EP1063275B1 - Process for hydrotreatment of middle distillate in two stages with intermediate stripping - Google Patents

Description

The present invention relates to the hydrotreatment of hydrocarbon fractions and, for example, middle distillates to produce hydrocarbon fractions with a low content of sulfur, nitrogen and aromatic compounds which can be used in particular in the field of fuels for internal combustion engines. These hydrocarbon fractions include jet fuel, diesel fuel, kerosene and gas oils. The invention relates more particularly to the manufacture of a fuel for compression ignition engine. In this field, the invention relates to a process for converting a middle distillate, and more particularly a diesel fuel fraction, in order to produce a fuel with a high cetane number, deflavored and desulphurized.

At the present time, the gasoil cuts, whether they come from the direct distillation of a crude oil or from the catalytic cracking process, still contain significant amounts of aromatic compounds, nitrogen compounds and sulfur compounds. Under the current legislative framework of most industrialized countries, engine fuel must contain less than about 500 parts per million by weight (ppm) of sulfur. In the vast majority of these countries, there are currently no standards imposing a maximum level of aromatics and nitrogen. However, several countries or states, such as Sweden and California, are planning to limit the aromatics content to less than 20% by volume, or even less than 10% by volume, and some experts believe that This content could be limited to 5% by volume. In Sweden, in particular, certain classes of diesel fuel already have to meet very stringent specifications. Thus, Class II diesel fuel in this country must not contain more than 50 ppm of sulfur and more than 10% by volume of aromatic compounds, and that of Class I must contain more than 10 ppm sulfur and 5% by weight. volume of aromatic compounds. Currently in Sweden Class III diesel fuel must contain less than 500 ppm sulfur and less than 25% by volume of aromatic compounds. Similar limits are also to be respected for the sale of this type of fuel in California.

Meanwhile, engine manufacturers in several countries are pushing for legislation to force tankers to produce and sell fuel with a cetane number that has a minimum value. Currently, French legislation requires a minimum cetane number of 49, but it is foreseeable that in the near future this minimum index will be at least 50 (as is already the case for Class I fuel in Sweden) and even presumably at least 55 and more likely between 55 and 65.

Many specialists are seriously considering the possibility of having in the future a standard imposing a lower nitrogen content for example at about 200 ppm and even certainly less than 100 ppm. Indeed a low nitrogen content makes it possible to obtain a better stability of the products and will generally be sought by both the seller of the product and the manufacturer.

It is therefore necessary to develop a reliable and efficient method for obtaining from conventional diesel straight-run or catalytic cracking (LCO-cut) or other conversion processes (coking, visbreaking, hydroconversion of residue etc.) a product having improved characteristics both as regards the cetane number and the contents of aromatics, sulfur and nitrogen. It is of particular importance and this is one of the advantages of the process of the present invention of producing the minimum amount of gaseous hydrocarbon compounds so as to obtain a directly and fully recoverable effluent as a very high quality fuel cut. Furthermore, the process of the present invention allows production over a long period of time without the need for regeneration of the employed catalysts which have the advantage of being very stable over time. Furthermore, according to one embodiment of the invention, another advantage lies in the recycling of hydrogen. Indeed according to this embodiment only a fraction of the gas containing hydrogen is sent to a drying-desulfurizing zone before recycling, which allows a saving on the size of the dryer-desulfurizer and the amount of material required to perform this operation. According to another variant of the invention the start of the unit is facilitated when using an oven to adjust the temperature of the feedstock into the second reactor.

US-A-5,114,562 discloses a process for the hydrotreatment of a middle distillate in at least two consecutive steps in order to produce desulfurized and deflavored hydrocarbon cuts comprising a first hydrodesulfurization step whose effluent is sent to a stripping zone with hydrogen in order to eliminate the hydrogen sulphide that it contains, then the desulfurized liquid fraction obtained is sent to a second so-called hydrogenation zone comprising at least two reactors operating in series in which the aromatic compounds are hydrogenated. According to this teaching, the hydrogen used in the stripping zone is the additional hydrogen necessary for the process, and the hydrocarbon compounds entrained during stripping are, after condensation by cooling, reintroduced into the first hydrodesulfurization stage. The gas, separated from the hydrocarbon compounds at the condensation stage, is treated by washing with an amine solution which makes it possible to eliminate the hydrogen sulphide that it contains and is then sent to the second so-called hydrogenation zone, then the effluent leaving the hydrogenation zone is separated into a desired liquid fraction and a gas fraction that is mixed with the fresh feedstock at the inlet of the first hydrodesulfurization step. This way of operating has several disadvantages. Thus the hydrocarbons entrained at the top of the stripper which are light compounds and that the hydrodesulfurization step is recycled in this step and thus cause a decrease in the hydrogen partial pressure which is not favorable to good hydrodesulfurization. Moreover another disadvantage is the need to have a recycling pump which increases the cost of the investment and operation of the whole.

US-A-5,110,444 discloses a process comprising hydrotreating a middle distillate in at least three distinct steps. The effluent from the first hydrodesulphurization stage is sent to a hydrogen stripping zone to remove the hydrogen sulphide that it contains, and then the desulphurized liquid fraction obtained is sent to a first hydrogenation zone whose The effluent is in turn sent to a second stripping zone distinct from the stripping zone following the hydrodesulfurization step. Finally, the liquid portion from the second stripping zone is sent to a second hydrogenation zone. The light hydrocarbons entrained at the top of the first stripper with hydrogen are recycled in the hydrodesulfurization step which is unfavorable to the good efficiency of this step since these compounds by vaporizing reduce the partial pressure of hydrogen. Moreover, this recycling involves the compulsory use of a recycling compressor which increases the cost of the investment and operation of the whole.

