U.S. patent application number 14/352363 was filed with the patent office on 2014-10-09 for process for conversion of petroleum feed comprising an ebullated bed hydroconversion step in a fixed bed hydrotreatment step for the production of low sulphur content fuel.
This patent application is currently assigned to IFP Energies nouvelles. The applicant listed for this patent is IFP Energies nouvelles. Invention is credited to Jerome Majcher, Wilfried Weiss.
Application Number | 20140299515 14/352363 |
Document ID | / |
Family ID | 47071326 |
Filed Date | 2014-10-09 |
United States Patent
Application |
20140299515 |
Kind Code |
A1 |
Weiss; Wilfried ; et
al. |
October 9, 2014 |
PROCESS FOR CONVERSION OF PETROLEUM FEED COMPRISING AN EBULLATED
BED HYDROCONVERSION STEP IN A FIXED BED HYDROTREATMENT STEP FOR THE
PRODUCTION OF LOW SULPHUR CONTENT FUEL
Abstract
Process for conversion of petroleum feed for production of low
sulphur content fuel comprising the following steps: a step of
ebullated bed hydroconversion of the feed in the presence of a
supported catalyst, a separation step allowing a residual fraction
to be obtained, a step of fixed bed hydrotreatment of the residual
fraction using an upstream system of permutable reactors.
Inventors: |
Weiss; Wilfried; (Valencin,
FR) ; Majcher; Jerome; (Lyon, FR) |
|
Applicant: |
Name |
City |
State |
Country |
Type |
IFP Energies nouvelles |
Rueil-Malmaison Cedex |
|
FR |
|
|
Assignee: |
IFP Energies nouvelles
Rueil-Malmaison Cedex
FR
|
Family ID: |
47071326 |
Appl. No.: |
14/352363 |
Filed: |
September 28, 2012 |
PCT Filed: |
September 28, 2012 |
PCT NO: |
PCT/FR2012/000385 |
371 Date: |
April 17, 2014 |
Current U.S.
Class: |
208/390 ; 208/44;
208/66; 208/97 |
Current CPC
Class: |
C10G 2300/205 20130101;
C10C 3/023 20130101; C10G 2300/701 20130101; C10G 2400/08 20130101;
C10G 69/08 20130101; C10G 2300/4006 20130101; C10G 65/04 20130101;
C10G 69/04 20130101; C10G 2300/1077 20130101; C10G 2300/4012
20130101; C10G 1/002 20130101; C10G 2300/107 20130101; C10G
2300/4018 20130101; C10G 65/12 20130101; C10G 2400/04 20130101;
C10G 2300/202 20130101; C10G 2300/301 20130101; C10G 2400/02
20130101 |
Class at
Publication: |
208/390 ; 208/66;
208/97; 208/44 |
International
Class: |
C10G 69/08 20060101
C10G069/08; C10C 3/02 20060101 C10C003/02; C10G 1/00 20060101
C10G001/00 |
Foreign Application Data
Date |
Code |
Application Number |
Oct 20, 2011 |
FR |
11/03.218 |
Claims
1. Method of conversion of a hydrocarbon-containing feed comprising
at least one fraction of hydrocarbons with a sulphur content of at
least 0.1% weight, an initial boiling temperature of at least
340.degree. C. and a final boiling temperature of at least
440.degree. C., comprising the following steps: a) a step of
hydroconversion in the presence of hydrogen in at least one reactor
containing a supported catalyst in an ebullated bed, b) a step of
separating the effluent obtained from step a) into at least one
light fraction of hydrocarbon fuel bases and a heavy fraction
containing predominantly compounds boiling at minimum 350.degree.
C., c) a step of fixed bed hydrotreatment of at least part of the
heavy fraction from step b) in which, under hydrotreatment
conditions, the heavy fraction and hydrogen are passed over a
hydrotreatment catalyst and in which the hydrotreatment step
comprises one or a plurality of fixed bed hydrotreatment zones
preceded by at least two storage zones also with fixed bed
hydrotreatment, arranged in series to be used in a cyclic manner
comprising the successive repetition of steps c'') and c''')
defined below: c') a step in which the storage zones are used
together for a duration at most equal to the deactivation and/or
clogging time of one of them, c'') a step during which the
deactivated and/or clogged storage zone is short-circuited and the
catalyst it contains is regenerated and/or replaced by fresh
catalyst, and during which the other storage zone(s) is/are used,
and c''') a step during which the storage zones are all used
together, wherein the storage zone of which the catalyst has been
regenerated during the preceding step is reconnected and said step
is continued for a duration equal at most to the deactivation
and/or clogging time of one of the storage zones.
2. Method according to claim 1 in which the hydrotreatment step
comprises a first hydrodemetallation step comprising one or a
plurality of fixed bed hydrodemetallation zones preceded by at
least two of said hydrotreatment storage zones and a second
subsequent hydrodesulphuration step comprising one or a plurality
of fixed bed hydrodesulphuration zones in which, during the first
hydrodemetallation step, under hydrodemetallation conditions, the
feed of hydrocarbons and hydrogen is passed over a
hydrodemetallation catalyst, then during the second subsequent
step, under hydrodesulphuration conditions, the effluent from the
first step is passed over a hydrodesulphuration catalyst.
3. Method according to claim 1, in which during step a) the
treatment in the presence of hydrogen is carried out under an
absolute pressure of 2.5 to 35 MPa, at a temperature of 330 to
550.degree. C. with an hourly spatial velocity of 0.1 to 10
h.sup.-1, and the quantity of hydrogen mixed in the feed is from 50
to 5000 Nm.sup.3/m.sup.3.
4. Method according to claim 1, in which the hydrotreatment step c)
is performed under an absolute pressure of around 2 to 35 MPa, at a
temperature of 300 to 500.degree. C. with an hourly spatial
velocity of 0.1 to 5 h.sup.-1, and a quantity of hydrogen mixed in
the feed is from 100 to 5000 Nm.sup.3/m.sup.3.
5. Method according to claim 1, in which the hydrocarbon-containing
feed is selected from the atmospheric residues, the vacuum residues
from direct distillation, crude petroleum, topped crude petroleum,
deasphalted oil, deasphalted resin, asphalts or deasphaltation
pitch, residues from the conversion processes, aromatic extracts
from the production chain of lubricant bases, bituminous sands or
their derivatives, oil shales or their derivatives, taken alone or
in combination.
6. Method according to claim 1, in which the separation step b) is
performed without decompression, the effluent from step a) is sent
to the fractioning section with a cutting point between 200 and
400.degree. C. so as to obtain a light fraction and said heavy
fraction, said heavy fraction being sent to the hydrotreatment step
while the light fraction is subjected to atmospheric distillation
to give a gaseous fraction, at least one light fraction of
hydrocarbons of the type naphtha, kerosene and/or diesel, and a
vacuum distillate fraction, the latter being at least partly sent
to the hydrotreatment step c).
7. Method according to claim 1, in which the separation step b) is
performed with decompression, the effluent from step a) is sent to
a fractioning section with a cutting point between 200 and
400.degree. C. to give a light fraction and said heavy fraction,
and in which the heavy fraction is fractioned by atmospheric
distillation into at least one atmospheric distillate fraction
containing at least one light fraction of hydrocarbons of the type
naphtha, kerosene and/or diesel, and an atmospheric residue
fraction, said atmospheric residue fraction being at least partly
fractioned by vacuum distillation into a vacuum distillate fraction
containing vacuum gas oil and a vacuum residue fraction, at least
part of said atmospheric residue fraction and/or vacuum residue
fraction being sent to the hydrotreatment step c).
8. Method according to claim 1, in which the heavy fraction
obtained in step b) is subjected, before being sent to the
hydrotreatment step, to separation of the sediment and catalyst
fines using at least one rotating filter or also at least one
basket filter or also a centrifuging system such as a hydrocyclone
associated with filters or in-line decantation, or in which the
heavy fraction obtained in step b) at the inlet to each storage
zone passes through a filtering plate situated upstream from the
catalytic bed(s) contained in the storage zone.
9. Method according to claim 1, in which at least part of the
effluent obtained in step c) is sent to a separation step, called
step d), comprising an atmospheric distillation and a vacuum
distillation and in which the effluent from the hydrotreatment step
is fractioned by atmospheric distillation into a gaseous fraction,
at least one atmospheric distillate fraction containing fuel bases
(naphtha, kerosene and/or diesel) and an atmospheric residue
fraction, at least part of the atmospheric residue is then
fractioned by vacuum distillation into a vacuum distillate fraction
containing vacuum gas oil and a vacuum residue fraction.
