FR3089518A1 - IMPROVED PROCESS FOR CONVERTING A HEAVY LOAD INTO MEDIUM DISTILLATES USING A LINKAGE OF HYDROCRACKING, VAPOCRACKING AND OLIGOMERIZATION UNITS - Google Patents

Abstract

The present invention relates to a process for the production of middle distillates comprising a hydrotreatment step followed by the hydrocracking of a VGO type hydrocarbon feed, separation of the effluent from the hydrocracking step, and the dispatch separate light effluent having a final boiling point below 250 ° C in an olefin production stage, that is to say a steam cracking stage. At least part of the effluent from the olefin production step, that is to say steam cracking, is then sent to an oligomerization step followed by a step for hydrogenation of the olefins obtained to produce middle distillates. Figure 1 to be published

Description

Description

Title of the invention: IMPROVED PROCESS FOR CONVERTING A HEAVY LOAD INTO MEDIUM DISTILLATES USING A LINK OF HYDROCRACKING, SPRAYING AND OLIGOMERIZATION UNITS

Technical area

Hydrocracking of heavy petroleum fractions is a key refining process which makes it possible to produce, from excess heavy and little recoverable charges, the fractions or fractions, lighter such as naphthas, kerosene and gas oils which serve basic to the production of commercial fuels, such as gasoline, jet fuels and diesel (Diesel according to English terminology) that the refiner seeks to adapt its production to demand. Certain hydrocracking processes also make it possible to obtain a highly purified residue which can constitute excellent bases for oils. One of the effluents particularly targeted by the hydrocracking process is the middle distillate (fraction which contains the diesel cut and the kerosene cut), that is to say cuts with an initial boiling point of at least 150 ° C. and final up to before the initial boiling point of the residue, for example less than 340 ° C, or even less than 370 ° C. Preferably, the middle distillate corresponds to the cut 150-370 ° C, and more preferably to a cut having a boiling point T5> 150 ° C and a T95 <370 ° C, that is to say that 95% by volume of the compounds present in said cut have a boiling point higher than 150 ° C and that 95% by volume of the compounds present in said cut have a boiling point lower than 370 ° C.

Hydrocracking is a process which derives its flexibility from three main elements which are, the operating conditions used, the types of catalysts used and the fact that hydrocracking of hydrocarbon feedstocks can be carried out in one or in two stages.

Indeed, hydrocracking is a process which can be declined in different versions which are for the main ones:

Hydrocracking in one step, which comprises first er and generally a thorough hydrotreatment which aims to achieve a hydrodenitrogenation and desulphurization of the feed before it is sent to the catalyst hydrocracking proper, in particular in the case where it comprises a zeolite. This advanced hydrotreatment of the feed causes only a limited conversion of the feed into lighter fractions, which remains insufficient and must therefore be completed on the more active hydrocracking catalyst. However, it should be noted that there is no separation between the two types of catalysts. All of the effluent at the outlet of the reactor is injected onto the hydrocracking catalyst proper and it is only afterwards that a separation of the products formed is carried out. This version of the hydrocracking one stage also called “Once Through” has a variant which presents a recycling of the fraction not converted towards the entry of the hydrotreatment stage or the hydrocracking stage with a view to a further conversion of the load.

Two-stage hydrocracking comprises a first stage which aims, as in the one-stage process, to carry out the hydrorefining of the feedstock but also to achieve a conversion of the latter in the order of generally 40 to 60%. The effluent from the first stage then undergoes separation (distillation), most often called intermediate separation, which aims to separate the conversion products from the unconverted fraction. In the second step of a 2-step hydrocracking process, only the fraction of the charge that is not converted during the first step is treated. This separation allows a two-step hydrocracking process to be more selective in diesel than a one-step process. In fact, the intermediate separation of the conversion products avoids their "over-cracking" into naphtha and gas in the second step on the hydrocracking catalyst. Furthermore, it should be noted that the unconverted fraction of the feed treated in the second step generally contains very low contents of NH3 as well as organic nitrogen compounds, in general less than 20 ppm by weight or even less than 10 ppm by weight. .

Conventionally, the 2-step process can be carried out with either an intermediate separation after the hydrorefining, in a process comprising a hydrorefining reactor and a hydrocracking reactor, or with an intermediate separation between the first and the second hydrocracking reactor in a process comprising in series the hydrorefining reactors, 1st hydrocracking, 2nd hydrocracking.

In particular, the hydrocracking of vacuum distillates or DSV makes it possible to produce light cuts (Naphtas, Kerosenes, Diesel oils, ...) more valuable than the DSV itself. This catalytic process does not entirely transform the DSV into light cuts. After fractionation, there remains a more or less significant proportion of fraction of unconverted DSV called UCO or Unconverted Oil according to the English terminology. To increase the conversion, this unconverted fraction can be recycled at the inlet of the hydrotreating reactor or at the inlet of the hydrocracking reactor in the case where it is a 1 step hydrocracking process or at the inlet of a second hydrocracking reactor treating the fraction not converted to Tissue from the fractionation stage, in the case where it is a 2-stage hydrocracking process. The fractionation at the outlet of the unit makes it possible to separate this heavy cut from naphtha and light gaseous effluents (light ends according to English terminology) as well as higher value cuts such as diesel and kerosene (called middle distillates).

The catalysts used in hydrocracking are all of the bifunctional type combining an acid function with a hydrogenating function. The acid function is provided by supports of large surfaces (150 to 800 m2.gl generally) having a high acidity, such as halogenated aluminas (chlorinated or fluorinated in particular), combinations of boron and aluminum oxides, amorphous silica-aluminas and zeolites. The hydrogenating function is provided either by one or more metals from group VIII of the periodic table, or by a combination of at least one metal from group VIB of the periodic table and at least one metal from group VIII, used in the presence of sulfur. The balance between the two acid and hydrogenating functions governs the activity and the selectivity of the catalyst. However, even with a very selective hydrocracking catalyst, it is difficult to combine activity and selectivity, for example catalysts based on silica-alumina are known to be more selective in middle distillates than zeolites but they are also much less active. In addition, whatever the catalyst, as soon as the DSV conversions are high, an overcracking phenomenon operates on the catalyst leading to an increase in petrol-oriented products.

The present invention relates to a process for producing middle distillates comprising a hydrotreatment step followed by the hydrocracking of a VGO type hydrocarbon feed, a separation of the effluent from the hydrocracking step, and sending the separate light effluent having a boiling point between 50 and 250 ° C in an olefin production stage, that is to say a steam cracking stage. At least part of the effluent from the olefin production step, that is to say steam cracking, is then sent to an oligomerization step followed by a step for hydrogenation of the olefins obtained to produce middle distillates.

The present invention allows the production of middle distillates with a yield and a maximum selectivity compared to a conventional hydrocracking process one or two stages.

The present invention also makes it possible to produce a middle distillate cut and in particular diesel which respects the commercial specifications of road diesel in terms of sulfur content, density and cetane index, which is not the case for the implementation of a sequence of an FCC step followed by an oligomerization unit according to the prior art.

