Process for producing a jet fuel, comprising a step of converting an alcohol stream in a fluidized bed, associated jet fuel and plant

Abstract

Disclosed is a process for producing a jet fuel, comprising a step of converting an alcohol stream in a fluidized bed, a jet fuel and a plant associated with said process. The process involves the following steps: (a) converting a C1 to C6 alcohol stream (14) to produce a mixture (16) containing paraffins, olefins, aromatics, and water; (b) separating water (40) from the mixture (16); (c) oligomerizing olefins; and (d) alkylating aromatics from the mixture (19); (e) forming a stream (24) of hydrocarbons to be hydrogenated; (f) hydrogenating the stream (24) of hydrocarbons to be hydrogenated; (g) recovering at least one jet fuel fraction (34) from the hydrogenated hydrocarbon stream (30). Conversion step (a) is carried out in a reaction zone (111a) comprising at least one fluidized catalytic bed. In the mixture (16) of paraffins, olefins, aromatics and water produced in conversion step (a) , the ratio of the mass of C3+ olefins to the total mass of olefins is greater than or equal to 0.8.

Description

[0001] TITLE: Process for manufacturing a jet fuel comprising a step of converting an alcohol stream in a fluidized bed, associated jet fuel and installation TECHNICAL FIELD The present invention relates to a process for manufacturing a jet fuel, comprising the following steps: (a) converting a C1 to C6 alcohol stream to produce a mixture containing paraffins, olefins, aromatics,and water; (b) separating water from the mixture to form a water-depleted mixture; (c) oligomerizing olefins from the water-depleted mixture; (d) alkylating aromatics from the water-depleted mixture; (e) forming a stream of hydrocarbons to be hydrogenated from at least a portion of the olefins oligomerized in step (c) and at least a portion of the aromatics alkylated in step (d); (f) hydrogenating the stream of hydrocarbons to be hydrogenated to form a stream (30) of hydrogenated hydrocarbons; (g) recovering at least one jet fuel fraction from the stream of hydrogenated hydrocarbons. The present invention relates to the field of the preparation and use of liquid fuels, in particular those of the jet fuel or renewable aviation fuel type. Due to the scarcity of fossil resources and growing environmental concerns,particularly with the aim of reducing greenhouse gas emissions, the use of alternative molecules, with a smaller carbon footprint, is increasingly sought to replace molecules of fossil origin. Renewable fuels derived from biological matter or from carbon dioxide transformed in the presence of decarbonized or electrolytic hydrogen (referred to by the English term "E-fuels") are an alternative to conventional fossil fuels. Conventional jet fuels can be blended with bases derived from renewable feedstocks as provided for by the D7566-21 standard, thus allowing the production of alternative aviation fuels. Aviation fuel bases derived from renewable feedstocks that can be incorporated into fossil jet fuel are, for example: - synthetic paraffinic kerosenes [SPK],from processes such as the Fischer-Tropsch process; - synthetic paraffinic kerosenes produced by the "Alcohol-to-Jet" route (transformation of alcohol into isoparaffinic kerosene) [ATJ-SPK]; - synthetic isoparaffins produced by hydrotreatment of iso-olefin intermediates, produced from fermented sugars [SIP-HFS]; - synthetic aromatic kerosenes obtained by alkylation of light aromatics from non-petroleum sources [SPK / A]; - synthetic kerosenes obtained from the hydrothermal conversion of fatty acid esters and fatty acids; - paraffinic kerosenes [SPK] obtained from hydrotreated hydrocarbons, esters and fatty acids. Currently, these renewable aviation fuel bases cannot for the most part be used alone,due to their composition being very different from fossil jet fuels. This difference in composition poses in particular problems of compatibility with the materials of the elements with which the fuel is in contact. On this point, the absence of aromatic compounds in the majority of available renewable aviation fuel bases can cause compatibility problems with the materials, and in particular with certain seals. STATE OF THE ART EP2123736 describes a process for producing a diesel fuel using a feedstock in the form of C1 to C5 alcohols, which may be totally or partially of biogenic origin, in which, from a mixture of olefinic hydrocarbons obtained at least partially by dehydration of C1 to C5 alcohols, with a proportion of odd olefins and iso-olefins,a synthetic hydrocarbon is oligomerized or hydrogenated. After subsequent hydrogenation and rectification, an aviation fuel is formed with a freezing point of -47°C or lower. US20210078921 describes the conversion of methanol to gasoline which may be carried out using a heavy gasoline treatment, followed by a separation operation. US4543435 describes a process for converting an oxygenated feedstock comprising methanol, dimethyl ether or the like into liquid hydrocarbons, comprising contacting the feedstock with a zeolite catalyst in a primary catalyst stage at elevated temperature and moderate pressure to convert the feedstock into hydrocarbons comprising C2-C4 olefins and C5+ hydrocarbons. EP1844125 relates to a process for producing synthetic fuels, wherein in a first step,a gas mixture comprising methanol and / or dimethyl ether and / or another oxygenated molecule, as well as water vapor is converted into olefins preferably having between 2 and 8 carbon atoms, at temperatures between 300 and 500 °C and in a second step, the resulting olefin mixture is oligomerized at higher pressure, into higher olefins having essentially more than 5, preferably between 10 and 20 carbon atoms. According to this process, a) the production of olefins in the first step is carried out in the presence of a gas stream consisting essentially of saturated hydrocarbons, separated from the product stream of the second step and returned to the first step and b) the production of olefins in the second step is carried out in the presence of a stream of water vapor,which is separated from the product stream of the first stage of the process and is returned to the first stage of the process. EP2147082 describes a process for producing synthetic fuels from a mixture, containing hydrogen and oxygenated compounds such as methanol and / or dimethyl ether, in a first stage, the mixture is reacted on a catalyst, to obtain a hydrocarbon product containing olefins preferably comprising 2 to 8 carbon atoms, and, in a second stage, the oligomerization of the hydrocarbon product thus obtained, into long-chain olefins, from which it is possible to obtain products which are gasoline and diesel. WO2011061198 describes a process for producing hydrocarbons in the form of gasoline, by conversion of syngas, to obtain an oxygen-containing compound, such as methanol and / or dimethyl ether, in a first converter,and by further conversion to hydrocarbons in a second converter. EP2940103 describes a process for preparing biofuels using ethanol by converting ethanol in admixture with hydrocarbons, in a catalytic process on a zeolite-type aluminosilicate bed, preferably in the presence of a hydrogen form of the zeolite catalyst. EP2720990 describes a process for converting an alcohol to a hydrocarbon, the process comprising contacting said alcohol, as a component of an aqueous solution at a concentration of not more than 20%, with a metal-loaded zeolite catalyst at a temperature of at least 100°C and up to 550°C, wherein said alcohol is producible by a fermentation process and is selected from ethanol, butanol, isobutanol, or a combination thereof, said metal includes vanadium,and said metal-loaded zeolite catalyst is catalytically active to convert said alcohol to said hydrocarbon. EP3795658 describes methods that reduce energy and water consumption in processes for producing fuel from renewable alcohol-containing feedstocks. The alcohol is converted directly to hydrocarbon transportation fuels by a catalytic process, with heat being transferred between intermediate process liquids to reduce thermal energy consumption. Overall water consumption is reduced by recovering catalytic process water and reducing the water temperature to reduce evaporative losses. US20160090333 describes methods for producing aviation-range hydrocarbons from biorenewable sources, such as oligomerization of biorenewable C3-C8 olefins, for example,derived from C3-C8 alcohols produced by biomass fermentation. Aviation range hydrocarbon production is increased by using an additional oligomerization zone to oligomerize gasoline separated from the effluent of a primary oligomerization zone in which biorenewable C3-C8 olefins have first been subjected to oligomerization. WO201145535 describes a process for producing distillate from a feedstock of heteroatomic organic compounds comprising at least one heteroatom chosen from oxygen, sulfur, halogen, alone or in combination, in which the treatment of the feedstock comprises at least one step of converting the heteroatomic organic compounds into olefins carried out in a first conversion zone, and, in at least a second oligomerization zone, a step of oligomerization of olefins originating at least in part from the conversion zone, in the presence of at least 0.5% by mass of oxygenated compounds,in order to produce a distillate. This process improves the distillate yield, by making it possible to obtain a higher oligomerization rate compared to the oligomerization of the same feedstock under the same reaction conditions. WO2022 / 063994 describes a process for obtaining a jet fuel comprising a step of converting a stream of oxygenated compounds at reduced temperatures (for example lower than 350°C), pressures of the order of 5 bar to 10 bar and high hourly space velocities (6 h, -1 at 10 a.m. -1) followed by a joint oligomerization and hydrogenation step in the same reactor. This process tends to keep the ethylene and aromatic contents produced during the conversion very low. None of these processes makes it possible to reproducibly optimize the content and type of compounds, particularly aromatics, in a final jet fuel composition in order to improve combustion quality while ensuring good compatibility with existing systems. SUMMARY OF THE INVENTION An aim of the invention is therefore to provide a process for manufacturing a jet fuel derived exclusively or at least in part from renewable feedstocks, which is efficient and productive, while meeting the requirements for use in the aeronautical field. To this end, the subject of the invention is a manufacturing process of the aforementioned type, in which the conversion step (a) is carried out in a reaction zone comprising at least one fluidized catalytic bed, and in which,in the mixture of paraffins, olefins, aromatics and water produced in conversion step (a), the ratio of the mass of C3+ olefins to the total mass of olefins is greater than or equal to 0.8. The process according to the invention may comprise one or more of the following characteristics, taken in isolation or in any technically possible combination: - the jet fuel fraction comprises between 2% by volume and 30% by volume of C8+ aromatics, preferably between 8% by volume and 25% by volume of C8+ aromatics, the mass content of aromatics in the water-depleted mixture being preferably greater than 6% by mass, in particular between 6% and 10% by mass; - the process comprises the separation of at least a portion of the oligomerized olefins in step (c) and / or at least a portion of the alkylated aromatics in step (d) into a fraction of C7- hydrocarbons, and into a fraction of C8+ hydrocarbons,the C7- hydrocarbon fraction being at least partially recycled to step (c) of oligomerization of olefins and / or to step (d) of alkylation of aromatics, the stream of hydrocarbons to be hydrogenated being formed by at least part of the C8+ hydrocarbon fraction; - step (c) of oligomerization of olefins and step (d) of alkylation of aromatics are carried out jointly in the same reactor; - the process comprises a step of separating the water-depleted mixture into a C1-C2 hydrocarbon fraction and a C3+ hydrocarbon fraction, at least part of the C1-C2 hydrocarbon fraction being conveyed to a steam cracker to extract an ethylene stream therefrom, at least part of the C3+ hydrocarbon fraction being sent to step (c) of olefin oligomerization and to step (d) of aromatic alkylation,a portion of the C1-C2 hydrocarbon fraction being optionally recycled to step (a) of converting the C1 to C6 alcohol stream; - the process comprises separating the C3+ hydrocarbon fraction, to form a C3- hydrocarbon fraction and a C4+ hydrocarbon fraction, the C4+ hydrocarbon fraction being sent to step (c) of olefin oligomerization and to step (d) of aromatic alkylation; - step (c) of olefin oligomerization is carried out in an oligomerization reactor, and step (d) of aromatic alkylation is carried out in an alkylation reactor, separately from step (c) of olefin oligomerization; - the process comprises: separating the water-depleted mixture into a C1-C2 hydrocarbon fraction, a C3-C5 hydrocarbon fraction, and a C6+ hydrocarbon fraction,the C1-C2 hydrocarbon fraction and the C6+ hydrocarbon fraction being sent at least in part to the alkylation step (d) in the alkylation reactor, the C3-C5 hydrocarbon fraction being sent to the oligomerization step (c) in the oligomerization reactor or comprising: separating the water-depleted mixture into a C3- hydrocarbon fraction, a C4-C5 hydrocarbon fraction, and a C6+ hydrocarbon fraction, the C3- hydrocarbon fraction and the C6+ hydrocarbon fraction being sent at least in part to the alkylation step (d) in the alkylation reactor,the C4 to C5 hydrocarbon fraction being sent at least in part to the oligomerization step (c) in the oligomerization reactor; - a product from the oligomerization reactor containing the oligomerized olefins and a product from the alkylation reactor containing the alkylated aromatics are separated into the C8+ hydrocarbon fraction and the C7- hydrocarbon fraction, the C7- hydrocarbon fraction being recycled at least in part to step (c) in the oligomerization reactor, at least a part of the C8+ hydrocarbon fraction forming the hydrocarbon stream to be hydrogenated; - the product of the oligomerization reactor and the product of the alkylation reactor containing the alkylated aromatics are separated into a C7- hydrocarbon fraction, a C8 to C16 hydrocarbon fraction and a C17+ hydrocarbon fraction, at least a portion of the C8 to C16 hydrocarbon fraction forming the hydrocarbon stream to be hydrogenated,the C17+ hydrocarbon fraction being at least partly recycled to step (a) of conversion of the C1 to C6 alcohol stream; - between 10% and 90% by mass of the aromatics contained in the hydrocarbon stream to be hydrogenated, preferably between 30% by mass and 80% by mass of the aromatics contained in the hydrocarbon stream to be hydrogenated are hydrogenated to cycloparaffins in step (f); - the C1 to C6 alcohol stream is introduced into conversion step (a) at a temperature at least 5°C higher than the bubble point of the C1 to C6 alcohol stream, and lower than the temperature of the conversion reaction carried out in step (a), in particular at least 50°C lower than the temperature of the conversion reaction carried out in step (a),advantageously at least 100°C lower than the temperature of the conversion reaction carried out in step (a); - the method comprises stripping at least a portion of the water separated in separation step (b) to produce a stream of extracted hydrocarbons, advantageously recycled to separation step (b), and a stream