
feed, at a temperature in the range 80.degree. C. to 150.degree.
C., at a total pressure in the range 0.7 to 2.5 MPa, and at an
hourly space velocity in the range 0.2 to 2.5 h.sup.-1.
[0050] Step a) of the invention may advantageously be carried out
in accordance with two distinct implementations, namely either in a
single reactor 4 over a single ion exchange resin, advantageously
used to simultaneously carry out esterification of alcohols and
carboxylic acids to esters and capture of the metals dissolved in
the feed, or in two different reactors on two ion exchange resins 3
and 4 of different natures, one having the specific function of
esterification of alcohols and carboxylic acids and the other the
capture of the metals dissolved in the feed.
[0051] In a first implementation, step a) advantageously consists
of passing said C5+ liquid paraffinic fraction over a single ion
exchange resin in a single reactor 4 to simultaneously carry out
the esterification of alcohols and carboxylic acids to esters and
the capture of metals dissolved in the feed.
[0052] Preferably, said resin is used at a temperature in the range
100.degree. C. to 150.degree. C. and preferably in the range
100.degree. C. to 130.degree. C., at a pressure in the range 1 to 2
MPa and preferably in the range 1 to 1.5 MPa, and at an hourly
space velocity in the range 0.5 to 2 h.sup.-1, preferably in the
range 0.5 to 1.5 h.sup.-1.
[0053] In this case, oxygen-containing compounds, carboxylic acids
and alcohols are adsorbed onto the active sites of said resin and
are esterified and the cationic and metallic compounds present in
said C5+ liquid paraffinic fraction are eliminated by adsorption or
by ion exchange. Said resin, which can simultaneously carry out the
esterification of alcohols and carboxylic acids to esters and the
capture of metals dissolved in the feed, advantageously comprises
sulphonic acid groups and is prepared by polymerization or
co-polymerization of aromatic vinyl groups followed by
sulphonation, said aromatic vinyl groups being selected from
styrene, vinyl toluene, vinyl naphthalene, vinyl ethyl benzene,
methyl styrene, vinyl chlorobenzene and vinyl xylene, said resin
having a degree of cross-linking in the range 20% to 35%,
preferably in the range 25% to 35%, preferably equal to 30%, and an
acid strength, assayed by potentiometry during neutralization with
a KOH solution, of 0.2 to 6 mmol H+ equivalent/g, preferably in the
range 0.2 to 2.5 mmol H+ equivalent/g.
[0054] Said acid ion exchange resin advantageously contains in the
range 1 to 2 terminal sulphonic groups per aromatic group.
Preferably, said resin has a size in the range 0.15 to 1.5 mm The
size of a resin is the diameter of the sphere encompassing the
resin particle. The resin size categories are measured by screening
through suitable screens, in accordance with a technique which is
known in the art.
[0055] A preferred resin is a resin constituted by co-polymers of
monovinyl aromatics and polyvinyl aromatics, and highly preferably,
a copolymer of divinyl benzene and polystyrene with a degree of
cross-linking in the range 20% to 35%, preferably in the range 25%
to 35%, and more preferably equal to 30%, and an acid strength,
representing the number of active sites of said resin, assayed by
potentiometry during neutralization using a KOH solution, in the
range 0.2 to 6 mmol H+ equivalent/g, preferably in the range 0.2 to
2.5 mmol H+ equivalent/g.
[0056] Another preferred resin which can simultaneously allow the
esterification of alcohols and carboxylic acids and the capture of
metals to be carried out is a resin constituted by a polysiloxane
grafted with alkylsulphonic type acid groups (of the
--CH.sub.2--CH.sub.2--CH.sub.2--SO.sub.3H type), with a size in the
range 0.5 to 1.2 mm and with an acid strength, representing the
number of active sites of said resin and assayed by potentiometry
during neutralization with a KOH solution, of 0.4 to 1.5 mmol H+
equivalent/g.
[0057] During said step and under these conditions, 95% of the
carboxylic acids are esterified. The acid conversion is analyzed by
the potassium hydroxide titration difference between the feed and
the effluent using a technique which is known to the skilled
person. The ASTM methods D 664, D 3242 or D 974 can be cited, for
example, as methods for carrying out said analysis.
[0058] This resin may advantageously be used in a fixed bed between
screens placed in an upflow or downflow tube reactor. Preferably,
said resin is used in an upflow bed reactor, the liquid being
injected into the bottom of the reactor at a sufficient surface
velocity to allow the bed of resin to expand without, however,
either transporting or fluidizing it. This implementation, compared
with a fixed bed, can attenuate the effects of clogging materials
and substantially increase the service life of the resin.
[0059] In accordance with a second implementation, step a)
advantageously consists of passing said C5+ liquid paraffinic
fraction 1 into two different reactors 3 and 4 shown in FIG. 2,
over two distinct ion exchange resins, of different natures, one
having the specific function of esterification of alcohols and
carboxylic acids and the other that of capturing the metals
dissolved in the feed.
[0060] Preferably, the reactor 3 containing the ion exchange resin
which can capture metals is used upstream of the reactor 4
containing the ion exchange resin which can carry out the
esterification of the alcohols and carboxylic acids.
[0061] In this case, the cationic and metallic compounds present in
said C5+ liquid paraffinic fraction are eliminated by adsorption or
by ion exchange on a first ion exchange resin. This first resin,
which is specific for capturing metals, advantageously comprises
sulphonic acid groups and is advantageously prepared by
polymerization or co-polymerization of aromatic vinyl groups
followed by sulphonation, said aromatic vinyl groups advantageously
being selected from styrene, vinyl toluene, vinyl naphthalene,
vinyl ethyl benzene, methyl styrene, vinyl chlorobenzene and vinyl
xylene, said resin having a degree of cross-linking in the range 1%
to 20%, preferably in the range 2% to 8%, and an acid strength,
representing the number of active sites in said resin, assayed by
potentiometry during neutralization with a KOH solution, in the
range 1 to 15 mmol H+ equivalent/g, preferably in the range 2.5 to
10 mmol H+ equivalent/g.
