US20080093262A1

US20080093262A1 - Process and installation for conversion of heavy petroleum fractions in a fixed bed with integrated production of middle distillates with a very low sulfur content

Info

Publication number: US20080093262A1
Application number: US11/585,338
Authority: US
Inventors: Andrea Gragnani; Frederick Morel
Original assignee: IFP Energies Nouvelles IFPEN
Current assignee: IFP Energies Nouvelles IFPEN
Priority date: 2006-10-24
Filing date: 2006-10-24
Publication date: 2008-04-24

Abstract

This invention relates to a process and an installation for treatment of a heavy petroleum feedstock, of which at least 80% by weight has a boiling point of greater than 340° C., whereby the process comprises the following stages:

- (a) Hydroconversion in a fixed-bed reactor operating with an upward flow of liquid and gas, whereby the net conversion in products boiling below 360° C. is from 10 to 99% by weight;
- (b) Separation of the effluent obtained from stage (a) into a gas containing hydrogen and H₂S, a fraction comprising the gas oil, and optionally a fraction that is heavier than the gas oil and a naphtha fraction;
- c) Hydrotreatment by contact with at least one catalyst of at least the fraction comprising the gas oil obtained in stage (b);
- d) Separation of the effluent obtained at the end of stage (c) into a gas containing hydrogen and at least one gas oil fraction having a sulfur content of less than 50 ppm, preferably less than 20 ppm, and more preferably still less than 10 ppm,
  the hydroconversion stage (a) being conducted at a pressure P1 and the hydrotreatment stage (c) being conducted at a pressure P2, the difference ΔP=P1−P2 being at least 2 MPa, the hydrogen supply for the hydroconversion (a) and hydrotreatment (c) stages being ensured by a single compression system with n stages.

Description

FIELD OF THE INVENTION

The invention relates to an improved process for conversion of heavy petroleum fractions in a fixed bed with integrated production of gas oil fractions with very low sulfur content, and the installation allowing the implementation of said process.
This invention relates to a process and an installation for the treatment of heavy hydrocarbon feedstocks containing sulfur-containing, nitrogen-containing and metallic impurities. It relates to a process making it possible to convert at least partially such a hydrocarbon feedstock, for example a direct-distillation vacuum distillate, a vacuum distillate that is obtained from a conversion process or oils that are deasphalted with solvent, into gas oils that meet sulfur specifications, i.e., having less than 50 ppm of sulfur, preferably less than 20 ppm, and even more preferably less than 10 ppm, and one or more heavy products that advantageously can be used as a feedstock for catalytic cracking (such as fluidized-bed catalytic cracking), as a feedstock for hydrocracking (such as high-pressure catalytic hydrocracking), as a fuel oil with high or low sulfur content, or as a feedstock for a carbon rejection process (such as the coker).

TECHNOLOGICAL BACKGROUND OF THE INVENTION

Until 2000, the authorized sulfur content in diesel fuel was 350 ppm. Much more stringent values have been imposed since 2005 since this maximum content is not to exceed 50 ppm. This maximum value will next be revised downward and should not exceed 10 ppm in a few years.
It is therefore necessary to develop processes meeting these requirements without prohibitively increasing the cost of production.
The gasolines and the gas oils that are obtained from the conversion process, such as, for example, hydroconversion, are very refractory in hydrotreatment compared to gas oils that are obtained directly from the atmospheric distillation of crude oils.
To obtain very low sulfur contents, it is necessary to convert the most refractory radicals, especially di- and trialkylated dibenzothiophenes, or those having a greater degree of alkylation, for which access of the sulfur atom to the catalyst is limited by the alkyl groups. For this family of compounds, the route of hydrogenation of an aromatic cycle before the desulfurization by breaking the Csp3-S bond is faster than the direct desulfurization by breaking the Csp2-S bond.
It is also necessary to obtain a major reduction of the nitrogen content by conversion especially of the most refractory radicals, especially benzacridines and benzocarbazoles, whereby the acridines are not only refractory, but also inhibit hydrogenation reactions.
Conversion gas oils therefore require very rigorous operating conditions to obtain the desired sulfur specifications.
A process for conversion of heavy petroleum fractions including a fixed bed for producing middle distillates with a low sulfur content has been described in particular in Patent Application US 2003/0089637. This process, however, makes it possible to reduce sulfur levels below 50 ppm only under very rigorous pressure conditions, which considerably increases the cost of the gas oil that is ultimately obtained.
There is therefore a genuine need for a process making it possible to hydrotreat conversion gas oils under less rigorous operating conditions allowing a reduction in investment costs while maintaining a reasonable cycle duration of the hydrotreatment catalyst and allowing sulfur contents of less than 50 ppm, preferably less than 20 ppm, and even more preferably less than 10 ppm, to be obtained.
Values in ppm are all expressed by weight.