of treated water; - the fluidized catalytic bed receives, in addition to the C1 to C6 alcohol stream, an exothermicity control stream comprising recycled water separated from the mixture in separation step (b) and / or C4- hydrocarbons obtained from the water-depleted mixture (19), the mass ratio of recycled water in the exothermicity control stream to the C1 to C6 alcohol stream in the feedstock supplied to conversion step (a) being advantageously between 0 and 1, preferably between 0.05 and 0.5; - after step (f) of hydrogenation of the hydrocarbon stream to be hydrogenated,the process comprises separating the hydrogenated hydrocarbon stream into at least the jet fuel fraction and a diesel fraction; - the C2 to C6 alcohol stream is obtained by fermentation of biomass or synthesis gas, and / or the C1 to C6 alcohol stream is obtained by catalytic conversion of carbohydrates,of carbon monoxide or carbon dioxide in the presence of hydrogen; - step (f) of hydrogenation of the hydrocarbon stream to be hydrogenated is carried out separately and downstream of step (c) of oligomerization of olefins from the water-depleted mixture and of step (d) of alkylation of aromatics from the water-depleted mixture; - step (a) of conversion is carried out using at least one catalyst comprising molecular sieves containing at least pores with 10 oxygen atoms (10-MR) or larger in their microporous structure; - step (a) of conversion is carried out in the presence of a conversion catalyst comprising a phosphorus-modified zeolite partially having an ALPO structure or in the presence of a conversion catalyst comprising a B-modified zeolite; - the catalyst for carrying out conversion step (a) comprises a phosphorus-modified zeolite, in particular having a P content of at least 0,05% by mass and preferably between 0.3% by mass and 7% by mass, a composite catalyst comprising at least 0.1% by mass of silicate, or a molecular sieve modified with phosphorus and with an alkaline earth or rare earth metal (MP modified molecular sieve), advantageously having a molar ratio M / P in the molecular sieve of less than 1. - the catalyst for carrying out the conversion step (a) comprises a zeolite modified with phosphorus having an ALPO structure advantageously characterized by a 27Al NMR signature of between 35 ppm and 45 ppm; - the catalyst for carrying out the conversion step (a) has been modified by adding one or more metals selected from the metals of group IIB, in particular Zn, of group IIIB, in particular Ga, the transition metals of group VIIIB, in particular Fe and / or Ni and / or Pt, of group VIB, in particular Mo,of group IB in particular Cu and / or Ag or of the lanthanide group in particular La; - the catalyst for carrying out conversion step (a) comprises a zeolite with pores with 10 oxygen atoms (10-MR) or more, modified by adding B before, after or simultaneously with the final catalyst formulation step; - the catalyst for carrying out conversion step (a) has an Si / Al atomic ratio, measured by chemical analysis in particular by NMR, taking into account only the Al which are part of the molecular sieve network structure of between 4 and 500, preferably between 5 and 200, or more preferably between 12 and 150; - the process comprises stripping at least part of the water separated in separation step (b) to produce a stream of extracted hydrocarbons, advantageously recycled to separation step (b),and a stream of treated water. The invention also relates to a use of the jet fuel fraction produced by implementing the manufacturing method as defined above, (i) pure, or (ii) in a mixture with a jet fuel resulting from a distillation of a petroleum and / or another renewable source, to supply at least one aircraft engine. Advantageously, the jet fuel fraction produced comprises between 2% by mass and 30% by mass of aromatics, in particular between 6% by mass and 20% by mass of aromatics. In particular, the jet fuel fraction comprises between 2% by mass and 30% by mass of aromatics having at least 8 carbon atoms, in particular between 6% by mass and 20% by mass of aromatics having at least 8 carbon atoms. Preferably,more than 50% by mass of the aromatics contained in the jet fuel fraction are monoaromatics having 8 to 14 carbon atoms. It advantageously comprises between 5% by mass and 20% by mass of cycloparaffins and at least 50% by mass of isoparaffins. The process according to the invention thus makes it possible to obtain a renewable aviation fuel that is totally substitutable ("drop-in") or compatible with the fuel and engine systems of current aircraft. The invention also relates to a jet fuel manufacturing plant comprising: - a stage for converting a C1 to C6 alcohol stream to form a mixture containing paraffins, olefins, aromatics,and water; - a stage for separating water from the mixture to form a water-depleted mixture; - a stage for oligomerizing olefins from the water-depleted mixture; - a stage for alkylating aromatics from the water-depleted mixture; - a stage for forming a stream of hydrocarbons to be hydrogenated from at least a portion of the olefins oligomerized in the oligomerization stage and at least a portion of the aromatics alkylated in the alkylation stage; - a stage for hydrogenating the stream of hydrocarbons to be hydrogenated (24) to form a stream of hydrogenated hydrocarbons; - a fractionation stage for recovering at least one jet fuel fraction from the stream of hydrogenated hydrocarbons; wherein the conversion stage comprises a reaction zone comprising at least one fluidized catalytic bed, the conversion stage being configured to produce a mixture containing paraffins, olefins, aromatics,and water in which the ratio of the mass of C3+ olefins to the total mass of olefins is greater than or equal to 0.8. In a variant, the installation comprises at least one recycling conduit to the conversion stage, of at least a portion of the water from the mixture separated at the water separation stage. DETAILED DESCRIPTION The terms "comprising" and "comprises" as used herein are synonymous with "including", "includes" or "contains", "containing", and are inclusive or without limits and do not exclude additional features, elements or steps of methods not specified. The expressions % by mass and % by mass have an equivalent meaning and refer to the proportion of the mass of a product relative to 100 g of a composition comprising it. The boiling points as mentioned herein are measured at atmospheric pressure,unless otherwise stated. An initial boiling point (hereinafter "IBP") is defined as the temperature value from which a first vapor bubble is formed. A final boiling point (hereinafter "FBP") is the highest temperature achievable during a distillation. At this temperature, no more vapor can be transported to a condenser. The determination of the initial and final points uses techniques known in the art and several methods adapted according to the distillation temperature range are applicable, for example NF EN 15199-1 (2020 version) or ASTM D2887 for the measurement of boiling points of petroleum fractions by gas chromatography, ASTM D7169 for heavy hydrocarbons, ASTM D7500, D86 or D1160 for distillates. By default, the expression stream, current, fraction, etc. "in Cn to Cm" designates a stream, a current, a fraction,etc. having a majority quantity (e.g., more than 50 mol%) of compounds having between n and m carbon atoms. The expression "Cn+" flow, stream, fraction, etc. means a flow, stream, fraction, etc. having a majority quantity (e.g., more than 50 mol%) of compounds having n or more carbon atoms. The expression "Cn-" flow, stream, fraction, etc. means a flow, stream, fraction, etc. having a majority quantity (e.g., more than 50 mol%) of compounds having n or fewer carbon atoms. Unless otherwise indicated, the percentages used are percentages by mass, and the pressures are absolute pressures. Obtaining the C1 to C6 alcohol stream The C1 to C6 alcohol stream mainly contains alcohols such as methanol, ethanol, propanols (n-propanol, i-propanol), butanols (n-butanol, i-butanol), pentanols (n-pentanols,i pentanol) and hexanols. It may comprise, in a minor part, C6+ alcohols, and / or oxygenated compounds such as methyl ethyl ether; dimethyl ether; diethyl ether; diisopropyl ether; formaldehyde; dimethyl carbonate; dimethyl ketone; acetic acid; furans, tetrahydrofurans and mixtures thereof. The C1 to C6 alcohol stream forming the feedstock of the process advantageously comprises more than 80% by mass of C1 to C6 alcohols, preferably more than 90% by mass of C1 to C6 alcohols. It advantageously comprises more than 50% by mass of methanol, in particular more than 80% by mass of methanol. In a variant, a C2 to C6 alcohol stream is further advantageously added between two conversion reaction zones of conversion step (a),as will be described below. The ratio of the mass flow rate of the C2 to C6 alcohol stream added between two reaction zones relative to the mass flow rate of the C1 to C6 alcohol stream entering the first reaction zone is for example less than 0.5, in particular between 0.05 and 0.5. Preferably, the C1 to C6 alcohol stream and advantageously the additional C2 to C6 alcohol stream are obtained from a renewable source such as biomass (including its constituents and derivatives) or from carbon monoxide or carbon dioxide, possibly captured, and hydrogen advantageously produced from renewable energies such as solar energy, wind, geothermal energy, waves or currents and / or energy whose production does not generate carbon dioxide such as nuclear energy. The production routes for the alcohols intended to form the C1 to C6 alcohol stream are for example,in a non-limiting manner: - anaerobic fermentation of sugars from biomass, in particular to obtain ethanol; - catalytic reaction of hydrogen with carbon dioxide or carbon monoxide to obtain in particular methanol, ethanol and other alcohols; - catalytic reaction of hydrogen with carbohydrates to obtain in particular C1 to C6 alcohols; - ABE fermentation (ethanol, acetone, butanol), to obtain ethanol and n-butanol; - anaerobic fermentation of sugars from biomass in particular to obtain propanol (iso or n), butanol (iso or n) or isoamyl alcohol; - anaerobic fermentation of a mixture containing at least carbon monoxide, carbon dioxide and hydrogen to obtain in particular ethanol, propanol (iso or n), butanol (iso or n) or isoamyl alcohol; For obtaining alcohols, in particular ethanol of renewable origin by fermentation in one embodiment, the alcohol,in particular, ethanol of renewable origin can be obtained by ethanolic fermentation in a bioreactor containing a culture of one or more microorganisms. Ethanol of renewable origin can then advantageously be obtained by: - ​​anaerobic fermentation of a substrate rich in sugars from biomass, or - anaerobic fermentation of a gas comprising CO, which may or may not be from biomass. For anaerobic fermentation, ethanol can thus be produced by anaerobic fermentation of a substrate rich in sugars from biomass. Sugars are composed of chains of 6 or 5 carbons, such as glucose, sucrose (dimer of glucose and fructose), xylose and arabinose. This substrate may for example comprise, or come directly from agri-food plants, sugar cane, sugar beet, sugar sorghum, or by depolymerization of starch from corn, wheat, barley, rye, sorghum, triticale, potato,sweet potato, cassava, and / or cellulose and hemicellulose from lignocellulosic biomass. The sugar-rich substrate may also be derived from lignocellullosic biomass by a treatment comprising (i) a step of separating the lignin, cellulose and hemicellulose contained in the lignocellullosic biomass, followed by (ii) a step of converting the cellulose and / or hemicellulose into sugars. Obtaining this type of substrate from lignocellulosic biomass is well known to those skilled in the art. Lignocellulosic biomass consists essentially of cellulose,of hemicellulose and lignin. This biomass comes from agricultural and forestry residues or by-products of wood or crop processing, whether woody plants or herbaceous plants. This lignocellulosic biomass may also include distillers' grains and allow the production of ethanol as described in document EP2675778. The first step (i) is a pretreatment step which makes it possible to separate the lignocellulosic matrix and release the cellulose and hemicellulose from the complex formed with the lignin by means of one or more pretreatments. Pretreatments include steam pretreatment (or steam explosion), hot water pretreatment (hydrothermal), ammonia explosion (AFEX), pretreatment in an acid medium or alkaline pretreatment. The steam explosion treatment consists of treating the biomass, preferably previously shredded or ground,with high-pressure saturated steam at temperatures of about 160–240°C and pressures of 0.7–4.8 MPa. The efficiency of steam treatment can be improved by adding H2SO4, CO2, or SO2 as a catalyst. In AFEX pretreatment, biomass is contacted with anhydrous liquid ammonia feedstock in a ratio of 1:1 to 2:1 (1–2 kg ammonia / kg dry biomass) for 10–60 min at 60–90°C and pressures above 3 MPa. Hydrothermal pretreatment is similar to steam explosion, but uses liquid water at elevated temperatures instead of steam. In the case of acid pretreatment, typically in the presence of dilute acid, an aqueous suspension of the cellulosic substrate is heated to the desired temperature and pretreated with preheated sulfuric acid (concentrations <4% mass) in a stainless steel reactor,The treatment is carried out at a temperature of 140 to 215°C. The residence time varies from a few seconds to a few minutes depending on the treatment temperature. Lime pretreatment is an inexpensive physicochemical alkaline treatment that improves the digestibility of cellulosic biomass. Using 0.1g of Ca(OH)2 / g of biomass, the treatment can be carried out over a wide temperature range from 25 to 130°C. The Organosolv process, which is a process for the delignification and / or saccharification of cellulosic materials and plant cultures, can also be used. In general, the Organosolv process involves the use of a mixture of water and a solvent such as alcohols or ketones and sometimes other solvents of a non-polar nature,as well as an acidic compound to facilitate hydrolysis. A process of this type is described for example in document US4470851A. Step (ii) is a step of converting cellulose and / or hemicellulose into sugars. It is also well known to those skilled in the art. It is typically a hydrolysis which can be catalyzed by acid or by enzymes such as cellulases, produced for example by the strain Trichoderma reesei, xylanases, xylosidases and arabinofuranosidases. The sugar-rich substrate is then subjected to fermentation. For example, this fermentation can be carried out using specialized microorganisms, and in particular yeasts, which make it possible to optimize the profitability of the production process, in particular the following yeasts: Ethanol Red® (Fermentie), Thermosacc® (Lallemand)), Angel Super Alcohol® (Angel®) and Fali® (AB Mauric)), the Saccharomyces cerevisiae yeast strains described in document FR3015985,Candida Shehatae or Pichia stipitis yeast strains, or any other suitable microorganism. For anaerobic fermentation, ethanol may be produced by anaerobic fermentation of a gas comprising CO. The substrate is then a gaseous substrate (a gas) containing CO. This gaseous substrate may be a by-product of an industrial process. In certain embodiments, the industrial process is selected from the group consisting of the manufacture of ferrous metal products, in particular steel mills, the manufacture of non-ferrous products, petroleum refining processes, the gasification of coal and / or biomass or biochar, the production of electric power, the production of carbon black, the production of ammonia, the production of methanol, the manufacture of coke, catalytic cracking (in particular during the regeneration of the catalyst carbon monoxide is produced) and the reforming of methane. In these embodiments,The gaseous substrate may be captured from the industrial process before it is emitted into the atmosphere, using any suitable method. Depending on the composition of the gas so captured, it may also be desirable to treat it to remove any unwanted impurities, such as dust particles, before introducing it into the fermentation. For example, the gas may be filtered or scrubbed by known methods. In other embodiments, the gaseous substrate may be derived from the gasification of biomass. The gasification process involves partial combustion of the biomass in a restricted supply of air or oxygen. The resulting gas generally comprises primarily CO and H2, with minor volumes of CO2, methane, ethylene, and ethane. For example, biomass by-products obtained during the extraction and processing of food products,such as sugar from sugarcane or starch from corn or cereals, or non-food biomass waste generated by the forestry industry, may be gasified to produce a CO-containing gas that may be used in the present invention. The gaseous substrate used typically has a substantial proportion of CO. The CO content of the gaseous substrate is typically 15% to 100% by volume, 15% to 95% by volume, 40% to 95% by volume, 40% to 60% by volume, and 45% to 55% by volume or is within any two of these limits. Advantageously, the CO-containing gas may comprise 25%, 30%, 35%, 40%, 45%, 50%, 55% or 60% CO by volume. Gases with lower CO contents, such as 6% by volume, may also be suitable, particularly when H2 and CO2 are also present. If the gaseous substrate contains CO, it is not necessary for the gaseous substrate to contain hydrogen,but this is not considered detrimental to ethanol production. The gaseous substrate may also contain CO2, for example, in a proportion of 1% to 80% by volume, or 1% to 30% by volume, or 5% to 10% by volume, or any two of these limits. Typically, carbon monoxide is added to the fermentation reaction in the gaseous state, or alternatively in the liquid state. For example, carbon monoxide can be supplied in a liquid by saturation. For example, a liquid can be saturated with a gas containing carbon monoxide, and then this liquid can be added to a bioreactor. This can be achieved using standard methodology. As an example, a microbubble dispersion generator (Hensirisak et. al. Scale-up of microbubble dispersion generator for aerobic fermentation; Applied Biochemistry and Biotechnolo RV Volume 101, Number 3 / October, 2002) could be used. Furthermore,It is often desirable to increase the CO concentration of the gas (or the partial pressure of CO in the gas) and thus increase the efficiency of fermentation reactions using CO as a substrate. Increasing the partial pressure of CO in the gas increases the mass transfer of CO into a fermentation medium. The composition of the gas streams used to feed a fermentation reaction can have a significant impact on the efficiency and / or costs of that reaction. For example, O2 can reduce the efficiency of an anaerobic fermentation process. Processing unwanted or unnecessary gases into the steps of a fermentation process before or after fermentation can increase the load on those steps (e.g., when the gas stream is compressed before entering a bioreactor, unnecessary energy may be used to compress gases that are not needed for fermentation). Therefore,It may be desirable to treat substrate streams, particularly substrate streams derived from industrial sources, to remove undesirable components and increase the concentration of desirable components. Any microorganism capable of fermenting a gaseous substrate comprising CO to produce ethanol may be used in the present invention. For example, microorganisms of the genera Moorella, Clostridia, Ruminococcus, Acetobacterium, Eubacterium, Butyribacterium, Oxobacter, Methanosarcina, Methanosarcina, and Desulfotomaculum may be used. For example, microorganism(s) of the genus Clostridium, including strains of Clostridium ljungdahlii, Clostridium carboxydivorans, Clostridium ragsdalei and Clostridium autoethanogenum, of the genus Moorella, including Moorella sp HUC22-1, of the genus Carboxydothermus, Moorella thermoacetica, Moorella thermoautotrophica, Ruminococcus productus, Acetobacterium woodii,Eubacterium limosum, Butyribacterium methylotrophicum, Oxobacter pfennigii, Methanosarcina barkeri, Methanosarcina acetivorans, or Desulfotomaculum kuznetsovii. Other specific examples of microorganisms are carboxydotrophic anaerobic bacteria. Examples of usable strains are described in WO201226833. It should be noted that the invention can be applied to a mixed culture of two or more microorganisms. With regard to the fermentation medium and conditions, regardless of the nature of the substrate (gaseous or not) used, for ethanol fermentation to occur by growth of one or more microorganisms, a suitable nutrient medium will have to be introduced into the bioreactor in addition to a substrate, under appropriate conditions. A nutrient medium will contain components, such as vitamins and minerals,sufficient to allow growth of the microorganism used. Reaction conditions to be considered are temperature, medium flow rate, pH, redox potential of the medium, stirring rate (if using a continuously stirred reactor), inoculum level, maximum substrate concentrations and substrate feed rates into the bioreactor to ensure that the substrate level does not become limiting, and maximum product concentrations to avoid product inhibition. Optimal reaction conditions will depend in part on the particular microorganism used. Methods for culturing microorganisms are known in the art and those skilled in the art will know how to optimize the culture conditions for each microorganism, depending on its nature. Examples of fermentation conditions suitable for anaerobic fermentation of a substrate comprising CO are detailed in WO2007 / 117157, WO2008 / 115080,WO2009 / 022925 and WO02 / 08438. The fermentation reactions can be carried out in any suitable bioreactor. In some embodiments of the invention, the bioreactor may comprise a first growth reactor in which the microorganisms are grown, and a second fermentation reactor, into which the broth from the growth reactor is introduced and in which most of the fermentation product (e.g., ethanol) is produced. The fermentation will result in a fermentation broth comprising a desirable product (ethanol) and / or one or more by-products (such as acetate and butyrate when the substrate is a CO-containing gas) as well as microorganism cells,in a nutrient medium. The ethanol recovery may include continuously removing a portion of the broth and recovering ethanol from the removed portion of the broth. For example, the removed portion of the ethanol-containing broth may be passed through a separation unit to separate, for example by filtration, the bacterial cells from the broth and produce a cell-free ethanol-containing permeate, and returning the microorganism cells to the bioreactor. In some embodiments, the recovery of ethanol and / or one or more other products or by-products produced in the fermentation reaction includes continuously removing a portion of the broth and separately recovering ethanol and one or more other products from the removed portion of the broth. For example, ethanol may be recovered from the fermentation broth using methods such as filtration,fractional distillation or evaporation, pervaporation, and extractive fermentation. Distillation of ethanol from a fermentation broth yields an azeotropic mixture of ethanol and water (i.e., 95% ethanol and 5% water). Anhydrous ethanol can then be obtained by using molecular sieve ethanol dehydration technology, which is also well known in the art. Renewable ethanol can also be obtained from biomass by converting a CO / H2-rich synthesis gas, the synthesis gas being derived from biomass. For example, biomass can be gasified to produce a CO / H2-rich synthesis gas (or syngas),this synthetic gas is then converted into methanol in the presence of a catalyst. A process of this type is for example described in document WO2012003901. The biomass used to produce syngas may in particular include wood fuels from natural forests and woodlands (e.g., sawdust), agricultural residues (e.g., rice husks, straw manure), energy crops that are grown exclusively for energy production (e.g., corn and oil palm), municipal waste (e.g., wood waste, rice, straw manure), energy crops that are grown exclusively for energy production (e.g., corn and oil palm), municipal waste (e.g., municipal solid waste and wastewater) and waste-derived biomass fuel (e.g.,wood pellets). It is also possible to obtain methanol from renewable sources. Methanol from renewable sources means methanol obtained from biomass. Biomass may include, but is not limited to, wood fuels from natural forests and woodlands (e.g., sawdust), agricultural residues (e.g., rice husks, straw manure), energy crops that are grown exclusively for energy production (e.g., corn and oil palm), municipal waste (e.g., wood waste, rice, straw manure), energy crops that are grown exclusively for energy production (e.g., corn and oil palm), municipal waste (e.g., municipal solid waste and wastewater), and waste-derived biomass fuel (e.g.,wood pellets). Methanol of renewable origin can in particular be obtained by converting a synthetic gas rich in CO / H2, this synthetic gas being derived from biomass. The biomass can for example be gasified to produce a synthetic gas (or "syngas" in English) rich in CO / H2, this synthetic gas then being converted into methanol in the presence of a catalyst. A process of this type is for example described in document WO2018134853A1 A synthetic gas suitable for subsequent conversion into methanol can also be obtained by partial oxidation in the presence of dioxygen of a biogas containing methane and CO2, this biogas resulting for example from the anaerobic digestion of biomass in the presence of one or more microorganisms. A process of this type is for example described in document WO2019060988A1. The alcohols intended to form the C1 to C6 alcohol stream and possibly the additional C2 to C6 alcohol stream,can also be obtained from carbon dioxide, in particular captured carbon dioxide. Several transformation routes exist. One example is the catalytic conversion of carbon dioxide into methanol in the presence of hydrogen. Another route consists of converting carbon dioxide into carbon monoxide by electroconversion or by inverse reaction of gas with water in the presence of hydrogen. Carbon monoxide is then converted by catalytic conversion into methanol, in the presence of hydrogen. The hydrogen used for the various operations described above is obtained in particular by steam reforming of methane, by reaction of gas with water, or is produced by electrolysis from renewable energies such as solar energy, wind, geothermal energy, waves or currents. Step (a) Conversion of the alcohol stream into C1 to C6 The conversion of the alcohol stream into C1 to C6 includes for example dehydration,carbon-carbon coupling and / or aromatization of at least one C1-C6 alcohol. The dehydration of C1-C6 alcohols, the carbon-carbon coupling and / or their aromatization can be carried out simultaneously. For the dehydration of methanol, it is generally converted into dimethyl ether, which is dehydrated to produce olefins having at least two carbon atoms. In this case of methanol and / or dimethyl ether, a carbon-carbon coupling takes place during dehydration. C2-C6 alcohols can be dehydrated to produce olefins containing the same number of atoms. The dehydration reactions of alcohols to produce alkenes have been known for a long time (J. Catal.7, p.163, 1967 and J. Am. Chem. Soc. 83, p. 2847, 1961). Many available solid acid catalysts can be used for the dehydration of alcohols ((Stud. Surf. Sci. Catal.51, p.260, 1989), EP0150832, Bulletin of the Chemical Society of Japan, vol 47(2),424-429 (1974)). However, γ-aluminas are most commonly used, especially for longer chain alcohols (with three or more carbon atoms). Indeed, catalysts with higher acidity, such as silica-aluminas, zeolites, heteropoly acids, or resin catalysts, can promote double bond displacement, backbone isomerization, and other olefin interconversion reactions. In addition, aromatization of C1-C6 alcohols occurs through oligomerization of olefin intermediates, cyclization of a chain with at least 6 carbons, and dehydrogenation of cycloparaffins to the corresponding aromatics. The mechanism of reactions during the conversion of C1-C6 alcohols is via acid catalysis. The catalyst can provide a proton to activate the molecules,alcohols and / or olefins via protonated intermediates. The stability of protonated intermediates or alkylcarbenium ions depends on the inductive effect of the substituents. Tertiary carbenium ions are the most stable, with primary carbenium ions being the least stable. Thus,Tertiary carbenium ions are most readily formed, and reactions involving the formation of primary carbenium ions are slow. Primary carbenium ions tend to transform into secondary or tertiary carbenium ions. The addition of carbenium ions to alkenes is the key step in the oligomerization of alkenes, and the addition of carbenium ions to aromatic hydrocarbons is the basis for the alkylation of aromatics with alkenes. Hydride transfer provides a route to convert a neutral molecule into a carbenium ion, or successive hydride transfer from the alkene to the carbenium ion results in the formation of aromatic compounds. The following reactions demonstrate this mechanism for propylene: Protonation: CH3-CH=CH2 + H, + → CH3-CH + -CH3 Dimerization: CH3-CH=CH2+ CH3-CH + -CH3→ CH3-CH(CH3)-CH2-CH + -CH3Deprotonation:CH3-CH(CH3)-CH2-CH + -CH3 → CH3-CH(CH3)-CH=CH-CH3 + H +Transfert H-: CH3-CH(CH3)-CH=CH-CH3 → CH3-C + (CH3)-CH=CH-CH3 + H- H- + R1 + → H-R1Cyclisation: CH3-C + (CH3)-CH=CH-CH3 → CH3-(-C=CH2-CH2-CH-CH2-) + + H + (méthylcyclopentyl cation) Isomérisation: CH3-(-C=CH2-CH2-CH-CH2-) + → (-CH2-CH2-CH2-CH2-CH2-CH-) + (cyclohexyl cation) (-CH2-CH2-CH2-CH2-CH2-CH-) + → (-CH2-CH2-CH2-CH2-CH = CH-) + H + (cyclohexène) Transfert H-: (-CH2-CH2-CH2-CH2-CH = CH-) → (-CH2-CH2-CH + -CH2-CH=CH-) + H- H- + R1 + → H-R1 (-CH2-CH2-CH-CH2-CH = CH-) + → (-CH2-CH2-CH=CH-CH = CH-) + H + (cyclohexadiene) (-CH2-CH2-CH=CH-CH = CH-) → (-CH + -CH2-CH=CH-CH = CH-) + H- H- + R1 + → H-R1 (-CH + -CH2-CH=CH-CH = CH-) → (-CH=CH2-CH=CH-CH = CH-) + H +(benzene) Overall, for the formation of an aromatic nucleus, three molecules of dihydrogen (six atoms) must be removed and these three hydrogen molecules will, following the hydride transfer mechanism, form three paraffins from the olefins. The hydride transfer mechanism typically occurs on catalysts with only an acid function and often requires severe conditions that are also conducive to coke formation. The production of aromatics from olefins (produced from alcohols) occurs on acid catalysts essentially by hydride transfer from one olefin to another. This produces more unsaturated molecules (and ultimately aromatics) and paraffins. These paraffins are not transformable because they are too inert to be converted under the optimal operating conditions for the conversion of C1-C6 alcohols. It is possible to add a dehydrogenating catalytic function,allowing aromatization by producing molecular hydrogen. If the catalyst also has a dehydrogenating function, the reactive