The present invention presents a solution that makes it possible to overcome, to a large extent, the disadvantages of the processes of the prior art. However information of the interior art and in particular that of the documents cited in the text of the present invention is an integral part of the knowledge of those skilled in the art of which all the characteristics must be considered as included in the present description.

In its broadest formulation, the present invention therefore relates to a process for the hydrotreatment of a hydrocarbon fraction such as, for example, a middle distillate, and in particular for the conversion of a diesel fuel fraction to produce a fuel with a high cetane number, deflavored and desulfurized in at least two successive stages. It also relates to the fuel obtained by said process. As used herein, the term "middle distillate" refers to hydrocarbon fractions boiling in the range of about 130 ° C to about 385 ° C, often about 140 ° C to about 375 ° C and most often about 150 ° C to about 300 ° F. about 370 ° C determined according to the appropriate ASTM method. The process of the present invention may also find application in the treatment of hydrocarbon fractions having a boiling point in the naphtha range. This process can be used to produce hydrocarbon cuts that can be used as a solvent, as an additive or as fuels preferably containing a reduced content of aromatic compounds. For the purposes of this description, the term kerosene means a fraction hydrocarbon boiling in the range 130 ° C to 250 ° C. The term "diesel fuel" as used herein means a hydrocarbon fraction boiling in the range from 230 ° C. to 385 ° C. The term "naphtha" as used herein means a hydrocarbon fraction ranging from C5 to a final boiling point of about 210 ° C. For the purposes of the present description, the term "diesel" denotes a hydrocarbon fraction boiling in the range 230 ° C. to 420 ° C. or heavier fractions boiling in the range 420 ° C. to 525 ° C. For the purposes of the present description, the term "jet fuel" denotes a hydrocarbon fraction boiling in the range 130 ° C. to 290 ° C. The hydrocarbon fraction which is preferably used in the present process is an initial boiling point fraction greater than about 150 ° C and having a boiling point of 90% distilled fraction, most often less than about 370 ° C. vs. This fraction usually contains nitrogen as organo-nitrogen compounds in an amount most often from about 1 ppm to about 1% by weight. It also contains sulfur in the form of sulfur-containing organic compounds in an amount usually greater than about 0.1% by weight and often from about 0.15 to about 5% by weight and most often from about 0.5 to about 3.5% by weight. The content of mono and / or polynuclear aromatic compounds is usually greater than about 10% by volume and often greater than about 20% by volume and usually less than about 60% by volume and often less than about 50% by volume.

• a) au moins une première étape dans laquelle on fait passer ladite charge et de l'hydrogène dans une zone d'hydrodésulfuration contenant au moins un catalyseur d'hydrodésulfuration comprenant sur un support minéral au moins un métal ou composé de métal du groupe VIB de la classification périodique des éléments et au moins un métal ou composé de métal du groupe VIII de ladite classification périodique, la dite zone étant maintenue dans des conditions d'hydrodésulfuration comprenant une température d'environ 260 °C à environ 450 °C et une pression d'environ 2 MPa à environ 20 MPa ,
• b) au moins une deuxième étape dans laquelle la charge partiellement désulfuré issu de l'étape d'hydrodésulfuration est envoyé dans une zone de stripage dans laquelle il est purifié par stripage à contre-courant par au moins un gaz contenant de l'hydrogène sous une pression sensiblement identique à celle régnant dans la première étape et à une température d'environ 100 °C à environ 350 °C dans des conditions de formation d'un effluent gazeux de stripage contenant de l'hydrogène et de l'hydrogène sulfuré et un effluent liquide ne contenant sensiblement plus d'hydrogène sulfuré,
• c) au moins une troisième étape dans laquelle la charge liquide issu de l'étape de stripage est, après addition d'hydrogène d'appoint sensiblement pur et le plus souvent d'hydrogène de recyclage, envoyé dans une zone d'hydrotraitement contenant un catalyseur d'hydrotraitement comprenant sur un support minéral au moins un métal noble ou composé de métal noble, du groupe VIII de préférence. Cet effluent, après avoir été porté, par échange direct de chaleur, à une température d'environ 220 °C à environ 360 °C et une pression d'environ 2 MPa à environ 20 MPa dans la dite zone, est maintenue dans des conditions d'hydrotraitement permettant d'obtenir un effluent partiellement désaromatisé.

More specifically, the present invention relates to a process for hydrotreatment of a hydrocarbon feedstock, containing sulfur compounds, nitrogen compounds and aromatic compounds, comprising the following steps:

• a) at least a first step in which said feedstock and hydrogen are passed into a hydrodesulfurization zone containing at least one hydrodesulfurization catalyst comprising on a mineral support at least one metal or group VIB metal compound of the periodic classification of the elements and at least one metal or group VIII metal compound of said periodic classification, said zone being maintained under hydrodesulfurization conditions comprising a temperature of about 260 ° C to about 450 ° C and a pressure from about 2 MPa to about 20 MPa,
• b) at least a second step in which the partially desulphurized feedstock from the hydrodesulfurization step is sent to a stripping zone in which it is purified by countercurrent stripping with at least one hydrogen-containing gas a pressure substantially identical to that prevailing in the first step and at a temperature of about 100 ° C to about 350 ° C under conditions of formation of a stripping gaseous effluent containing hydrogen and hydrogen sulfide and a liquid effluent containing substantially no more hydrogen sulfide,
• c) at least a third step in which the liquid charge resulting from the stripping step is, after addition of substantially pure auxiliary hydrogen and most often of recycle hydrogen, sent to a hydrotreatment zone containing a hydrotreatment catalyst comprising on a mineral support at least one noble metal or noble metal compound, preferably group VIII. This effluent, after being heated by direct heat exchange, to a temperature of about 220 ° C. to about 360 ° C. and a pressure of about 2 MPa to about 20 MPa in said zone, is maintained under conditions hydrotreating to obtain a partially deflavored effluent.