SUMMARY OF THE INVENTION

The present inventors have found that it is possible to minimize investment costs by optimizing operating pressures used while obtaining gas oils of good quality having such limited sulfur contents.

DETAILED DESCRIPTION OF THE INVENTION

Thus, the process of the invention is a process for treatment of a heavy petroleum feedstock of which at least 80% by weight has a boiling point of greater than 340° C., which comprises the following stages:

- (a) Hydrocracking in a fixed bed with at least one catalyst at a temperature of about 300 to about 500° C. and often from about 350 and 450° C., a pressure of at least 4 MPa and less than or equal to 17 MPa, an hourly space velocity of from 0.1 to 10 h⁻¹and in the presence of 50 to 5000 Nm³of hydrogen per m³of feedstock, the net conversion into products boiling below 360° C. being from 10 to 99% by weight,
- (b) Separation of the effluent obtained from stage (a) into a gas containing hydrogen and H₂S, a fraction comprising the gas oil, and optionally a fraction that is heavier than the gas oil and a naphtha fraction;
- c) Hydrotreatment by contact with at least one catalyst of at least the fraction comprising the gas oil obtained in stage (b) at a temperature of from 200 to 500° C., at a liquid hourly space velocity relative to the catalyst volume of 0.1 to 10 h⁻¹in the presence of 100 to 5000 Nm³of hydrogen per m³of feedstock;
- d) Separation of the effluent that is obtained at the end of stage (c) into a gas containing hydrogen and at least one gas oil fraction having a sulfur content of less than 50 ppm, preferably less than 20 ppm, and even more preferably less than 10 ppm,
  the hydroconversion stage (hydrocracking) (a) being conducted at a pressure P1 and the hydrotreatment stage (c) being conducted at a pressure P2, the difference ΔP=P1−P2 being at least 2 MPa, generally from 4 to 8 MPa, and preferably from 5 to 7 MPa, and the hydrogen supply for the hydroconversion (a) and hydrotreatment (c) stages being ensured by a single compression system with n stages, n being greater than or equal to 2, generally between 2 and 5, preferably between 2 and 4, and particularly preferably equal to 3.

The liquid hourly space velocity (LHSV) corresponds to the ratio of the feedstock liquid flow rate in m³/h per volume of catalyst in m³.
According to the process of the invention, the pressure P1 implemented in the catalytic hydroconversion stage (a) in a fixed bed is between 6 and 17 MPa and preferably between 8 and 12 MPa.
The pressure P2 implemented in the hydrotreatment stage (c) is between 4 and 8 MPa and preferably between 4.5 and 6 MPa.
Thus, in the process according to the invention, pressures that are completely different for each of the hydroconversion and hydrotreatment stages can be used, which makes it possible in particular to significantly limit the investment costs.
In the process according to the invention, the use of the pressure that is optimum for each particular stage is made possible by implementing a single, multistage hydrogen supply system.
Thus, the hydroconversion stage is supplied with hydrogen originating from delivery from the last compression stage, and the hydrotreatment stage is supplied with hydrogen originating from delivery from an intermediate compression stage, i.e., at a lower total pressure.
According to one particular embodiment, the process of the invention implements a single, 3-stage hydrogen compressor in which the delivery pressure of the first stage is between 4 and 5 MPa, preferably between 4.5 and 5 MPa, the delivery pressure of the second stage is between 8 and 12 MPa, preferably between 9 and 11 MPa, and the delivery pressure of the third stage is between 12 and 17 MPa, preferably between 13 and 15 MPa.
In one particular embodiment, the hydrogen originating from the delivery from the second compression stage feeds the hydrotreatment reactor.
According to one particular embodiment, the partial hydrogen pressure in the hydrotreatment reactor P2_H2is between 3.4 and 8 MPa and preferably between 4 and 6 MPa.
These elevated partial hydrogen pressure values are made possible by the fact that all of the make-up hydrogen necessary to the process is brought into stage (c). In this invention, the “make-up hydrogen” is mentioned in contrast to the recycled hydrogen. The hydrogen purity is generally between 84 and 100% and preferably between 95 and 100%.
According to another embodiment, the hydrogen supplying the last compression stage can be recycled hydrogen originating from the separation stage (d) and/or the separation stage (b).
This recycled hydrogen can optionally supply an intermediate stage of the compressor that has stages. In this case, it is preferred that said hydrogen has been purified before its recycling.
According to another embodiment, the delivery hydrogen from the initial compression stage and/or from an intermediate stage can, moreover, supply a unit for hydrotreatment of gas oil that is obtained directly from atmospheric distillation, called “straight-run gas oil.” In a conventional way, the straight-run gas oil hydrotreatment unit is operated at a pressure of between 3 and 6.5 MPa and preferably between 4.5 and 5.5 MPa.