10. Method according to claim 9, in which the light fraction
obtained without decompression in the separation step b) is sent to
the separation step d).
11. Method according to claim 9, in which part of the vacuum
residue fraction is recycled in the hydroconversion step a).
12. Method according to claim 9, in which at least part of the
vacuum distillate fraction and/or the vacuum residue fraction
is/are sent to a catalytic cracking section, called step e), in
which it is/they are treated under conditions allowing production
of a gaseous fraction, a petrol fraction, a diesel fraction and a
residual fraction.
13. Method according to claim 12, in which at least part of the
residual fraction obtained in catalytic cracking step e) is
recycled to the inlet of step e) and/or a) and/or c).
14. Method according to claim 9, in which the atmospheric residue
and/or the vacuum distillate and/or the vacuum residue are mixed
with flux bases selected from the light cycle oils from catalytic
cracking, heavy cycle oils from catalytic cracking, the residue
from catalytic cracking, kerosene, gas oil, vacuum distillate
and/or a decanted oil.
15. Method according to claim 14, in which the flux base is
selected from kerosene, gas oil and/or vacuum distillate obtained
from separation step b) of the process after hydroconversion, or
gas oil and/or a fraction of the residual fraction obtained in the
catalytic cracking step e).
Description
[0001] The present invention concerns the refining and conversion
of heavy hydrocarbon fractions containing amongst others
sulphurated impurities. More particularly it concerns a conversion
method for heavy petroleum feeds for the production of fuel bases
of the type vacuum distillate, atmospheric residue and vacuum
residue, in particular bunker fuel bases with low sulphur content.
The method according to the invention also allows the production of
atmospheric distillates (naphtha, kerosene and diesel), vacuum
distillates and light gases (C1 to C4).
[0002] Whereas land-based industry has been subjected to stringent
regulations on the sulphur content contained in fuel bases (petrol,
diesel) in recent decades, the sulphur content in marine fuel has
so far been subjected to few constraints. In fact the fuels
presently available on the market contain up to 4.5% weight
sulphur. As a result ships have become the main source of sulphur
dioxide emissions (SO.sub.2).
[0003] In order to reduce these emissions, the International
Maritime Organisation has submitted recommendations in terms of
specifications concerning marine fuels. These recommendations are
given in the 2010 version of standard ISO 8217 (Annex VI of the
MARPOL Convention). The specification concerning sulphur henceforth
applies to SO.sub.x emissions and is reflected by a recommendation
of a sulphur content equivalent to less than or equal to 0.5%
weight. Another highly restrictive recommendation is the content of
sediment after ageing according to ISO 10307-2 which must be less
than or equal to 0.1%.
[0004] The present invention allows the production of fuel bases,
in particular bunker fuel bases, complying with the recommendations
of the MARPOL agreement.
[0005] U.S. Pat. No. 6,447,671 describes a method for converting
heavy petroleum fractions comprising a first ebullated bed
hydroconversion step, a step of eliminating catalyst particles
contained in the hydroconversion effluent, then a fixed bed
hydrotreatment step. This method is designed to give fuel bases
(petrol and diesel) having in particular a low sulphur content. The
method, however, also leads to a heavy fuel with sulphur contents
higher than the new recommendations.
[0006] The present invention adapts and improves the conversion
method described by the applicant in U.S. Pat. No. 6,447,671 for
the production of fuel bases by integrating in this method a system
of permutable reactors (or storage zones) functioning in fixed beds
in the hydrotreatment step.
[0007] More particularly the invention concerns a method converting
a hydrocarbon-containing feed comprising at least one fraction of
hydrocarbons with a sulphur content of at least 0.1% weight, an
initial boiling temperature of at least 340.degree. C. and a final
boiling temperature of at least 440.degree. C., comprising the
following steps: [0008] a) a step of hydroconversion in the
presence of hydrogen in at least one reactor containing a catalyst
supported in an ebullated bed, [0009] b) a step of separating the
effluent obtained from step a) into at least one light fraction of
hydrocarbon fuel bases and a heavy fraction containing
predominantly compounds boiling at minimum 350.degree. C., [0010]
c) a step of fixed bed hydrotreatment of at least part of the heavy
fraction from step b) in which, under hydrotreatment conditions,
the heavy fraction and hydrogen are passed over a hydrotreatment
catalyst and in which the hydrotreatment step comprises one or a
plurality of fixed bed hydrotreatment zones preceded by at least
two storage zones also with fixed bed hydrotreatment, arranged in
series to be used in a cyclic manner comprising the successive
repetition of steps c'') and c'') defined below: [0011] c') a step
in which the storage zones are used together for a duration at most
equal to the deactivation and/or clogging time of one of them,
[0012] c'') a step during which the deactivated and/or clogged
storage zone is short-circuited and the catalyst it contains is
regenerated and/or replaced by fresh catalyst, and during which the
other storage zone(s) is/are used, and [0013] c''') a step during
which the storage zones are all used together, wherein the storage
zone of which the catalyst has been regenerated during the
preceding step is reconnected and said step is continued for a
duration equal at most to the deactivation and/or clogging time of
one of the storage zones.
[0014] The system of permutable reactors is known from patents FR
2681871, FR 2784687 and EP 1343857. It is conventionally used to
extend the operating cycle of a hydrotreatment unit by allowing
replacement of a deactivated and/or clogged catalyst in the
permutable reactors without stopping the entire unit.
[0015] The present invention is targeted at the production of fuel
bases complying with the recommendations of the International
Maritime Organisation and atmospheric distillates and/or vacuum
distillates.
[0016] The hydroconversion step allows partial conversion of the
heavy feed in order to produce atmospheric distillates and/or
vacuum distillates. Although the ebullated bed technology is known
to be suitable for heavy feeds with impurities, the ebullated bed
by its nature produces catalyst fines and sediments which must be
removed before the subsequent fixed bed treatment in order to
prevent clogging of the catalytic bed(s). The fines come
principally from wear on the catalyst in the ebullated bed. The
sediment can be precipitated asphaltenes. Initially the
hydroconversion conditions and in particular the temperature to
which they are subjected causes them to undergo reactions in the
feed (dealkylation, polymerisation) leading to their
precipitation.
[0017] The permutable reactors in the method according to the
invention solve this problem of fines and sediments at the same
time as the problem of sulphur content, without pushing excessively
the operating conditions in the hydrotreatment step.
[0018] In fact the permutable reactors receiving the heavy fraction
from the hydroconversion limit the content of fines and sediment by
their filter function (fixed bed).
[0019] In order to achieve the recommendation of a sulphur content
of less than or equal to 0.5% weight, the operating conditions of
the hydrotreatment section must be severe, which is conventionally
reflected by an increase in temperature in the fixed beds. However
this increase in temperature promotes the deposit of coke,
accelerating the process of clogging and deactivation. The
permutable reactors are intended to protect the hydrotreatment
section and thus relieve the load on the main hydrotreatment
reactors sufficiently to perform the necessarily extensive
desulphuration of the heavy fraction obtained after
hydroconversion, without the operating conditions becoming
excessively harsh.
[0020] In other words, the integration of the permutable reactors
in a sequence of ebullated bed hydroconversion then fixed bed
hydrotreatment allows the obtaining of fuel bases complying with
future specifications. The fuel bases produced by the method
according to the invention have a sulphur content of less than or
equal to 0.5% weight and a sediment content after ageing of less
than or equal to 0.1%. Furthermore if the base fuel is constituted
by vacuum distillate from the process, the sulphur content can be
less than or equal to 0.1% weight.
[0021] Another important point in the process is the partial
conversion of the feed allowing production, in particular by
ebullated bed hydroconversion, of low sulphur content distillates
(naphtha, kerosene and diesel). In the context of the present
invention, the rate of conversion of the feed into lighter
fractions lies habitually between 10 and 90% and more often between
25 and 60% or even limited to around 50%. The conversion rate
mentioned above is defined as being the mass fraction of organic
compounds with a boiling point higher than 350.degree. C. at the
entry to the reaction section, minus the mass fraction of organic
compounds with a boiling point higher than 350.degree. C. at the
outlet from the reaction section, all divided by the mass fraction
of organic compounds with a boiling point higher than 350.degree.
C. at the entry to the reaction section.
[0022] In order to constitute the fuel meeting the requirements of
viscosity, the fuel bases obtained by the present method may be
mixed with flux bases so as to achieve the target viscosity of the
desired fuel grade. The viscosity of bunker fuel must be less than
or equal to 380 cSt (50.degree. C.).