In said process the gaseous effluent formed in the stripping stage containing gaseous hydrocarbons under the conditions of said stripping zone, hydrogen and hydrogen sulfide is sent to a cooling means, in which it is cooled at a temperature sufficient to form a hydrocarbon liquid fraction which is sent at least partly into the stripping zone and a hydrocarbon-depleted gas fraction that is sent, mixed with all the partially deflavored effluent from step c, in an elimination zone of the hydrogen sulphide it contains and from which purified hydrogen and a partially desulphurized and deflavored hydrocarbon liquid fraction are recovered.

Various embodiments of the present invention are feasible either separately from each other or in combination. In an advantageous embodiment, the stripping gas is a fraction of the makeup hydrogen used in the process of the invention. This fraction of the makeup hydrogen usually represents less than 90% by volume of the total makeup hydrogen used in the process, often less than 60% and most often from about 1% to about 50%.

According to a particular embodiment of the invention, at least a portion of the purified hydrogen recovered from the hydrogen sulfide removal zone contained in the hydrocarbon-depleted gaseous fraction obtained at the end of the process. stripping zone, is sent to a drying-desulfurizing zone from which substantially pure and substantially dry hydrogen is recovered and the other part is recovered without further treatment.

The purification of hydrogen from the gaseous mixture containing hydrogen and hydrogen sulfide from the stripping zone is usually carried out according to one or the other of the conventional techniques well known to those skilled in the art and in particular by pretreating said gaseous mixture with a solution of at least one amine under conditions permitting the removal of hydrogen sulfide by absorption, said amine being most often selected from the group consisting of monoethanolamine, diethanolamine, diglycolamine, diisopropylamine, and methyldiethanolamine. In a particular embodiment of this absorption, the gaseous mixture will be brought into contact with a basic solution, preferably an aqueous solution of an amine selected from the group mentioned above, which forms, with hydrogen sulfide, a compound of addition which makes it possible to obtain a purified gas containing proportions of hydrogen sulphide well below 500 ppm by weight and often below about 100 ppm by weight. Most often the amount of hydrogen sulfide remaining is less than about 50 ppm by weight and very often less than about 10 ppm by weight. This method of purifying the gas mixture is a conventional method well known to those skilled in the art and widely described in the literature. It is for example briefly described for the treatment of natural gas containing hydrogen sulphide for example in the encyclopedia Ullmann's volume A12 pages 258 and following. In the context of the present invention treatment with an aqueous amine solution is usually carried out at a temperature of about 10 ° C to about 100 ° C and often about 20 to about 70 ° C. Usually the amount of amine used is such that the molar ratio of hydrogen sulfide to amine is from about 0.1: 1 to about 1: 1 and often from about 0.3: 1 to about 0.8: 1 and The pressure at which absorption by the amine solution of the hydrogen sulfide is carried out is usually from about 0.1 MPa to about 50 MPa, often from about 1 MPa to about 25 MPa and most often from about 1 MPa to about 10 MPa. The regeneration of the amine solution is conventionally carried out by variation of pressure. To obtain a drier gas and further elimination of the hydrogen sulphide initially present in the gaseous mixture, it is also possible to provide at least a part of this gaseous mixture with additional treatment such as, for example, a treatment of the gas leaving the gas mixture. absorption step, in a hydrogen sulphide adsorption zone comprising at least one reactor and often at least two adsorption reactors containing for example a preferably regenerable sieve or for example zinc oxide and operating by at a temperature of from about 10 ° C to about 400 ° C, and often from about 10 ° C to about 100 ° C and most often from about 20 ° C to about 50 ° C under a total pressure of 100 ° C. about 0.1 MPa to about 50 MPa, often from about 1 MPa to about 25 MPa and preferably from about 1 MPa to about 10 MPa. According to this embodiment, when the adsorption zone comprises two reactors, one reactor is used to treat the gas while the other is in the course of regeneration or replacement of the material that it contains allowing the drying and the desulfurization of the mixture. gaseous entering said zone. At the outlet of this additional treatment, the content of hydrogen sulphide in the gas is usually less than 1 ppm by weight and often of the order of a few tens of ppb by weight.

According to one embodiment, the gaseous effluent formed in the stripping step is cooled by at least one cooling means located inside the stripping zone near the outlet of said gaseous effluent from said stripping zone.

According to another advantageous embodiment, the gaseous effluent formed in the stripping step is cooled by at least one cooling means located outside the stripping zone and is at least partially condensed, the liquid obtained is then at least less partly returned to the stripping zone.

According to another embodiment, the gaseous effluent formed in the stripping step is cooled by at least one cooling means, at least part of the liquid hydrocarbon fraction obtained is returned to the stripping zone and optionally at least another part is sent in admixture with the hydrocarbon feedstock in step a) hydrodesulfurization.