Prior art

Patent LR 2,968,010 B1 describes a chaining process comprising a catalytic cracking step and an oligomerization step. This application describes a process for converting a heavy charge making it possible to improve the selectivity to middle distillate. The process uses a catalytic cracking unit followed by one or more oligomerization units of olefins with a number of carbon atoms ranging from C2 to C9 making it possible to preferentially produce an additional medium distillate cut. The light part of the oligomerate produced and which cannot be incorporated into the middle distillate cut is partially recycled,

- Either at the oligomerization stage to be transformed into middle distillais by reaction with the light olefins of the feed, as described in patent LR 2 871 167.

- Either at the LCC to be cracked into light olefins which return to the oligomerization units in addition to the olefins in the feed in order to preferably form heavy oligomers which can be incorporated into the middle distillate cut.

Patent application WO 13/093227 describes an improved process for converting a heavy load making it possible to improve the selectivity in middle distillates. The invention consists in the implementation of a hydrotreatment or hydroconversion process known as of pretreatment of the heavy feed upstream of the whole catalytic cracking unit followed by an oligomerization unit which makes it possible to increase the net production of cuts C3, C4 and light petrol at the end of the catalytic cracking unit (LCC in English), and consequently very significantly increase the production of middle distillate after the oligomerization step , as well as the selectivity for middle distillate on petrol. Thus, the charge of the catalytic cracking unit consists of the unconverted charge of the pretreatment unit which has substantially the same distillation range as the heavy input charge, i.e. a temperature d 'boiling above about 340 ° C.

The research carried out by the applicant led him to discover that the implementation of a specific sequence made it possible to significantly improve the yield of medium distillates and the selectivity of medium distillate compared to the gasoline cut.

Summary of the invention

In particular, the present invention relates to a process for the production of middle distillates comprising at least the following steps:

A) Hydro-treatment (HDT) of hydrocarbon feeds containing at least 20% by weight and preferably at least 80% by weight of boiling compounds above 340 ° C., in the presence of hydrogen and at least a hydrotreating catalyst, at a temperature between 200 and 450 ° C, under a pressure between 2 and 18

MPa, at a space speed between 0.1 and 6 h -1 and at a quantity of hydrogen introduced such that the volume ratio of liter of hydrogen / liter of hydrocarbon is between 100 and 2000 NL / L,

B) Hydrocracking (HCK) of at least part of the effluent from step a), step b) of hydrocracking operating, in the presence of hydrogen and at least one hydrocracking catalyst and at a temperature between 250 and 480 ° C, under a pressure between 2 and 25 MPa, at a space velocity between 0.1 and 6 h -1 and at an amount of hydrogen introduced such that the volume ratio of liter of hydrogen / liter of hydrocarbon is between 80 and 5000 NL / L,

C) Separation of the effluent (SEP) from step b) to obtain at least one naphtha cut having a distillation range between 20 ° C and 220 ° C, a medium distillate cut and an effluent unconverted liquid with an initial boiling point above 340 ° C,

D) sending at least part of the naphtha cut having a distillation range of between 20 ° C and 220 ° C in a steam cracking step to produce at least one olefinic effluent having a boiling point final below 220 ° C, said steam cracking step being carried out at a reactor outlet temperature between 700 ° C and 900 ° C, at a pressure between 0.1 and 0.5 MPa during a residence time included between 0.005 and 2 s and at a water vapor and charge mass flow ratio (H ₂ O / HC) is between 0.1 and 2,

E) selective hydrogenation (SHU) of at least part of the olefinic effluent having a final boiling point below 220 ° C from step d) to convert the diolefins to mono-olefins , said step f) being implemented at a temperature between 0 ° C and 200 ° C, a pressure between 0.1 and 5 MPa and at a WH between 0.5 and 8 h -1,

F) the oligomerization (OLG) of at least a part of the olefinic effluent from step e) to produce at least one olefinic hydrocarbon effluent having a number of carbon atoms greater than or equal to 8 , said oligomerization step f) being carried out at a temperature between 100 ° C and 350 ° C, at a pressure between 1 and 10 MPa and in the presence of an oligomerization catalyst,

G) Total hydrogenation (HYD) of at least part of the middle distillate cut from step c) to transform the olefins into alkanes of the middle distillate cut, said step g) being implemented at a temperature between 100 ° C and 400 ° C, a pressure between 1 and 10 MPa, a WH between 0.5 and 10 h-1 and with a volume ratio of hydrogen / hydrocarbon is between 100 and 2000 NL / L and in the presence of a hydrogenation catalyst.

A synergy resulting from the specific sequence according to the present invention is to allow both, the production of medium distillates with a yield and a selectivity maximized compared to a hydrocracking process one or two conventional stages and respecting the commercial specifications of road diesel in terms of sulfur content, density and cetane number, which is not the case with the implementation of a sequence comprising an FCC stage as a production stage of olefin according to the prior art.

Another advantage of the present invention is to allow, in addition to very clearly improving the selectivity of middle distillate compared to the gasoline cut, to allow obtaining a gasoline cut with reduced content of impurities such as sulfur, nitrogen as well as in diolefins and which is then sent to the oligomerization stage in order to be converted into olefins of higher molecular weight corresponding to a middle distillate cut, which has an undeniable advantage in terms of performance and lifetime of the oligomerization catalyst.

Throughout the text, a naphtha cut corresponds to a hydrocarbon cut with a distillation interval between 20 ° C and 220 ° C (corresponding to a C5-C12 cut having a number of carbon atoms between 5 and 12) preferably between 20 ° C and 160 ° C (corresponding to a C5-C9 cut having a number of carbon atoms between 5 and 9).

A middle distillate cut corresponds to a hydrocarbon cut with a distillation interval between 130 ° C and 380 ° C, and preferably between 150 ° C -370 ° C.

The residue cut here called UCO corresponds to a hydrocarbon cut with a distillation start point greater than 370 ° C., preferably greater than 380 ° C. Description of the embodiments

Charges

The present invention relates to a hydrocracking process in one step of hydrocarbon feeds called mother feed (stream 1), containing at least 20% volume and preferably at least 80% volume of compounds boiling above 340 ° C. , preferably above 350 ° C and preferably between 340 and 580 ° C (that is to say corresponding to compounds containing at least 15 to 20 carbon atoms).

Said hydrocarbon feedstocks can advantageously be chosen from VGOs (Vacuum gas oil) according to English terminology or vacuum distillates (DSV) such as for example heavy gas oils obtained from the direct distillation of crude oil or from conversion such as FCC such as LCO or Light Cycle Oil according to English terminology, coking or visbreaking as well as feedstocks originating from aromatic extraction units from lubricating oil bases or from dewaxing at solvent for lubricating oil bases, or still distillates from desulphurization or hydroconversion of RAT (atmospheric residues) and / or RSV (vacuum residues), or else the filler can advantageously be a deasphalted oil, or fillers from biomass or any mixture of the charges mentioned above and preferably VGOs.

Paraffins from the Fischer-Tropsch process are excluded.

In general, said fillers have a boiling point T5 greater than 340 ° C, and better still greater than 370 ° C, that is to say that 95% of the compounds present in the charge have a boiling point greater than 340 ° C, and better still greater than 370 ° C.