The reaction conditions of each of the stages will now be described in greater detail, especially in conjunction with the drawings in which:

FIG. 1 shows the diagram of the installation allowing implementation of one embodiment of the process according to the invention;

FIG. 2 shows the diagram of the installation allowing implementation of another embodiment of the process according to the invention.

The process according to the invention is very particularly suitable for treatment of heavy feedstocks, i.e., feedstocks of which at least 80% by weight has a boiling point of greater than 340° C. Their initial boiling point is generally established at least 340° C., often at least 370° C. or even at least 400° C. These are, for example, direct-distillation vacuum distillates, vacuum distillates that are obtained from conversion processes such as, for example, those that originate from coking, from a fixed-bed hydroconversion (such as those obtained from the HYVAHL® processes for treatment of heavy products developed by the applicant) or processes for hydrotreatment of heavy products in a boiling bed (such as those obtained from H-OIL® processes), or else oils that are deasphalted with solvent (for example with propane, with butane or with pentane) coming from deasphalting of direct-distillation vacuum residue or residues that are obtained from the HYVAHL® or H-OIL® processes. The feedstocks can also be formed by mixing these various fractions. They can also contain fractions originating from the process that is the object of this invention and those recycled for its supply. They can also contain gas-oil fractions and heavy gas oils originating from the catalytic cracking that has a distillation interval of about 160° C. to about 500° C. They can also contain aromatic extracts and paraffins that are obtained within the framework of the production of lubricating oils. According to this invention, the feedstocks that are treated are preferably vacuum distillates.
The sulfur content of the feedstock is highly variable and is nonlimiting. The content of metals such as nickel and vanadium is generally between about 1 ppm and 30 ppm, but it is without any technical limitation.
The feedstock is treated first of all in a hydroconversion section (II) in the presence of hydrogen originating from the hydrogen compression zone (I). Then, the treated feedstock is separated into the separation zone (III) from where, among other fractions, a gas oil fraction is recovered that then supplies the hydrotreatment zone (IV) from where the remaining sulfur is removed.
Each of these reaction zones is shown in FIGS. 1 and 2. The different physical reactions or transformations carried out in each of these zones will be described below.
Zone (I) represents the compression of hydrogen in several stages (three in the figures). In this zone, the make-up hydrogen is treated, optionally mixed with the flows of purified recycling hydrogen, to raise its pressure to the level required by stage (a). Said single compression system generally includes at least two compression stages that are generally separated by compressed gas cooling systems, liquid and vapor phase separation units and optionally inputs of the purified recycling hydrogen flows. The breakdown into several stages therefore makes available hydrogen at one or more intermediate pressures between that of the input and that of the output of the system. This (these) intermediate pressure level(s) can supply hydrogen to at least one catalytic hydrotreatment or hydrocracking unit.
More specifically, the make-up hydrogen that is necessary to the operation of zones (II) and (IV) achieves a pressure of between 1 and 3.5 MPa, and preferably between 2 and 2.5 MPa by a pipe (4) in a zone (I) where it is compressed, optionally with other recycling hydrogen flows, in a multistage compression system. Each compression stage (1, 2 and 3), three in the figures, is separated from the next by a liquid-vapor separation and cooling system (33), (34) and (35) allowing the gas temperature and the amount of liquid carried to the following compression stage to be reduced. The pipes allowing the evacuation of this liquid are not shown in the figures.
Between the first and the last stage, and more often between the second and the third stage, one pipe (7) routes at least part, and preferably all, of the compressed hydrogen to the hydrotreatment zone (IV). The hydrogen leaving the zone (IV) through the pipe (8) is sent to the following compression stage, more often the third and last. The pipe (14) routes the hydrogen to zone (II).
The feedstock to be treated (such as defined above) enters the hydroconversion zone (II) in a fixed bed via a pipe (10). The effluent that is obtained in the pipe (11) is sent into the separation zone (III).
This zone (II) comprises at least one fixed-bed reactor that contains at least one catalyst.
The operation is usually carried out under an absolute pressure of 6 to 17 MPa, and most often of 8 to 12 MPa, at a temperature of about 300° C. to about 500° C., and often from about 350 to about 450° C. The liquid hourly space velocity (LHSV) relative to the volume of catalyst and the partial pressure of hydrogen are important factors that one skilled in the art knows to select based on the characteristics of the feedstock to be treated and the desired conversion. Most often, the LHSV, relative to the catalyst volume, is located in a range that goes from about 0.1 h⁻¹to about 10 h⁻¹, and preferably from about 0.2 h⁻¹to about 5 h⁻¹. The amount of hydrogen mixed with the feedstock is usually from about 50 to about 5000 normal cubic meters (Nm³) per cubic meter (m³) of the liquid feedstock, and most often from about 100 to about 2000 Nm³/m³, and preferably from about 200 to 1500 Nm³/m³.
The net conversion into products boiling below 360° C. of the fraction having a boiling point of greater than 540° C. is between 10 and 99% by weight, most often between 10 and 70% by weight, and advantageously between 15 and 45%.