[0023] In a variant of the method, the effluent obtained after
hydrotreatment may be subjected to a separation step (step d) from
which normally at least one fraction of fuel bases (naphtha,
kerosene, diesel) and heavy fractions such as vacuum distillate or
vacuum residue are recovered.
[0024] According to a variant of the method, at least part of one
of these heavy fractions can be sent to a catalytic cracking
section (step e) in which it is processed under conditions allowing
production amongst others of a light cycle oil (LCO) and a heavy
cycle oil (HCO). These oils can be used as flux for mixing the fuel
bases resulting from the method according to the invention to
constitute the fuel, so as to achieve the target viscosity.
DETAILED DESCRIPTION
Feed
[0025] The feeds processed in the method according to the invention
are advantageously selected from atmospheric residues, vacuum
residues from direct distillation, crude petroleum, topped crude
petroleum, deasphalted oils, resin from deasphaltation, asphalt or
deasphaltation pitch, residues from conversion processes, aromatic
extracts from production chains of lubricant bases, bituminous
sands or their derivatives, oil shales or their derivatives, taken
alone or in combination. These feeds can advantageously be used as
such or diluted with a hydrocarbonated fraction or a mixture of
hydrocarbonated fractions which can be selected from products
resulting from fluid catalytic cracking (FCC), a light cycle oil
(LCO), a heavy cycle oil (HCO), a decanted oil (DO), an FCC
residue, or can come from distillation, the gas oil fractions in
particular being those obtained by atmospheric or vacuum
distillation such as for example vacuum gas oil. The heavy feeds
can also advantageously comprise cuts from the coal or biomass
liquefaction process, aromatic extracts or any other
hydrocarbonated cuts or even non-petroleum feeds such as pyrolysis
oil.
[0026] Said heavy feeds generally have a sulphur content of at
least 0.1% weight, an initial boiling temperature of at least
340.degree. C. and a final boiling temperature of at least
440.degree. C.
[0027] According to the present invention the feeds treated are
preferably atmospheric residues or vacuum residues or mixtures of
these residues.
Ebullated Bed Hydroconversion (Step a)
[0028] According to the present invention the feed is subjected to
a hydroconversion step performed in at least one reactor containing
a supported catalyst in an ebullated bed functioning with liquid
and gas upflow. The main objective of hydroconversion is to convert
the heavy fraction into lighter cuts while partially refining the
feed.
[0029] Ebullated bed technology is widely known so only the main
operating conditions are described here. Ebullated bed technologies
use supported catalysts in the form of extrusions, the diameter of
which is generally of the order of 1 mm or less than 1 mm. The
catalysts remain inside the reactors and are not evacuated with the
products. The temperature levels are high in order to obtain high
conversion while minimising the quantities of catalysts used. The
catalytic activity can be held constant thanks to in-line
replacement of the catalyst. It is therefore not necessary to stop
the unit in order to replace exhausted catalyst nor to increase the
reaction temperatures throughout the cycle in order to compensate
for deactivation. Furthermore, by working under constant operating
conditions, constant yields and product quality are obtained
throughout the cycle. Thus because the catalyst is held in
agitation by significant liquid recycling, the load loss over the
reactor remains low and constant.
[0030] The conditions of the feed processing in step a) in the
presence of hydrogen are normally conventional conditions of
ebullated bed hydroconversion of a liquid hydrocarbonated fraction.
This is operated normally at an absolute pressure of 2.5 to 35 MPa,
often 5 to 25 MPa, and most often 6 to 20 MPa, at a temperature of
330 to 550.degree. C. and often 350 to 500.degree. C. The hourly
spatial velocity (VVH) and the hydrogen partial pressure are
significant factors which are selected as a function of
characteristics of the product to be processed and the desired
conversion. The VVH is defined as the volumetric flow of the feed
divided by the total volume of the reactor, and generally lies in a
range from 0.1 h.sup.-1 to 10 h.sup.-1 and preferably 0.2 h.sup.-1
to 5 h.sup.-1. The quantity of hydrogen mixed into the feed is
normally 50 to 5000 standard cubic metres (Nm.sup.3) per cubic
metre (m.sup.3) of liquid feed and most often 100 to 1000
Nm.sup.3/m.sup.3 and preferably 200 to 500 Nm.sup.3/m.sup.3.
[0031] A conventional granular hydroconversion catalyst can be used
comprising, on an amorphous support, at least one metal or metal
compound with a hydro-dehydrogenating function. This catalyst may
be a catalyst comprising metals of group VIII, for example nickel
and/or cobalt, most often in association with at least one metal of
group VIB, for example molybdenum and/or tungsten. For example a
catalyst can be used comprising 0.5 to 10% weight nickel and
preferably 1 to 5% weight nickel (expressed as nickel oxide NiO)
and 1 to 30% weight molybdenum, preferably 5 to 20% weight
molybdenum (expressed as molybdenum oxide MoO.sub.3) on an
amorphous mineral support. This support is selected for example
from the group formed from alumina, silica, silica-aluminas,
magnesia, clays and mixtures of at least two of these minerals. The
support can also contain other compounds and for example oxides
selected from the group formed by boron oxides, zircon, titanium
oxide, phosphoric anhydride. Usually an alumina support is used,
and very often a support of alumina doped with phosphorus and where
applicable boron. When phosphoric anhydride P.sub.2O.sub.5 is
present, its concentration is usually less than 20% by weight and
most often less than 10% by weight. The concentration of boron
trioxide B.sub.2O.sub.3 is normally 0 to 10% by weight. The alumina
used is normally a gamma or eta alumina. This catalyst is most
often present in the form of extrusions. The total content of
oxides of metals of groups VI and VIII is often 5 to 40% by weight
and generally 7 to 30% by weight, and the weight ratio expressed in
metal oxide between the metal (or metals) of group VI to the metal
(or metals) of group VIII is generally 20 to 1 and most often 10 to
2.
[0032] The exhausted catalyst is partly replaced by fresh catalyst,
generally by extraction from the bottom of the reactor and
introduction of fresh or new catalyst at the top of the reactor at
regular intervals i.e. for example in batches or
quasi-continuously. The catalyst can also be introduced at the
bottom and extracted at the top of the reactor. For example fresh
catalyst can be introduced every day. The rate of replacement of
exhausted catalyst with fresh catalyst may for example be around
0.05 kilograms to around 10 kilograms per cubic metre of feed. This
extraction and replacement are performed using devices allowing
continuous function of this hydroconversion step. The unit normally
comprises a recirculation pump allowing the catalyst to be
maintained in the ebullated bed by continuous recycling of at least
part of the liquid extracted at the head of the reactor and
re-injected at the base of the reactor. It is also possible to send
the exhausted catalyst extracted from the reactor to a regeneration
zone in which the carbon and sulphur it contains are eliminated,
before it is re-injected into hydroconversion step a).
[0033] Most often this hydroconversion step a) is implemented under
the conditions of the H-OIL.RTM. process as described for example
in U.S. Pat. No. 6,270,654.
[0034] The ebullated bed hydroconversion can be carried out in a
single reactor or in several reactors (generally two) arranged in
series. By using at least two ebullated bed reactors in series,
products of better quality can be obtained with a better yield,
thus limiting the requirements for energy and hydrogen in any
post-treatment. Also hydroconversion in two reactors allows
improved operability in relation to the flexibility of the
operating conditions and the catalytic system. Generally the
temperature of the second reactor is preferably at least 10.degree.
C. higher than that of the first ebullated bed reactor. The
pressure of the second reactor is 0.1 to 1 MPa lower than for the
first reactor to allow the flow of at least part of the effluent
from the first step without pumping being necessary. The different
operating conditions in terms of temperature in the two
hydroconversion reactors are selected in order to be able to
control the hydrogenation and conversion of the feed into the
desired products in each reactor.
[0035] Where applicable the effluent obtained at the end of the
first hydroconversion step is subjected to separation of the light
fraction, and at least part, preferably all, of the residual
effluent is treated in the second hydroconversion step. This
separation is advantageously performed in an inter-stage separator
described in U.S. Pat. No. 6,270,654 and in particular allows
avoidance of over-cracking of the light fraction in the second
hydroconversion reactor.
[0036] It is also possible to transfer all or part of the exhausted
catalyst extracted from the reactor in the first hydroconversion
step, operating at lower temperature, directly to the reactor of
the second step, operating at higher temperature, or to transfer
all or part of the exhausted catalyst extracted from the reactor in
the second step directly to the reactor of the first step. This
cascade system is described in U.S. Pat. No. 4,816,841.