The substantially pure hydrogen recovered after the stripping step is recycled to the stripping zone at at least one point of introduction located between the point of introduction of a portion of the hydrogen-containing gas used for stripping. and the point of introduction of the effluent of hydrodesulfurization step a) into said stripping zone.

The substantially pure hydrogen recovered after the stripping step is recycled directly and / or after mixing with the feedstock in the hydrodesulfurization zone of step a).

The substantially pure hydrogen, and preferably previously dried and deeply desulphurized, recovered after the stripping step is recycled directly and / or after mixing with the liquid effluent of the stripping zone and with the additional hydrogen in the hydrotreatment zone of step c).

According to a particular embodiment of the process of the invention, if it is not desired to carry out drying on all of the substantially pure hydrogen recovered from the hydrogen sulfide removal zone, it is advantageous to carry out drying and deep desulphurization of the hydrogen that is desired to be recycled in step c) of hydrotreatment.

In a preferred embodiment of the invention, the operating conditions of steps a) and c) are chosen as a function of the characteristics of the feed which may be a straight-run diesel fuel cut, a diesel fuel cut from catalytic cracking or a cut. diesel fuel from the coking or the viscosity of residues or a mixture of two or more of these cuts. They are usually chosen so as to obtain a product at the leaving step a) containing less than 100 ppm of sulfur and less than 200 ppm of nitrogen, preferably less than 100 ppm of nitrogen and most often less than 50 ppm of nitrogen and the conditions of the step c) are chosen so as to obtain a product, at the outlet of said step c), containing less than 20% by volume of aromatic compounds. These conditions may be tightened so as to obtain after the second stage a fuel containing less than 10% by volume of aromatic compounds or even less than 5% by volume of aromatic compounds, less than 50 ppm or even less than 10 ppm of sulfur, less 50 ppm or even less than 20 ppm nitrogen or even less than 10 ppm and having a cetane number of at least 50 and even at least 55 and most often between 55 and 60.

To obtain such results, the conditions of step a) comprise a temperature of about 300 ° C to about 450 ° C, a total pressure of about 2 MPa to about 20 MPa and a global hourly space velocity of liquid charge d. about 0.1 to about 4 and that of step b) a temperature of about 200 ° C to about 400 ° C, a total pressure of about 3 MPa to about 15 MPa and a global hourly space velocity of charge liquid from about 0.5 to about 10.

The catalyst employed in step a) contains on a mineral support at least one metal or metal compound of group VIB of the periodic table of elements in an amount expressed by weight of metal relative to the weight of the finished catalyst, usually about 0.5 to 40%, at least one metal or group VIII metal compound of said periodic classification in an amount expressed by weight of metal relative to the weight of the finished catalyst usually from about 0.1 to 30%. Often the catalyst used will contain at least one element selected from the group consisting of silicon, phosphorus and boron or compounds of this or these elements. The catalyst will contain, for example, phosphorus or at least one phosphorus compound in an amount expressed by weight of phosphorus pentoxide with respect to the weight of the support of approximately 0.001 to 20%. The amount of metal or group VIB metal compound expressed in weight of metal relative to the weight of the finished catalyst will preferably be about 2 to 30% and most often about 5 to 25% and that of the metal or the Group VIII metal compound will preferably be from about 0.5 to 15% and most often from about 1 to 10%.

When it is desired to remain in a relatively low pressure range while wishing to obtain excellent results, it is possible to perform a first step a1) under conditions making it possible to reduce the sulfur content of the product to a value of about 500 to 800 ppm and then send this product to a subsequent step a2) in which the conditions will be chosen to reduce the sulfur content to less than about 100 ppm, preferably less than about 50 ppm and the product from this step a2) is then sent to step b). In this embodiment, the conditions of step a2) are milder than when for a given load one operates in a single step a) since the product sent in this step a2) already has a greatly reduced sulfur content. In this embodiment, the catalyst of step a1) can be a conventional catalyst of the prior art such as for example that described in the text of the patent applications on behalf of the applicant FR-A-2197966 and FR-A- A-2538813 and that of step a2) is that described above for step a). It is not within the scope of the present invention using the same catalyst in steps a1) and a2).

In these steps a), a1), a2) the mineral support of the catalyst is preferably chosen from the group formed by alumina, silica, silica-aluminas, zeolites and mixtures of at least two of these compounds. minerals. Alumina is very commonly used.

In a preferred embodiment of the invention, the catalyst of these steps a), a1), a2) will comprise at least one metal or a metal compound selected from the group consisting of molybdenum and tungsten and at least one metal or a metal compound selected from the group consisting of nickel, cobalt and iron. Most often this catalyst contains molybdenum or a molybdenum compound and at least one metal or a metal compound selected from the group consisting of nickel and cobalt.