In this hydroconversion stage, it is possible to use any standard catalyst, in particular a granular catalyst comprising, on an amorphous substrate, at least one metal or metal compound with a hydro-dehydrogenating function. This catalyst can be a catalyst comprising metals of group VIII, for example nickel and/or cobalt, most often in combination with at least one metal of group VIB, for example molybdenum and/or tungsten. It is possible, for example, to use a catalyst comprising from 0.5 to 10% by weight of nickel and preferably from 1 to 5% by weight of nickel (expressed in terms of nickel oxide NiO), and from 1 to 30% by weight of molybdenum and preferably from 5 to 20% by weight of molybdenum (expressed in terms of molybdenum oxide MoO₃) on an amorphous mineral substrate. This substrate will be selected from, for example, the group that is formed by alumina, silica, silica-aluminas, magnesia, clays and mixtures of at least two of these minerals. This substrate can likewise contain other compounds, and, for example, oxides that are selected from the group that is formed by boron oxide, zirconia, titanium oxide, and phosphoric anhydride. Most often, an alumina substrate is used, and very often an alumina substrate that is doped with phosphorus and optionally boron is used. The concentration of phosphoric anhydride P₂O₅is usually less than about 20% by weight and most often less than about 10% by weight. This concentration of P₂O₅is usually at least 0.001% by weight. The concentration of boron trioxide B₂O₃is usually from about 0 to about 10% by weight. The alumina that is used is usually a γ- or η-alumina. The total content of oxides of metals of groups VI and VIII is often from about 5 to about 40% by weight and in general from about 7 to 30% by weight, and the ratio by weight expressed in terms of metal oxide between the metal (or metals) of group VI to the metal (or metals) of group VII is in general from about 20 to about 1 and most often from about 10 to about 2.
Another type of usable catalyst comprises at least one metal of group VIII and at least one metal from the group VIB and one silica-alumina.
Another type of usable catalyst is a catalyst that contains at least one matrix, at least one Y zeolite and at least one hydro-dehydrogenating metal. The matrices, metals, and additional elements described above can also be part of the composition of this catalyst. Advantageous Y zeolites are described in the Patent Applications WO-00/71641, EP-911 077 as well as U.S. Pat. Nos. 4,738,940 and 4,738,941.
The hydrocracked effluent obtained from stage (a) is then separated in stage (b). It is introduced by a pipe (11) into at least one separator (15) that separates, on the one hand, a gas containing hydrogen (gaseous phase) in the pipe (16) and, on the other hand, a liquid effluent in the pipe (17). It is possible to use a hot separator followed by a cold separator. A series of hot and cold separators at medium and low pressure can also be present.
The liquid effluent is sent into a separator (18) that is preferably composed of at least one distillation column, and it is separated into at least one distillate fraction that includes a gas oil fraction and that is located in the pipe (21). It is also separated into at least one fraction that is heavier than the gas oil that is evacuated via the pipe (23).
At the level of the separator (18), the acid gas can be separated in a pipe (19), the naphtha can be separated in an additional pipe (20), and the fraction that is heavier than the gas oil can be separated in a vacuum distillation column into a vacuum residue discharging via the pipe (23) and one or more pipes (22) that correspond to vacuum gas oil fractions.
The fraction from the pipe (23) can be used as an industrial fuel oil with a low sulfur content or can advantageously be sent to a carbon rejection process, such as, for example, coking.
Naphtha (20), obtained separately, optionally with the naphtha (29) separated in zone (IV) added, is advantageously separated into heavy and light gasolines, the heavy gasoline being sent into a reforming zone and the light gasoline being sent to a zone where paraffin isomerization is carried out.
The vacuum gas oil (22) may optionally be sent, alone or in a mixture with similar fractions of different origins, into a catalytic cracking process in which these fractions are advantageously treated under conditions that make it possible to produce a gaseous fraction, a gasoline fraction, a gas oil fraction and a fraction that is heavier than the gas oil fraction that is often called the slurry fraction by ones skilled in the art. They can also be sent into a catalytic hydrocracking process in which they are advantageously treated under conditions making it possible to produce in particular a gaseous fraction, a gasoline fraction, or a gas oil fraction.
In FIGS. 1 and 2, the separation zone (III) formed by the separators (15) and (18) is shown by dotted lines.
For distillation, the conditions are, of course, selected based on the initial feedstock. If the initial feedstock is a vacuum gas oil, the conditions will be more rigorous than if the initial feedstock is an atmospheric gas oil. For an atmospheric gas oil, conditions are generally selected such that the initial boiling point of the heavy fraction is from about 340° C. to about 400° C., and for a vacuum gas oil, they are generally selected such that the initial boiling point of the heavy fraction is from about 540° C. to about 700° C.
For naphtha, the final boiling point is between about 120° C. and about 180° C.
The gas oil is between the naphtha and the heavy fractions.
The fraction points given here are indicative, but the operator will choose the fraction point based on the quality and the quantity of the desired products, as is generally practiced.
At the outlet of stage (b), the gas oil fraction most often has a sulfur content of between 100 and 10,000 ppm, and the gasoline fraction most often has a sulfur content of at most 1000 ppm. The gas oil fraction thus does not meet 2005 sulfur specifications.