Separation of the Hydroconversion Effluent (Step b)
[0037] The effluent obtained after the ebullated bed
hydroconversion undergoes at least one separation step, where
applicable supplemented by other further separation steps allowing
separation of at least one light hydrocarbon fraction containing
fuel bases and a heavy fraction containing predominantly compounds
boiling at minimum 350.degree. C. The separation step can
advantageously be implemented by any method known to the person
skilled in the art, such as for example the combination of one or a
plurality of high- and/or low-pressure separators and/or
distillation steps and/or high- and/or low-pressure stripping,
and/or liquid/liquid extraction steps and/or solid/liquid
separation steps and/or centrifuging steps. Preferably the
separating step b) leads to the acquisition of a gaseous phase, at
least one light hydrocarbon fraction of the type naphtha, kerosene
and/or diesel, a vacuum distillate fraction and a vacuum residue
fraction and/or an atmospheric residue fraction.
[0038] According to a first embodiment the effluent from the
hydroconversion undergoes a separation step with decompression
between the ebullated bed hydroconversion and the fixed bed
hydrotreatment. This configuration can be described as a
non-integrated system.
[0039] According to the non-integrated system, separation is
preferably performed in a fractioning section which can firstly
comprise a high-pressure, high-temperature separator (HPHT) and
where applicable a high-pressure, low-temperature separator (HPLT)
and/or an atmospheric distillation and/or a vacuum distillation.
The effluent obtained from step a) is separated (generally in an
HPHT separator) into a light fraction and a heavy fraction
containing predominantly compounds boiling at minimum 350.degree.
C. Separation is not made following a precise cutting point, it
rather resembles a flash. If we have to define a cutting point, we
could say that this lies between 200 and 400.degree. C.
[0040] Preferably the so-called heavy fraction is then fractioned
by atmospheric distillation into at least one atmospheric
distillate fraction containing at least one light fraction of
hydrocarbons of the type naphtha, kerosene and/or diesel, and an
atmospheric residue fraction. At least part of the atmospheric
residue fraction can be sent for hydrotreatment. The atmospheric
residue can also at least be partly fractioned by vacuum
distillation into a vacuum distillate fraction containing vacuum
gas oil and a vacuum residue fraction. Said vacuum residue fraction
is advantageously sent at least partly to the hydrotreatment step
c). Part of the vacuum residue can also be recycled in
hydroconversion step a).
[0041] According to a second embodiment the effluent from the
ebullated bed hydroconversion undergoes a separation step without
decompression between hydroconversion and hydrotreatment. This
configuration can be called an integrated system.
[0042] According to the integrated system, the effluent obtained
after step a) is separated (generally in an HPHT separator) into a
light fraction and a heavy fraction containing predominantly
compounds boiling at minimum 350.degree. C. Separation is not made
according to a precise cutting point but is similar to a flash. If
we have to define a cutting point, we could say that this lies
between 200 and 400.degree. C.
[0043] Preferably the heavy fraction is then sent to the
hydrotreatment step. The light fraction can undergo further
separation steps, where applicable in the presence of the light
fraction from the inter-stage separator between the two
hydroconversion reactors. Advantageously it is subjected to
atmospheric distillation giving a gaseous fraction, at least one
light fraction of hydrocarbons of the type naphtha, kerosene and/or
diesel, and a vacuum distillate fraction, wherein the latter can be
sent at least partly to the hydrotreatment step c). Another part of
the vacuum distillate can constitute part of a fuel as a flux
agent. Another part of the vacuum distillate can be refined by
hydrocracking and/or by fluid catalytic cracking.
[0044] Separation according to the integrated system allows better
thermal integration and is reflected by a saving in energy and
equipment. Also this system has technical and economic advantages
given that the high-pressure flux will not require an increase in
pressure in view of the subsequent hydrotreatment. This embodiment
therefore, by its simplified intermediate fractioning, allows a
reduction in the consumption of utilities and hence the investment
costs.
[0045] The gaseous fractions from the separation step (either the
integrated or non-integrated system) preferably undergo
purification treatment to recover the hydrogen and recycle it
towards hydroconversion and/or hydrotreatment reactors. The same
applies to gaseous effluent from hydrotreatment. The gas phase from
the optional inter-stage separator can also be added.
[0046] The separation step in the non-integrated or integrated
system allows the arrangement of two independent hydrogen circuits,
one linked to hydroconversion and the other to hydrotreatment, and
which depending on need can be linked together. The hydrogen can be
added to one or the other or to both reaction sections, the
recycling gas can supply one or the other or both reaction
sections. The compressor may where applicable be common to both
sections. The possibility of linking the two hydrogen circuits
allows optimisation of hydrogen management and limits the
investment in terms of compressors and/or purifications units for
the gaseous effluent. The various embodiments of hydrogen
management which can be used in the present invention are described
in application FR 2957607.
[0047] The processing of the various fuel base cuts (LPG, naphtha,
kerosene, diesel and/or vacuum gas oil) obtained in the
non-integrated or integrated system is not the subject of the
present invention and these methods are well known to the person
skilled in the art. The products obtained can be integrated in fuel
reservoirs (also called fuel pools) or undergo additional refining
steps. The naphtha, kerosene, gas oil and vacuum gas oil fractions
may be subjected to one or a plurality of treatments
(hydrotreatment, hydrocracking, alkylation, isomerisation,
catalytic reforming, catalytic cracking or thermal cracking or
other) to bring these to the required specifications (sulphur
content, smoke point, octane, cetane etc.) either separately or
combined.
[0048] Advantageously the vacuum distillate from the ebullated bed,
after separation, can undergo hydrotreatment. This hydrotreated
vacuum distillate can be used as a flux for the fuel pool having a
sulphur content of less than or equal to 0.5% weight or be used
directly as fuel with a sulphur content of less than or equal to
0.1% weight. Part of the atmospheric residue, the vacuum distillate
and/or vacuum residue can undergo other additional refining steps
such as hydrocracking or fluid catalytic cracking.
Hydrotreatment (Step c)
[0049] The heavy fraction containing compounds boiling at minimum
350.degree. C. from separation step b) is then sent to the
hydrotreatment step comprising one or a plurality of fixed bed
hydrotreatment zones preceded by at least two storage zones also
with fixed bed hydrotreatment.
[0050] Hydrotreatment (HDT) in particular means reactions of
hydrodesulphuration (HDS), hydrodenitration (HDN) and
hydrodemetallation (HDM), but also hydrogenation,
hydrodeoxygenation, hydrodearomatisation, hydroisomerisation,
hydrodealkylation, hydrocracking, hydrodeasphaltation and Conradson
carbon reduction.
[0051] According to a preferred variant, the hydrotreatment step
comprises a first hydrodemetallation step comprising one or a
plurality of fixed bed hydrodemetallation zones preceded by at
least two of said hydrotreatment storage zones, and a second
subsequent hydrodesulphuration step comprising one or a plurality
of fixed bed hydrodesulphuration zones and in which, during the
first said hydrodemetallation step, under hydrodemetallation
conditions, the feed of hydrocarbons and hydrogen is passed over a
hydrodemetallation catalyst, then during the second subsequent
step, under hydrodesulphuration conditions, the effluent from the
first step is passed over a hydrodesulphuration catalyst. This
process known under the name HYVAHL-F.TM. is described in U.S. Pat.
No. 5,417,846.
[0052] The person skilled in the art will easily understand that in
the hydrodemetallation step, predominantly hydrodemetallation
reactions are performed but in parallel also part of the
hydrodesulphuration reactions. Similarly in the hydrodesulphuration
step predominantly hydrodesulphuration reactions are performed but
in parallel also part of the hydrodemetallation reactions.
[0053] The catalytic beds in the fixed bed reactors, in particular
the upper parts of the catalytic beds and more particularly the
upper parts of the first catalytic bed in contact with the feed,
are susceptible to clogging fairly quickly due to the asphaltenes
and sediments contained in the feed, which is reflected initially
by an increase in the load loss and sooner or later requires a
stoppage of the hydrotreatment unit for replacement of the catalyst
in the main hydrotreatment reactors.
[0054] The function of the permutable reactors is to filter the
sediments and the catalyst fines from the hydroconversion step,
which prevents clogging in the main reactors.