intermediates can be converted into the corresponding aromatics by dehydrogenation. This is particularly the case with bifunctional catalysts, having an acid function and a dehydrogenating function. The dehydrogenating function can be provided by metals from groups VIB, VIIIB and IB and IIB and mixtures, preferably gallium, zinc or mixtures. In general, olefin dehydrogenation reactions are thermodynamically limited and require high temperatures. In order to favor the dehydrogenating pathway, another reactant can be added to conversion step (a) to shift the thermodynamic equilibrium. Advantageously, this reactant is carbon dioxide which can be converted into carbon monoxide and water: CO2 + H2 → CO + H2O The CO2 molecules,CO and H2 can be separated from the other hydrocarbons and recycled to the synthesis of alcohols, by fermentation of synthesis gas and by catalytic conversion of the synthesis gas to make methanol. The process then comprises the addition, in step (a) of conversion of the C1 to C6 alcohol stream, of a stream containing carbon dioxide, and the joint conversion of the carbon dioxide into carbon monoxide during step (a) of conversion of the C1 to C6 alcohol stream. Advantageously, the stream containing carbon dioxide comprises more than 5% by mass of carbon dioxide,in particular more than 10% by mass of carbon dioxide. The mass ratio in the feedstock supplied to the conversion step (a) of the carbon dioxide supplied in the carbon dioxide stream to the C1 to C6 alcohols supplied in the C1 to C6 alcohol stream is between 5% and 75%. The stream containing carbon dioxide is added for example as a mixture with the C1 to C6 alcohol stream or if several reaction zones are arranged in series to carry out the conversion step, between two reaction zones or in a given reaction zone. Preferably, more than 2 mol%, in particular more than 5 mol% of the carbon dioxide is converted to carbon monoxide together with the conversion of the C1 to C6 alcohol stream in the conversion step (a). The main product of the acid-catalyzed dehydration of ethanol and / or methanol is ethylene and / or propylene and water. More generally a mixture containing paraffins, olefins,aromatics, and water is produced. Paraffins include n-paraffins, i-paraffins and cycloparaffins. According to the invention, in the mixture of paraffins, olefins, aromatics and water produced in conversion step (a), the ratio of the mass of C3+ olefins to the total mass of olefins is greater than or equal to 0.80, in particular greater than 0.82, preferably greater than or equal to 0.85, the ratio being calculated on the dry stream, after separation of the water. Advantageously, excluding water circulating in a water recycle, the mixture produced containing paraffins, olefins, aromatics, and water contains more than 10% by mass of water, in particular between 10% and 60% by mass of water depending on the alcohol composition in the mixture of C1 to C6 alcohols. In the case where only methanol is present, the water content of the mixture produced (excluding recycle) is between 55% by mass and 60% by mass. Advantageously,on a dry basis excluding water and any recycle, the mixture produced containing paraffins, olefins, aromatics, and water contains more than 2% by mass of aromatics, in particular more than 6% by mass of aromatics, in particular between 6% and 30% by mass of aromatics. On a dry basis, excluding water, it advantageously contains: - less than 5% by mass of methane, in particular between 0.1% by mass and 4% by mass of methane; - less than 5% by mass of dimethyl ether, in particular less than 1% by mass of dimethyl ether; - less than 5% by mass of residual C1 to C6 alcohols, in particular less than 1% by mass of residual C1 to C6 alcohols; - less than 15% by mass of ethylene, in particular between 5% by mass and 10% by mass of ethylene; - more than 30% by mass of propylene, in particular between 35% by mass and 60% by mass of propylene; - less than 15% by mass of paraffins, in particular between 3% and 8% by mass of paraffins; - more than 10% by mass of C4 to C7 olefins,in particular between 25% and 40% of C4 to C7 olefins; - more than 6% of aromatics, in particular between 6% and 10% of aromatics, in particular less than 3% of C9+ aromatics. The production of light olefins (ethylene and propylene) from a mixed alcohol feedstock in an oxygenate-to-olefin process is for example described in US Patent 7,288,689. Said patent proposes various processes for producing C1 to C4 alcohols, optionally in a mixed alcohol stream, and optionally for converting the alcohols into light olefins. The dehydration and aromatization conversion produces light olefins, having at least 2 carbon atoms and aromatics from C1-C6 alcohols using a composite catalyst comprising the following steps: a) providing a catalyst comprising molecular sieves containing at least 10 oxygen atom (10-MR) or larger pores in their microporous structure,b) carrying out the conversion in a reaction zone having a fluidized bed, providing a reaction zone and a catalyst regeneration zone, said catalyst circulating in both zones, such that at least a portion of the regenerated catalyst passes into the reaction zone and at least a portion of the catalyst in the reaction zone passes into the regeneration zone; c) contacting the C1 to C6 alcohols in the reactor with the catalyst under conditions effective to convert at least a portion of the feedstock to form a reactor effluent comprising a mixture of paraffins, olefins, aromatics, and water. The catalyst may be a mixture of two or more catalysts and optionally a binder. It is desirable to have substantially 100% conversion of the alcoholic compound in the reactor. This conversion rate is adjusted by optimizing the contact time,the reaction temperature and the catalyst regeneration frequency. In a specific embodiment, the weight hourly space velocity (WHSV) of the alcohol in the reaction zone is about 0.5 h, -1 at about 10 a.m. -1 , advantageously around 1 hour -1 at about 6 a.m. -1The molecular sieves used in the composition of the catalyst are selected from the list of molecular sieves with crystalline structure MFI, MOR, MEL, clinoptilolite, FER, FAU, MWW, BETA, MCM-41, ZSM-21, ZSM-22, ZSM-23, ZSM-42, ZSM-57, LTL or a mixture thereof. Preferably, the selected molecular sieve is a zeolite, a crystalline aluminosilicate selected from the group MFI, MOR, MEL, clinoptilolite, FER or a mixture thereof. More preferably, in the case of MFI, the molecular sieve is preferably a ZSM-5 zeolite. In another embodiment, the molecular sieve is preferably obtained without the addition of a structuring agent. Other examples are described by the International Zeolite Association (Atlas of Zeolite Structure Types, 1987, Butterworths).Crystalline silicates are microporous crystalline inorganic polymers based on XO4 tetrahydric networks linked to each other by sharing oxygen ions, where X can be trivalent (e.g. Al, B, ... ) or tetravalent (e.g. Ge, Si, ... ). The crystal structure of a crystalline silicate is defined by the specific order in which a network of tetrahedral units are linked together. The size of the pore openings of the crystalline silicate is determined by the number of tetrahedral units, or, alternatively, oxygen atoms, required to form the pores and the nature of the cations that are present in the pores. They possess a unique combination of the following properties: high internal surface area; uniform pores with one or more discrete sizes; ion exchangeability; good thermal stability; and the ability to adsorb organic compounds.Since the pores of these crystalline aluminosilicates are similar in size to many organic molecules of practical interest, they control the entry and exit of reactants and products, resulting in particular selectivity in catalytic reactions. Crystalline aluminosilicates with the MFI structure possess a bidirectional intersecting pore system with the following pore diameters: a straight channel along

[0010] : 0.53-0.56 nm and a sinusoidal channel along

[0100] : 0.51-0.55 nm. Crystalline aluminosilicates with the MEL structure possess a bidirectional intersecting straight pore system with straight channels along

[0100] having pore diameters of 0.53-0.54 nm.56 nm and a sinusoidal channel along

[0100] : 0.51-0.55 nm.The molecular sieves used in the present invention (in H+ or NH4+ form) have an initial Si / Al ratio advantageously between 4 and 500, preferably from 4 to 100, or more preferably from 4 to 30. The conversion to the H+ or NH4+ form is known per se and described in US3911041 and US5573990. The Si / Al atomic ratio is measured by chemical analysis, for example by NMR. It includes only the Als which are part of the lattice structure of the molecular sieve.According to a first embodiment, said zeolite is a phosphorus-modified zeolite manufactured by a method comprising in this order: - Selection of a molecular sieve as described in the list above; - Introduction of P under conditions effective to advantageously introduce at least 0.05% by mass of P by adding an aqueous solution containing a phosphorus precursor; - Optional separation of the solid from the aqueous liquid; - Optional washing or optional drying or optional drying followed by washing; - Calcination. Optionally, the method for manufacturing said phosphorus-modified zeolite comprises the steps of steam heat treatment and leaching. The method consists of steam heating / steam heat treatment followed by leaching.It is generally known to those skilled in the art that steam treatment of zeolites results in aluminum leaving the zeolite lattice and residing as aluminum oxides inside and outside the pores of the zeolite. This transformation is known as zeolite dealumination and this term will be used throughout the text. In the steam treatment step, the temperature is preferably 400°C to 870°C, more preferably 480°C to 760°C. The pressure is preferably atmospheric pressure and the partial pressure of water may range from 13 kPa to 100 kPa. The steam atmosphere preferably contains from 5% to 100% by volume of steam with from 0% to 95% by volume of an inert gas, preferably nitrogen. The steam treatment is preferably carried out for a period of 0.01 hours to 200 hours, advantageously 0.05 hours to 200 hours, more preferably 0.05 hours to 50 hours.Steam treatment tends to reduce the amount of tetrahedral aluminum in the crystalline silicate lattice, forming alumina. Treatment of steam-treated zeolite with an acid solution results in the dissolution of extra-lattice aluminum oxides. This transformation is known as leaching, and this term will be used throughout the text. A leach may be made with an organic acid such as citric acid, formic acid, oxalic acid, tartaric acid, malonic acid, succinic acid, glutaric acid, adipic acid, maleic acid, phthalic acid, isophthalic acid, fumaric acid, nitrilotriacetic acid, hydroxyethylenediaminetetraacetic acid, ethylenediaminetetraacetic acid, trichloroacetic acid, trifluoroacetic acid or a salt of such an acid (e.g. the sodium salt) or a mixture of two or more of these acids or salts.The other inorganic acids may include an inorganic acid such as nitric acid, hydrochloric acid, methanesulfuric acid, phosphoric acid, phosphonic acid, sulfuric acid or a salt of such an acid (e.g. sodium or ammonium salts) or a mixture of two or more such acids or salts. The leaching with an acidic aqueous solution containing a phosphorus source is advantageously carried out under reflux conditions, i.e. at the boiling temperature of the solution. The amount of said acidic solution is advantageously between 2 liters and 10 liters per kg of molecular sieve. A typical leaching period is about 0.5 hours to 24 hours. Advantageously, the acidic aqueous solution containing the P source in the leaching step has a pH of 3, advantageously 2, or lower.Advantageously, said aqueous acid solution is a solution of phosphorus acids, a mixture of phosphorus acids and organic or inorganic acids or mixtures of salts of phosphorus acids and organic or inorganic acids. The phosphorus acids or the corresponding salts may be phosphate ([PO4]. 3- , being basic), of phosphite type ([HPO3] 2-, being dibasic) or hypophosphite ([H2PO2]-, being monobasic). Unexpectedly, a larger amount of phosphorus than would be expected from the typical pore volume of the molecular sieve and assuming that the pores of the molecular sieves are filled with the phosphorous acid solution used, remains in the solid molecular sieve material. Both factors, namely dealumination and P retention, stabilize the lattice aluminum in the zeolite lattice, thus avoiding further dealumination. This leads to higher hydrothermal stability, tuning of the molecular sieve properties, and adjustment of the acidic properties, thus increasing the selectivity of the molecular sieve. Then the zeolite is separated, advantageously by filtration, and optionally washed. A drying step can be considered between the filtration and washing steps.The solution after washing can either be separated, for example, by filtration from the solid, or evaporated. P can be introduced by any means or, for example, according to the recipe described in US 3,911,041, US 5,573,990 and US 6,797,851. The separation of the liquid from the solid is advantageously carried out by filtration at a temperature between 0°C-90°C, centrifugation at a temperature between 0°C-90°C, evaporation or equivalent. Optionally, the zeolite can be dried after separation before washing. Advantageously, said drying is carried out at a temperature between 40°C-60°C, advantageously for 1h-10h. This drying can be carried out either under static conditions or in a gas flow. Air, nitrogen or any inert gas can be used.The washing step can be carried out either during filtration (separation step) with a portion of cold (<40°C) or hot (>40 but <90°C) water or the solid can be subjected to an aqueous solution and treated under reflux conditions for 0.5 h to 10 h followed by evaporation or filtration. The final calcination step is advantageously carried out at a temperature of 400°C-700 °C either under static conditions or in a gas flow. Air, nitrogen or any inert gas can be used.According to one embodiment of the invention, the phosphorus-modified zeolite is manufactured by a process comprising in this order: - Selection of a molecular sieve as described in the list above; - Heat treatment with steam at a temperature ranging from 400°C to 870°C for 0.01 -200h; - Leaching with an aqueous acid solution under conditions effective to remove a substantial portion of Al from the zeolite; - Introduction of P with an aqueous solution containing the P source under conditions effective to advantageously introduce at least 0.05% by mass of P - Separation of the solid from the liquid; - An optional washing step or an optional drying step or an optional drying step followed by a washing step; - A calcination step.Optionally, between the steam heat treatment step and the leaching step, there is an intermediate step such as, for example, contacting with silica powder and drying. Advantageously, the final P content is at least 0.05% by mass and preferably between 0.3% by mass and 7% by mass. Advantageously, at least 10% by mass of Al, relative to the parent zeolite MFI, MEL, FER, MOR and clinoptilolite, has been extracted and removed from the zeolite by leaching. Then, the zeolite is either separated from the washing solution or dried without separation of the washing solution. Said separation is advantageously carried out by filtration. Then the zeolite is calcined, for example, at 400°C for 2 hours to 10 hours. The residual P content is adjusted by the P concentration in the aqueous acid solution containing the P