The gas oil fraction is then sent (alone or optionally with an external naphtha and/or gas oil fraction added to the process) into a hydrotreatment zone (IV) provided with at least one fixed bed of a hydrotreatment catalyst in order to reduce the sulfur content to below 50 ppm, preferably below 20 ppm, and even more preferably below 10 ppm. It is also necessary to significantly reduce the nitrogen content of the gas oil to obtain a desulfurized product with a stable color.
It is possible to add to said gas oil fraction a fraction that is produced outside of the process according to the invention, which normally cannot be directly incorporated into the gas oil pool. This hydrocarbon fraction can be selected from, for example, the group that is formed by the LCO (light cycle oil) originating from fluidized-bed catalytic cracking as well as a gas oil that is obtained from a high-pressure hydroconversion process of a vacuum distillation gas oil.
Usually, an operation proceeds at a total pressure of about 4 to 8 MPa, and preferably from about 4.5 to 6 MPa. The temperature in this stage is usually from about 200 to about 500° C., preferably from about 330 to about 410° C. This temperature is usually adjusted based on the desired level of hydrodesulfurization and/or saturation of aromatic compounds and should be compatible with the desired cycle duration. The liquid hourly space velocity or LHSV and the partial hydrogen pressure are selected based on the characteristics of the feedstock to be treated and the desired conversion. Most often, the LHSV is in a range of from about 0.1 h⁻¹to about 10 h⁻¹and preferably 0.1 h⁻¹-5 h⁻¹, and advantageously from about 0.2 h⁻¹to about 2 h⁻¹.
The total amount of hydrogen mixed with the feedstock depends largely on the hydrogen consumption from stage b) as well as the recycled purified hydrogen gas sent to stage a). It is, however, usually from about 100 to about 5000 normal cubic meters (Nm³) per cubic meter (m³) of the liquid feedstock and most often from about 150 to 1000 Nm³/m³.
The operation of stage d) in the presence of a large amount of hydrogen makes it possible to usefully reduce the partial pressure of ammonia. In the preferred case of this invention, the partial pressure of ammonia is generally less than 0.5 MPa.
An operation is likewise usefully carried out with a reduced partial hydrogen sulfide pressure that is compatible with the stability of the sulfide catalysts. In the preferred case of this invention, the partial hydrogen sulfide pressure is generally less than 0.5 MPa.
In the hydrodesulfurization zone, the ideal catalyst should have a strong hydrogenation capacity so as to accomplish thorough refinement of the products and to obtain a major reduction of sulfur and nitrogen. According to the preferred embodiment of the invention, the hydrotreatment zone operates at a relatively low temperature, which points in the direction of thorough hydrogenation and therefore an improvement of the content of aromatic compounds of the product and its cetane number and limitation of coking. The scope of this invention would not be exceeded by using a single catalyst or several different catalysts simultaneously or in succession in the hydrotreatment zone. Usually, this stage is carried out industrially in one or more reactors with one or more catalytic beds and with downward liquid flow.
In the hydrotreatment zone, at least one fixed bed of the hydrotreatment catalyst comprising a hydro-dehydrogenating function and an amorphous substrate is used. A catalyst is preferably used whose substrate is selected from, for example, the group that is formed by alumina, silica, silica-aluminas, magnesia, clays and mixtures of at least two of these minerals. This substrate can also contain other compounds and, for example, oxides selected from the group that is formed by boron oxide, zirconia, titanium oxide, and phosphoric anhydride. Most often, an alumina substrate is used and, better, η- or γ-alumina. The hydrogenating function is ensured by at least one metal of group VIII, for example nickel and/or cobalt, optionally in combination with a metal of group VIB, for example molybdenum and/or tungsten. Preferably, a catalyst based on NiMo will be used. For gas oils that are difficult to hydrotreat and for very high levels of hydrodesulfurization, one skilled in the art knows that the desulfurization of an NiMo-based catalyst is superior to that of a CoMo catalyst because the first has a greater hydrogenating function than the second. For example, it is possible to use a catalyst that comprises from 0.5 to 10% by weight of nickel and preferably from 1 to 5% by weight of nickel (expressed in terms of nickel oxide NiO), and from 1 to 30% by weight of molybdenum and preferably from 5 to 20% by weight of molybdenum (expressed in terms of molybdenum oxide (MoO₃)) on an amorphous mineral substrate. In an advantageous case, the total content of oxides of metals of groups VI and VIII is often from about 5 to about 40% by weight and in general from about 7 to 30% by weight, and the ratio by weight expressed in terms of metal oxide between the metal (metals) of group VI to the metal (or metals) of group VIII is in general from about 20 to about 1 and most often from about 10 to about 2.
The catalyst can also contain an element such as phosphorus and/or boron. This element may have been introduced into the matrix or may have been deposited on the substrate. It is also possible to deposit silicon on the substrate, alone or with phosphorus and/or boron. The concentration of said element is usually less than about 20% by weight (calculated in terms of oxide) and most often less than about 10% by weight, and it is usually at least 0.001% by weight. The concentration of boron trioxide B₂O₃is usually from about 0 to about 10% by weight.
Preferred catalysts contain silicon deposited on a substrate (such as alumina), optionally with P and/or B also deposited, and also containing at least one metal of group VIII (Ni, Co) and at least one metal of group VIB (W, Mo).
The hydrotreated effluent that is obtained exits via the pipe (25) to be sent into the separation zone (V) shown in a diagram in dotted lines in FIGS. 1 and 2.