[0055] Another problem traditionally found when using fixed beds is
deactivation of the catalyst due to the high deposits of metals
occurring during the hydrotreatment reactions. To compensate for
this deactivation, the reactor temperature is then increased.
Nonetheless this increase in temperature promotes coke deposits.
The deposition of impurities is supplemented by that of coke, the
whole then tending rapidly to deactivate and clog the catalytic
system. These phenomena lead to stoppages of the hydrotreatment
units for replacement of catalysts and to overconsumption of
catalysts, which the person skilled in the art wishes to
minimise.
[0056] The function of the permutable reactors is to protect the
main hydrotreatment reactors downstream by preventing clogging
and/or deactivation. The permutable reactors are thus used to
extend the operating cycle of the hydrotreatment unit by allowing
replacement of deactivated and/or clogged catalyst in the
permutable reactors working in a cyclic fashion, without stopping
the entire unit for a specific period. According to the method of
the invention, extended hydrotreatment of the heavy fraction thus
becomes possible in order in particular to obtain fuel bases with
known sulphur content.
[0057] The function of the storage zones in the HDM section is
described in FIG. 1 comprising two storage zones (or permutable
reactors) Ra and Rb. This hydrotreatment process consists of a
series of cycles each comprising four successive steps: [0058] a
first step (step c') during which the feed passes successively
through reactor Ra then reactor Rb, [0059] a second step (step c'')
during which the feed passes only through reactor Rb, reactor Ra
being short-circuited for regeneration and/or replacement of the
catalyst, [0060] a third step (step c''') during which the feed
passes successively through reactor Rb then reactor Ra, [0061] a
fourth step (step c'''') during which the feed passes only through
reactor Ra, reactor Rb being short-circuited for regeneration
and/or replacement of the catalyst.
[0062] Preferably after regeneration and/or replacement of the
catalyst of one reactor, this reactor is reconnected downstream of
the running reactor.
[0063] During step c') of the method, the effluent from the step of
ebullated bed hydroconversion is introduced via the conduit (96)
and the conduit (98) which comprises an open valve V1 towards the
conduit (100), and the storage reactor Ra containing a fixed bed A
of catalyst. During this period the valves V3, V4 and V5 are
closed. The effluent from the reactor Ra is sent via the conduit
(102), the conduit (104) which comprises an open valve V2 and
conduit (106) into the storage reactor Rb containing a fixed bed B
of catalyst. The effluent from the reactor Rb is sent via the
conduits (108) and (110) comprising an open valve V6 and the
conduit (112) to the main hydrotreatment section which will be
described below.
[0064] During step c''') of the method, the valves V1, V2, V4 and
V5 are closed and the feed is introduced via the conduit (96) and
the conduit (114) comprising an open valve V3 towards the conduit
(106) and the reactor Rb. During this period the effluent from the
reactor Rb is sent via the conduits (108) and (110) comprising an
open valve V6 and conduit (112) to the main hydrotreatment
section.
[0065] During step c''') the valves V1, V2 and V6 are closed and
the valves V3, V4 and V5 are open. The feed is introduced via the
conduit (96) and the conduits (114) and (116) to the reactor Rb.
The effluent from reactor Rb is sent via the conduit (108), the
conduit (116) containing an open valve V4 and the conduit (100) to
the storage reactor Ra. The effluent from reactor Ra is sent via
the conduits (102) and (118) comprising an open valve V5 and the
conduit (112) to the main hydrotreatment section.
[0066] During step c''''), the valves V2, V3, V4 and V6 are closed
and the valve V1 and V5 are open. The feed is introduced via the
conduit (96) and the conduits (98) and (100) to the reactor Ra.
During this period the effluent from the reactor Ra is sent via the
conduits (102) and (118) comprising an open valve V5 and the
conduit (112) to the main hydrotreatment section. The cycle then
begins again.
[0067] The deactivation and/or clogging time varies depending on
the feed, the operating conditions of the hydroconversion step, the
operating conditions of the hydrotreatment step and the catalyst(s)
used. It is generally expressed in a fall in catalytic performance
(an increase in concentration of metals and/or other impurities in
the effluent), an increase in the temperature necessary to maintain
a catalytic activity or--in the specific case of clogging--via a
significant increase in the load loss. The load loss .DELTA.p
expressing a degree of clogging is measured permanently throughout
the cycle over each of the zones and can be defined by a pressure
increase resulting from a partially blocked passage of the outflow
through the zone. Similarly the temperature is measured permanently
throughout the cycle on each of the two zones.
[0068] In order to define a deactivation and/or clogging time, the
person skilled in the art will first define a maximum permitted
value for the load loss .DELTA.p and/or temperature as a function
of the feed to be treated, the selected operating conditions and
catalysts, and above which the storage zone must be disconnected.
The deactivation and/or clogging time is thus defined as the time
when the limit value for load loss and/or temperature is reached.
In the case of a heavy fraction hydrotreatment process, the limit
value for load loss is generally between 0.3 and 1 MPa (3 and 10
bar), preferably between 0.5 and 0.8 MPa (5 and 8 bar). The limit
value for temperature is generally between 400 and 430.degree. C.,
where the temperature here and throughout the text corresponds to
the mean temperature measured in the catalytic bed.
[0069] In a preferred embodiment a catalyst conditioning section is
used allowing switching of these storage zones during operation
i.e. without stopping the operation of the unit. Firstly a system
which operates at a moderate pressure (10 to 50 bar but preferably
15 to 25 bar) allows the performance of the following operations on
the disconnected storage reactor: washing, stripping, cooling,
before discharge of worn catalyst; then heating and sulphuration
after loading of fresh catalyst; then another system of
pressurisation/depressurisation and slide valves of appropriate
technology effectively allows switching of these storage zones
without stopping the unit, i.e. without affecting the operating
factor, since all operations of washing, stripping, discharge of
worn catalyst, reload of fresh catalyst, heating and sulphuration
are performed on the disconnected reactor or storage zone.
Alternatively a pre-sulphurated catalyst can be used in the
conditioning section so as to simplify the process of switching
during operation.
[0070] The effluent from the permutable reactors is then sent into
the main hydrotreatment reactors.
[0071] Each hydrotreatment zone or hydrotreatment storage zone
contains at least one catalytic bed (for example 1, 2, 3, 4 or 5
catalytic beds). Preferably each storage zone contains one
catalytic bed. Each catalytic bed contains at least one catalytic
layer containing one or a plurality of catalysts, where applicable
preceded by at least one inert layer, for example alumina or
ceramic in the form of extrusions, balls or discs. The catalyst
used in the catalytic bed(s) can be identical or different.
[0072] In another variant, the heavy fraction obtained in step b)
can be subjected, before being sent to the hydrotreatment step, to
separation of the sediments and catalyst fines using at least one
rotating filter or at least one basket filter, or a centrifuging
system such as a hydrocyclone associated with filters or in-line
decantation.
[0073] In another variant the heavy fraction obtained in step b),
at the inlet to each storage zone, passes through a filtering plate
located upstream of the catalytic bed(s) contained in the storage
zone. This filtering plate described in patent FR 2889973 can trap
the clogging particles contained in the feed by means of a specific
distributor plate comprising a filtering medium.
[0074] According to a variant, a co-feed can be introduced with the
residual fraction in the hydrotreatment step. This co-feed can be
selected from atmospheric residues, vacuum residues from direct
distillation, deasphalted oils, aromatic extracts from the
production chains of lubricant bases, hydrocarbonated fractions or
a mixture of hydrocarbonated fractions which can be selected from
the products resulting from a fluid catalytic cracking process: a
light cycle oil (LCO), a heavy cycle oil (HCO), a decanted oil, or
which can come from distillation, the gas oil fractions in
particular being those obtained by atmospheric or vacuum
distillation, such as for example vacuum gas oil.
[0075] The hydrotreatment step can advantageously be implemented at
a temperature between 300 and 500.degree. C., preferably 350 to
420.degree. C. and at an absolute pressure advantageously between 2
MPa and 35 MPa, preferably between 10 and 20 MPa. The temperature
is normally adjusted as a function of the desired hydrotreatment
level. Usually the global VVH lies in a range from 0.1 h.sup.-1 to
5 h.sup.-1 and preferably 0.1 h.sup.-1 to 2 h.sup.-1. The quantity
of hydrogen mixed with the feed is normally 100 to 5000 standard
cubic metres (Nm.sup.3) per cubic metre (m.sup.3) of liquid feed
and most often 200 to 2000 Nm.sup.3/m.sup.3 and preferably 300 to
1500 Nm.sup.3/m.sup.3. Normally this hydrotreatment step is
performed industrially in one or a plurality of liquid downflow
reactors.