source, the drying conditions and a washing procedure if any.A drying step may be considered between the filtration and washing steps. The phosphorus-modified zeolite catalyst may be the phosphorus-modified zeolite itself or it may be the phosphorus-modified zeolite formulated into a catalyst by combining with other materials that provide additional hardness or catalytic activity to the finished catalyst product. According to a second embodiment, the catalyst of the process is a composite catalyst manufactured by a process comprising the following steps: a) Selecting a molecular sieve from the list as defined above b) Contacting the molecular sieve with a metal silicate comprising at least one alkaline earth metal, such that the composite comprises at least 0.1% by mass of silicate.The molecular sieve is preferably brought into contact with the metal silicate by one of the following two methods: - During the catalyst formulation step by mechanically mixing the molecular sieve with the metal silicate forming a precursor to be used in the formulation step; - Physically mixing the previously formulated molecular sieve and the previously formulated metal silicate in situ in the reaction medium intended for carrying out the conversion. Said molecular sieve and / or said composite catalyst containing the molecular sieve and the metal silicate may be post-treated by calcinations, reductions or hydrothermal steam treatment. In the case of the use of zeolites as molecular sieve components, the phosphorus may be introduced before, simultaneously or after mixing with the metal silicate.In a particular embodiment of the invention, the molecular sieve may be modified either before or after the introduction of the metal silicate. Preferably, the molecular sieve has undergone some form of modification before the introduction of the metal silicate. By modification is meant herein that the molecular sieve may have undergone steam heat treatment, leaching (e.g., acid leaching), washing, drying, calcination, impregnation, or some form of ion exchange. This means that at least some of the cations initially comprised in the crystal structure may be replaced by a wide variety of other cations according to techniques well known in the art. The replacing cations may be hydrogen, ammonium, or other metal cations, including mixtures of such cations.The selected molecular sieve is then formulated into a composite catalyst to comprise at least 10% by mass of a molecular sieve as described herein and at least one metal silicate comprising at least one alkaline earth metal, such that the composite comprises at least 0.1% by mass of silicate. At least one of the metal silicates comprised in the composite catalyst comprises at least one alkaline earth metal, preferably Ca. Metal silicates are insoluble in water and alkaline earth metal ions, particularly calcium, are polyvalent and have a large radius in the hydrated state. Thus, without wishing to be bound by theory, it is believed that the ion exchange reaction with the molecular sieve occurs very slowly, because the alkaline earth metal ion must lose many of its highly coordinated water molecules to enter the micropores of the sieve.As a result, the alkaline earth ions only expose the acidic sites located on the outer surface of the molecular sieve, thereby increasing the catalyst selectivity. Furthermore, without wishing to be bound by theory, the presence of silicate anions is believed to further enhance the catalytic properties of the composite catalyst. Silicate anions, for example, can provide silicon atoms to repair defects in the molecular sieve. This can thus lead to further stabilization of the catalyst under severe hydrothermal conditions. As a result, the metal silicate acts as a catalyst promoter. The metal silicate may comprise more than one alkaline earth metal selected from Ca, Mg, Sr, and Ba. The metal silicates may also comprise other metals selected from one or more of the following: Ga, Al, Ce, In, Cs, Sc, Sn, Li, Zn, Co, Mo, Mn, Ni, Fe, Cu, Cr, Ti, and V.Preferably, the other metal is selected from one or more of Al, Mg, Ce, Co and Zn or mixtures thereof. These bi-, tri- or polymetallic silicates can be synthesized by any method known in the art. This can be for example by ion exchange in solution or in the solid state (Labhsetwar et al., Reactivity of Solids, vol. 7, issue 3, 1989, 225-233). The silicate anion can be present in any form in the solid metal silicate. Examples include SiO3. 2- , SiO4 4- , Si2O7 6- , Si3O 10 8- and the like. The preferred catalyst promoter is a calcium silicate with a very open and accessible porous structure. An even more preferred catalyst promoter comprises a synthetic crystalline hydrated calcium silicate having a chemical composition of Ca6Si6O 17(OH)2which corresponds to the known mineral xonotlite (having a molecular formula 6CaO.6SiO2.H2O ) Generally, a synthetic hydrated calcium silicate is synthesized hydrothermally under autogenous pressure. A particularly preferred synthetic hydrated calcium silicate is commercially available from the company Promat of Ratingen in Germany under the brand name Promaxon. Other examples of metal silicates comprising alkaline earth metals include CaAI2Si2O8, Ca2AI2SiO7, CaMg(Si2O6)x as well as mixtures thereof. Before mixing with the molecular sieve, said metal silicate compounds may be modified by calcination, steam treatment, ion exchange, impregnation or phosphating. Said metal silicates may be an individual compound or may be part of mixed compounds.The metal silicate may be contacted with the molecular sieve by a step of simultaneously formulating a mixture of the metal silicate with the molecular sieve or an in situ mixture of separately formulated materials in the reaction medium before carrying out the conversion. Said contact may be achieved by mechanically mixing the molecular sieve with the metal silicate comprising an alkaline earth metal. This may be achieved by any known mixing method. The mixing may last for a period of time ranging from 1 minute to 24 hours, preferably from 1 minute to 10 hours. If not carried out in the in situ conversion reactor, it may be carried out in a batch mixer or in a continuous process, such as in an extruder, for example a single or twin screw extruder at a temperature of 20°C to 300°C under vacuum or high pressure. Said contact may be carried out in an aqueous or non-aqueous medium.Prior to the formulation step, other compounds facilitating the formulation may be added, such as thickening agents or polyelectrolytes improving the cohesion, dispersion and flow properties of the precursor. In the case of oil drop drying or spray drying, a rather liquid liquid (high water content) is prepared. In another embodiment, the contacting is carried out in the presence of phosphorus-containing compounds. In a particular embodiment, the contacting is carried out in an aqueous medium at a pH of less than 5, more preferably less than 3.According to a third embodiment, the catalyst of the process is a molecular sieve modified with phosphorus (P) and with an alkaline earth or rare earth metal (M) (MP-modified molecular sieve) manufactured by a process comprising the following steps: a) selecting at least one molecular sieve selected from - a P-modified molecular sieve which contains at least 0.3% by mass of P - a P-modified molecular sieve before or during step b) introducing at least 0.3% by mass of P b) contacting said molecular sieve with a compound containing an alkaline earth or rare earth metal (M-containing compound) to introduce at least 0.05% by mass of the alkaline earth or rare earth metal M. Optionally, contacting the molecular sieve with the P-containing compound and the M-containing compound may be carried out simultaneously.The introduction of the alkaline earth or rare earth metal (M) is achieved by contacting the molecular sieve with a solution of one or more compounds containing M. Said solution may contain a higher concentration of alkaline earth or rare earth metal than that found in the final MP-modified molecular sieve. The molecular sieve is selected from the list described previously. Prior to P-modification and / or alkaline earth or rare earth metal modification (M-modification), the molecular sieve may undergo further treatments, including steam heat treatment, leaching (e.g., acid leaching), washing, drying, calcination, impregnation, or ion exchange. In addition or alternatively, these steps may also be performed during or after P-modification.By ion exchange steps is meant herein that at least a portion of the cations initially comprised in the crystal structure are replaced by a wide variety of other cations according to techniques well known in the art. The replacing cations may be hydrogen, ammonium or other metal cations, including mixtures of such cations. The modification of molecular sieves with phosphorus is known per se. This modification is carried out by treating molecular sieves with P compounds in aqueous or non-aqueous media, by chemical vapor deposition of organic P compounds or by impregnation. The catalyst may be pre-formulated with or without a binder. Preferred P compounds typically used for this purpose may be selected from the group of phosphoric acid, NH4H2PO4 or (NH4)2HPO4. The M-containing compound may be selected from organic compounds, salts, hydroxides and oxides.These compounds may also contain phosphorus. It is essential that these compounds are present in solubilized form, before bringing them into contact with the molecular sieve or by forming a solution upon contact with the molecular sieve. The final molar ratio M / P in the molecular sieve MP is preferably less than 1.According to a particular embodiment of the invention, the molecular sieve may be modified with phosphorus according to the method comprising the following steps, in the order indicated: - steam heat treatment of the molecular sieve at a temperature ranging from 400°C to 870°C for 0.01 h -200 h; - leaching with an acidic aqueous solution containing the P source under conditions effective to remove a substantial portion of Al from the molecular sieve and to introduce at least 0.3% of phosphorus by mass of the molecular sieve; further modification may then be carried out according to the following steps, in the order indicated: - separation of the solid from the liquid; - an optional washing step or an optional drying step or an optional drying step followed by a washing step; - a calcination step.Preferably, the separation, the optional washing and drying steps and the calcination are carried out after the introduction of the M-containing compound into the molecular sieve. The metal M can be any alkaline earth or rare earth metal. Preferably, the alkaline earth metal is Ca. However, it is also possible to use Mg, Sr and Ba. Possible rare earth metals include La and Ce. Advantageously, the final P content of the molecular sieve is at least 0.3% by mass and preferably between 0.3% by mass and 7% by mass. Advantageously at least 10% by mass of Al has been extracted and removed from the molecular sieve by leaching. The residual P content is adjusted by the P concentration in the leaching solution, separating the conditions during the separation of the solid from the liquid and / or the optional washing procedure during which impregnation and / or adsorption can also take place.A drying step may be considered between the separation and / or washing steps. The molecular sieve is then either separated from the washing solution or dried without separation of the washing solution. Said separation is advantageously carried out by filtration. Then the molecular sieve is calcined, for example, at 400 °C for 2 hours to 10 hours. The M-modification of the molecular sieve is carried out either on an already P-modified molecular sieve or during / after the P-modification process. The P-modification may be carried out as described above in which the sieve is dealuminated by steam heat treatment and then leached with an acidic P-containing solution. In this case, advantageously, the treatment of the molecular sieve with the M-containing solution is carried out after the leaching or washing step, i.e. after the phosphorous compound has been added and the P-modification has taken place and before the separation step.However, the introduction of M into the molecular sieve can also be considered: - during the leaching step; - before the washing step but after leaching and drying; - on calcined molecular sieves brought into contact with P; - on molecular sieve which has not been leached to introduce P but which has been brought into contact with P during the washing step. The introduction of M onto the molecular sieves can be carried out either by impregnation or by adsorption from an aqueous solution of compounds containing M. The introduction of the compound containing M can be carried out at temperatures ranging from room temperature up to the boiling point of the solution. The concentration of the compound containing M in the solution is at least 0.05-M, preferably between 0.05 and 1.0 M.The amount of alkaline earth metal or rare earth (M) in the molecular sieves MP may vary from at least 0.05% by mass, preferably from 0.05% by mass to 7% by mass, more preferably from 0.1% by mass to 4% by mass. Before the formulation of the composite catalyst, the molecular sieve may undergo further treatments including steam treatment, leaching (e.g. acid leaching), washing, drying, calcination, impregnation and ion exchange steps. In addition or alternatively, these steps may also be carried out after formulation of the catalytic composite. The alkaline earth metal or rare earth M is preferably selected from one or more of: Mg, Ca, Sr, Ba, La, Ce. More preferably, M is an alkaline earth metal. More preferably, M is Ca. Particularly in the case of P modification by vaporization and leaching, M may be a rare earth metal such as La and Ce.The M-containing compound is preferably in the form of an organic compound, a salt, a hydroxide or an oxide. The compound is preferably in a solubilized form upon contacting it with the molecular sieve. Alternatively, the solution of the M-containing compound may be formed after contacting the molecular sieve with said compound. Possible M-containing compounds include metal M compounds such as sulfate, formate, nitrate, acetate of metal M, halides, oxyhalides, borates, carbonate, hydroxide, oxide and mixtures thereof. These may be, for example, calcium sulfate, formate, nitrate, acetate, halides, oxyhalides, borates, carbonate, hydroxide, oxide and mixtures thereof. The compound containing M may also comprise other metals selected from one or more of Mg, Sr, Ba, Ga, Al, Ce, In, Cs, Sc, Sn, Li, Zn, Co, Mo, Mn, Ni, Fe, Cu, Cr, Ti and V.The M-containing compounds may also further comprise phosphorus. These M-containing compounds, which are sparingly soluble in water, may be dissolved to form a well-solubilized solution by heating and / or changing the pH of the solution by adding phosphoric, acetic, or nitric acid or the corresponding ammonium acid. salts of said acids. The concentration of the M-containing compound is at least 0.05 M. Alkaline earth and rare earth metals M, especially Ca, have a large hydration sphere radius in the hydrated state. Thus, without wishing to be bound by theory, it is believed that the ion exchange reaction with the acid sites located within the microporous structures of the molecular sieve occurs very slowly. As a result, the metal M chosen exposes only the acid sites located on the outer surface of the molecular sieve, thereby increasing the selectivity of the catalyst.In the case of P-modified molecular sieves, the M modification leads