Here, it comprises a separator (26), preferably a cold separator, where a gaseous phase exiting via the pipe (8) and a liquid phase exiting via the pipe (27) are separated.
The liquid phase is sent into a separator (31), preferably a stripper, to remove the hydrogen sulfide exiting into the pipe (28), most often mixed with naphtha. A gas oil fraction is drawn off by the pipe (30), a fraction that meets sulfur specifications, i.e., having less than 50 ppm of sulfur, and generally less than 20 ppm of sulfur, or even less than 10 ppm. The H₂S—naphtha mixture is then optionally treated to recover the purified naphtha fraction. Separation can also be carried out at the level of the separator (31), and the naphtha can be drawn off via the pipe (29).
The process according to the invention also advantageously comprises a hydrogen recycling loop for the 2 zones (II) and (IV) that can be independent for the two zones, but is preferably shared, and that is now described based on FIG. 1.
The gas containing the hydrogen (gaseous phase of the pipe (16) that is separated in the zone (III)) is treated to reduce its sulfur content and optionally to eliminate the hydrocarbon compounds that have been able to pass during the separation.
Advantageously and according to FIG. 1, the gaseous phase of the pipe (16) enters a purification and cooling system (36). It is sent into a cooling tower after having been washed by injected water and partially condensed by a recycled hydrocarbon fraction from the low-temperature section downstream from the cooling tower. The effluent from the cooling tower is sent into a separation zone where water, a hydrocarbon fraction and a gaseous phase are separated.
A portion of the recycled hydrocarbon fraction is sent into the separation zone (III), and advantageously into the pipe (37).
The gaseous phase that is obtained and from which hydrocarbon compounds have been removed is sent, if necessary, into a treatment unit to reduce the sulfur content. Advantageously, this is a treatment with at least one amine.
In certain cases, it is enough that only a portion of the gaseous phase is treated. In other cases, all of it should be treated.
The gas that contains the hydrogen that has thus optionally been purified is then sent to a purification system that makes it possible to obtain hydrogen with a purity comparable to make-up hydrogen.
A membrane purification system offers an economical means of separating hydrogen from other light gases based on a permeation technology. An alternative system could be purification by adsorption with regeneration by pressure variation known under the term Pressure Swing Adsorption (PSA). A third technology or a combination of several technologies could also be considered.
At the outlet of the purification system, one or more pipes (5) and (6) allow recycling of purified hydrogen to the zone (I), normally at one or more pressure levels. Direct recycling to the feed (38) of the zone (II) can also be considered, and in this case, purification of this flow by membranes or PSA is no longer necessary.
One particular embodiment has been described here for separating the entrained hydrocarbon compounds; any other method known to one skilled in the art is suitable.
In the preferred embodiment of FIG. 1, all of the make-up hydrogen is introduced via the pipe (7) at the level of the zone (IV).
According to another embodiment, a pipe bringing solely a portion of the hydrogen at the level of zone (IV) can be provided.
The zone (IV) being able to benefit from a high flow rate of high-purity hydrogen operates at a partial hydrogen pressure that is very near the total pressure and for the same reason at very low partial pressures of hydrogen sulfide and ammonia. This advantageously makes it possible to reduce the total pressure and the amounts of catalyst necessary to obtain specifications for the gas oil that is produced and overall to minimize investment costs.
The process of the invention is implemented in an installation comprising the following reaction zones:
A single hydrogen compression zone that consists of n compression stages arranged in series, n being between 2 and 6, preferably between 2 and 5, preferably between 2 and 4 and being more preferably equal to 3,
A catalytic hydroconversion zone (II) that consists of at least one fixed-bed reactor that is fed with hydrogen via the last compression stage and is connected via the pipe (11) to a separation zone (III) that consists of at least one separator (15) and at least one distillation column (18), the separator allowing the separation of a hydrogen-rich gas via the pipe (16) and a liquid phase that is brought via the pipe (17) to the distillation column (18); the pipe (21) drawing off the distilled gas oil fraction is connected to
A hydrotreatment zone (IV) that consists of a fixed-bed hydrotreatment reactor that is supplied with hydrogen by an intermediate compression stage, and of which the effluent pipe (25) is connected to
A separation zone (V) that makes it possible to evacuate the hydrogen to the last compression stage.
Thus, according to one embodiment of the invention, the installation is such as that shown in a diagram in FIG. 1.
The detail of the various reaction zones is such as has been described above in conjunction with the description of the process.
According to one particular embodiment, in the installation according to the invention, an intermediate compression stage, the first one in FIG. 2, is connected to a straight-run gas oil hydrotreatment reactor (40).
The invention also relates to the use in an installation for conversion of a heavy petroleum feedstock in a fixed bed of a single multistage hydrogen compressor.
The invention will be illustrated using the following example that is not limiting.
In an installation according to the invention (as illustrated in FIG. 1) with a single, three-stage compression system, the conversion of a mixture of vacuum gas oil (VGO) and a heavy coker gas oil (HCGO) that is obtained from a Middle-Eastern crude oil is performed.
The properties of the mixture are as follows:


	VGO + HCGO

	Density at 15° C.	0.945
	Sulfur, % by Weight	3.4
	Nitrogen, ppm by Weight	1554
	Conradson Carbon, % by Weight	<1
	Nickel + Vanadium, ppm by Weight	<2
	C₇Insolubles, % by Weight	<0.05
	ASTM Distillation, ASTM D1160, ° C.
	T
5%	368
	T 50%	456
	T 95%	558

The VGO+HCGO mixture described in the table above is sent into a mild hydrocracking unit (MHC) that operates under the following conditions:

- Total pressure: 9.5 MPa
- VVH: 0.7 h⁻¹
- Catalyst: NiMo on alumina such as the HR-548 catalyst that is marketed by the Axens Company
- Mean reactor temperature: 370° C.

Under these conditions, a conversion of the VGO+HCGO feedstock of about 25% by weight is obtained. The yield of the diesel fraction is 25.7% by volume. The sulfur content of this diesel fraction is 150 ppm. This product therefore does not meet the international specifications that limit the sulfur of the diesel fuels to less than 10 ppm.
This diesel fraction is sent into a fixed-bed hydrotreatment unit that operates under the following conditions:

- Total pressure: 5.7 MPa
- VVH: 0.85 h⁻¹
- Catalyst: CoMo on alumina such as the HR-526 marketed by the Axens Company,
- Mean reactor temperature: 345° C.

The addition of hydrogen from the diesel hydrotreatment unit is taken from the outlet of the 2^ndstage of the compressor. The delivery pressure of the 2^ndstage of the make-up compressor is 6.5 MPa. The high-pressure purging of the hydrotreatment unit is recycled to the intake of the 3^rdstage of the compressor. The delivery pressure of the 3^rdstage of the make-up compressor is 10.2 MPa. The diesel hydrotreatment unit does not comprise a hydrogen recycling compressor.
At the outlet of the hydrotreatment reactor, the yields and the qualities of products are illustrated in the following table. The diesel that is produced has a sulfur content that is less than 10 ppm by weight, which duly meets future international specifications.


	Yields, % by Volume
	Naphtha	1.4
	Diesel	25.7
	Hydrotreated VGO	75.7
	H₂Consumption, % by Weight	1.3
	Diesel Properties
	Density at 15° C.	0.870
	Sulfur, ppm	<10
	Cetane	48
	Hydrotreated VGO Properties
	Density at 15° C.	0.890
	Sulfur, ppm	<300
	Hydrogen, % by Weight	13.0

Without further elaboration, it is believed that one skilled in the art can, using the preceding description, utilize the present invention to its fullest extent. The preceding preferred specific embodiments are, therefore, to be construed as merely illustrative, and not limitative of the remainder of the disclosure in any way whatsoever.
In the foregoing and in the examples, all temperatures are set forth uncorrected in degrees Celsius and, all parts and percentages are by weight, unless otherwise indicated.
The preceding examples can be repeated with similar success by substituting the generically or specifically described reactants and/or operating conditions of this invention for those used in the preceding examples.
From the foregoing description, one skilled in the art can easily ascertain the essential characteristics of this invention and, without departing from the spirit and scope thereof, can make various changes and modifications of the invention to adapt it to various usages and conditions.