[0076] The operating conditions of the permutable reactors are
generally identical to those of the main hydrotreatment reactors.
The VVH value of each permutable reactor in operation is preferably
0.25 to 4 h.sup.-1 and more often 1 to 2 h.sup.-1. The global VVH
value of the permutable reactors and that of each reactor is
selected to achieve the maximum HDM while controlling the reaction
temperature (limitation of exothermicity).
[0077] The hydrotreatment catalysts used are preferably known
catalysts and are generally granular catalysts comprising on a
support at least one metal or metal compound with a
hydro-dehydrogenating function. These catalysts are advantageously
catalysts comprising at least one metal from group VIII, selected
generally from the group formed by nickel and/or cobalt, and/or at
least one metal from group VIB, preferably molybdenum and/or
tungsten. For example a catalyst is used comprising 0.5 to 10%
weight nickel and preferably 1 to 5% weight nickel (expressed in
nickel oxide NiO) and 1 to 30% of weight molybdenum, preferably 5
to 20% of weight molybdenum (expressed as molybdenum oxide
MoO.sub.3) on a mineral support. This support is for example
selected from the group formed by alumina, silica, silica-aluminas,
magnesia, clays and mixtures of at least two of these minerals.
Advantageously the support contains other doping compounds, in
particular oxides selected from the group formed by boron oxide,
zircon, cerine, titanium oxide, phosphoric anhydride and a mixture
of these oxides. Most often an alumina support is used, and very
often an alumina support doped with phosphorus and where applicable
boron. The concentration of phosphoric anhydride P.sub.2O.sub.5 is
normally between 0 or 0.1% and 10% by weight. The concentration of
boron trioxide B.sub.2O.sub.5 is conventionally between 0 or 0.1%
and 10% by weight. The alumina used is normally a .gamma. or .eta.
alumina. The catalyst is most often in the form of extrusions. The
total content of metal oxides of the groups VIB and VIII is often 5
to 40% by weight and generally 7 to 30% by weight, and the weight
ratio expressed in metal oxide between the metal (or metals) of
group VIB to the metal (or metals) of group VIII is generally 20 to
1 and most often 10 to 2.
[0078] In the case of a hydrotreatment step including an HDM step
then an HDS step, most often specific catalysts are used adapted to
each step.
[0079] Catalysts which can be used in the HDM step are for example
indicated in patents EP 113297, EP 113284, U.S. Pat. No. 5,221,656,
U.S. Pat. No. 5,827,241, U.S. Pat. No. 7,119,045, U.S. Pat. No.
5,622,616 and U.S. Pat. No. 5,089,463. Preferably HDM catalysts are
used in the permutable reactors.
[0080] The catalysts which can be used in the HDS step are for
example indicated in patents EP 113297, EP 113284, U.S. Pat. No.
6,589,908, U.S. Pat. No. 4,818,743 or U.S. Pat. No. 6,332,976.
[0081] Also a mixed catalyst can be used which is active in HDM and
HDS both for the HDM section and for the HDS section as described
in patent FR 2940143.
[0082] Before injection of the feed, the catalysts used in the
method according to the present invention are preferably subjected
to sulphuration treatment (in situ or ex situ).
Separation of the Hydrotreatment Effluent (Step d)
[0083] In a current form of implementation of the invention, the
effluent obtained in step c) is at least partly and often fully
sent to a separation step d) comprising an atmospheric distillation
and a vacuum distillation. The effluent from the hydrotreatment
step is fractioned by atmospheric distillation into a gaseous
fraction, at least one atmospheric distillate fraction containing
fuel bases (naphtha, kerosene and/or diesel) and an atmospheric
residue fraction. At least part of the atmospheric residue can then
be fractioned by vacuum distillation into a vacuum distillate
fraction containing vacuum gas oil and a vacuum residue
fraction.
[0084] The vacuum residue fraction and/or vacuum distillate
fraction and/or atmospheric residue fraction can constitute at
least partly the basis for fuel with low sulphur content, with a
sulphur content of less than or equal to 0.5% weight and a sediment
content after ageing of less than or equal to 0.1%. The vacuum
distillate fraction can constitute a fuel base with a sulphur
content of less than or equal to 0.1% weight. Part of the vacuum
residue can also be recycled into hydroconversion step a).
[0085] The separation stage of step b) in the integrated system,
according to which the heavy fraction is recovered for
hydrotreatment and a light fraction, may be separate from the
separation step d), but preferably the light fraction obtained
without decompression in separation step b) is sent to separation
step d).
[0086] Although the fuel bases produced according to the present
method contain very little sediment, a supplementary filtration
step may be provided.
Catalytic Cracking (Step e)
[0087] According to a variant at least part of the vacuum
distillate fraction and/or vacuum residue fraction resulting from
step d) is sent to a catalytic cracking section known as step e) in
which it is treated under conditions allowing production of a
gaseous fraction, a petrol fraction, a diesel fraction and a
residual fraction.
[0088] In one embodiment of the invention at least part of the
residual fraction obtained in catalytic cracking step e), often
called the slurry fraction by the person skilled in the art, is
recycled to the inlet of step e) and/or a) and/or c). The residual
fraction can thus be sent at least partly or fully to the heavy
fuel storage zone of the refinery.
[0089] In a particular embodiment of the invention, part of the gas
oil fraction (or LCO) and/or part of the residual fraction
(containing the HCO) obtained during step e) can be used to
constitute flux bases which will be mixed with the fuel bases
obtained by the present method.
[0090] The catalytic cracking step e) is most often a fluid
catalytic cracking step, for example using the method developed by
the applicant known as R2R. This step can be performed in a
conventional manner known to the person skilled in the art under
adequate cracking conditions with a view to obtaining
hydrocarbon-containing products of lower molecular weight. This
step can use processes and devices for thermal exchange, in
particular of solid particles, with a view to reducing the catalyst
temperature at the entry to the reaction zone. Descriptions of the
function and catalysts which can be used in the context of fluid
cracking in this step e) are given for example in the documents of
patents U.S. Pat. No. 4,695,370, EP-B-184517, U.S. Pat. No.
4,959,334, EP-B-323297, U.S. Pat. No. 4,965,232, U.S. Pat. No.
5,120,691, U.S. Pat. No. 5,344,554, U.S. Pat. No. 5,449,496,
EP-A-485259, U.S. Pat. No. 5,286,690, U.S. Pat. No. 5,324,696,
EP-B-542604 and EP-A-699224, the descriptions of which are regarded
as incorporated in the present description by this reference.
[0091] The fluid catalytic cracking reactor can function with
upflow or downflow. Although not a preferred form of embodiment of
the present invention, it is also possible to perform catalytic
cracking in a mobile bed reactor. Particularly preferred catalysts
for catalytic cracking are those which contain at least one zeolite
normally mixed with an appropriate matrix such as for example
alumina, silica or silica-alumina.
Fluxing
[0092] In order to constitute a fuel meeting the viscosity
recommendations which must be less than or equal to 380 cSt
(50.degree. C.), the fuel bases obtained by the present method can
be mixed if necessary with flux bases to achieve the target
viscosity of the desired fuel grade.
[0093] The flux bases can be selected from light cycle oils (LCO)
from catalytic cracking, heavy cycle oils (HCO) from catalytic
cracking, residue from catalytic cracking, kerosene, gas oil,
vacuum distillate and/or decanted oil (DO).
[0094] Preferably kerosene, gas oil and/or vacuum distillate
obtained from the separation step b) of the method after
hydroconversion, or gas oil and/or a fraction of the residual
fraction obtained in catalytic cracking step e) is used.
DESCRIPTION OF THE FIGURES
[0095] The following figures show advantageous embodiments of the
invention. The installation and the method according to the
invention are described in essence. The operating conditions
described above are not repeated.
[0096] FIG. 2 shows the process according to the invention with
intermediate separation with decompression (non-integrated
system).