The nitrogen content of the mother feeds treated in the process according to the invention is usually greater than 500 ppm by weight, preferably between 500 and 10,000 ppm by weight, more preferably between 700 and 4,000 ppm by weight and again more preferred between 1000 and 4000 ppm weight. The sulfur content of the mother feeds treated in the process according to the invention is usually between 0.01 and 5% by weight, preferably between 0.2 and 4% by weight and even more preferably between 0.5 and 3% weight.

The load can optionally contain metals. The cumulative nickel and vanadium content of the charges treated in the process according to the invention is preferably less than 1 ppm by weight.

The asphaltenes content is generally less than 3000 ppm by weight, preferably less than 1000 ppm by weight, even more preferably less than 200 ppm by weight.

The charge may optionally contain asphaltenes. The asphaltenes content is generally less than 3000 ppm by weight, preferably less than 1000 ppm by weight, even more preferably less than 200 ppm by weight.

In the case where the filler contains compounds of the resins and / or asphaltenes type, it is advantageous to pass the filler beforehand over a bed of catalyst or adsorbent different from the hydrocracking or hydrotreating catalyst.

A) Hydrotreating stage (HDT)

According to the invention, the method comprises a step a) of hydrotreating said charges (flow 1) in the presence of hydrogen (flow 2) and at least one hydrotreatment catalyst, at a temperature between 200 and 450 ° C, under a pressure between 2 and 18 MPa, at a space speed between 0.1 and 6 h-1 and at a quantity of hydrogen introduced such as the volume ratio liter of hydrogen / liter d 'hydrocarbon is between 100 and 2000 NL / L.

The operating conditions such as temperature, pressure, hydrogen recycling rate, hourly space velocity, can be very variable depending on the nature of the feed, the quality of the desired products and the facilities available to the refiner.

Said step a) can advantageously be implemented in a fixed bed or in bubbling ht.

Preferably, step a) of hydrotreatment according to the invention operates at a temperature between 250 and 450 ° C, very preferably between 300 and 430 ° C, under a pressure between 5 and 16 MPa , at a space speed between 0.2 and 5 h -1, and at a quantity of hydrogen introduced such that the volume ratio liter of hydrogen / liter of hydrocarbon is between 300 and 1500 NL / L.

Conventional hydrotreatment catalysts can advantageously be used, preferably which contain at least one amorphous support and at least one hydro-dehydrogenating element chosen from at least one element from the non-noble groups VIB and VIII, and most often at at least one element from group VIB and at least one element from group VIII which is non-noble.

Preferably, the amorphous support is alumina or alumina silica.

Preferred catalysts are chosen from NiMo, NiW or CoMo catalysts on alumina and NiMo or NiW on silica alumina.

The effluent from the hydrotreatment step and entering step a) of hydrocracking generally comprises a nitrogen content preferably less than 300 ppm by weight and preferably less than 50 ppm by weight.

B) Hydrocracking stage (HCK)

According to the invention, the method comprises a step b) of hydrocracking at least part of the effluent (stream 3) from step a), and preferably all of said step b ) operating, in the presence of hydrogen and at least one hydrocracking catalyst, at a temperature between 250 and 480 ° C, under a pressure between 2 and 25 MPa, at a space speed between 0.1 and 6 h -1 and at a quantity of hydrogen introduced such that the volume ratio liter of hydrogen / liter of hydrocarbon is between 80 and 5000 NL / L.

The operating conditions such as temperature, pressure, hydrogen recycling rate, hourly space velocity, may be very variable depending on the nature of the feed, the quality of the desired products and the facilities available to the refiner.

Said step b) can advantageously be implemented in a fixed bed or in a bubbling bed.

Preferably, step b) of hydrocracking according to the invention operates at a temperature between 320 and 450 ° C, very preferably between 330 and 435 ° C, under a pressure between 3 and 20 MPa , at a space speed between 0.2 and 4 h -1, and at a quantity of hydrogen introduced such that the volume ratio of liter of hydrogen / liter of hydrocarbon is between 200 and 2000 NL / L.

These operating conditions used in the process according to the invention generally make it possible to achieve conversions by pass, into products having boiling points below 340 ° C, and better still below 370 ° C, greater than 15%. weight and even more preferably between 20 and 95% by weight.

Step b) covers the pressure and conversion domains ranging from mild hydrocracking to high pressure hydrocracking. Mild hydrocracking is understood to mean hydrocracking leading to moderate conversions, generally less than 40%, and operating at low pressure, preferably between 2 MPa and 6 MPa. High pressure hydrocracking is generally carried out at higher pressures between 5MPa and 20MPa, so as to obtain conversions greater than 50%.

Step a) of hydrotreating and step b) of hydrocracking can advantageously be carried out in the same reactor or in different reactors. In the case where they are produced in the same reactor, the reactor comprises several catalytic beds, the first catalytic beds comprising the hydrotreating catalyst (s) and the following catalytic beds comprising the hydrocracking catalyst (s).

Catalyst for step b) of hydrocracking

Or the hydrocracking catalysts used in hydrocracking step b) are conventional hydrocracking catalysts, of bifunctional type combining an acid function with a hydrogenating function and optionally at least one binder matrix.

Preferably, the hydrocracking catalyst (s) comprise at least one group VIII metal chosen from iron, cobalt, nickel, ruthenium, rhodium, palladium and platinum and preferably cobalt and nickel and / or at least one metal from group VIB chosen from chromium, molybdenum and tungsten, alone or as a mixture and preferably from molybdenum and tungsten.

Hydrogenating functions of the NiMo, NiMoW, NiW type are preferred. Catalysts known to those skilled in the art may be suitable.

Preferably, the group VIII metal content in the hydrocracking catalyst (s) is advantageously between 0.5 and 15% by weight and preferably between 2 and 10% by weight, the percentages being expressed as a percentage by weight d 'oxides.

Preferably, the metal content of group VIB in the hydrocracking catalyst (s) is advantageously between 5 and 25% by weight and preferably between 15 and 22% by weight, the percentages being expressed as a percentage by weight of oxides .

The catalyst or catalysts can also optionally at least one promoter element deposited on the catalyst and chosen from the group formed by phosphorus, boron and silicon, optionally at least one element from group VIIA (chlorine, fluorine preferred), and optionally at least one element of group VIIB (preferred manganese), optionally at least one element of group VB (preferred niobium).

Preferably, the hydrocracking catalyst (s) comprise a zeolite chosen from US Y zeolites, alone or in combination, with other zeolites from beta zeolites, ZSM-12, IZM-2, ZSM-22, ZSM-23, SAPO-11, ZSM-48, ZBM-30, alone or as a mixture. Preferably the zeolite is the USY zeolite alone.

The hydrocracking catalyst (s) may optionally comprise at least one porous or poorly crystallized mineral matrix of the oxide type chosen from aluminas, silicas, silica-aluminas, aluminates, boron alumina-oxide, magnesia, silica-magnesia, zirconia, titanium oxide, clay, alone or as a mixture and preferably alumina and silica-alumina.

An organic agent can be present on the catalyst of step b).