[0062] Said acid ion exchange resin advantageously contains between
1 and 2 terminal sulphonic acid groups per aromatic group.
Preferably, said resin has a size in the range 0.15 to 1.5 mm The
size of the resin is the diameter of the sphere encompassing the
resin particle. The size classes for the resin are measured by
screening through suitable screens using a technique which is known
to the skilled person.
[0063] Preferably, the first resin is a resin constituted by
copolymers of monovinyl aromatics and polyvinyl aromatics; more
preferably, a copolymer of divinyl benzene and polystyrene with a
degree of cross-linking in the range 1% to 20% and an acid
strength, representing the number of active sites of said resin and
assayed by potentiometry during neutralization with a KOH solution,
of 1 to 15 mmol H+equivalent/g, preferably in the range 2.5 to 10
mmol H+ equivalent/g.
[0064] Preferably, said first resin is used at a temperature in the
range 80.degree. C. to 110.degree. C., at a pressure in the range 1
to 2 MPa and preferably in the range 1 to 1.5 MPa, and at an hourly
space velocity in the range 0.2 to 1.5 h.sup.-1, preferably in the
range 0.5 to 1.5 h.sup.-1.
[0065] The effluent from the reactor 3 containing said first resin
which is specific for the capture of metals is then introduced into
a second reactor 4 located downstream of the first reactor and
containing a second resin with a different nature and which is
specific to esterification of the alcohols and carboxylic acids
contained in said effluent.
[0066] The oxygen-containing compounds, carboxylic acids and
alcohols are adsorbed onto the active sites of said second resin
and are esterified and the cationic and metallic compounds present
in the effluent from reactor 3 are eliminated by adsorption or by
ion exchange. Said second resin, which can simultaneously carry out
the esterification of alcohols and carboxylic acids to esters and
the capture of metals dissolved in the feed, advantageously
comprises sulphonic acid groups and is advantageously prepared by
polymerization or copolymerization of aromatic vinyl groups
followed by sulphonation. The aromatic vinyl groups are
advantageously selected from styrene, vinyl toluene, vinyl
naphthalene, vinyl ethyl benzene, methyl styrene, vinyl
chlorobenzene and vinyl xylene, said second resin having a degree
of cross-linking, i.e. a ratio of the mass of copolymer/mass of
polymer, which is advantageously in the range 20% to 35%,
preferably in the range 25% to 35% and more preferably 30%, and an
acid strength, representing the number of active sites of said
resin, assayed by potentiometry during neutralization with a KOH
solution, in the range 0.2 to 6 mmol H+equivalent/g, preferably in
the range 0.2 to 6 mmol H+ equivalent/g.
[0067] Said second acid ion exchange resin advantageously contains
1 to 2 terminal sulphonic groups per aromatic group. Preferably,
said second resin has a size in the range 0.15 to 1.5 mm.
[0068] A preferred second resin is a resin constituted by
copolymers of monovinyl aromatics and aromatic polyvinyls, and more
preferably, a copolymer of divinyl benzene and polystyrene with a
degree of cross-linking in the range 20% to 35%, preferably in the
range 25% to 35% and more preferably 30%, and an acid strength,
representing the number of active sites of said resin, assayed by
potentiometry during neutralization with a KOH solution, in the
range 0.2 to 6 mmol H+ equivalent/g and preferably in the range 0.2
to 6 mmol H+ equivalent/g.
[0069] Another preferred resin which can simultaneously allow
esterification of alcohols and carboxylic acids and capture of
metals to be carried out is a resin constituted by a polysiloxane
grafted with alkylsulphonic type acid groups (of the
--CH.sub.2--CH.sub.2--CH.sub.2--SO.sub.3H type), with a size in the
range 0.5 to 1.2 mm and with an acid strength, representing the
number of active sites of said resin and assayed by potentiometry
during neutralization with a KOH solution, of 0.4 to 1.5 mmol H+
equivalent/g.
[0070] Preferably, said second resin is used at a temperature in
the range 100.degree. C. to 150.degree. C. and preferably in the
range 100.degree. C. to 130.degree. C., at a pressure in the range
1 to 2 MPa and preferably in the range 1 to 1.5 MPa, and at an
hourly space velocity in the range 0.5 to 2 h.sup.-1, preferably in
the range 0.5 to 1.5 .sup.-1.
[0071] During this step and under these conditions, 95% of the
carboxylic acids are esterified. Analysis of the conversion of the
acids is given by the difference in the potassium hydroxide
titration between the feed and the effluent. The ASTM methods D
664, D 3242 or D 974 can be cited, for example, as methods for
carrying out said analysis.
[0072] These resins may advantageously be used in a fixed bed
between screens placed in an upflow or downflow tube reactor.
Preferably, said resin is used in an upflow bed reactor, the liquid
being injected into the bottom of the reactor at a sufficient
surface velocity to allow the bed of resin to expand without,
however, either transporting or fluidizing it. This implementation,
compared with a fixed bed, can attenuate the effects of clogging
materials and substantially increase the service life of the
resin.
[0073] In the case in which said C5+ liquid paraffinic fraction
passes into two different reactors over two distinct ion exchange
resins of different natures, one principally carrying out metals
capture, the other principally carrying out esterification,
intermediate re-heating is preferably employed between the two
steps. Preferably, the water formed during the step for
esterification of the acids and alcohols over the first resin
principally carrying out the capture of metals is removed in order
to intensify the esterification reaction over the second resin.
Simultaneously adding intermediate re-heating and water separation
boosts the overall conversion of the carboxylic acids present in
the feed.
[0074] The reaction for esterification of the organic acids by the
alcohols present in said C5+ liquid paraffinic fraction, termed a
heavy fraction, with an initial boiling point in the range
15.degree. C. to 40.degree. C., produced by the Fischer-Tropsch
synthesis, produces water which is a compound that inhibits the
hydrotreatment and hydrocracking catalysts, necessitating an
increase in the severity of the operating conditions.
Step b)
[0075] In accordance with the invention, the effluent derived from
step a) then undergoes a step for eliminating at least a portion of
the water formed during said step a), preferably all of the water
formed, in a separator 6.