[0097] In FIG. 2 the feed (10) preheated in the enclosure (12) and
mixed with the recycled hydrogen (14) and heated in the enclosure
(16) is introduced via the conduit (18) at the bottom of the first
ebullated bed reactor (20) working with liquid and gas upflow and
containing at least one hydroconversion catalyst. The reactor (20)
normally comprises a recirculation pump (22) allowing the catalyst
to be maintained in the ebullated bed by continuous recycling of at
least part of the liquid extracted from the upper part of the
reactor and re-injected at the base of the reactor. The hydrogen
can also be introduced into the oven (12), thus eliminating the
enclosure (16). The hydrogen is supplied via the hydrogen recycled
from the process (64), supplemented with added hydrogen (24). Fresh
catalyst can be added at the top of the reactor (not shown). The
catalyst can be added periodically or continuously. The exhausted
catalyst can be extracted from the base of the reactor (not shown)
to be either eliminated or regenerated to eliminate the carbon and
sulphur before being re-injected at the top of the reactor. The
partly exhausted catalyst extracted at the bottom of the first
reactor can also be transferred directly to the top of the second
hydroconversion reactor (40) (cascading) (not shown). Where
applicable the converted effluent (26) from the reactor (20) can
also be subjected to separation of the light fraction (28) in an
inter-stage separator (30).
[0098] All or part of the effluent (32) from the first
hydroconversion reactor (20) is advantageously mixed with
additional hydrogen (34), if required first preheated (not shown).
This mixture is then injected via the conduit (36) into a second
ebullated bed hydroconversion reactor (40) functioning with liquid
and gas upflow and containing at least one hydroconversion
catalyst. The operational conditions, in particular temperature, in
this reactor are selected to achieve the desired conversion level
as already described. The addition and extraction of catalyst is
performed in the same way as described for the first reactor. The
reactor (40) normally comprises a recirculation pump (38) allowing
maintenance of the ebullated bed catalyst by continuous recycling
of at least part of the liquid extracted from the upper part of the
reactor and re-injected at the base of the reactor. The effluent
treated in the reactor (40) is sent via the conduit (42) to a
high-pressure, high-temperature separator (HPHT) (44) from which a
gaseous fraction (46) and a liquid fraction (48) are recovered. The
cutting point is generally between 200 and 400.degree. C. The
gaseous fraction (46) is sent, optionally mixed with the light
fraction (28) from the optional inter-stage separator (30) between
the two hydroconversion reactors, generally via an exchanger (not
shown) or a cooling tower (50) for cooling, to a high-pressure,
low-temperature separator (HPLT) (52) from which a gaseous fraction
(54) is recovered containing the gases (H.sub.2, H.sub.2S,
NH.sub.3, C1-C4 hydrocarbons etc.) and a liquid fraction (56).
[0099] The gaseous fraction (54) from the high-pressure,
low-temperature separator (HPLT) (52) is treated in the hydrogen
purification unit (58) from which the hydrogen (60) is recovered to
be recycled via the compressor (62) and the conduit (64) to the
reactors (20) and/or (40). The gases containing the undesirable
nitrated and sulphurated compounds are evacuated from the
installation (flux (66)).
[0100] The liquid fraction (56) from the high-pressure,
low-temperature separator (HPLT) (52) is decompressed in the device
(68) then sent to the fractioning system (70). Optionally a
medium-pressure separator (not shown) may be installed after the
decompressor (68) to recover a gaseous fraction which is sent to
the purification unit (58), and a liquid phase which is brought to
the fractioning section (70).
[0101] The liquid fraction (48) from the high-pressure,
high-temperature separator (HPHT) (44) is decompressed in the
device (72) then sent to the fractioning system (70). Evidently the
fractions (56) and (48) can be sent together, after decompression,
to the system (70). Depending on a variant not shown, the fractions
(70) and (172) can be common and treat all the light fractions
including that from the inter-stage separator. The fractioning
system (70) comprises an atmospheric distillation system to produce
a gaseous effluent (74), at least one so-called light fraction (76)
containing in particular naphtha, kerosene and diesel, and an
atmospheric residue fraction (78). Part of the atmospheric residue
fraction can be sent via the conduit (80) to the hydrotreatment
reactors. All or part of the atmospheric residue fraction (78) is
sent to a vacuum distillation column (82) to recover a fraction
(84) containing the vacuum residue containing the fines and
sediments, and a vacuum distillates phase (86) containing vacuum
gas oil.
[0102] The vacuum residue fraction (84), where applicable mixed
with part of the atmospheric residue fraction (80), is mixed with
the recycled hydrogen (88), possibly supplemented by added hydrogen
(90). It passes optionally through an oven (92) and a filter (94).
Optionally a co-feed (198) may be introduced. The effluent thus
heated and filtered is introduced via the conduit (96) at the top
of the permutable reactor system of the hydrotreatment step. The
operation of the permutable reactors has been described with
reference to FIG. 1.
[0103] The effluent leaving the storage reactors is optionally
re-mixed with the hydrogen (91) arriving through the conduit (112)
in a main HDM reactor (120) which contains a fixed bed catalyst
(122). For reasons of legibility, a single HDM and HDS reactor is
shown on the figure but the HDM and HDS section conventionally
comprises a plurality of HDS and HDM reactors in series. If
necessary the recycled and/or added hydrogen can also be introduced
into the hydrotreatment reactors between the various catalytic beds
(quench) (not shown).
[0104] The effluent from the HDM reactor is extracted via the
conduit (124) then sent to the first HDS reactor (130) where it
passes through a fixed bed catalyst (132).
[0105] The effluent is sent via the conduit (134) into a
high-pressure, high-temperature separator (HPHT) (136) from which a
gaseous fraction (138) and a liquid fraction (140) are recovered.
The gaseous fraction (138) is generally sent via an exchanger (not
shown) or a cooling tower (142) for cooling, to a high-pressure,
low-temperature separator (HPLT) (144) from which a gaseous
fraction (146) is recovered containing the gases (H.sub.2,
H.sub.2S, NH.sub.3, C1-C4 hydrocarbons etc.) and a liquid fraction
(148).
[0106] The gaseous fraction (146) from the high-pressure,
low-temperature separator (HPLT) (144) is treated in the hydrogen
purification unit (150) from which the hydrogen (152) is recovered
to be recycled via the compressor (154) and the conduit (156) to
the hydrotreatment section. The gases containing the undesirable
nitrated and sulphurated compounds are evacuated from the
installation (flux (158)).
[0107] The liquid fraction (148) from the high-pressure,
low-temperature separator (HPLT) (144) is decompressed in the
device (160) then sent to the fractioning system (172). Optionally
a medium-pressure separator (not shown) can be installed after the
decompressor (160) to recover a vapour phase which is sent to the
purification unit (150) and a liquid phase which is brought to the
fractioning section (172).
[0108] The liquid fraction (140) from the high-pressure,
high-temperature separator (HPHT) (136) is decompressed in the
device (174) then sent to the fractioning system (172). Evidently
fractions (148) and (140) can be sent together, after
decompression, to the system (172). The fractioning system (172)
comprises an atmospheric distillation system to produce a gaseous
effluent (176), at least one so-called light fraction (178)
containing in particular naphtha, kerosene and diesel, and an
atmospheric residue fraction (180). Part of the atmospheric residue
fraction (180) can be extracted via the conduit (182) to constitute
the desired fuel bases. All or part of the atmospheric residue
fraction (180) is sent to a vacuum distillation column (184) to
recover a fraction containing the vacuum residue (186), and a
vacuum distillates fraction (188) containing vacuum gas oil. At
least part of the vacuum residue fraction (186) is preferably
recycled via the conduit (190) to the hydroconversion step, not in
order to increase the yield.
[0109] FIG. 3 describes the method according to the invention with
an intermediate separation without decompression (integrated
system). In reference to FIG. 3, essentially the differences of the
installation and process (separation step) are described in
relation to the process of FIG. 2, the steps of hydroconversion,
hydrotreatment and separation after hydrotreatment (and their
reference symbols) being strictly identical to those in FIG. 2.
[0110] The effluent treated in the hydroconversion reactors is sent
via the conduit (42) to a high-pressure, high-temperature separator
(HPHT) (44) from which a lighter fraction (46) and a residual
fraction (48) are recovered. The cutting point is generally between
200 and 400.degree. C.
[0111] The residual fraction (48) is sent directly to the
hydrotreatment section, following any possible passing through a
filter (94).
[0112] The lighter fraction (46) is sent, optionally mixed with the
gaseous fraction (28) from the optional inter-stage separator (30)
between the two hydroconversion reactors, generally via an
exchanger (not shown) or a cooling tower (50) for cooling, to a
high-pressure, low-temperature separator (HPLT) (52) from which are
recovered a gaseous fraction (54) containing the gases (H.sub.2,
H.sub.2S, NH.sub.3, C1-C4 hydrocarbons etc.) and a liquid fraction
(56).