to the formation of mixed M-Al-phosphates on the outer surface. Considering that phosphorus is more strongly bound to the alkaline earth or rare earth M than to Al, this modification leads to a stabilization of phosphorus on the outer surface of the molecular sieve where phosphorus is most labile. However, it is essential that all M atoms located on the outer surface are saturated with phosphorus. This can be ensured in the presence of excess phosphorus and by the presence of M in solution form, which serves for example to wash out excess phosphorus avoiding clogging of the micropore entrance. Before, after or simultaneously with the formulation step to form the composite, other components can optionally be mixed with the modified molecular sieve, whether MP or not.In a particular embodiment, the modified or unmodified molecular sieve (MP) may be combined with other materials that impart additional hardness or catalytic activity to the finished catalyst product. The materials, which may be mixed with the molecular sieve, may be various inert or catalytically active matrix materials and / or various binder materials. Such materials include clays, quartz, alumina or alumina sol, silica or silica sol, and / or metal oxides such as titanium oxide, zirconia, and mixtures thereof. In one embodiment, certain binder materials may also serve as diluents to control the feed-to-product conversion rate and therefore improve selectivity. In one embodiment, the binders also improve catalyst attrition under industrial operating conditions.Natural clays suitable as binders are, for example, clays of the kaolin family or the montmorillonite family. Such clays may be used in the raw state as extracted or they may be subjected to various treatments before use, such as calcination, acid treatment or chemical modification. In addition to the above, other materials which may be included in the composite catalyst of the invention include various forms of metals including rare earth or alkaline earth metals, phosphates (e.g. metal phosphates, where the metal is selected from one or more of Ca, Ga, Al, Ca, Ce, In, Cs, Sr, Mg, Ba, Sc, Sn, Li, Zn, Co, Mo, Mn, Ni, Fe, Cu, Cr, Ti and V).Examples of possible phosphates include amorphous calcium phosphate, monocalcium phosphate, dicalcium phosphate, dehydrated dicalcium phosphate, α- or β-tricalcium phosphate, octacalcium phosphate, hydroxyapatite, etc. zirconia, silica-thoria, silica-beryllium, silica-titanium, calcium-alumina. Examples of ternary binder compositions include, for example, calcium-silica-alumina or silica-alumina-zirconia. These components are effective in increasing catalyst density and increasing the strength of the formulated catalyst. The catalyst for use in fluidized bed reactors has a substantially spherical shape, generally formed by atomization. Generally when using a fluidized bed as a reactor, the catalyst particle size can vary from about 20 µm to 500 µm, more preferably from 30 µm to 100 µm.The crystal size of the molecular sieve contained in the composite catalyst is preferably less than about 10 μm, more preferably less than about 5 μm and most preferably less than about 2 μm. According to another embodiment, unmodified molecular sieves have first been formulated with a binder and matrix materials and then modified with phosphorus and alkaline earth metal silicates. According to another particular embodiment, the molecular sieves have optionally been dealuminated and then modified with phosphorus during the formulation step. The introduction of the alkaline earth metal silicate can be carried out during the formulation step or on the formulated solid. According to a preferred embodiment, the molecular sieves have first been optionally dealuminated and modified with phosphorus and then formulated.The introduction of the metal is carried out simultaneously with the phosphorus modification step and / or on the already formulated catalyst. After formulation, the composite catalyst may undergo further treatments including further steaming, leaching, washing, drying, calcination, impregnation and ion exchange steps. If the molecular sieve has not been modified with phosphorus prior to the mixture formulation step, i.e. the step of introducing the metal silicate into the molecular sieve, this may be carried out after such a step. According to a feature of this embodiment, the molecular sieve is a phosphorus-modified (P-modified) zeolite. Said phosphorus-modified (P-modified) zeolite has already been described above. According to another embodiment, the unmodified molecular sieve has first been formulated with a binder and a matrix material and then modified with phosphorus and metals.According to a particular embodiment, the molecular sieves have optionally been dealuminated and then modified with phosphorus during the formulation step. The introduction of the metal can be carried out during the formulation step or on the formulated solid. According to a preferred embodiment, the molecular sieve has first optionally been dealuminated and modified with phosphorus and then formulated. The introduction of the metal is carried out simultaneously with a phosphorus modification step and / or on a formulated catalyst. The final catalyst containing a phosphorus-modified zeolite advantageously has a 27Al NMR signature of between 35 ppm and 45 ppm, characteristic of the presence of an ALPO structure. The mass content of said AlPO4 structure in the catalyst can be up to 99% by mass and is advantageously between 10% and 98% by mass. The presence of this ALPO structure is characterized by the following method, illustrated in Figure 9.The measurement is performed by magic angle spinning (MAS) solid-state NMR performed on the Bruker Avance 500 spectrometer, with a 4 mm zirconia MAS probe at a rotation speed of 15 kHz. In order to obtain quantitative MAS spectra, single-pulse excitation was applied using a short pulse length of 0.6 μsec. Each spectrum results from 5000 scans separated by a 0.5 second delay. The chemical shifts of the 27Al spectra were referenced to AlCl3 solution (0.1M, (0 ppm)). In the case where there is only one zeolitic aluminum source in the catalyst, the content of the AlPO4 phase is estimated directly by a ratio of signal area at 35 ppm -45 ppm (centered at 39 ppm in Figure 9) in the 27Al MAS versus a total spectrum area between -50 ppm and 100 ppm.In the case where the binder contains aluminum and phosphorus, the content of the AlPO4 phase in the zeolite is estimated by a signal area ratio at 35 ppm-45 ppm in the 27Al MAS relative to the total area of ​​the spectrum between -50 ppm and 100 ppm after subtracting the signal intensities of the binders. To improve the selectivity towards the formation of aromatics, the catalysts as described above can be further modified by adding one or more metals selected from the metals of group IIB (e.g. Zn), IIIB (e.g. Ga), transition metals of group VIIIB (e.g. Fe and / or Ni and / or Pt), group VIB (e.g. Mo), group IB (e.g. Cu and / or Ag) or the lanthanide group (e.g. La).The introduction of the metal(s) can be done using various methods known to those skilled in the art, such as, but not limited to, ion exchange, dry or equilibrium impregnation, grafting, chemical vapor deposition (CVD). The introduction of the metal(s) is done from one or more solutions containing the metals in the form of salts. The metal salts are previously dissolved in the treatment solution; the counter-ion of the metals is chosen from sulfates, nitrates, carbonates, hydroxides, phosphates, carboxylates (formate, acetate, propionate for example), dicarboxylates (oxalate, malonate, succinate for example). After introduction of the metal(s), various treatments can be applied, including drying, calcination, steam heat treatment. The metal content is generally between 0% by mass and 5% by mass, preferably between 0% by mass and 2.5% by mass.The selected metal can typically be Ga. 3+ or Zn 2+, in the presence / absence of other metals such as Pt. The modification of the above-mentioned catalyst can also be carried out by adding B, at a rate of 0.1% by mass to 5% by mass, preferably from 0.1% by mass to 1% by mass, more preferably from 0.1% by mass to 0.5% by mass, in the presence or absence of the other above-mentioned metals. In another embodiment, the oxygenate conversion process is carried out on a catalyst comprising a zeolite with pores of 10 oxygen atoms (10-MR) or more, modified by adding B either before, after or simultaneously with the stage of formulating the final catalyst. Preferably, the catalyst may be, for example, a ZSM-5 modified with B. The B content in the final catalyst is between 0.1% by mass and 5% by mass, preferably between 0.1% by mass and 1% by mass, more preferably between 0.1% by mass and 0.5% by mass.Advantageously, the zeolite contained in the final catalyst has an Si / Al atomic ratio, measured by chemical analysis (for example by NMR), taking into account only the Als which are part of the network structure of the molecular sieve of between 4 and 500, preferably between 5 and 200, or more preferably between 12 and 150. With regard to the conversion stage in which the conversion step (a) is carried out, the C1-C6 alcohol stream is brought into contact with the catalyst described above in a reaction zone of at least one reactor under operating conditions to produce the mixture containing paraffins, olefins, aromatics, and water as defined above. In this step (a), converting alcohols, the mixture can generally be produced in a temperature range of 300°C to 600°C, especially between 330°C and 550°C, in particular between 350°C and 500°C or between 410°C and 580°C.The pressure can also vary over a wide range. Preferred pressures are in the range of about 100 kPa to about 5 MPa, the most preferred range being about 150 kPa to about 1.0 MPa. The foregoing pressures refer to the partial pressure of the oxygen-containing organic compounds. The conversion step (a) can be carried out in a single reaction zone or in several reaction zones arranged in series or in parallel. After a certain operating time, the catalyst must be regenerated. In particular, several reactors can be used so that the exothermicity of the reaction is controlled so as to avoid excessive temperatures. Preferably, the maximum temperature difference within a single reactor will not exceed 100°C and preferably 75°C. The reactor(s) can be of the isothermal or adiabatic fluidized bed type.The conversion reaction can also be carried out continuously in a configuration comprising a fluidized bed that forms a reaction zone where the reaction takes place and a fluidized bed that forms a regeneration zone where the regeneration takes place (for example by controlled combustion in the presence of oxygen) or fluidized beds connected in series, in which the feedstock passes from one bed to the other with cooling between them and in which the catalyst is mobile and circulates between the reactor and the catalyst regeneration zone. Fluidized beds offer significant advantages when the reactions are particularly highly exothermic. Once the solids in the bed are fluidized, the solids within the bed behave like a liquid. The size, shape, formation, rate of rise and coalescence of gas bubbles in fluidized beds have a quantitative similarity to those of gas bubbles in liquids.The liquid-like behavior of a fluidized bed therefore allows solids to be handled like a fluid, and supplying and / or extracting solids therefore becomes possible. Rigorous mixing in a fluidized bed allows for uniform temperature, even for highly exothermic reactions, and therefore allows for more flexible reactor control. Rigorous mixing also improves contact between solids and fluids, and improves heat and mass transfer. There are many variations of fluidized beds, which are, for example, described in available technical manuals (e.g., Handbook of fluidization and fluid-particle systems, Taylor&Francis Group LLC, 2003). Fluidization phenomena in gas-solid systems depend greatly on the types of powders used. There are several classifications, all based on Geldart's original work.Many catalysts used in fluidized bed systems are group A particles, characterized by dense phase expansion after minimal fluidization and before bubbling begins. Gas bubbles appear at the minimum bubbling velocity. Fluidization regimes can be broadly classified into two categories: particulate (smooth) and aggregative (bubbling) fluidization. In particulate fluidization, solid particles typically disperse relatively uniformly throughout the fluidization medium, without readily identifiable bubbles. Thus, particulate fluidization is sometimes also referred to as homogeneous fluidization. In heterogeneous or aggregative fluidization, voids (bubbles) containing no solids are typically formed and are observed in a bubbling fluidized bed or in a bed exhibiting slugging or "slugging."For gas-solid systems, there are several distinct fluidization regimes: fixed bed, particulate fluidization, bubble fluidization, slugging fluidization, and turbulent fluidization; criteria are available for each of these. When the operating speed is higher than the transport speed, so that recycling of entrained particles is necessary to maintain a bed, additional fluidization regimes are possible. Particle regime: Umf ≤ U < Umb For group A powders, the fixed bed expands homogeneously (particle fluidization) above the minimum fluidization velocity (Umf), and no bubbles are observed as long as the velocity remains below the minimum bubbling velocity (Umb). Bubbling regime: Umb ≤ U < Ums Bubbles appear when the gas velocity is increased above the minimum bubbling velocity (Umb).Gas bubbles form above the distributor, coalesce, and grow. The bubbling regime is characterized by the coexistence of a bubble phase and a dense / emulsion phase. The majority of the fluidizing gas is present in the form of bubbles, and therefore the gas velocity through the dense phase is very low. Slugging regime: Ums ≤ U < Uc With large bed height-to-diameter ratios, the bed provides sufficient time for the bubbles to coalesce into larger ones. When the bubbles reach approximately the size of the bed cross-section, the bed enters the slugging regime with a periodic passage of large bubbles forming slugs and a large, smooth fluctuation in the bed pressure drop. The velocity Uc corresponds to the bed operating conditions where the slugs reach their maximum diameter and the amplitude of the pressure fluctuation is greatest.Transition to the turbulent regime: Uc ≤ U < Uk When the gas velocity is continuously increased beyond this velocity Uc, large bubbles begin to fragment into smaller bubbles with a lower pressure fluctuation. This velocity is denoted Uk, and characterizes the transition from the bubbling regime to the turbulent regime. Turbulent regime: Uk ≤ U < Utr Up to the transport velocity (Utr), the bed is in the turbulent regime. Bubbles or voids are still present, although they are less distinguishable in the dense suspension. In this regime, interactions between gas voids and the dense phase / emulsion are vigorous and ensure efficient gas-solid contact. Rapid fluidization regime: U > Utr Beyond the transport velocity (Utr), particles begin to be entrained and continuous operation is no longer possible without the replacement or recycling of entrained and transported particles.Fast fluidized beds are typically characterized by a dense phase region at the bottom, close to the distributor, coexisting with a dilute phase region at the top. Particle velocity increases with elevation in the bed and thus the bed density decreases. Pneumatic transport: U >> Utr All particles introduced at the bottom