[0076] This water is acidic in nature as it advantageously contains
protons exchanged during capture of the metals by the upstream
specific cation exchange resin or by the only resin allowing
simultaneous esterification and capture of metals. This water may
also contain a fraction of dissolved CO and CO.sub.2 originating
from the Fischer-Tropsch synthesis. The water is eliminated via the
line 7.
[0077] It is also advantageous to add in said step b) gas of the
nitrogen (N.sub.2) or hydrogen (H.sub.2) type, 5, to eliminate more
dissolved CO and CO.sub.2 by stripping.
[0078] In the case when hydrogen is added, this advantageously acts
as a makeup gas for the hydrotreatment step.
[0079] This step can also eliminate products of the light ether
type formed during the reaction of alcohols with themselves.
[0080] The water may be eliminated using any of the methods and
techniques known to the skilled person, for example by drying,
passage over a dessicant, flash, decantation, etc.
[0081] The effluent from water elimination step b) constitutes at
least part and preferably the whole of the feed for hydrogenation
step c) of the process of the invention.
Step c)
[0082] Step c) of the process of the invention is a step for
hydrogenation of the unsaturated olefinic type compounds of at
least a portion and preferably the whole of the effluent derived
from step b) of the process of the invention, in the presence of
hydrogen and a hydrogenation catalyst.
[0083] The effluent from step b) of the process of the invention is
admitted in the presence of hydrogen (line 8) into a hydrogenation
zone 9 containing a hydrogenation catalyst which is intended to
saturate the unsaturated olefinic type compounds present in said
effluent.
[0084] Preferably, the catalyst used in step c) of the invention is
a non-cracking or low cracking hydrogenation catalyst comprising at
least one metal from group VIII of the periodic table of the
elements and comprising at least one support based on a refractory
oxide.
[0085] Preferably, said catalyst comprises at least one metal from
group VIII selected from nickel, cobalt, ruthenium, indium,
palladium and platinum and comprising at least one support based on
refractory oxide selected from alumina and silica-alumina.
[0086] Preferably, the metal from group VIII is selected from
nickel, palladium and platinum; highly preferably, it is selected
from palladium and platinum.
[0087] In accordance with a preferred implementation of step c) of
the process of the invention, the metal from group VIII is selected
from palladium and/or platinum and the quantity of this metal is
advantageously in the range 0.1% to 5% by weight, preferably in the
range 0.2% to 0.6% by weight with respect to the total catalyst
weight.
[0088] In accordance with a highly preferred implementation of step
c) of the process of the invention, the metal from group VIII is
palladium.
[0089] According to another preferred implementation of step c) of
the process of the invention, the metal from group VIII is nickel
and the quantity of this metal is advantageously in the range 5% to
25% by weight, preferably in the range 7% to 20% by weight with
respect to the total catalyst weight.
[0090] The support for the catalyst used in step c) of the process
of the invention is a support based on a refractory oxide,
preferably selected from alumina and silica-alumina, more
preferably alumina.
[0091] When the support is an alumina, it has a BET specific
surface area which can limit polymerization reactions at the
surface of the hydrogenation catalyst, said surface area being in
the range 5 to 140 m.sup.2/g.
[0092] When the support is a silica-alumina, the support contains a
percentage of silica in the range 5% to 95% by weight, preferably
in the range 10% to 80%, more preferably in the range 20% to 60% by
weight and highly preferably in the range 30% to 50%, a BET
specific surface area in the range 100 to 550 m.sup.2/g, preferably
in the range 150 to 500 m.sup.2/g, more preferably less than 350
m.sup.2/g and still more preferably less than 250 m.sup.2/g.
[0093] The hydrogenation step c) of the process of the invention is
preferably carried out in one or more fixed bed reactors.
[0094] In hydrogenation zone 9, the feed is brought into contact
with the hydrogenation catalyst in the presence of hydrogen and at
operating temperatures and pressures which allow hydrogenation of
the unsaturated olefinic type compounds present in the feed.
[0095] The operating conditions for hydrogenation step c) of the
process of the invention are advantageously as follows: the
temperature in said hydrogenation zone 9 is in the range
100.degree. C. to 180.degree. C., preferably in the range
120.degree. C. to 165.degree. C., the total pressure is in the
range 0.5 to 6 MPa, preferably in the range 1 to 5 MPa and more
preferably in the range 2 to 5 MPa. The flow rate of the feed is
such that the hourly space velocity (ratio of the hourly volume
flow rate at 15.degree. C. for fresh liquid feed to the volume of
charged catalyst) is in the range 1 to 50 h.sup.-, preferably in
the range 2 to 20 h.sup.-1 and more preferably in the range 4 to 20
h.sup.-1. The hydrogen which supplies the hydrotreatment zone is
introduced at a flow rate such that the hydrogen/hydrocarbon volume
ratio is in the range 5 to 300 Nl/l/h, preferably in the range 5 to
200, more preferably in the range 10 to 150 Nl/l/h, and still more
preferably in the range 10 to 50 Nl/l/h.
[0096] Under these conditions, the unsaturated olefinic type
compounds are more than 50%, preferably more than 75% and more
preferably more than 85% hydrogenated.
[0097] The effluent from step c) optionally undergoes a step for
elimination of at least a portion of the water formed during
hydrogenation step c), preferably all of the water which is
formed.
[0098] This water may also contain a fraction of dissolved CO and
CO.sub.2 originating from the Fischer-Tropsch synthesis. This step
for eliminating water takes place in the separator 11 and water is
eliminated via the line 12.
[0099] It may also be advantageous to add to said step for
elimination of at least a portion of the water a nitrogen (N.sub.2)
or hydrogen (H.sub.2) type gas (line 10) to eliminate more
dissolved CO and CO.sub.2 by stripping.
[0100] This step can also eliminate light ether type products
formed during the reaction of alcohols on themselves.
[0101] The water may be eliminated using any of the methods and
techniques known to the skilled person, for example drying, passage
over a dessicant, flash, decantation, etc.