In a particular and preferred embodiment of the invention the catalyst of these steps a), a1), a2) will comprise boron or at least one boron compound preferably in an amount expressed by weight of boron trioxide relative to the weight of the support from about 0 to 10%. Other embodiments are also often employed and in this case the catalyst will comprise for example silicon or a silicon compound, or a combination of silicon and boron or compounds of each of these elements possibly associated with phosphorus or to a phosphorus compound. As non-limiting examples of specific combinations containing these elements or compounds of these elements, the following combinations may be mentioned: Ni-Mo-P, Ni-Mo-PB, Ni-Mo-Si, Ni-Mo-Si -B, Ni-Mo-P-Si Ni-Mo-Si-BP, Co-Mo-P, Co-Mo-PB, Co-Mo-Si, Co-Mo-Si-B, Co-Mo-P- Si, Co-Mo-Si-BP, Ni-WP, Ni-WPB, Ni-W-Si, Ni-W-Si-B, Ni-WW-P-Si, Ni-W-Si-BP, WP, Co-WPB, Co-W-Si, Co-W-Si-B, Co-WP-Si, Co-W-Si-BP, Ni-Co-Mo-P, Ni-Co-Mo-PB, Ni-Co-Mo-Si, Ni-Co-Mo-Si-B, Ni-Co-Mo-Si-P.

The catalyst employed in step c) contains on a mineral support at least one noble metal or noble metal compound of group VIII of the periodic table of elements in an amount expressed by weight of metal relative to the weight of the finished catalyst of about 0.01 to 20% and preferably at least one halogen. The inorganic support of the catalyst employed in step c) is chosen independently of the support used for the catalyst of step a). Most often the catalyst of step c) comprises at least one metal or a noble metal compound selected from the group consisting of palladium and platinum.

The inorganic support of the catalyst employed in step c) is usually selected from the group consisting of alumina, silica, silica-aluminas, zeolites and mixtures of at least two of these mineral compounds. This support will preferably comprise at least one halogen chosen from the group formed by chlorine, fluorine, iodine and bromine and preferably from the group formed by chlorine and fluorine. In an advantageous embodiment, this support will comprise chlorine and fluorine. The amount of halogen will most often be from about 0.5 to about 15% by weight based on the weight of the support. The most commonly used support is alumina. Halogen is usually introduced onto the support from the corresponding acid halides and platinum or palladium from aqueous solutions of their salts or compounds such as for example hexachloroplatinic acid in the case of platinum.

• La figure 1 explique brièvement, à titre illustratif, un mode de mise en oeuvre du procédé seloninvention.
• La figure 2 présente un de mode de réalisation comparatif dans lequel la fraction gazeuse appauvrie en hydrocarbures, issue du moyen de refroidissement après stripage est envoyée dans la zone d'élimination de l'hydrogène sulfuré, en mélange avec une partie seulement de l'effluent partiellement désaromatisé issu de l'étape c.
• La figure 3 illustre un procédé selon la technique antérieure.

The amount of metal of this catalyst of step c) will preferably be from about 0.01 to 10%, often from about 0.01 to 5% and most often from about 0.03 to 3% expressed. in weight of metal relative to the weight of the finished catalyst.

• FIG. 1 briefly explains, by way of illustration, one embodiment of the method according to the invention.
• FIG. 2 shows a comparative embodiment in which the gaseous fraction depleted in hydrocarbons resulting from the cooling means after stripping is sent to the hydrogen sulfide elimination zone, mixed with only a part of the effluent. partially deflavored from step c.
• Figure 3 illustrates a method according to the prior art.

In these figures, similar members are designated by the same numbers and letters of reference.

The hydrocarbon feedstock arriving via line 1 is mixed with substantially pure hydrogen arriving via lines 42, 42a and 26, and this mixture is introduced via line 6 into reactor R1 after exchanging heat with the effluent. of the reactor R1 in the exchanger EC1 and have been preheated in the furnace F1. Hydrogen arriving via the lines 42 and 42a is introduced into the reactor R1 via the line 27 as cooling gas (quench). The effluent leaving the reactor R1 is sent after heat exchange in the exchanger EC1 with the hydrocarbon feedstock via line 7 in the extractor S1 (denominated by those skilled in the art by the Anglo-Saxon term stripper) in which it is purified by a portion of the additional hydrogen arriving via lines 2 and 4. The substantially pure hydrogen recycling is also introduced in the extractor S1 by the lines 42 and 25. The gaseous effluent leaving the extractor S1 via the line 28 exchanges heat in the exchanger EC4 located on the line 5 with a mixture of hydrogen recycled by the line 24 and of additional hydrogen introduced via lines 2, 3 and 5. Said mixture is then introduced via line 9 into reactor R2. The condensed hydrocarbon liquid fraction entering via the line 28 into the separator flask B1 exits via the line 29 and is returned using the pump P1 via the line 30 in part in the extractor S1 and possibly via the line 36 through the valve V36 and the line 6 partly in the reactor R1. The liquid fraction exiting the extractor S1 passes through the level regulating valve V1, then is sent via lines 8 and 9, after mixing with hydrogen arriving via line 5, exchanging heat in the exchanger EC2 with the effluent leaving reactor R2 by line 10 and after having been reheated in furnace F2 enters reactor R2. If the pressure of the fluid exiting the extractor S1 through the line 8 is less than that prevailing in the reactor R2, a pump will be used to adjust this pressure to a level at least equal to that of the pressure in the reactor R2 . Hydrogen arriving via the lines 22, 22a, 22b and 23 is introduced into the reactor R2 as a cooling gas (quench). The gaseous effluent leaving the separator tank B1 via the line 11 passes through the pressure regulating valve V2 and is then mixed with the effluent from the reactor R2 arriving via line 10, before being sent via lines 12 and 14 after mixing. with washing water arriving via line 13 and heat exchange in exchanger EC3 in a separating flask B2 from which an aqueous fraction is recovered via line 15, line 16 a hydrocarbon liquid fraction constituting the partially desulphurized effluent and deflavored desired and line 17 a gaseous fraction containing hydrogen and hydrogen sulfide, a portion of which may optionally be purged by the line 18 which comprises a valve V18 (not shown in the figures, allowing the the purge flow rate, the hydrogen leakage is the minimum purge counted on line 18) and another part, or all of it when there is no purge, is sent by the line 19 in the absorber S2 of at least partial removal of the hydrogen sulphide in which an absorption solution is introduced via the line 20 and recovered by the line 21. By lines 22 and 22a is recovered at least a portion purified hydrogen that can be sent through the flow control valve V3 in the drying-desulfurizing zone SE1 and then recycled to the reactor R2 via the lines 22b and 23. Another part of the purified hydrogen recovered by the line 22 is sent via the lines 42 and 25 through the flow control valve V4 into the extractor S1 and / or through the lines 42, 42a, 26 and 6 into the reactor R1. Although this is not shown in the figures, the line 42a usually comprises a valve V40 (not shown in the figures) for regulating the flow rate of the purified hydrogen that is sent into the reactor R1.