of the fluidized bed are transported in the dilute phase with a concentration varying along the bed height. A typical example of a reaction zone is the riser fluidized bed used in fluid catalytic cracking (FCC) applications. Risers are vertical pipes with a high height-to-diameter ratio (>10) and the ideal riser approximates plug flow conditions, such that the catalyst and fluid phase pass through the riser with minimal remixing.In a transport fluidized bed reactor (fast fluidization or pneumatic transport), core-annular flow can occur in which a high-speed dilute core is surrounded by a denser, slower-moving annulus. When circulating mass fluxes are low, solids in the annulus flow downward at the wall. When circulating mass fluxes are high, solids in the annulus flow upward along the wall. This non-uniform flow phenomenon results in inefficient gas-solid contact and suboptimal catalyst performance, and significant remixing of gas and solids will occur, especially when there is downward flow in the wall region.For rapid fluidization, internals are used to redistribute the axial and radial gas-solid flow structure, i.e., to improve the uniformity of the gas-solid flow structure in space and thus promote radial gas-solid exchange. Fluidized transport reactors require recirculation of catalyst particles to the bottom of the reactor. This offers the possibility of controlling the catalyst density in the fluidized bed by recirculating more or less catalyst. At the bottom of the fluidized bed, the feed fluid is distributed homogeneously over the cross-section of the reactor vessel. At the top of the reaction zone, the reaction vapors are separated from the entrained catalyst by means of baffles, a disengagement zone, and cyclones.The catalyst is collected, freed from remaining hydrocarbons and is advantageously returned to the bottom of the fluidized bed zone by vertical collectors ("standpipe") and valves. For the exothermic reaction like the conversion carried out in step (a), it is preferable to have a homogeneous temperature across the catalyst bed (radially and axially) in order to avoid hot spots and to properly control the catalytic reaction. This can be accomplished by rapid recirculation and possibly remixing of the catalyst in the reactor vessel. Means of controlling the average reaction temperature are to introduce the feedstock into the reaction zone at a temperature lower than the average bed temperature and / or to remove heat from the catalyst bed by heat exchange.This heat exchange can be achieved by internal heat exchange tubes through which a cooling medium circulates and removes heat from the reactor vessel or by external heat exchange by circulating the hot catalyst, collected at the top of the reactor, around the heat exchanger tubes and recirculating the cooled catalyst back into the reactor vessel. Regarding catalyst regeneration, the C1 to C6 alcohol conversion reactor also has a regeneration zone (or regenerator) whose main purpose is to remove coke deposits on the catalyst by combustion with oxygen. Regenerators are fast fluidized bed systems. Typically, the regenerator includes a dense catalyst bed at the bottom of the vessel and a more dilute bed near the top of the vessel.There are two types of regenerators, which operate in either partial combustion or full combustion mode. In partial combustion mode, less than stoichiometric air is supplied to the regenerator. Most of the carbon is converted to carbon monoxide and only some to carbon dioxide. Ideally, all the oxygen is consumed and no oxygen is present in the flue gases. The CO / CO2 ratio in the flue gases is usually between 0.5 and 2.0. In full combustion mode, excess air is supplied to the regenerator. Ideally, all the carbon in the coke is converted to carbon dioxide, and no carbon monoxide is present in the flue gases. The residual oxygen content in the flue gases is between 1.0% by volume and 3.0% by volume on a dry basis.Partial combustion regenerators have several advantages over total combustion regenerators, especially when the catalyst is sensitive to high temperatures and the steam environment: (i) more coke can be burned for a given air flow rate, since the required amount of air is less than the stoichiometric amount, and (ii) the heat of combustion released is lower, allowing moderate temperature control and better preservation of catalytic activity in the presence of steam produced by hydrogen combustion. A potential disadvantage of the partial combustion regenerator is the higher amount of coke remaining on the regenerated catalyst. In the case of total combustion regeneration, the carbon remaining on the catalyst is low and the restoration of catalytic activity is higher.The potential disadvantage of total combustion regenerators is a higher heat release due to the total combustion reaction and therefore a more irreversible loss of catalytic activity. Using two-stage regeneration can reduce catalyst deactivation. In two-stage regeneration, the first stage operates at a moderate temperature to burn mainly the hydrogen present in the coke, which has a higher reaction rate, as well as some of the carbon. In the second stage, using excess air, the remaining carbon is burned at a higher temperature to carbon dioxide, and thanks to the absence of water vapor in the second-stage regenerator, high-temperature catalyst deactivation can be minimized.The use of fluidized beds allows for very precise control of the reaction exothermicity while providing continuous catalyst regeneration, promoting productivity and simplifying operations. One or more diluents may be present in the C1-C6 alcohol stream fed to the reaction zone, for example, in an amount of 1 mol% to 95 mol%, based on the total moles of all feed and diluent components introduced into the reaction zone. Typical diluents include, but are not limited to, helium, argon, nitrogen, hydrogen, water (optionally recycled), paraffins, alkanes (especially methane, ethane, and propane), aromatics, and mixtures thereof. Preferred diluents are water and nitrogen. Water may be injected as a liquid or vapor. The use of a diluent can offer two advantages.The first advantage is to reduce the alcohol partial pressure and therefore improve the selectivity to light olefins, mainly propylene. Generally, the lower the alcohol partial pressure, the higher the selectivity to light olefins and conversely, the higher the partial pressure, the higher the selectivity to heavy olefins such as butenes and pentenes. There is an optimum for the yield of light olefins depending on the partial pressure, the reaction temperature, the hourly space velocity of the feedstock and the properties of the catalyst. The second advantage of using a diluent is that it can act as a heat sink for the exothermic alcohol conversion. Thus, the higher the specific molar heat capacity, the more heat can be absorbed by the diluents.It is preferred that the diluents can be easily separated from the light olefin products, preferably by simple phase separation. Therefore, a preferred diluent is water. Diluents may be added in an amount of 1 mol% to 95 mol% of the combined feed (C1-C6 alcohol stream + diluents), preferably 10 mol% to 75 mol%. Step (b) of separating water from the mixture The water from the mixture containing paraffins, olefins, aromatics, and water produced in conversion step (a) is separated from the mixture at a water separation stage, to form a water-depleted mixture. The water separation step is preceded by a step of cooling the effluent from step (a), condensing the water as well as a portion of the hydrocarbons. The implementation temperature for this step is generally between 20°C and 100°C.The separation is based, for example, on the difference in density and solubility between water and the rest of the hydrocarbons. It is generally implemented in a three-phase separator tank for separating a water-rich aqueous phase (hereinafter the water separated from the mixture), a liquid hydrocarbon phase and a gaseous hydrocarbon phase. The water-depleted mixture contains less than 5% by mass, preferably less than 1% by mass of the water present in the mixture produced in step (a). The water separated from the mixture advantageously contains less than 10% by mass of hydrocarbons. The water separated from the mixture is optionally at least partially recycled to the conversion step (a). Since the conversion step (a) is implemented using at least one fluidized bed, the recycled water is optionally reinjected into the fluidized bed.In this case, the mass ratio of recycled water to the C1 to C6 alcohol stream in the feedstock supplied to conversion step (a) is advantageously between 0 and 1, preferably between 0.05 and 0.5. In both previous cases, the recycled water forms a stream that controls the exothermicity of the reaction, reduces the hydrocarbon partial pressures and modifies the acidity of the catalyst, which improves the olefin selectivity. The water separated but not recycled in conversion step (a) is advantageously treated by stripping in a stripping column to separate the hydrocarbons it contains in a stream of extracted hydrocarbons. The extracted hydrocarbon stream is reinjected into separation step (b). In the case where conversion step (a) is carried out using at least one fluidized bed, a stream of C4- hydrocarbons is optionally added to the fluidized bed in addition to or as a substitute for the recycled water.This stream is for example formed by at least a portion of a C1-C2 hydrocarbon fraction separated from the water-depleted mixture which will be described below. This stream also controls the exothermicity of the reaction. Advantageously, the C1 to C6 alcohol stream is introduced into conversion step (a) at a temperature at least 5°C higher than the bubble point of the C1 to C6 alcohol stream, and preferably lower than the temperature of the conversion reaction carried out in step (a), for example at least 50°C lower than the temperature of the conversion reaction, advantageously at least 100°C lower than the temperature of the conversion reaction. The reheating of this stream absorbs the calories released by the conversion of the alcohols.Optional steps for separating C1-C2 hydrocarbons and C3 hydrocarbons Advantageously, the water-depleted mixture is introduced into a separation stage comprising at least one distillation column (hereinafter referred to as a deethanizer) to separate the C1-C2 hydrocarbons (methane, ethane, ethylene) and other gaseous molecules lighter than C2 such as CO, CO2 and hydrogen from the rest of the water-depleted mixture. The distillation column operates, for example, at a pressure greater than 20 barg and preferably greater than 30 barg. A fraction of C1-C2 hydrocarbons is extracted from the top of the column. It contains more than 50% by mass of C1-C2 hydrocarbons and other gaseous molecules such as CO, CO2 and hydrogen.The C1-C2 hydrocarbon fraction preferably contains more than 90% by mass, in particular more than 95% by mass of the C1-C2 hydrocarbons and other gaseous molecules such as CO, CO2 and hydrogen contained in the water-depleted mixture. The C1-C2 hydrocarbon fraction is at least partly conveyed to an ethylene recovery unit, for example within a steam cracker. Thus, ethylene, even produced in a minor quantity during conversion step (a), can be recovered. In the variant in which a stream containing carbon dioxide is added to conversion step (a), the carbon dioxide, carbon monoxide and hydrogen present in the C1-C2 hydrocarbon fraction are optionally separated from the other hydrocarbons, in particular by distillation, membrane separation or pressure swing adsorption and combinations thereof.These compounds are then advantageously recycled to a preliminary alcohol synthesis step, in particular by synthesis gas fermentation and by catalytic conversion of the synthesis gas to produce in particular methanol. In particular, the methanol or ethanol thus produced are then advantageously recycled to form a part of the C1 to C6 alcohol stream. In an advantageous variant, in particular applicable to the case of a conversion step (a) implemented using at least one fluidized bed, a part of the C1-C2 hydrocarbon fraction is recycled to step (a) of conversion of the C1 to C6 alcohol stream in the form of a recycle stream. For example, the ratio of the mass flow rate of the part of the C1-C2 hydrocarbon fraction recycled to the conversion step (a) to the mass flow rate of the C1-C2 hydrocarbon fraction drawn from the distillation column is less than 1 and is in particular between 0.1 and 0.8.The C3+ hydrocarbon fraction is recovered at the bottom of the column. It comprises more than 90% by mass of the C3+ hydrocarbons contained in the water-depleted mixture. On a dry basis, it advantageously contains: - less than 5% by mass of C1-C2 hydrocarbons, - less than 15% by mass of paraffins, in particular between 3% and 10% by mass of paraffins, - more than 40% by mass of C3 to C7 olefins, in particular between 50% and 80% of C3 to C7 olefins, - less than 5% by mass of C8+ olefins, in particular between 0.1% by mass and 4.0% by mass of C8+ olefins; and / or - more than 6% of aromatics, in particular between 6% and 20% of aromatics. C3 to C7 olefins generally contain propylene. The C3+ hydrocarbon fraction contains in some cases more than 30% by mass of propylene. In one embodiment, the C3+ hydrocarbon fraction is sent directly to the oligomerization and alkylation step.Alternatively, in a particular embodiment, an additional step of separating the C3- hydrocarbons, in particular propylene, is implemented in a second distillation column. This separation allows the recovery of the propylene. The second distillation column operates for example at a pressure greater than 5 barg, and preferably greater than 10 barg. It produces at the top a fraction of C3- hydrocarbons, containing more than 50% by mass of propylene, and at the bottom, a fraction of C4+ hydrocarbons. The fraction of C3- hydrocarbons preferably contains more than 90% by mass, in particular more than 95% by mass of the C3- hydrocarbons contained in the fraction of C3+ hydrocarbons from the first distillation column. The fraction of C4+ hydrocarbons comprises more than 90% by mass of the C4+ hydrocarbons contained in the water-depleted mixture.On a dry basis, it advantageously contains: - less than 5% by mass of C3- hydrocarbons, - less than 25% by mass of paraffins, in particular between 10% and 15% by mass of paraffins, - more than 15% by mass of C4 to C7 olefins, in particular between 25% and 40% of C4 to C7 olefins, - less than 2% by mass of C8+ olefins, in particular between 0.5% by mass and 1.0% by mass of C8+ olefins; and / or - more than 6% of aromatics, in particular between 7% and 20% of aromatics, in particular less than 5% of C9+ aromatics. Step c) of oligomerization of olefins from the water-depleted mixture and step (d) alkylation of aromatics from the water-depleted mixture In a first embodiment, step (c) of oligomerization of olefins from the water-depleted mixture and step (d) of alkylation of aromatics from the water-depleted mixture are carried out jointly in the same reactor or the same reactors of the same reaction stage.Advantageously, the hydrocarbon feedstock formed from the C3+ hydrocarbon fraction or the C4+ hydrocarbon fraction described above is oligomerized for its olefins and alkylated for its aromatics by contact with an acid catalyst. For example, a multi-reactor plant may be used in which the exothermicity of the reaction may be controlled so as to avoid excessive temperatures. Preferably, the maximum temperature difference within the same reactor will not exceed 100°C and preferably 75°C. The