[0113] The gaseous fraction (54) from the high-pressure,
low-temperature separator (HPLT) (52) is treated in the hydrogen
purification unit (58) from which the hydrogen (60) is recovered
for recycling via the compressor (156) and the conduits (88) and
(64) to the hydrotreatment section and/or hydroconversion section.
The gases containing the undesirable nitrated, sulphurated and
oxygenated compounds are evacuated from the installation (flux
(66)). In this configuration a single compressor (156) is used to
supply all reactors requiring hydrogen.
[0114] The liquid fraction (56) from the high-pressure,
low-temperature separator (HPLT) (52) is decompressed in the device
(68) then sent to the fractioning system (70). Optionally a
medium-pressure separator (not shown) can be installed after the
decompressor (68) to recover a gaseous fraction which is sent to
the purification unit (58), and a liquid fraction which is brought
to the fractioning section (70).
[0115] The fractioning system (70) comprises an atmospheric
distillation system to produce a gaseous effluent (74), at least
one so-called light fraction (76) containing in particular naphtha,
kerosene and diesel, and an atmospheric residue fraction (192).
Part of the atmospheric residue fraction can be sent via the
conduit (192) to the hydrotreatment reactors while another part of
the atmospheric residue fraction (194) can be sent to another
process (hydrocracking or FCC or hydrotreatment).
[0116] The following examples illustrate the invention without
however limiting its scope.
Example 1
[0117] A vacuum residue (RSV Oural), see table 1, is treated. The
feed is subjected to a hydroconversion step in two successive
ebullated bed reactors. The operating conditions and the yield, the
sulphur content and the viscosity of the effluent leaving the
hydroconversion section are given in tables 2 and 3.
TABLE-US-00001 TABLE 1 Characteristics of the feed Cut RSV Oural
530+ Density 15/4 1.002 Sulphur % mass 2.55 Conradson carbon 16
Asphaltenes C7 (% mass) 3.8 NI + V ppm 253
TABLE-US-00002 TABLE 2 Operating conditions for ebullated bed
section Catalyst NiMo on alumina T ebullated bed R1 (.degree. C.)
415 T ebullated bed R2 (.degree. C.) 417 Pressure, MPa 18 VVH*
(h.sup.-1) 0.243 H.sub.2 inlet (Nm.sup.3/m.sup.3 feed) 348
TABLE-US-00003 TABLE 3 Yield, sulphur content and viscosity of
ebullated bed section (% weight/feed) Yield (% S (% Viscosity at
Product weight) weight) 50.degree. C. (Cst) NH.sub.3 0.26 0
H.sub.2S 2.11 94.12 C1-C4 (gas) 2.06 0 Naphtha (IBP-180.degree. C.)
7.30 0.05 Diesel (180-350.degree. C.) 19.80 0.11 Vacuum distillates
(350-480.degree. C.) 24.50 0.55 44 Vacuum residue (480+.degree. C.)
45.60 0.89 5000 Residue (350.degree. C.+) - fixed bed feed 70.10
0.77 400 with H.sub.2 consumed representing 1.63% of the feed.
[0118] The effluent from the hydroconversion is then subjected to
separation. The heavy fraction (fraction 350.degree. C.+) is then
sent to the hydrotreatment section including two permutable
reactors. The operating conditions and the yield, the sulphur
content and the viscosity of the effluent leaving the
hydrotreatment section are given in tables 4 and 5.
TABLE-US-00004 TABLE 4 Operating conditions of fixed bed Catalyst
CoMoNi on alumina T (.degree. C.) 380 Pressure MPa 15 VVH
(Sm.sup.3/h intermediate feed/m.sup.3 cata) 0.15 H.sub.2/HC
(Nm.sup.3/h H2/Sm.sup.3 intermediate feed) 1000
TABLE-US-00005 TABLE 5 Yield, sulphur content and viscosity of
fixed bed section (% weight/hydroconversion effluent) Yield (% S (%
Viscosity at Products weight) weight) 50.degree. C. (Cst) NH.sub.3
0.48 0 H.sub.2S 0.77 94.12 C1-C4 (gas) 1.07 0 Naphtha
(IBP-150.degree. C.) 1.82 0.01 Diesel (150-375.degree. C.) 6.98
0.02 Vacuum distillates (375-550.degree. C.) 47.40 0.09 120 Vacuum
residue (550+.degree. C.) 42.68 0.46 7000 with H.sub.2 consumed
representing 1.20% of the intermediate feed.
[0119] Table 6 summarises the performance of yield, the sulphur
content and the viscosity of the effluent over the complete
hydroconversion and hydrotreatment chain.
TABLE-US-00006 TABLE 6 Yield, sulphur content and viscosity of
complete chain (% weight/feed) Viscosity Yield (% S (% 50.degree.
C. Products weight) weight) (Cst) NH.sub.3 0.59 0 H.sub.2S 2.65
94.12 C1-C4 (gas) 2.81 0 Naphtha (IBP-180.degree. C.) ex ebullated
bed 7.30 0.05 Naphtha (IBP-150.degree. C.) ex fixed bed 1.28 0.01
Gas oil (180-350.degree. C.) ex ebullated bed 19.80 0.11 Gas oil
(150-375.degree. C.) ex fixed bed 4.89 0.02 Vacuum distillates
(375-550.degree. C.) ex 33.23 0.09 120 fixed bed Vacuum residue
(550+.degree. C.) ex fixed 29.92 0.46 7000 bed with H.sub.2
consumed representing 2.47% of the feed.
[0120] To obtain a bunker fuel compliant with specifications, the
fuel is made up of 50% vacuum distillates ex fixed bed, 48% vacuum
residue ex fixed bed and 2% gas oil ex fixed bed. It has a
viscosity of 380 cSt at 50.degree. C., a sulphur content of 0.27%
weight and a sediment content of <0.1% weight. Not all the
vacuum distillate ex fixed bed is used and it could directly
constitute a fuel with 0.1% S.
Example 2
Comparative
TABLE-US-00007 [0121] TABLE 7 Operating conditions, fines content,
load loss and operating factor in case 1 (according to invention
fixed bed with PRS) and case 2 (comparative fixed bed without PRS)
Case 1 (according to invention) 2 (comparative) Configuration of
fixed bed section With PRS Without PRS Catalyst CoMoNi on CoMoNi on
alumina alumina T (.degree. C.) 380 380 Pressure MPa 15 15 H2/HC
(Nm3/h H2/Sm3 intermediate 1000 1000 feed) Duration of cycle
(months) 11 5 Number of cycles 1 2 Number of permutable reactors 2
-- Number of permutations per cycle 3 -- VVH HDM 0.25 0.25 VVH HDS
0.38 0.38 VVH global fixed bed 0.15 0.15 % HDM section HDM 70 70
Sediment at outlet of ebullated bed 0.5 0.5 section (% weight)
Sediment at outlet of PRS (case with 0.2 0.5 PRS) or entry to fixed
bed (case without PRS) % weight Sediment at outlet of fixed bed
section <0.1 <0.1 (% weight) Catalyst fines at outlet of
ebullated bed 70 70 section (ppm weight) Catalyst fines at outlet
from PRS (case 25 70 with PRS) or entry to fixed bed (case without
PRS) (ppm weight) Catalyst fines of outlet of fixed bed 10 10
section (ppm weight) Delta P section HDM start of cycle or 1.5 1.5
permutation (in bar) Delta P section HDM end of cycle or 5.0 7.5
permutation (in bar) Operational factor (%) 0.92 0.83
[0122] The comparative example is intended to show the benefit of
using the permutable reactors between the ebullated bed section and
the remainder of the fixed bed section in order to capture the
majority of catalyst fines and sediments from the first ebullated
bed step.
[0123] Case 1 corresponds to example 1 according to the invention
which, thanks to the permutable reactors (PRS), allows a high
operating factor and a lower catalyst quantity (size of reactor).
This operational factor corresponds to a stoppage of 1 month and 11
months' operation.
[0124] Case 2 corresponds to a case without PRS where the quantity
of catalysts installed is the same as case 1. Because of the
catalyst fines and the sediments from the ebullated bed section,
the load loss of the fixed bed section increases greatly to a point
where it is no longer possible to operate, and the unit must then
be stopped and the catalyst changed. The operational factor is
degraded because with this installation it is only possible to
perform two cycles of 5 months in one year. Furthermore, to
overcome these additional load losses, the feed pump, the
additional hydrogen compressor and the recycling gas compressor
must be over-dimensioned, which leads to greater investment costs
and greater consumption of utilities.
* * * * *