Claims

1. Process for treatment of a heavy petroleum feedstock of which at least 80% by weight has a boiling point of greater than 340° C., which comprises the following stages:

(a) Hydrocracking in a fixed bed with at least one catalyst at a temperature of 300-500° C., a pressure of at least 4 MPa and less than or equal to 17 MPa, an hourly space velocity of 0.1 to 10 h⁻¹and in the presence of 50 to 5000 Nm³of hydrogen per m³of feedstock, the net conversion into products boiling below 360° C. being from 10 to 99% by weight,

(b) Separation of the effluent that is obtained from stage (a) into a gas containing hydrogen and H₂S, a fraction comprising the gas oil, and optionally a fraction that is heavier than the gas oil and a naphtha fraction;

c) Hydrotreatment by contact with at least one catalyst of at least the fraction comprising the gas oil obtained in stage (b), at a temperature of 200 to 500° C., at a liquid hourly space velocity relative to the catalyst volume of 0.1 to 10 h⁻¹in the presence of 100 to 5000 Nm³of hydrogen per m³of feedstock;

d) Separation of the effluent that is obtained at the end of stage (c) into a gas containing hydrogen and at least one gas oil fraction that has a sulfur content of less than 50 ppm,

the hydroconversion stage (a) being conducted at a pressure P1 and the hydrotreatment stage (c) being conducted at a pressure P2, the difference ΔP=P1−P2 being at least 2 MPa, the hydrogen supply for the hydroconversion (a) and hydrotreatment (c) stages being ensured by a single compression system with n stages, n being greater than or equal to 2.

2. Process according to claim 1, in which n is between 2 and 6.

3. Process according to claim 2, in which n is between 2 and 5.

4. Process according to claim 3, in which n is between 2 and 4.

5. Process according to claim 4, characterized by the fact that n is equal to 3.

6. Process according to claim 1, in which ΔP is from 4 to 8 MPa.

7. Process according to claim 6, in which ΔP is from 5 to 7 MPa.

8. Process according to claim 1, in which in stage (d), a gas oil whose sulfur content is less than 20 ppm is separated.

9. Process according to claim 8, in which in stage (d), a gas oil whose sulfur content is less than 10 ppm is separated.

10. Process according to claim 1, in which the pressure P1 implemented in the fixed-bed catalytic hydroconversion stage (a) is between 6 and 17 MPa.

11. Process according to claim 10, in which the pressure P1 is between 8 and 12 MPa.

12. Process according to claim 1, in which the pressure P2 implemented in the hydrotreatment stage (c) is between 4 and 8 MPa.

13. Process according to claim 12, in which the pressure P2 is between 4.5 and 6 MPa.

14. Process according to claim 1, in which n=3 and the delivery pressure of the first compression stage is between 4 and 5 MPa, the delivery pressure of the second compression stage is between 8 and 12 MPa, and the delivery pressure of the third compression stage is between 12 and 17 MPa.

15. Process according to claim 15, in which n=3 and the delivery pressure of the first compression stage is between 4.5 and 5 MPa, the delivery pressure of the second compression stage is between 9 and 11 MPa, and the delivery pressure of the third compression stage is between 13 and 15 MPa.

16. Process according to claim 1, in which n=3 and in which the delivery hydrogen from the second compression stage supplies the hydrotreatment reactor.

17. Process according to claim 1, in which the partial hydrogen pressure in the P2_H2hydrotreatment reactor is between 3.4 and 8 MPa.

18. Process according to claim 18, in which P2_H2is between 4 and 6 MPa.

19. Process according to claim 1, according to which the hydrogen supplying the last compression stage is the recycled hydrogen originating from the separation stage (d) or from the separation stage (b).

20. Process according to claim 1, according to which the delivery hydrogen from an intermediate compression stage can, moreover, supply a hydrotreatment unit of gas oil obtained directly from atmospheric distillation, called “straight-run gas oil,” at a pressure of between 3 and 6.5 MPa.

21. Installation for treatment of a heavy petroleum feedstock comprising the following reaction zones:

a single hydrogen compression zone that consists of n compression stages arranged in series, n being greater than or equal to 2,

a catalytic hydroconversion zone (II) that consists of at least one fixed-bed reactor that is supplied with hydrogen via the last compression stage, and connected via the pipe (11) to

a separation zone (III) that consists of at least one separator (15) and at least one distillation column (18), the separator allowing the separation of a hydrogen-rich gas via the pipe (16) and a liquid phase that is brought via the pipe (17) to the distillation column (18); the pipe (21) drawing off the distilled gas oil fraction is connected to

a hydrotreatment zone (IV) that consists of a fixed-bed hydrotreatment reactor that is supplied with hydrogen by an intermediate compression stage, and whose pipe of the effluent (25) is connected to

a separation zone (V) allowing evacuation of hydrogen to the last compression stage.

22. Installation according to claim 22, in which n is preferably between 2 and 6.

23. Installation according to claim 22, in which n is preferably between 2 and 5.

24. Installation according to claim 23, in which n is preferably between 2 and 4.

25. Installation according to claim 24, in which n is equal to 3.

26. Installation according to claim 21, in which the delivery from an intermediate compression stage feeds a straight-run gas oil hydrotreatment reactor (40).