A preferred catalyst comprises and preferably consists of at least one group VI metal and / or at least one non-noble group VIII metal, a USY zeolite and an alumina binder.

An even more preferred catalyst comprises and preferably consists of nickel, molybdenum, phosphorus, a USY zeolite and alumina.

Another preferred catalyst comprises and preferably consists of nickel, tungsten, a USY zeolite and alumina or silica alumina.

In general, the catalyst (s) used in hydrocracking step b) advantageously contains:

0.1 to 60% by weight of zeolite,

- 0.1 to 40% by weight of at least one element from groups VIB and VIII (% oxide)

- 0.1 to 99.8% by weight of matrix (% oxide)

- 0 to 20% by weight of at least one element chosen from the group formed by P, B, Si (% oxide), preferably 0.1-20%

- 0 to 20% by weight of at least one element of group VIIA, preferably 0.1 to 20%

- 0 to 20% by weight of at least one element of group VIIB, preferably 0.1 to 20%

- 0 to 60% by weight of at least one element of the group VB, preferably 0.1 to 60%;

The percentages being expressed as a percentage by weight relative to the total mass of catalyst, the sum of the percentages of the elements constituting said catalyst being equal to 100%

The hydrocracking catalyst (s) used according to the invention are preferably subjected to a sulphurization treatment making it possible to transform, at least in part, the metallic species into sulphide before they are brought into contact with the charge to be treated. This activation treatment by sulfurization is well known to those skilled in the art and can be carried out by any method already described in the literature.

C) Separation step (SEP)

According to the invention, the method comprises a step c) of separation of at least the effluent from step b) (flow 4) to obtain at least one naphtha cut (flow 7) having an interval distillation between 20 ° C and 220 ° C, a middle distillate cut (flow 8) having a distillation range between 130 ° C and 380 ° C and an unconverted liquid effluent (flow 9) having a boiling point initial greater than 340 ° C.

The separation step generally comprises a primary separation of the gas and liquid effluents, a hydrogen recycling section associated with a compression step, as well as distillations for the fractionation of the various hydrocarbon fractions. Preferably, the gas / liquid separation is implemented in a high pressure high temperature separator operating at a temperature between 50 and 450 ° C, preferably between 100 and 400 ° C, even more preferably between 200 and 300 ° C, and a pressure corresponding to the outlet pressure from hydrocracking stage b) minus the pressure drops. It allows the separation of a gaseous effluent (stream 5) comprising at least hydrogen and a hydrocarbon effluent.

Preferably, the method comprises sending the gaseous effluent comprising at least hydrogen in a compression step before its recycling in at least step b) of hydrocracking. This step allows the gas to be recycled upstream, that is to say at higher pressure. The gaseous effluent comprising at least hydrogen can advantageously be mixed with make-up hydrogen before or after its introduction into the compression step, preferably via a make-up or make-up hydrogen compressor according to Anglo-Saxon terminology.

According to a variant, part of the gaseous effluent comprising at least compressed hydrogen can also advantageously be sent in steps b) of hydrocracking, a) of hydrotreatment, e) of selective hydrogenation and / or g) total hydrogenation.

Preferably, said separation step c) comprises a separation means such as for example a separator flask or a steam stripper preferably operating at a pressure between 0.5 and 2 MPa, which has for purpose of carrying out a separation of hydrogen sulfide (H ₂ S) from at least one hydrocarbon effluent produced during step b) of hydrocracking. The hydrocarbon effluent from this first separation can advantageously undergo atmospheric distillation, and in certain cases the combination of atmospheric distillation and vacuum distillation. The purpose of distillation is to separate the converted hydrocarbon products, that is to say generally having boiling points below 380 ° C, preferably below 370 ° C and preferably below 340 ° C and an unconverted liquid fraction (residue) (LJCO) having boiling points above 340 ° C, preferably above 370 ° C and preferably above 380 ° C. According to another variant, the fractionation step consists only of an atmospheric distillation column.

The converted hydrocarbon products having boiling points below 380 ° C, preferably below 370 ° C and preferably below 360 ° C are advantageously distilled at atmospheric pressure to obtain several converted fractions and preferably at minus a light gas cut C1-C4 (flow 6) having a boiling point below 20 ° C, a naphtha cut (flow 7) having a distillation range between 20 ° C and 220 ° C, and preferably between 20 ° C and 160 ° C, a medium distillate cut (flow 8) and an unconverted liquid effluent (flow 9) having an initial boiling point greater than 340 ° C, preferably greater than 370 ° C and preferably greater at 380 ° C also called UCO or "unconverted oil" according to English terminology.

The naphtha cut (flow 7) from step c) of separation is a hydrocarbon cut with a distillation interval between 50 ° C and 220 ° C (C5-C12), preferably between 50 ° C and 160 ° C (C5-C9).

The middle distillery cut (flow 8) separated in step c) can advantageously be separated into a diesel and kerosene cut to be sent to the fuel pools (jet fuel and diesel).

According to the invention, at least part and preferably all of the naphtha cut having a distillation range between 20 ° C and 220 ° C from step c) is sent to step d ) steam cracking to produce olefins.

In a variant of the invention, at least part of the unconverted liquid effluent (stream 9c) in step b) of hydrocracking and separated in step c), having a boiling point initial higher than 340 ° C preferably higher than 370 ° C and preferably higher than 380 ° C also called UCO or "unconverted oil" according to English terminology can also be sent in step d) of steam cracking. In this case, the feed from step d) of steam cracking comprises and preferably consists of at least part of the liquid effluent not converted in step b) of hydrocracking and separated in step c) , having an initial boiling point greater than 340 ° C in admixture with at least part and preferably the whole of the naphtha fraction having a distillation range of between 20 ° C and 220 ° C.

In another variant of the invention, at least part and preferably all of the unconverted liquid effluent (stream 9b) from step c) can be recycled in step b) of hydrocracking.

In another variant of the invention, at least part and preferably all of the unconverted liquid effluent (stream 9b) from step c) can be recycled in step a) of hydrotreating.

In another variant of the invention, at least part and preferably all of the unconverted liquid effluent (stream 9b) from step c) can be sent in a second hydrocracking step b ') (not shown in Figure 1) operating in the presence of hydrogen and at least one second hydrocracking catalyst, at a temperature between 250 and 480 ° C, under a pressure between 2 and 25 MPa, a space velocity between 0.1 and 6 h -1 and at a quantity of hydrogen introduced such that the volume ratio liter of hydrogen / liter of hydrocarbon is between 100 and 2000 NL / L.

Preferably, said second hydrocracking step is carried out under identical or different conditions to those of said first hydrocracking step b) and in the presence of a hydrocracking catalyst identical or different from that used. work in said step b).

Preferably, said second hydrocracking step d) operates, in the presence of hydrogen and at least one hydrocracking catalyst, at a temperature between 320 and 450 ° C and very preferably between 330 and 435 ° C, under a pressure between 3 and 20 MPa, at a space velocity between 0.2 and 3 h ¹ and at a quantity of hydrogen introduced such that the volume ratio liter of hydrogen / liter of hydrocarbon is between 200 and 2000 NL / L.