Although it is not shown in the figures the possible recycling of the hydrogen leaving the adsorber S2 to be sent into the reactor R2 through the lines 22, 22a, 22b and 23 and / or 22, 22a, 24 and 5 requires the use of a compressor to adjust the pressure to a level at least equal to the pressure level prevailing in this reactor R2. It is the same for the hydrogen leaving the absorber S2 which is optionally recycled in the extractor S1 and / or in the reactor R1. It is therefore possible to use a single compressor to obtain a gas at the pressure required for the various recycling considered. In this case, this compressor will be located near the adsorber S2 on the line 22. It is also possible to provide the use of two compressors, one located on line 22b and the other located on line 42.

FIG. 2 illustrates the comparative case where only a part of the effluent leaving reactor R2 via line 10 is cooled in heat exchanger EC2; another part is sent via line 31 after mixing with the part having exchanged heat by line 32 in the hot separator B3 from which a gas is recovered via the line 34 which is mixed with the gas arriving via the line 11 and the line 33 a hydrocarbon liquid fraction constituting part of the partially desulfurized effluent and desired deflavored that is mixed with the hydrocarbon fraction leaving the separator tank B2 by line 16 to obtain the desired partially desulphurized and deflavored effluent that is recovered via line 35.

Figure 3 illustrating the prior art is to be directly compared to the diagram of Figure 1 illustrating the present invention which it differs in two essential points. The extractor S1 used does not include a recovery tank of a liquid fraction and therefore it is not reintroduced in this extractor heavy hydrocarbons that are entrained in the gas leaving the top of the extractor S1. Since there is no recovery of the heavy hydrocarbons entrained at the top of the extractor S1, there is consequently no introduction of a fraction of these heavy hydrocarbons into the reactor R1.

The examples which follow illustrate the invention without limiting its scope. The data common to the three examples are extracted from computer simulations created from PGGC software based on results from pilot units:

Charge :

Debit t / h 100 Density at 15 ° C kg / m3 852 sulfur content % mass 1.44 nitrogen content % mass 0,011 aromatic content % mass 30 distillation 5% volume ° C 253 50% volume ° C 306 95% volume ° C 386 Composition of the booster: Hydrogen: 99.9% vol Methane: 0.1% vol

Operating conditions :

Absolute pressure Temperature MPa ° C R1 reactor outlet 6.4 350 S1 power supply 255 R2 reactor outlet 5.7 310 B2 ball 5.5 40 Space velocity R1 reactor: 1 m ³ / (m ³ / h) Reactor R2: 2 m ³ / (m ³ / h)

Example 1 (according to the invention)

For flow rates in the different flows, see the summary table after the description of the three examples.
According to the diagram of FIG. 1, the feedstock arriving via line 1 is mixed with the recycling gas of line 26, the whole being introduced into reactor R1 via line 6, at a pressure of 6.6 MPa and a temperature of 330.degree. C, after being heated in the exchanger EC1 and the furnace F1. The increase in temperature in the reactor is limited to a range of 20 ° C using the hydrogen quench gas arriving through the line 27. The reactor effluent R1 is sent via line 7, after cooling in exchanger EC1, to the head plate of the hydrogen stripper S1.
The overhead product of the hydrogen extractor is cooled, the cooling interval is 55 ° C, and is passed through line 28 into the flask B1. The liquid of this flask is returned, through the lines 29 and 30 and the pump P1, to the head plate of the hydrogen extractor while the steam is sent through the line 11, through the pressure regulating valve V2 to the flask B2 after mixing in line 12 with the effluent from the reactor R2. This overhead vapor contains about 2295 kg / hr of diesel fuel product which will not be treated in the hydrogenation reactor.
Towards the middle of the hydrogen extractor, stripping hydrogen is injected through line 25 to extract most of the hydrogen sulfide and make-up hydrogen through line 4 in the bottom. the extractor to complete the extraction of hydrogen sulfide in the liquid feed that is sent to the hydrogenation reactor R2 via line 8.

The flow rate of this liquid charge is controlled by the level control valve V1, then it is mixed with the additional hydrogen of the line 3 and the recycling hydrogen of the line 24. This mixture is sent to the reactor R2 through the exchanger EC2 and the furnace F2 via line 9 to reach the required temperature at the reactor inlet.
The temperature increase in the reactor R2 is controlled by the injection of a quench hydrogen gas via the line 23.