[0102] At the end of step c) of the process of the invention, at
least a portion and preferably all of the liquid hydrogenated
effluent is sent to a hydrocracking/hydroisomerization zone 14.
Step d)
[0103] In accordance with step d) of the process of the invention,
at least a portion and preferably all of the liquid hydrogenated
effluent from hydrogenation step c) of the process of the invention
is sent to the hydroisomerization/hydrocracking zone 14 containing
the hydroisomerization/hydrocracking catalyst, preferably at the
same time as a stream of hydrogen.
[0104] The operating conditions in which
hydroisomerization/hydrocracking step d) of the process of the
invention is carried out are preferably as follows:
[0105] The pressure is generally maintained between 0.2 and 15 MPa,
preferably in the range 0.5 to 10 MPa and advantageously in the
range 1 to 9 MPa; the hourly space velocity is generally in the
range 0.1 h.sup.-1 to 10 h.sup.-1, preferably in the range 0.2 to 7
h.sup.-1 and advantageously in the range 0.5 to 5.0 h.sup.-1. The
hydrogen ratio is generally in the range 100 to 2000 normal litres
of hydrogen per litre of feed per hour, preferably in the range 150
to 1500 normal litres of hydrogen per litre of feed per hour.
[0106] The temperature used in this step is generally in the range
200.degree. C. to 450.degree. C., preferably in the range
250.degree. C. to 450.degree. C., advantageously in the range
300.degree. C. to 450.degree. C., and more advantageously more than
320.degree. C. or, for example, in the range 320.degree. C. to
420.degree. C.
[0107] Hydroisomerization and hydrocracking step d) of the process
of the invention is advantageously carried out under conditions
such that the conversion per pass of products with a boiling point
of 370.degree. C. or more into products with boiling points of less
than 370.degree. C. is more than 80% by weight, and more preferably
at least 85%, preferably more than 88%, in order to obtain middle
distillates (gas oil and kerosene) with sufficiently good cold
properties (pour point, freezing point) so that they satisfy
specifications in force for this type of fuel.
The Hydroisomerization/Hydrocracking Catalysts
[0108] The majority of the catalysts in current use in
hydroisomerization are bi-functional in type, associating an acid
function with a hydrogenating function. The acid function is
supplied by supports with large surface areas (generally of 150 to
800 m.sup.2/g) and with a superficial acidity, such as halogenated
aluminas (in particular chlorinated or fluorinated),
phosphorus-containing aluminas, combinations of oxides of boron and
aluminium, and silica-aluminas. The hydrogenating function is
generally supplied either by one or more metals from group VIII of
the periodic table of the elements such as iron, cobalt, nickel,
ruthenium, rhodium, palladium, osmium, iridium or platinum, or by a
combination of at least one metal from group VI, such as chromium,
molybdenum or tungsten, and at least one group VIII metal.
[0109] In the case of bi-functional catalysts, the balance between
the two functions, acid and hydrogenating, is the fundamental
parameter which governs the activity and selectivity of the
catalyst. A weak acid function and a strong hydrogenating function
produces less active catalysts which are also less selective as
regards isomerization, while a strong acid function and a weak
hydrogenating function produces catalysts which are highly active
and selective as regards cracking. A third possibility is to use a
strong acid function and a strong hydrogenating function to obtain
a catalyst which is highly active but also highly selective as
regards isomerization. Thus, by carefully selecting each of the
functions, it is possible to adjust the activity/selectivity
balance of the catalyst.
[0110] Advantageously, the hydroisomerization/hydrocracking
catalysts are bi-functional catalysts comprising an amorphous acid
support (preferably a silica-alumina) and a metallic
hydro-dehydrogenating function which is preferably provided by at
least one noble metal. The support is termed amorphous, i.e. free
of molecular sieve and in particular zeolite, as is the catalyst.
The amorphous acid support is advantageously a silica-alumina, but
other supports may be used. When it is a silica-alumina, the
catalyst preferably contains no added halogen other than that which
may be introduced for impregnation with the noble metal, for
example.
[0111] More generally and preferably, the catalyst contains no
added halogen, for example fluorine. In general and preferably, the
support has not undergone impregnation with a silicon compound.
[0112] In accordance with a first preferred implementation, the
hydroisomerization/hydrocracking catalyst contains at least one
hydrodehydrogenating element selected from noble group VIII metals,
preferably platinum and/or palladium, and at least one amorphous
refractory oxide support, preferably silica-alumina.
[0113] A preferred hydroisomerization/hydrocracking catalyst used
in step d) of the process of the invention comprises up to 3% by
weight of a metal of at least one hydro-dehydrogenating element
selected from noble metals from group VIII, preferably deposited on
the support, and highly preferably, the noble group VIII metal is
platinum; and a support comprising (or preferably constituted by)
at least one silica-alumina, said silica-alumina having the
following characteristics: [0114] a weight content of silica,
SiO.sub.2, in the range 5% to 95%, preferably in the range 10% to
80%, more preferably in the range 20% to 60% and still more
preferably in the range 30% to 50% by weight; [0115] a Na content
of less than 300 ppm by weight, preferably less than 200 ppm by
weight; [0116] a total pore volume in the range 0.45 to 1.2 ml/g,
measured by mercury porosimetry; [0117] the porosity of said
silica-alumina being as follows: [0118] i) the volume of mesopores
with a diameter in the range 40 .ANG. to 150 .ANG. and with a mean
diameter in the range 80 to 140 .ANG., preferably in the range 80
to 120 .ANG., represents 20-80% of the total pore volume measured
by mercury porosimetry; [0119] ii) the volume of macropores with a
diameter of more than 500 .ANG., preferably in the range 1000 .ANG.
to 10000 .ANG., represents 20% to 80% of the total pore volume,
measured by mercury porosimetry; [0120] a specific surface area in
the range 100 to 550 m.sup.2/g, preferably in the range 150 to 500
m.sup.2/g, more preferably less than 350 m.sup.2/g and still more
preferably less than 250 m.sup.2/g.