The description of the catalyst used in said second hydrocracking step is identical to that of the catalyst used in said first hydrocracking step b).

Preferably, the catalyst used in step b ’) is different from that of step b). The effluent produced in the second hydrocracking step b ') can advantageously, in whole or in part and preferably in whole, be sent in the separation step c) so as to be separated into at least said three effluents according to step c).

A purge is advantageously carried out on the recycled heavy fraction. Said purging of part of the recycled heavy fraction can be carried out to limit the formation of heavy aromatic products or HPNA (Heavy PolyNuclear Aromatics in English), Said purge is between 0 to 5% by weight of the heavy fraction relative to the incoming DSV mother charge, and preferably between 0 to 3% by weight.

Said second step b ’) of hydrocracking can advantageously be implemented in a reactor or in separate reactors, each reactor advantageously comprising one or more catalytic beds.

The catalysts used in the different catalytic beds can advantageously be the same or different.

The operating temperatures of the various reactors and / or catalytic beds can also be identical or different.

D) Steam cracking stage (SC)

According to the invention, the method comprises sending at least part and preferably all of the naphtha cut (stream 7) having a distillation interval between 20 ° C and 220 ° C in a steam cracking step to produce at least one olefinic effluent (stream 11) having a final boiling point below 220 ° C, said steam cracking step being carried out at a reactor outlet temperature (ROT - Reactor Outlet temperature) between 700 ° C and 900 ° C, at a pressure between 0.1 and 0.5 MPa during a residence time between 0.005 and 2 s and at a water vapor and charge mass flow ratio (H ₂ O / HC) is between 0.1 and 2.

At least part and preferably all of the naphtha cut having a distillation range of between 20 ° C and 220 ° C from step c) is sent to the steam cracking step to produce olefins in accordance to the invention.

Preferably the steam cracking step is carried out at a pressure between 0.1 and 0.2 MPa during a residence time between 0.02 and 2 s.

According to the invention, step d) of steam cracking makes it possible to produce at least one olefinic effluent (stream 12) having a final boiling point below 220 ° C. Preferably, said olefinic effluent (12) is an effluent having a boiling point range of between -80 ° C and 220 ° C and preferably between -50 ° C and 160 ° C. Said effluent corresponds to the mixture of the C3-C4 gas cut and the naphtha cut consisting mainly of mono-olefins, aromatics and possibly diolefins and having a distillation range between 20 ° C and 220 ° C.

In particular, step d) of steam cracking advantageously allows, in addition to the production of said olefinic effluent (12) having a final boiling point below 220 ° C, the production of a dry gas cut (flow 10) consisting of hydrogen, methane, ethane, and carbon monoxide (CO), an ethylene cut (flow 11) and a middle distillate cut (flow 13).

The dry gas cut (flow 10) can advantageously be sent to the hydrogen recycling section of section c) or be used as fuel in the refinery.

The ethylene cut (stream 11) can advantageously be used as a primary material in the petrochemical industry.

The middle distillate cut (flow 13) can advantageously be sent in step g) of final hydrogenation.

E) Selective hydrogenation stage (SHU)

According to the invention, the method comprises the selective hydrogenation of at least part and preferably all of the olefinic effluent (stream 12) having a final boiling point below 220 ° C from step d) to convert the diolefins to olefins, said step e) being carried out at a temperature between 0 ° C and 200 ° C, a pressure between 0.1 and 5 MPa and at a WH between 0 , 5 and 8 a.m. ¹ .

Preferably, said step e) is implemented at a pressure between 0.5 and 3 MPa.

Preferably, said step e) is carried out in the presence of a conventional selective hydrogenation catalyst, preferably of the metallic type based on palladium, platinum or nickel supported on an alumina, a silica or a silicealumina.

F) Oligomerization step (OLG)

According to the invention, the method comprises the oligomerization of at least a part and preferably all of the olefinic effluent (flow 14) from step e) to produce at least one medium distillate effluent (flow 17), said step e) of oligomerization being carried out at a temperature between 100 ° C and 350 ° C, at a pressure between 1 and 10 MPa and in the presence of an oligomerization catalyst.

The operating temperature is preferably between 150 ° C and 270 ° C. The operating pressure is preferably between 2 and 6 MPa, more preferably between 4 and 5 MPa.

The oligomerization can be carried out in one or more stages, with one or more reactors arranged in parallel or in series, and one or more catalysts.

The following description of the catalyst and the operating conditions can be applied to any one of the stages and / or to any one of the reactors.

The solid oligomerization catalyst operating in the heterogeneous phase is chosen from those known from the prior art. The particular conditions of the various stages of the process according to the invention are described below in more detail.

The oligomerization catalysts used are preferably acid catalysts based on alumina silica or amorphous alumina-silica of crystallized zeolite or resins. Among the zeolites, preference will be given to those described in patent PR 2,894,850.

The object of the oligomerization step f) is to oligomerize the olefins at C3 to C12 which may be formed from several disjointed cuts or not such as for example a cut C3 to C4 and a naphtha cut C5-C12, corresponding to the hydrocarbon fraction with a distillation interval between 20 ° C and 220 ° C, alone or as a mixture to obtain a mixture of hydrocarbons containing mono-olefins and corresponding to the middle distillary fraction (flow 17). Typically, from C4 olefins, oligomers are obtained whose number of carbon atoms is largely less than or equal to 30, and for the most part between 8 and 22 and preferably between 10 and 22 .

Said step f) of oligomerization allows the production of at least one medium distillate hy drocarbon effluent (stream 17). In particular, said oligomerization step f) also produces, in addition to the middle distillery cut (flow 17), advantageously a C3 / C4 cut (flow 15), and a naphtha cut (flow 16) C5-C9 (corresponding to the hydrocarbon fraction with a distillation interval between 20 ° C. and 160 ° C.) which can advantageously be sent to the petrol pool.

The different sections can advantageously be separated downstream from the oligomerization step f).

Part of the C5-C12 naphtha cut (stream 16) produced by the oligomerization unit, and preferably light petrol C5-C9 can optionally be recycled to the oligomerization reactor to improve the yield in middle distillate.

G) Total hydrogenation stage (HYD)

According to the invention, the method comprises a step g) of total hydrogenation of at least part and preferably all of the middle distillate cut (flow 17) from step f) to produce a hydrogenated medium distillate cut, said step g) being carried out at a temperature between 100 ° C and 400 ° C, a pressure between 1 and 10 MPa, a WH between 0.5 and 10 h-1 and with a ratio volume of hydrogen / hydrocarbon is between 100 and 2000 NL / L and in the presence of a hydrogenation catalyst.

Preferably, the final hydrogenation catalyst is chosen from partially or totally sulfur-based catalysts based on Ni or NiMo or CoMo or even Ni CoMo and NiW, supported on a support of refractory oxide type (Al, Ti, Si) optionally with one or more promoters (F, P, Ca, Na).

Said step g) aims to transform the olefins into alkanes. It is carried out under the operating and catalyst conditions known to those skilled in the art.