• recyclage vers R1, ligne 26,
• refroidissement (quench) vers R1, ligne 27,
• gaz de stripage vers S1, ligne 25,

The effluent from the reactor R2 is withdrawn via the line 10, and then it is mixed with the steam of the hydrogen extractor head in the line 12. In this mixture, before the exchanger EC3, the line 13 is injected. a certain amount of wash water to remove the ammonium sulfide formed in the reactors. After partial condensation in the exchanger EC3, the three phases in the presence are separated in the flask B2.
The hydrocarbon phase, constituting the product of the process, is sent via line 16 to a subsequent treatment for extracting the residual hydrogen sulphide and removing the light fractions.
The aqueous phase is withdrawn via line 15 to be sent to the water treatment.
The vapor phase exits the flask B2 via the line 17 is sent to an amine wash S2 via the line 19 while the excess hydrogen can be purged by the line 18. Part of the hydrogen gas, purified hydrogen sulphurized by the amine wash, is separated into different streams:

• recycling to R1, line 26,
• quench to R1, line 27,
• stripping gas to S1, line 25,

• recyclage vers R2, ligne 24,
• refroidissement vers R2, ligne 23.

Another part of this gas is dried and desulphurized in the reactor SE1 and separated into different streams:

• recycling to R2, line 24,
• cooling to R2, line 23.

An amine solution, low in hydrogen sulfide, is sent to the top of the wash column through line 20, while the hydrogen sulfide rich solution is passed through line 21 to an amine regeneration section.

The final product has the following properties :

Debit t / h 98.6 Density at 15 ° C kg / m3 822 sulfur content % mass <0.0010 nitrogen content % mass <0.0010 aromatic content % mass 1.5 distillation 5% volume ° C 240 50 9o volume ° C 294 95% volume ° C 376

Example 2 (comparative)

For flow rates in the different flows, see the summary table after the description of the three examples.
According to the diagram of FIG. 2, the description is identical to the corresponding description of FIG. 1 until the exit of the effluent from the reactor R2, line 10, after cooling in the exchanger EC2. This effluent is sent via line 32 into the hot separator tank B3 at a pressure of 5.5 MPa and a temperature of 270 ° C.

The vapor phase of the flask B3, line 34, is mixed in line 12 with the overhead product of the hydrogen extractor, line 11. In this mixture, before the exchanger EC3, a certain amount of washing water to remove ammonium sulphide formed in the reactors. After partial condensation in the exchanger EC3, the three phases in the presence are separated in the flask B2.
The hydrocarbon phase, line 16, is mixed in the line 35 with the liquid of the flask B3, line 33, to constitute the product of the process which is sent to a subsequent treatment for the extraction of residual hydrogen sulphide and elimination of the fractions. light.

The aqueous phase is withdrawn via line 15 to be sent to the water treatment.
The vapor phase leaves the ball B2 by the line 17 it is treated as described in connection with Figure 1 ci in front.

The final product has the following properties :

Debit t / h 98.6 Density at 1.5 ° C kg / m3 823 sulfur content % mass <0.0010 nitrogen content % mass <0.0010 aromatic content % mass 1.8 distillation 5% volume ° C 240 50% volume ° C 294 95% volume ° C 376

Example 3 (comparative)

For flow rates in the different flows, see the summary table after the description of the three examples.
According to the diagram of FIG. 3 without condensation of heavy hydrocarbons at the top of the hydrogen extractor S1, the diagram is identical to that of example 1, with the exception of condensate recovery at the top of the extractor. hydrogen this recovery does not exist in Example 3.
Referring to the table indicating the flow rates in the various streams, it should be noted that, for this example, 3659 kg / h of heavy products are not recovered with respect to Example 1. These heavy products can not be treated in the R2 reactor, therefore, to obtain the same quality for the final product, the treatment in the reactor R2 should be more severe in Example 3 than in Example 1.

The final product has the following properties :

Debit t / h 98.6 Density at 15 ° C kg / m3 830 Sulfur content % mass <0.0010 Nitrogen content % mass <0.0010 Aromatic content % mass 2 Distillation 5% volume ° C 240 50% volume ° C 294 95% volume ° C 376

Flow rates of different flows :

Feed name flux example 1 example 2 example 3 kg / h kg / h kg / h Charge R1 1 100000 100000 100000 H2 recycling to R1 26 3452 4281 3451 Quench H2 to R1 27 3380 3772 3379 Add H2 to S1 3 277 277 277 H2 recycling to S1 25 556 556 556 Steam to B2 balloon 11 12355 13470 16150 Of which heavy towards B2 balloon 11 2294 2333 5953 S1 background (R2 load) 8 95310 95416 91513 Addition H2 to R2 3 1079 1144 1063 H2 recycling to R2 24 2239 2667 2318 Quench H2 to R2 23 1740 1933 1659 Steam B3 34 n / A. 10345 n / A. Liquid B3 33 n / A. 90815 n / A. B2 supply 14 112723 23815 112704 Wash water 13 1585 1585 1585 B2 Ball acidic water 15 1651 1651 1651 Hydrogen purge 18 158 185 158 B2 Balloon Hydrocarbons 16 99664 8652 99648 Product 35 99664 99467 99648

Claims (18)