[0121] A second preferred hydroisomerization/hydrocracking catalyst
used in step d) of the process of the invention comprises up to 3%
by weight of a metal of at least one hydro-dehydrogenating element
selected from noble metals from group VIII of the periodic table of
the elements, and preferably, the noble group VIII metal is
platinum; 0.01% to 5.5% by weight of oxide of a doping element
selected from phosphorus, boron and silicon; and a non-zeolitic
support based on silica-alumina containing a quantity of more than
15% by weight and 95% by weight or less of silica (SiO.sub.2), said
silica-alumina having the following characteristics: [0122] a mean
pore diameter, measured by mercury porosimetry, in the range 20 to
140 .ANG.; [0123] a total pore volume, measured by mercury
porosimetry, in the range 0.1 ml/g to 0.5 ml/g; [0124] a total pore
volume, measured by nitrogen porosimetry, in the range 0.1 ml/g to
0.6 ml/g; [0125] a BET specific surface area in the range 100 to
550 m.sup.2/g; [0126] a pore volume, measured by mercury
porosimetry, included in pores with a diameter of more than 140
.ANG., of less than 0.1 ml/g; [0127] a pore volume, measured by
mercury porosimetry, included in pores with a diameter of more than
160 .ANG., of less than 0.1 ml/g; [0128] a pore volume, measured by
mercury porosimetry, included in pores with a diameter of more than
200 .ANG., of less than 0.1 ml/g; [0129] a pore volume, measured by
mercury porosimetry, included in pores with a diameter of more than
500 .ANG., of less than 0.1 ml/g; [0130] an X ray diffraction
diagram which contains at least the principal characteristic peaks
of at least one of the transition aluminas included in the group
composed of alpha, rho, chi, eta, gamma, kappa, theta and delta
aluminas; [0131] a settled catalyst packing density of more than
0.55 g/cm.sup.3.
[0132] Advantageously, the characteristics associated with the
corresponding catalyst are identical to those of the silica-alumina
described above.
[0133] The two steps c) and d) of the process of the invention,
hydrogenation and hydroisomerization-hydrocracking, may
advantageously be carried out on the two types of catalysts in two
or more different reactors and/or in the same reactor.
[0134] In accordance with a second preferred implementation, the
hydroisomerization/hydrocracking catalyst contains at least one
hydrodehydrogenating element selected from non-noble group VIII
metals and metals from group VIB and at least one amorphous
refractory oxide support, preferably silica-alumina.
[0135] Preferably, the metal from group VIII is selected from
nickel and cobalt, and the metal from group VIB is selected from
molybdenum and tungsten.
[0136] Preferably, said catalyst is in the sulphide form.
[0137] A third preferred hydroisomerization/hydrocracking catalyst
used in step d) of the process of the invention comprises at least
one hydro-dehydrogenating element selected from non-noble metals
from group VIII and metals from group VIB of the periodic table of
the elements, preferably between 2.5% and 5% by weight of oxide of
the non-noble element from group VIII and between 20% and 35% by
weight of oxide of a group VIB element with respect to the weight
of the final catalyst; preferably, the non-noble group VIII metal
is nickel and the group VIB metal is tungsten; optionally 0.01% to
5.5% by weight of oxide of a doping element selected from
phosphorus, boron and silicon; preferably, 0.01% to 2.5% by weight
of oxide of a doping element and a non-zeolitic support based on
silica-alumina containing a quantity of more than 15% by weight and
95% by weight or less of silica (SiO.sub.2), preferably a quantity
of more than 15% by weight and 50% by weight or less of silica,
said silica-alumina having the following characteristics: [0138] a
mean pore diameter, measured by mercury porosimetry, in the range
20 to 140 .ANG.; [0139] a total pore volume, measured by mercury
porosimetry, in the range 0.1 ml/g to 0.5 ml/g; [0140] a total pore
volume, measured by nitrogen porosimetry, in the range 0.1 ml/g to
0.6 ml/g; [0141] a BET specific surface area in the range 100 to
550 m.sup.2/g; [0142] a pore volume, measured by mercury
porosimetry, included in pores with a diameter of more than 140
.ANG., of less than 0.1 ml/g; [0143] a pore volume, measured by
mercury porosimetry, included in pores with a diameter of more than
160 .ANG., of less than 0.1 ml/g; [0144] a pore volume, measured by
mercury porosimetry, included in pores with a diameter of more than
200 .ANG., of less than 0.1 ml/g; [0145] a pore volume, measured by
mercury porosimetry, included in pores with a diameter of more than
500 .ANG., of less than 0.1 ml/g; [0146] an X ray diffraction
diagram which contains at least the principal characteristic peaks
of at least one of the transition aluminas included in the group
composed of alpha, rho, chi, eta, gamma, kappa, theta and delta
aluminas; [0147] a settled catalyst packing density of more than
0.55 g/cm.sup.3.
[0148] Advantageously, the characteristics associated with the
corresponding catalyst are identical to those of the silica-alumina
described above.
[0149] When the third preferred hydroisomerization/hydrocracking
catalyst is used in step d) of the process of the invention, the
catalyst is sulphurized.
[0150] In accordance with a first preferred implementation of the
process of the invention, in hydrogenation step c), a catalyst is
used which contains palladium and in
hydroisomerization/hydrocracking step d), a catalyst containing
platinum is used.
[0151] In accordance with a second preferred implementation of the
process of the invention, in hydrogenation step c) a catalyst
containing palladium is used and in
hydroisomerization/hydrocracking step d), a sulphurized catalyst
containing at least one hydro-dehydrogenating element selected from
non-noble metals from group VIII and group VIB metals is used.
[0152] In a third preferred implementation of the process of the
invention, in hydrogenation step c) a catalyst containing at least
one non-noble hydro-dehydrogenating element from group VIII is used
and in hydroisomerization/hydrocracking step d), a sulphurized
catalyst containing at least one hydro-dehydrogenating element
selected from non-noble group VIII metals and group VIB metals is
used.
Step e)
[0153] In accordance with step e) of the process of the invention,
the effluent derived from step d) undergoes the separation of
unreacted hydrogen and light gases in a gas/liquid separator 15
then recycling of the unreacted hydrogen and said light gases to
the hydroisomerization/hydrocracking step d) (line 17).