List of Figures

FIG. 1 represents a diagram of the process according to the invention in which the units are noted:

[0132] (HDT) for 1 hydro treatment

[0133] (HCK) for hydrocracking

[0134] (SEP) for the fractionation of hydrocracking products

[0135] (SC) for the Steam Cracking unit

[0136] (SHU) for the selective diolefin hydrogenation unit,

(OLG) for said oligomerization unit and,

(HYD) for the total hydrogenation unit of the middle distillate cut resulting from steam cracking and oligomerization.

The process according to the present invention can be defined as a process for the production of middle distillate from a heavy load (1) of the gas oil type under vacuum using 7 steps in series a) a hydrotreatment step ( HDT) produce a vacuum distillate cut with a low nitrogen content (3) from the heavy feed (1) with an addition of hydrogen (2), b) a hydrocracking step (HCK) which converts the heavy cut hydrotreated (3) in a C1-C4 cut (6), a naphtha C5-160 ° C cut (7), and a middle distillate cut (8). At the end of this stage is also recovered a dirty gas cut (5) consisting mainly of hydrogen, hydrogen disulfide, ammonia and an unconverted fraction (9) which has substantially the same distillation range as the feed heavy entry. The naphtha cut (7) resulting from step c) is sent in a step d) of steam cracking to produce light olefins. The products of this step are a dry gas cut (10) consisting of hydrogen, methane, ethane and carbon monoxide (CO), an ethylene cut (11), and a cut having a point of final boiling below 220 ° C (12) comprising a C3-C4 cut and the naphtha cut (C5-160 ° C) consisting mainly of olefins, paraffins and possibly diolefins. A middle distillate cut (13) is also produced and sent to a total hydrogenation unit (HYD) to remove the olefins. The cut having a final boiling point below 220 ° C (12) from the steam cracking unit is sent in a selective hydrogenation (SHU) step e) to convert the diolefins to olefins. The olefinic effluent obtained (14) at the end of the selective hydrogenation unit (SHU) is then sent to an oligomerization step f) (OLG) which produces a C3 / C4 cut (15), a cut gasoline C5-160 ° C (16) and a middle distillate cut (17) which is sent to a total hydrogenation unit (HYD) to produce a middle distillate cut (18) without olefins.

Examples

Two examples are provided below to illustrate the improved performance of the process according to the invention compared to the process according to the prior art.

Comparative Example 1: One Step Hydrocracking

This first example constitutes the basic case and corresponds to a conventional hydrocracking unit known as in one step.

The charge of the process according to the invention is a vacuum distillate which has the characteristics given in Table 1.

[0144]

[Tables 1]

Charge Density 0.9237 g / ml Sulfur 2.41 % wt Nitrogen 836 ppm wt TMP 485 ° C % 370 + 93 % Amount 100 BASE 100

Table 1. Characteristic of the VGO load of the method of Example 1

The feed 1 is mixed with the hydrogen flow (flow 2) and then the mixture is introduced into the hydrocracking section A comprising the hydrotreating step (HDT) and hydrocracking step (HCK) with the following operating conditions:

[0147] [Tables2]

VGO hydrocracking stage (A) Hydrotreating section catalyst NiMo on alumina Conversion section catalyst NiMo + zeolite on alumina VGO charge rate 100 (BASE 100) t / h Partial pressure of outgoing hydrogen 130 bars g H2 / process load ratio 1280 L / L WH hydrotreating section 1.0 h-1 WH conversion section 1.0 h-1 Hydrotreatment section temperature 396 ° C Temperature section conversion 370 ° C

Table 2. Operating conditions of the HCK of Example 1

Under these conditions, the yields of the hydrocracked VGO output products relative to the process charge are given in the table below:

[0150]

[Tables3]

Destination Section C1-C4 2.4 % wt To products Naphtha (PI - 150 ° C) 14.0 % wt To products Medium distillais (150 ° C- 370 ° C) 50.0 % wt To products Hydrotreated VGO (370 ° C +) 33.2 % wt To products

Table 3. Yields of the products of the hydrocracking unit of Example 1

Example 2 according to the invention: Chain of a hydrocracking unit followed by a steam cracking unit (SC) with an oligomerization unit

In Example 2, the hydrocracking unit is supplied with the same load as that of Example 1 and operates under the same operating conditions. The yields of the products released from hydrocracking with respect to the process charge are identical to those of Example 1:

[0154] [Tables4]

Destination Section C1-C4 2.4 % wt To products Naphtha (PI - 150 ° C) 14.0 % wt To steam cracking Distillais middle (150 ° C370 ° C) 50.0 % wt To products Hydrotreated VGO (370 ° C +) 33.2 % wt To products

Table 4. Product yields of the hydrocracking unit of Example 2

For example 2, the naphtha cut (PI-150) is sent to a steam cracking unit (SC) and then to the selective hydrogenation unit (SHU) to produce the olefins necessary for the step of oligomerization. The operating conditions of the steam cracking and selective hydrogenation unit are presented in Table 5.

[0157]

[Tables5]

SC and SHU unit Fuel charge flow 14.0 (BASE 100) t / h Steam cracking Pressure 0.15 MPa H2O / load ratio 1.0 L / L Residence time 0.475 s Reactor outlet temperature 850 ° C Selective hydrogenation Catalyst LD271 (Axens) Temperature 50 ° C Pressure 10 bar g WH 5 h-1

Table 5. Operating conditions of the steam cracking and selective hydrogenation unit of the gasoline cut of Example 2

Under these conditions, the yields of the products obtained on the steam cracking and selective hydrogenation unit with respect to the process charge (flow 1) are given below:

[0160] [Tables6]

Destination C1-C2 4.3 % wt To products C3-C4 3.7 % wt Towards oligomerization Naphtha (PI - 150 ° C) 5.1 % wt Towards oligomerization Medium distillais (150 ° C- 370 ° C) 0.8 % wt To products

Table 6. Product yields of the steam cracking unit (SC) and of selective hydrogenation in Example 2

The C3-C4 cuts are mainly composed of olefinic species, such as propylene and butene as well as the naphtha cut (PI-150 ° C) are then sent to the oligomerization unit. The operating conditions of the oligomerization unit are presented in Table 7.

[0163]

[Paintings?]

Oligomerization unit (Section E) Catalysts IP811 (Axens) Load flow (C3-150 ° C) 9.2 (BASE 100) t / h Pressure 50 bar g Temperature 200 ° C Quantity of catalyst 25.0 t Catalysts IP811 (Axens)

Table 7. Operating conditions of the oligomerization step of Example 2

The middle distillate cut resulting from the oligomerization is then hydrogenated with a sulfurized nickel catalyst (LD 541) at 160 ° C. under 50 bar of hydrogen, and with a WH of 1.5 h-1 and a flow rate d hydrogen gas based on the liquid charge rate of 100 NL / L

The yields of the products from hydrocracking, steam cracking, selective hydrogenation and oligomerization in relation to the process charge are given below:

[0167] [Tables 8]

Origin C1-C2 4.7 % wt Hydrocracking and steam cracking C3-C4 2.6 % wt Hydrocracking and steam cracking + Oligomerization Naphtha (PI - 150 ° C) 5.3 % wt Steam cracking + Oligomerization Distillais middle (150 ° C370 ° C) 54.0 % wt Hydrocracking and steam cracking + Oligomerization Hydrotreated VGO (370 ° C +) 33.2 % wt Hydrocracking

Table 8. Yields of the products resulting from the sequence of procedures of Example 2

The sequencing of the hydrocracking unit, steam cracking and selective hydrogenation of gasolines with an oligomerization unit of cuts C3-150 ° C, improves the amount of average distillates produced.