1. A process for hydrotreating a hydrocarbon feed containing sulphur-containing compounds, nitrogen-containing compounds and aromatic compounds, comprising the following steps:

   a) at least one first step wherein said feed and hydrogen are passed into a hydrodesulphurisation zone containing at least one hydrodesulphurisation catalyst comprising, on a mineral support, at least one metal or compound of a metal from group VIB of the periodic table and at least one metal or compound of a metal from group VIII of said periodic table, said zone being maintained under hydrodesulphurisation conditions including a temperature of about 260°C to about 450°C and a pressure of about 2 MPa to about 20 MPa,

   b) at least one second step wherein the partially desulphurised feed from the first hydrodesulphurisation step is sent to a stripping zone wherein it is purified by stripping with a counter-current of at least one gas containing hydrogen at a pressure substantially identical to that prevailing in the first step and at a temperature of about 100°C to about 350°C under conditions for formation of a gaseous stripping effluent containing hydrogen and hydrogen sulphide and a liquid feed containing substantially no more hydrogen sulphide,

   c) at least one third step wherein, after adding substantially pure makeup hydrogen and recycled hydrogen, the liquid feed from the stripping step is sent to a hydrotreatment zone containing a hydrotreatment catalyst comprising, on a mineral support, at least one noble metal or compound of a noble metal from group VIII, said zone being maintained under hydrotreatment conditions to obtain a partially dearomatised effluent,

   wherein the gaseous effluent formed in the stripping step, containing hydrocarbons which are gaseous under the conditions in said stripping zone, hydrogen and hydrogen sulphide, is sent to a cooling means, wherein it is cooled to a temperature which is sufficient to form a liquid hydrocarbon fraction which is sent at least partly to the stripping zone, and a gaseous fraction which is depleted in hydrocarbons which is sent, in mixture with the whole of the partially dearomatized effluent from step c, to a zone for eliminating the hydrogen sulphide it contains and from which purified hydrogen and a liquid hydrocarbon fraction partially desulphurized and dearomatized are recovered.
2. A process according to claim 1, wherein, after adding substantially pure makeup hydrogen and recycled hydrogen, the liquid effluent from the stripping step is sent to a hydrotreatment zone containing a hydrotreatment catalyst comprising, on a mineral support, at least one noble metal or compound of a noble metal from group VIII after having been heated by direct heat exchange to a temperature of about 220°C to about 360°C and a pressure of about 2 MPa to about 20 MPa, said zone being maintained under hydrotreatment conditions to produce a partially dearomatised effluent.
3. A process according to claim 1 or claim 2, wherein a fraction of the makeup hydrogen is used in stripping step b) as a hydrogen-containing stripping gas.
4. A process according to any one of claims 1 to 3, wherein the operating conditions of step a) are selected so as to obtain a product containing less than 100 ppm of sulphur and less than 200 ppm of nitrogen and the conditions of step c) are selected so as to obtain a product containing less than 20% by volume of aromatic compounds.
5. A process according to any one of claims 1 to 4, wherein the gaseous effluent formed in the stripping step is cooled by at least one cooling means located inside the stripping zone close to the outlet for said gaseous effluent from said stripping zone.
6. A process according to any one of claims 1 to 5, wherein the gaseous effluent formed in the stripping step is cooled by at least one cooling means located outside the stripping zone and is at least partially condensed, the liquid obtained being sent to the stripping zone.
7. A process according to any one of claims 1 to 6, wherein at least a portion of the purified hydrogen obtained from the outlet from the zone for eliminating the hydrogen sulphide it contains is sent to a drying-desulphurising zone.
8. A process according to any one of claims 1 to 7, wherein at least a portion of the substantially pure hydrogen recovered after the stripping step is recycled to the stripping zone into at least one introduction point located between the point for introducing a portion of the hydrogen-containing gas used for stripping and the point for introducing effluent from hydrodesulphurisation step a) into said stripping zone.
9. A process according to any one of claims 1 to 8, wherein at least a portion of the substantially pure hydrogen recovered after the stripping step is recycled directly and/or after mixing with the feed to the hydrodesulphurisation zone of step a).
10. A process according to any one of claims 1 to 9, wherein at least a portion of the substantially pure hydrogen recovered after the stripping step is recycled directly and/or after mixing with the liquid effluent from the stripping zone and with the makeup hydrogen to the hydrotreatment zone of step c).
11. A process according to any one of claims 1 to 10, wherein the catalyst from step a) comprises at least one metal or compound of a metal selected from the group formed by molybdenum and tungsten and at least one metal or compound of a metal selected from the group formed by nickel, cobalt and iron.
12. A process according to any one of claims 1 to 11, wherein the catalyst from step a) further comprises at least one element selected from the group formed by silicon, phosphorous and boron or one or more compounds of this or these elements.
13. A process according to any one of claims 1 to 12, wherein the catalyst of step a) further comprises phosphorous or at least one phosphorous compound.
14. A process according to any one of claims 1 to 13, wherein the catalyst of step a) further comprises boron or at least one boron compound.
15. A process according to any one of claims 1 to 14, wherein the catalyst of step a) further comprises silicon or at least one silicon compound.
16. A process according to any one of claims 1 to 15, wherein the support for the catalysts used in step a) and in step c) are selected independently of each other from the group formed by alumina, silica, silica-aluminas, zeolites and mixtures of at least two of these mineral compounds.
17. A process according to any one of claims 1 to 16, wherein the support for the catalyst of step c) comprises at least one halogen, preferably selected from the group formed by chlorine and fluorine.
18. A process according to any one of claims 1 to 17, wherein the catalyst for step c) comprises at least one nobel metal or compound of a noble metal selected from the group formed by palladium and platinum.

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