[0154] Said light gases include light C1-C4 gases, carbon monoxide
(CO), carbon dioxide (CO.sub.2) and water in the vapour form.
[0155] Said gases are separated from the liquid effluent in one or
more flash drums 15, i.e. one or more drums which carry out
separation of the gases and the liquids introduced via the line
14b, at staged temperatures and pressures in order to increase the
recovery of hydrogen. This flash staging may advantageously be
accompanied by a heat exchanger aimed at recovering heat energy
and/or cooling the effluents from the separator drums in order to
minimize losses of hydrogen.
[0156] By using an ion exchange resin upstream of the
hydrotreatment and hydrocracking steps, the process of the
invention can reduce the total oxygen content of the feed and thus
limit the formation of carbon monoxide (CO) originating from the
decomposition of oxygen-containing compounds present in the feed in
the hydroisomerization/hydrocracking section. Carbon monoxide (CO)
is an inhibitor of the metallic compounds present on the
hydroisomerization/hydrocracking catalyst, and its content must be
minimized in order not to require an increase in temperature in
order to compensate for the low activity and maintain
conversion.
[0157] However, when the gaseous effluent 17 from said separation
has a high CO fraction, i.e. more than 10 ppm by volume, said
gaseous effluent is advantageously sent to a methanation reactor 18
in which the conversion of carbon monoxide (CO) and hydrogen into
methane is advantageously carried out, with the aim of limiting the
CO content. The principle of methanation, and the catalysts used
are known to the skilled person and their use in purifying
effluents containing H.sub.2 and CO is known.
[0158] A purge is advantageously carried out (line 20) in order to
eliminate the products formed during the methanation step 18. A
makeup of hydrogen (line 21) is then advantageously carried out in
order to compensate for that purge.
Step f)
[0159] In accordance with step f) of the process of the invention,
the effluent from step e) for separating hydrogen and the light
gases of the process of the invention is sent, to a distillation
train 22 via a line 16, which combines atmospheric distillation
with optional vacuum distillation, which is intended to separate
conversion products with a boiling point of less than 340.degree.
C. and preferably less than 370.degree. C. and in particular
including those formed during step d) in the
hydroisomerization/hydrocracking reactor 14, and to separate the
residual fraction with an initial boiling point which is generally
more than at least 340.degree. C. and preferably at least
370.degree. C. or higher. Of the converted and hydroisomerized
products, in addition to the light C1-C4 gases (line 23), at least
one gasoline (or naphtha) fraction is separated (line 24), and at
least one kerosene middle distillate fraction (line 25) and a gas
oil fraction (line 27) are separated. Preferably, the residual
fraction, with an initial boiling point which is generally over at
least 340.degree. C. and preferably at least 370.degree. C. is
recycled (line 28) to step d) of the process of the invention to
the head of the hydroisomerization and hydrocracking zone 14. In
accordance with another implementation of step f) of the process of
the invention, said residual fraction may supply excellent oil
bases.
[0160] It may also be advantageous to recycle (line 26) at least
part and preferably all of at least one of the kerosene and gas oil
cuts obtained to step d) (line 14). The gas oil and kerosene cuts
are preferably recovered separately or mixed, but the cut points
are adjusted by the operator as a function of requirements. It has
been shown that it is advantageous to recycle part of the kerosene
to improve its cold properties.
Products Obtained
[0161] The gas oil(s) obtained have a pour point of at most
0.degree. C., generally less than -10.degree. C. and usually less
than -15.degree. C. The cetane index is more than 60, generally
more than 65, and usually more than 70.
[0162] The kerosene(s) obtained have a freezing point of at most
-35.degree. C., generally less than -40.degree. C. The smoke point
is more than 25 mm, generally more than 30 mm In this process,
gasoline (unwanted) production is as low as possible. The gasoline
yield is always less than 50% by weight, preferably less than 40%
by weight, advantageously less than 30% by weight or 20% by weight
or even 15% by weight.
EXAMPLE 1
Implementation of the Process of the Invention
[0163] The C5+ paraffinic effluent derived from a Fischer-Tropsch
synthesis unit is described in Table 1.
TABLE-US-00001 TABLE 1 Composition of C5+ fraction of FT effluent
Units Effluent from FT unit Density @ 15.degree. C. -- 0.782
Paraffins content wt % 80 Olefins content wt % 15 Alcohols content
wt % 3 Acid content wt % 2 Ester content wt % 2 CO content ppm by
weight 30 CO.sub.2 content ppm by weight 390 Organometallics ppm 5
Simulated distillation Initial boiling point .degree. C. 25 5% by
weight .degree. C. 50 10% by weight .degree. C. 77 30% by weight
.degree. C. 200 50% by weight .degree. C. 300 70% by weight
.degree. C. 400 90% by weight .degree. C. 530 95% by weight
.degree. C. 575 End point .degree. C. 650 370.degree. C. + fraction
wt % 35 Water ppm by weight 276
Step a)
[0164] The C5+ fraction 1 passed over an ion exchange resin with
trade name Amberlyst 35 sold by Rohm & Haas, said resin
allowing simultaneous capture of metals dissolved in the feed and
esterification of alcohols and carboxylic acids to esters. Said
resin was constituted by divinyl benzene--polystyrene copolymers
with a degree of cross-linking of 20 and an acid strength, assayed
by potentiometry during neutralization with a KOH solution, of 4.15
mmol H+ equivalent/g.
[0165] Step a) was carried out at a temperature of 110.degree. C.,
a pressure of 1 MPa, and at an hourly space velocity of 1 h.sup.-1.
Under these conditions, 95% of the acids were esterified, the
analysis of the conversion of the acids being given by the
difference in potassium hydroxide titration between the feed and
the effluent using the ASTM D664 method. The composition of the
outlet effluent is given in Table 2.