In the basic case, this production of middle distillates represents 50.0% by weight of the process charge against 54.0% by weight in the configuration according to the present invention, an increase of 4%.

The naphtha cut, for its part, decreased by 8.7%, going from 14.0% by weight in the basic case to 5.3% by weight.

The selectivity in middle distillais compared to naphtha (average distillais / naphta ratio) is therefore improved, going from 3.6 to 10.2, which is a considerable improvement, given the evolution of the fuel market.

Example 3 according to the invention: Chain of a hydrocracking unit followed by a steam cracking unit with an oligomerization unit (charge 2)

In Example 3, the hydrocracking unit is supplied by the feed 2 which corresponds to the vacuum distillate used in Example 2 of patent WO 13093227. The properties are recalled in the table below:

[0175] [Tables9]

Charge Density 0.9149 g / cc Sulfur 1.22 % wt Nitrogen 0.127 % wt TMP 481 ° C % 360 + 80 % Amount 100 BASE 100

Table 9. Characteristic of the VGO load of the method of Example 3

Load 2 is mixed with the hydrogen flow and then the mixture is introduced into the hydrocracking section A comprising the hydrotreatment (HDT) and hydrocracking (HCK) stage with the operating conditions of l 'example 1. The yields of the hydrocracked products with respect to the process charge are given in the table below:

[0178] [Tables10]

Destination Section C1-C4 2.4 % wt To products Naphtha (PI - 160 ° C) 14.7 % wt To steam cracking Kerosene (160 ° C- 220 ° C) 11.6 % wt To products Diesel (220 ° C- 360 ° C) 35.2 % wt To products Hydrotreated VGO (360 ° C +) 36.5 % wt To products

Table 10. Yields of the products of the hydrocracking unit of Example 3

For example 3, the naphtha cut (PI-160) is sent to a steam cracking unit (SC) then to the selective hydrogenation unit (SHU) under the same operating conditions of example 2. In these conditions, the yields of the products obtained on the unit of

SC and selective hydrogenation in relation to the process charge are given below:

[0181] [Tablesll]

Destination C1-C2 4.6 % wt To products C3-C4 3.9 % wt Towards oligomerization Naphtha (PI - 160 ° C) 5.4 % wt Towards oligomerization Kerosene (160 ° C- 220 ° C) 0.8 % wt To products

Table 11. Product yields of the steam cracking unit (SC) and selective hydrogenation of Example 3

The C3-C4 sections and the naphtha section (PI-160 ° C) were then converted in the oligomerization unit under the same conditions of Example 2. The middle distillary section (kerosene + diesel) from the oligomerization is then hydrogenated with a sulfurized nickel catalyst (LD 541) at 160 ° C under 50 bar of hydrogen, and with a WH of 1.5h-1 and a flow of hydrogen gas compared to the flow of liquid charge from 100 NL / L

The yields of the products from the hydrocracking, steam cracking, selective hydrogenation and oligomerization sequence with respect to the process charge are given below:

[0185] [Tablesl2]

Origin C1-C2 5.0 % wt Hydrocracking and steam cracking C3-C4 2.6 % wt Hydrocracking and steam cracking + Oligomerization Naphtha (PI - 160 ° C) 5.8 % wt Steam cracking + Oligomerization Kerosene (160 ° C- 220 ° C) 13.5 % wt Hydrocracking and Steam cracking + Diesel (220 ° C- 360 ° C) 36.7 % wt Oligomerization Hydrotreated VGO (360 ° C ⁺ ) 36.7 % wt Hydrocracking

Table 12. Yields of the products resulting from the sequence of procedures of Example 3

The sequence of the hydrocracking unit and the steam cracking and selective hydrogenation of gasolines with an oligomerization unit of cuts C3-150 ° C, improves the amount of average distillates produced.

In the case of the hydrocracking unit alone (Table 10), this production of middle distillates (diesel + kerosene) represents 46.8% by weight of the process charge against 50.2% by weight in the configuration according to the present invention, an increase of 3.4%.

The naphtha cut, for its part, decreases by 8.9%, going from 14.7% by weight in the basic case to 5.8% by weight.

The selectivity in middle distillais compared to naphtha (average distillais / naphta ratio) is therefore improved, going from 3.2 to 8.7, which is a considerable improvement, having regard to the development of the fuel market.

Comparative Example 4: Sequence of a hydrocracking unit followed by an ECC unit with an oligomerization unit (charge 2)

Example 4 is carried out on the load 2 identical to the load used in example 3 and under the conditions of example 2 of application WO13 / 093227.

In Example 4, the ECC unit is preceded by a mild hydrocracking unit, supplied with the same charge as that of Example 3, that is to say a direct distillation VGO.

The operating conditions of the mild hydrocracking unit are presented in Table 13.

[Tables95]

Soft hydrocracking unit Example 4 VGO charge rate t / h 230 Temperature ° C 365 WH h-1 1.0 H2 / HC L / L 400 Pressure bars g 70

Table 13: Operating conditions of the mild hydrocracking unit - Example 4

Under these conditions, the yields with respect to the charge of the mild hydrocracked outlet products are given in Table 14.

[0198] [Tables14]

% weight Destination Naphtha + Kerosene (PI - 220 ° C) 2.0 to products Diesel (220 ° C- 360 ° C) 20 to products VGO hydro treated 78 to ECC TOTAL 100

Table 14: Yields of the products of the mild hydrocracking unit of Example 4

At the end of the catalytic cracking, the yields of the products with respect to the hydrotreated VGO charge evolve as illustrated in table 15 below:

[0201] [Tables 15]

% weight Dry gases 2.37 C3 Cup 9.92 Section C4 14.23 Naphtha (PI-160 ° C) 42.13 Kerosene (160-220 ° C) 10.15 Diesel (220 ° C- 360 ° C) 12.14 Slurry (360 ° C +) 4.45 Coke 4.61

Table 15: Yield of the products after the catalytic cracking of Example 4

The content of naphtha olefins PI-160- ° C is equal to 44.0%, and the nitrogen content is 20 ppm wt. The sulfur content is 150 ppm.

The C3 cut mainly composed of propylene and propane, as well as the C4 cut mainly composed of butenes and butane, as well as the naphtha cut (PI-160 ° C) are then sent to the oligomerization unit.

The purification and oligomerization conditions are given in table 16 below:

[0206] [Tables 16]

Charging flow (C3-160 ° C) (t / h) 119 Purification Unit Quantity of NaX molecular sieve t 8.8 Oligomerization unit Temperature ° C 160-250 * Pressure bar g 50 Quantity of catalyst IP 811 catalyst (Axens) t 240

Table 16: Operating conditions of the purification and oligomerization units The lowest content of impurities (nitrogen and sulfur) in the feed allows a very strong reduction in the amount of molecular sieve used in the purification unit located upstream of the oligomerization unit which goes from 22.1 tonnes to 8.8 tonnes.