TABLE-US-00002 TABLE 2 Composition of effluent derived from step a)
after esterification Units Effluent from FT unit Density @
15.degree. C. -- 0.782 Paraffins content wt % 80 Olefins content wt
% 15 Alcohols content wt % 1 Acid content wt % <0.1 Ester
content wt % 4 CO content ppm by weight 30 CO.sub.2 content ppm by
weight 390 Organometallics Ppm <1 Simulated distillation Initial
boiling point .degree. C. 25 5% by weight .degree. C. 50 10% by
weight .degree. C. 77 30% by weight .degree. C. 200 50% by weight
.degree. C. 300 70% by weight .degree. C. 400 90% by weight
.degree. C. 530 95% by weight .degree. C. 575 End point .degree. C.
650 370.degree. C. + fraction wt % 35 Water ppm by weight 2500
Step b)
[0166] The effluent derived from step a) then underwent the
elimination of the water formed during said step a), by
decanting/coalescence in suitable equipment known to the skilled
person.
Step c)
[0167] All of the effluent derived from water elimination step b)
then underwent a step for hydrogenation in the presence of hydrogen
and a hydrogenation catalyst with trade name LD265 sold by Axens,
said catalyst comprising 0.3% by weight of palladium deposited on
an alumina with a specific surface area of 69 m.sup.2/g.
[0168] Hydrogenation step c) was carried out at a reaction
temperature of 130.degree. C., at a pressure of 3.5 MPa, the
hydrogen was introduced at a flow rate such that the
hydrogen/hydrocarbon volume ratio was 32 Nl/l/h, and with an hourly
space velocity of 10 h.sup.-1. Under these conditions, the
conversion of olefins was 85% by weight.
[0169] The liquid effluent derived from hydrogenation step c) had
the composition described in Table 3:
TABLE-US-00003 TABLE 3 Composition of effluent derived from step c)
Units Effluent from FT unit Density @ 15.degree. C. -- 0.782
Paraffins content wt % 93 Olefins content wt % 2 Alcohols content
wt % 1 Acid content wt % <0.1 Ester content wt % 4 CO content
ppm by weight 0 CO.sub.2 content ppm by weight 390 Organometallics
Ppm <1 Simulated distillation Initial boiling point .degree. C.
25 5% by weight .degree. C. 50 10% by weight .degree. C. 77 30% by
weight .degree. C. 200 50% by weight .degree. C. 300 70% by weight
.degree. C. 400 90% by weight .degree. C. 530 95% by weight
.degree. C. 575 End point .degree. C. 650 370.degree. C. + fraction
wt % 35 Water ppm by weight 300
Step d)
[0170] All of the effluent from hydrogenation step c) underwent a
hydroisomerization/hydrocracking step in the presence of fresh
hydrogen and a hydroisomerization/hydrocracking catalyst, in which
the residual fraction with an initial boiling point of more than
370.degree. C., unreacted hydrogen and light gases were
recycled.
[0171] The hydroisomerization/hydrocracking catalyst comprised 0.6%
by weight of platinum and a support comprising 29.3% by weight of
silica, SiO.sub.2, and 70.7% by weight of alumina, Al.sub.2O.sub.3,
a Na content of 100 ppm by weight, a total pore volume comprising
0.69 ml/g measured by mercury porosimetry, a volume of mesopores
with a mean diameter of 80 .ANG. representing 78% of the total pore
volume measured by mercury porosimetry, a volume of macropores with
a diameter of more than 500 .ANG. representing 22% of the total
pore volume measured by mercury porosimetry and a specific surface
area of 300 m.sup.2/g.
[0172] The hydroisomerization/hydrocracking step was carried out
under the conditions described in Table 4.
[0173] The conversion per pass for products with a boiling point of
370.degree. C. or more into products with a boiling point of less
than 370.degree. C. was 85%.
TABLE-US-00004 TABLE 4 Operating conditions for
hydroisomerization/hydrocracking step Unit H.sub.2 partial pressure
MPa 5 Space velocity, HSV h.sup.-1 1.0 Reaction temperature
.degree. C. 345 Hydrogen ratio N1/1 600
Step e)
[0174] The effluent from the hydroisomerization/hydrocracking step
underwent separation, in a gas/liquid separator, of the unreacted
hydrogen and light gases which were recycled to the
hydroisomerization/hydrocracking step in order to recover a liquid
effluent. The carbon monoxide (CO) content generated per pass in
the gaseous effluent was limited to 1.1% by weight.
Step f)
[0175] The liquid effluent from the separation step e) was then
sent to a distillation train to separate the light products formed
during these steps: the gases (C1-C4), a gasoline cut, a gas oil
cut and a kerosene cut, and also a fraction, termed the residual
fraction, which had an initial boiling point equal to 370.degree.
C. which was recycled in its entirety to the inlet to the
hydroisomerization/hydrocracking reactor in order to maximize the
production of gas oil and kerosene.
[0176] The yields are given in Table 5.
TABLE-US-00005 TABLE 5 Yield of various cuts after separation Wt %
Boiling point C1-C4 1.9 -161.degree. C. to 35.degree. C. Naphtha
12.1 35.degree. C. to 150.degree. C. Kerosene 34.5 150.degree. C.
to 250.degree. C. Gas oil 51.6 250.degree. C. to 370.degree. C.
Example 2
Comparative
[0177] A process for producing middle distillates from a C5+ liquid
paraffinic fraction, termed a heavy fraction, with an initial
boiling point in the range 15.degree. C. to 40.degree. C. produced
by Fischer-Tropsch synthesis which was the same as that used in
Example 1 was carried out, comprising a step for hydrogenation
followed by a hydroisomerization/hydrocracking step, with no prior
step for passage over at least one ion exchange resin; this was
carried out for comparison purposes.
[0178] The C5+ paraffinic fraction derived from the Fischer-Tropsch
synthesis unit was described in Table 1 of Example 1.
Hydrogenation Step
[0179] The C5+ liquid paraffinic fraction underwent a step for
hydrogenation in the presence of hydrogen and a hydrogenation
catalyst with trade name LD265 sold by Axens, said catalyst
comprising 0.3% by weight of palladium deposited on an alumina with
a specific surface area of 69 m.sup.2/g.