The yields of the products from the mild hydrocracking sequence, catalytic cracking and oligomerization of the cuts C3 =, C4 =, petrol, with respect to the charge entering the soft hydrocracking unit are given in Table 17:

[0209] [Tables17]

% weight Origin Example 2 According to the invention Dry gases LCC 1.9 Raffinai Oligomerization 4.4 Naphtha + Kerosene (PI-220 ° C) Soft hydrocracking and Oligomerization and LCC 32.7 Diesel (220 ° C- 360 ° C) Soft hydrocracking and LCC and oligomerization 54.0 Slurry (360 ° C +) LCC 3.8 Coke LCC 3.6

Table 17: Yields of the products resulting from the soft hydrocracking sequence and catalytic cracking and oligomerization of Example 4

The sequence of the hydrocracking unit and the steam cracking and selective hydrogenation with an oligomerization unit according to Example 3 according to the invention produces diesel cuts which meet the commercial specifications of road diesel, unlike the diagram LCC + oligomerization of comparative example 4, as shown in the table below:

[0212] [Tablesl8]

Comparative example 4 Example 3 Specifications Road diesel Europe Rdt diesel (220-360 ° C) % wt 54.0 36.7 - sulfur ppm wt 500 <8 <10 density g / cm3 0.889 0.842 0.820-0.845 Cetane number calculated - 42 54 > 46

Table 18. Comparison of commercial specifications for road diesel between patents

FCC and example 3

The sequence according to the invention therefore makes it possible to obtain diesel fuel with specifications in terms of sulfur content, density and cetane index relative to the implementation of a sequence according to the prior art , while allowing the obtaining of average distillate yield higher than that of a conventional hydrocracking process.

Claims (1)

[Claim 1] Process for the production of middle distillais comprising at least the following steps: a) Hydrotreating (HDT) of hydrocarbon feeds containing at least 20% by weight and preferably at least 80% by weight of boiling compounds above 340 ° C., in the presence of hydrogen and at least one catalyst for hydrotreatment, at a temperature between 200 and 450 ° C, under a pressure between 2 and 18 MPa, at a space speed between 0.1 and 6 h-1 and at a quantity of hydrogen introduced such as the volume ratio liter of hydrogen / liter of hydrocarbon is between 100 and 2000 NL / L, b) Hydrocracking (HCK) of at least part of the effluent from step a), step b) of hydrocracking operating, in the presence of hydrogen and at least one catalyst for hydrocracking and at a temperature between 250 and 480 ° C, under a pressure between 2 and 25 MPa, at a space speed between 0.1 and 6 h-1 and at a quantity of hydrogen introduced such as the volume ratio liter of hydrogen / liter of hydrocarbon is between 80 and 5000 NL / L, c) Separation of the effluent (SEP) from step b) to obtain at least one naphtha cut having a distillation range between 20 ° C and 220 ° C, a medium distillate cut and an unconverted liquid effluent having an initial boiling point above 340 ° C, d) sending at least part of the naphtha cut having a distillation range of between 20 ° C and 220 ° C in a steam cracking step to produce at least one olefinic effluent having a final boiling point below 220 ° C, said steam cracking step being carried out at a reactor outlet temperature between 700 ° C and 900 ° C, at a pressure between 0.1 and 0.5 MPa during a residence time between 0.005 and 2 s and a water vapor and charge (H ₂ O / HC) mass flow ratio is between 0.1 and 2, e) selective hydrogenation (SHU) of at least part of the olefinic effluent having a final boiling point below 220 ° C. from step d) to convert the diolefins to mono-olefins, said step e) being implemented at a temperature between 0 ° C and 200 ° C, a pressure between 0.1 and 5 MPa and at a WH between 0.5 and 8 h -1, f) the oligomerization (OLG) of at least part of the effluent [Claim 2] [Claim 3] [Claim 4] [Claim 5] [Claim 6] [Claim 7] olefinic from step e) to produce at least one olefinic hydrocarbon effluent having a number of carbon atoms greater than or equal to 8, said oligomerization step f) being carried out at a temperature between 100 ° C and 350 ° C, at a pressure included between 1 and 10 MPa and in the presence of an oligomerization catalyst, g) Total hydrogenation (HYD) of at least part of the middle distillate cut from step c) to transform the olefins into alkanes of the middle distillate cut, said step g) being carried out at a temperature between 100 ° C and 400 ° C, a pressure between 1 and 10 MPa, a WH between 0.5 and 10 h-1 and with a volume ratio of hydrogen / hydrocarbon is between 100 and 2000 NL / L and in the presence of a hydrogenation catalyst. Process according to Claim 1, in which the hydrotreatment stage a) operates at a temperature between 300 and 430 ° C, under a pressure between 5 and 16 MPa, at a space speed between 0.2 and 5 h- 1, and at a quantity of hydrogen introduced such that the volume ratio liter of hydrogen / liter of hydrocarbon is between 300 and 1500 NL / L. Process according to either of Claims 1 and 2, in which the hydrocracking stage b) operates at a temperature between 330 and 435 ° C, under a pressure between 3 and 20 MPa, at a space speed between 0, 2 and 4 h ¹ , and at a quantity of hydrogen introduced such that the volume ratio of liter of hydrogen / liter of hydrocarbon is between 200 and 2000 NL / L. Process according to one of Claims 1 to 3, in which at least part of the liquid effluent which is not converted in hydrocracking stage b) and separated in stage c), having an initial boiling point greater than 340 ° C is also sent in step d) of steam cracking. Method according to one of Claims 1 to 3, in which at least part of the unconverted liquid effluent from step c) is recycled in step b) of hydrocracking. Method according to one of Claims 1 to 3, in which at least part of the unconverted liquid effluent from step c) is recycled in step a) of hydrotreatment. Method according to one of claims 1 to 3 in which at least part of the unconverted liquid effluent from step c) is sent in a second hydrocracking step b ') operating in the presence of hydrogen and d '' at least one second hydrocracking catalyst, a temperature between 250 and 480 ° C, under a pressure between 2 and 25 MPa, at a space speed between 0.1 and 6 h ¹ and at a quantity of hydrogen introduced such as the volume ratio liter of hydrogen / liter of hydrocarbon is between 100 and 2000 NL / L. [Claim 8] Method according to one of the preceding claims, in which said step d) of steam cracking is carried out at a pressure of between 0.1 and 0.2 MPa during a residence time of between 0.02 and 2 vs [Claim 9] Method according to one of the preceding claims, in which step d) of steam cracking allows, in addition to the production of said olefinic effluent having a final boiling point below 220 ° C, the production of a dry gas cut consisting of hydrogen, methane, ethane, and carbon monoxide, a medium ethylene cut and a distillate cut. [Claim 10] Process according to Claim 8 or 9, in which the middle distillate cut is sent in step g) of final hydrogenation. 1/1