[0180] In order to maintain a conversion into olefins of 85% by
weight, as for Example 1, the hydrogenation step was carried out at
a reaction temperature of 150.degree. C., at a pressure of 3.5 MPa,
the hydrogen was introduced at a flow rate such that the
hydrogen/hydrocarbon volume ratio was 32 Nl/l/h and at an hourly
space velocity of 8 h.sup.-1.
[0181] Under these conditions, the conversion into olefins was
maintained at 85% by weight.
[0182] The liquid effluent derived from hydrogenation step c) had
the composition described in Table 9:
TABLE-US-00006 TABLE 9 Composition of effluent derived from the
hydrogenation step Units Effluent from FT unit Density @ 15.degree.
C. -- 0.782 Paraffins content wt % 91 Olefins content wt % 2
Alcohols content wt % 3 Acid content wt % 2 Ester content wt % 2 CO
content ppm by weight 30 CO.sub.2 content ppm by weight 390
Organometallics ppm <1 Simulated distillation Initial boiling
point .degree. C. 25 5% by weight .degree. C. 50 10% by weight
.degree. C. 77 30% by weight .degree. C. 200 50% by weight .degree.
C. 300 70% by weight .degree. C. 400 90% by weight .degree. C. 530
95% by weight .degree. C. 575 End point .degree. C. 650 370.degree.
C. + fraction wt % 35 Water ppm by weight 276
[0183] Hydroisomerization/Hydrocracking Step
[0184] All of the effluent from the hydrogenation step underwent a
hydroisomerization/hydrocracking step in the presence of fresh
hydrogen and a hydroisomerization/hydrocracking catalyst, in which
the residual fraction with an initial boiling point of more than
370.degree. C., unreacted hydrogen and light gases were
recycled.
[0185] The hydroisomerization/hydrocracking catalyst comprised 0.6%
by weight of platinum and a support comprising 29.3% by weight of
silica, SiO.sub.2, and 70.7% by weight of alumina, Al.sub.2O.sub.3,
a Na content of 100 ppm by weight, a total pore volume comprising
0.69 ml/g measured by mercury porosimetry, a volume of mesopores
with a mean diameter of 80 .ANG. representing 78% of the total pore
volume measured by mercury porosimetry, a volume of macropores with
a diameter of more than 500 .ANG. representing 22% of the total
pore volume measured by mercury porosimetry and a specific surface
area of 300 m.sup.2/g.
[0186] The hydroisomerization/hydrocracking step was carried out
under the conditions described in Table 10.
[0187] To maintain the conversion per pass for products with a
boiling point of 370.degree. C. or more into products with a
boiling point of less than 370.degree. C. at 85%, the temperature
was increased and adjusted to 370.degree. C.
TABLE-US-00007 TABLE 10 Operating conditions for
hydroisomerization/hydrocracking step Unit H.sub.2 partial pressure
MPa 5 Space velocity, HSV h.sup.-1 1.0 Reaction temperature
.degree. C. 370 Hydrogen ratio N1/1 600
Separation Step
[0188] The effluent from the hydroisomerization/hydrocracking step
underwent separation, in a gas/liquid separator, of the unreacted
hydrogen and light gases which were recycled to the
hydroisomerization/hydrocracking step in order to recover a liquid
effluent. The carbon monoxide (CO) content generated per pass in
the gaseous effluent was limited to 2% by weight.
[0189] The presence of carbon monoxide (CO) derived from the
decomposition of oxygen-containing compounds in the hydrocracking
section not eliminated by passage over at least one ion exchange
resin prior to hydrogenation and to
hydroisomerization/hydrocracking, and being an inhibitor of the
activity of the hydroisomerization/hydrocracking catalyst activity,
necessitated an increase in the reaction temperature in order to
maintain the conversion.
Final Distillation Step
[0190] The liquid effluent from the separation step was then sent
to a distillation train to separate the light products formed
during these steps: the (C1-C4) gases, a gasoline cut, a gas oil
cut and a kerosene cut, and also a fraction, termed the residual
fraction, which had an initial boiling point equal to 370.degree.
C. which was recycled in its entirety to the inlet to the
hydroisomerization/hydrocracking reactor in order to maximize the
production of gas oil and kerosene.
[0191] The yields are given in Table 11.
TABLE-US-00008 TABLE 11 Yield of various cuts after separation Wt %
Boiling point C.sub.1-C.sub.4 2.6 -161.degree. C. to 35.degree. C.
Naphtha 15 35.degree. C. to 150.degree. C. Kerosene 35 150.degree.
C. to 250.degree. C. Gas oil 47.4 250.degree. C. to 370.degree.
C.
[0192] The presence of carbon monoxide (CO) in the gaseous effluent
which was an inhibitor of the hydrogenating function of the
hydrocracking catalyst modified not only the activity but also the
selectivity of said catalyst for middle distillates.
[0193] It can be seen that the absence of a prior step for passage
over at least one ion exchange resin, necessitating an increase in
temperature in order to maintain the conversion, resulted in an
increase in the production of an unwanted light gas and naphtha
fraction by over-cracking.
[0194] Thus, by using an ion exchange resin upstream of the
hydrotreatment and hydroisomerization/hydrocracking steps, the
process of the invention exemplified in Example 1 can reduce the
total oxygen content of the feed and thereby limit the formation of
carbon monoxide (CO) originating from the decomposition of
oxygen-containing compounds present in the feed in the
hydroisomerization/hydrocracking section. In fact, carbon monoxide
(CO) is an inhibitor of the metallic compounds present on the
hydroisomerization/hydrocracking catalyst and its content must be
minimized in order not to require an increase in temperature in
order to compensate for the drop in activity and to maintain
conversion.
[0195] Thus, it can be seen that the process of the invention can
reduce the production of carbon monoxide (CO) (1.1% by weight) by
carrying out step a) compared with a process which is not in
accordance with the invention which does not carry out said step a)
for passage over at least one ion exchange resin, and can reduce
the temperatures employed in the hydrogenation and
hydroisomerization/hydrocracking steps in order to obtain the same
conversion of 85% of products with a boiling point of 370.degree.
C. or more into products with a boiling point of less than
370.degree. C.
* * * * *