reactor(s) may be of the isothermal or adiabatic type with a fixed or moving bed.The olefin oligomerization and aromatic alkylation reaction may be carried out continuously in a configuration comprising a series of fixed beds connected in series, in at least one operating reactor in which the raw material passes from one bed to another with cooling between them and at least one similar reactor connected in parallel, which undergoes a catalyst regeneration operation. The olefin oligomerization and aromatic alkylation reaction may be carried out continuously in a configuration comprising a series of moving beds connected in series, in which the raw material passes from one bed to another with cooling between them and in which the catalyst is mobile and circulates between the reactor and the catalyst regeneration zone. Preferably, the abovementioned steps are carried out jointly by means of at least two successive reactors.In the case of fixed bed reactors, alternatively a single reactor can contain several catalytic beds with cooling systems between the beds or be equipped with a quench flow injection in order to lower the temperature between the beds. The reaction conditions of the first reactor are chosen so as to convert a part of the low carbon number olefinic compounds (C3-C8) into intermediate olefins (C8+) and the alkylation of aromatics by light olefins.Advantageously, the first reactor comprises a first catalytic zone and operates at high temperature, for example greater than or equal to 200°C, and preferably less than 350°C, and a pressure of between 25 bar and 60 bar. The second reactor preferably operates at temperatures and pressures chosen so as to promote the conversion of a portion of the low carbon number olefinic compounds (C3-C8) into intermediate olefins (C8+) and the alkylation of the aromatics by light olefins. The effluent from the first reactor, comprising the unreacted olefins, the intermediate olefins, aromatics, water and optionally other compounds such as paraffins and optionally a reducing gas, then undergoes oligomerization and / or alkylation in this second reactor comprising a second catalytic zone, which makes it possible to obtain an effluent of heavier hydrocarbons, rich in distillate.A cooling section is advantageously provided between two successive reactors and possibly a flash drum. The mass flow rate through the oligomerization reactor(s) is advantageously sufficient to allow a relatively high conversion, without being too low to avoid undesirable parallel reactions. The weight hourly space velocity (WHSV) of the feedstock is, for example, 0.1 h. -1 at 8 p.m. -1 , preferably 0.5 h -1 at 10 a.m. -1 , more preferably 0.8 h -1 at 5 a.m. -1. The temperature at the inlet of the reactor(s) is advantageously sufficient to allow a relatively high conversion, without being too high in order to avoid undesirable parallel reactions. The temperature at the inlet of the or each reactor is for example from 150°C to 400°C, preferably from 180°C to 350°C, more preferably from 200°C to 290°C. The pressure across the olefin oligomerization and aromatic alkylation reactor(s) is advantageously sufficient to allow a relatively high conversion, without being too low in order to avoid undesirable parallel reactions. The pressure across the or each reactor is for example from 8 bara to 100 bara, preferably from 10 bara-85 bara, more preferably from 25 bara to 75 bara (bars, absolute pressure).As regards the nature of the catalyst, a first family of catalysts used comprises an acid catalyst of either amorphous or crystalline aluminosilicate type, or a silicoaluminophosphate, in H+ form, chosen from the following list and containing or not alkaline elements or rare earths: MFI (ZSM-5, silicalite-1, boralite C, TS-1), MEL (ZSM-11, silicalite-2, boralite D, TS-2, SSZ-46), ASA (amorphous silica-alumina), MSA (mesoporous silica-alumina), FER (Ferrierite, FU-9, ZSM-35), MTT (ZSM-23), MWW (MCM-22, PSH-3, ITQ-1, MCM-49), TON (ZSM-22, Theta-1, NU-10), EUO (ZSM-50, EU-1), ZSM-48, MFS (ZSM-57), MTW, MAZ, BEA (zeolite Beta), MOR (mordenite), FAU (faujasite type zeolite) LTL (zeolite L), zeolite Omega and the family of microporous materials composed of silica, aluminum, oxygen and possibly boron.The zeolite can be subjected to different treatments before use, which can be: ion exchange, modification with metals, steaming treatment, acid treatments or any other dealumination method, surface passivation by silica deposition, or any combination of the above-mentioned treatments. The alkali or rare earth content is from 0.05% by mass to 10% by mass, preferably from 0.2% by mass to 5% by mass. Preferably, the metals used are Mg, Ca, Ba, Sr, La, Ce used alone or in mixture. A second family of catalysts used includes phosphorus-modified zeolites optionally containing an alkali or a rare earth.In this case, the zeolite can be chosen from the following list: MFI (ZSM-5, silicalite-1, boralite C, TS-1), MEL (ZSM-11, silicalite-2, boralite D, TS-2, SSZ-46), MSA (mesoporous silica-alumina), FER (Ferrierite, FU-9, ZSM-35), MTT (ZSM-23), MWW (MCM-22, PSH-3, ITQ-1, MCM-49), TON (ZSM-22, Theta-1, NU-10), EUO (ZSM-50, EU-1), MFS (ZSM-57), ZSM-48, MTW, MAZ, FAU, LTL, BEA (zeolite Beta), MOR. Zeolite can be subjected to various treatments before use, which can be: ion exchange, modification with metals, steam treatment, acid treatments or any other dealumination method, mesoporization treatments, surface passivation by silica deposition, or any combination of the above-mentioned treatments. The alkali or rare earth content is from 0.05% by mass to 10% by mass, preferably from 0.2% by mass to 5% by mass.Preferably, the metals used are Mg, Ca, Ba, Sr, La, Ce used alone or in a mixture. A third family of catalysts used includes bifunctional catalysts, comprising: - a support, from the following list: MFI (ZSM - 5, silicalite- 1, boralite C, TS-1), MEL (ZSM-11, silicalite-2, boralite D, TS- 2, SSZ-46), ASA (amorphous silica-alumina), MSA (mesoporous silica-alumina), FER (Ferrierite, FU-9, ZSM-35), MTT (ZSM-23), MWW (MCM-22, PSH-3, ITQ-1, MCM-49), TON (ZSM-22, Theta-1, NU-10), EUO (ZSM-50, EU- 1), MFS (ZSM-57), ZSM- 48, MTW, MAZ, BETA, FAU, LTL, MOR, and microporous materials of the ZSM-48 family consisting of silicon, aluminum, oxygen and optionally boron. MFI or MEL (Si / Al> 25), MCM-41, MCM-48, SBA- 15, SBA- 16, SiO2, Al203, hydrotalcite, or a mixture thereof.- a metallic phase (Me) at a level of 0.1% by mass, the metal being selected from the following elements: Zn, Mn, Co, Ni, Ga, Fe, Ti, Zr, Ge, Sn, and Cr used alone or in a mixture. These metal atoms can be inserted into the tetrahedral structure of the support. The incorporation of this metal can be carried out either by adding this metal during the synthesis of the support, or incorporated after synthesis by ion exchange or impregnation, the metals then being incorporated in the form of cations, and not integrated within the structure of the support. The zeolite can be subjected to different treatments before use, which can be: ion exchange, modification with metals, steam treatment, acid treatments or any other dealumination method, mesoporization treatments, surface passivation by silica deposition, or any combination of the above-mentioned treatments.The alkali or rare earth content is from 0.05% by mass to 10% by mass, preferably from 0.2% by mass to 5% by mass. Preferably, the metals used are Mg, Ca, Ba, Sr, La, Ce used alone or in a mixture. A fourth family of catalysts used includes amorphous solids such as silica-alumina, silica-phosphate, silica-borate, silica-titanium, silica-zirconia and / or mixtures. The catalyst can be a mixture of the materials described previously in the four catalyst families. In addition, the active phases can also be combined with other constituents (binder, matrix) giving the final catalyst increased mechanical strength, or improved activity. If the hydrocarbon feedstock is oligomerized in a plant comprising several reactors in series, the reactors in the series can be loaded with the same catalyst or different catalysts.In a variant, step (c) of oligomerization of the olefins is carried out in an oligomerization reactor, and step (d) of alkylation of the aromatics is carried out in an alkylation reactor, separately from step (c) of oligomerization of the olefins. Advantageously, the water-depleted mixture is then separated in the first column into the fraction of C1 - C2 hydrocarbons, taken for example at the top of the column, into a fraction of C3 to C5 hydrocarbons, for example withdrawn at an intermediate stage of the column, and into a fraction of C6+ hydrocarbons, for example taken at the bottom of the column. The C1-C2 hydrocarbon fraction and the C6+ hydrocarbon fraction are sent to alkylation step (d) in the alkylation reactor, the C3-C5 hydrocarbon fraction being sent to oligomerization step (c) in the oligomerization reactor.Alternatively, the water-depleted mixture is separated into a C3- hydrocarbon fraction, taken for example at the top of the column, into a C4-C5 hydrocarbon fraction, for example withdrawn at an intermediate stage of the column, and into a C6+ hydrocarbon fraction, for example taken at the bottom of the column. The C3- hydrocarbon fraction and the C6+ hydrocarbon fraction are sent at least in part to the alkylation step (d) in the alkylation reactor, the C4 to C5 hydrocarbon fraction being sent at least in part to the oligomerization step (c) in the oligomerization reactor. The product of the oligomerization reactor contains more than 50% by mass of C7+ olefins, in particular more than 60% by mass of C9 to C12 olefins. Alkylation takes place under temperature and pressure conditions effective to maintain more than 20% by mass of the feedstock in the alkylation zone in the liquid phase.Advantageously, the alkylation of aromatics with alkenes is carried out in the liquid phase, the aromatics being present mainly in the liquid phase. It can be carried out using solid acid catalysts. Zeolites and silica-alumina catalysts possessing shape selectivity are generally used. In this process, the reactor conditions are chosen so that the alkene introduced into the reactor is predominantly dissolved in the aromatic feedstock. This is generally achieved by an optimal combination of operating conditions, such as pressure, temperature and the choice of catalyst, having a sufficiently high catalytic activity. The alkenes present in the gas phase can cause rapid deactivation of alkylation catalysts typically used in the liquid phase.Examples of usable operating conditions are given in US 4,891,458 which describes the liquid phase synthesis of ethylbenzene with zeolite beta, while US 5,334,795 describes the use of MCM-22 in the liquid phase synthesis of ethylbenzene; US 7,649,122 describes the use of MCM-22 in the liquid phase synthesis of ethylbenzene in the presence of a maintained water content. US 4,549,426 describes the liquid phase synthesis of alkylbenzene with steam-stabilized zeolite Y. US 8,134,036 describes the liquid phase aromatic alkylation on at least one catalyst bed containing a first catalyst modified by the inclusion of a rare earth metal ion. The types of products that can preferably be produced correspond to the following generic chemical formulas: monoalkyl benzene, dialkyl benzene and trialkyl benzene. The alkyl chains (Rx) each have 2 to 10 carbon atoms, preferably 2 to 6 carbon atoms.These chains can be of equal or different lengths. The aromatic compounds produced during step (a) of conversion of C1 to C6 alcohols are typically monoaromatics, possibly alkylated (benzene, toluene, ethylbenzene and xylenes), the alkylating agent being olefins. The alkylation reaction is exothermic, hence it may be useful to inject part of the aromatic compounds and / or part of the olefins between the different beds of the reactor, if there are several beds. Aromatics, having less than 8 carbon atoms can be recycled, as well as olefins that are too short, for example, having less than 5 carbon atoms. The alkylation catalyst is for example in the form of beads, but it is most often in the form of extrudates. It consists of an acidic solid mixed with an amorphous phase. The acidic solid is shaped using a matrix, which is an amorphous phase.The acidic solid is preferably at least one zeolite, preferably chosen from zeolites of structural type FAU and more particularly zeolite Y, zeolites of structural type MOR (zeolite mordenite), zeolites of structural type EUO, (i.e. zeolites EU-1, ZSM-50, TPZ-3), zeolite NU-87 of structural type NES, zeolite NU-86 (described in EP 463768 A), zeolite NU-85 (described in EP 462745 A), zeolite NU-88 (described in FR 2752567), and zeolite IM-5 (described in FR 2 754809), zeolite Beta, zeolite MCM-22, zeolite MCM-36, zeolite MCM-49 or zeolite MCM-56. Preferably, the catalyst is a beta zeolite having a silica / alumina molar ratio (expressed as SiO2 / Al2O3) of about 10 to about 200 or about 20 to about 50. The beta zeolite may have a low sodium content, for example less than about 0.2% by mass expressed as Na2O, or less than about 0.02% by mass.The sodium content can be reduced by any method known to those skilled in the art, such as by ion exchange. These zeolites are at least partly in acid form (H. + ), but may also contain cations other than H +and such as alkaline earths or rare earths. The zeolite catalyst may be modified with a rare earth metal ion, such as lanthanum, cerium, neodymium or praseodymium, for example. The BET surface area of ​​the catalyst used is in particular between 50 m2 / g and 900 m2 / g, preferably between 100 m2 / g and 700 m2 / g. The Na / Al ratio of the final catalyst is less than 5 atomic % and preferably less than 2%. The zeolite content in the catalyst is in particular between 5% by mass and 95% by mass, preferably between 10% by mass and 90% by mass relative to the final catalyst. The overall Si / Al ratio of these zeolites is between 2.6 and 200, preferably between 5 and 100, and even more preferably between 5 and 80.The catalyst matrix is ​​a support selected from the group consisting of alumina, silica, silica-alumina, alumina-boron oxide, magnesia, silica-magnesia, zirconia, titanium oxide, clay, these compounds being used alone or in mixtures. Preferably, an alumina support is used. Preferably, the solid acid catalyst has shape selectivity to avoid the formation of alkylaromatics of too large a size, such as those having more than 16 carbon atoms. If the molecular size of the alkylaromatics is close to the size of micropores in the catalyst, the formation and diffusion out of the pores is still feasible, on the other hand, the formation of alkylaromatics that are too large to enter, reside or exit the pores are thus not produced.As indicated above, the reaction zone is operated at a temperature and pressure such that they maintain phase conditions preferably having more than 20% by mass of liquid. For the production of alkylaromatics, having at least 8 carbon atoms, the reaction temperature is in particular between 140°C and 320°C, and is generally between 160°C and 280°C. In one embodiment, the reaction temperature is in the range 190°C to 240°C. The alkylation pressure is generally maintained sufficiently high to ensure the presence of a liquid phase. In one embodiment, the pressures are in the range 20 barg to 100 barg, in particular from 30 barg to 50 barg. When operating under essentially liquid phase conditions, an upflow reactor mode is generally used.Flow rates can typically vary from the liquid hourly space velocity (LHSV) of about 1 h-1 to 100 h-1 per bed, preferably about 2 h-1 to 70 h-1 per bed. The aromatic / alkylating agent ratio is, for example, between 0.05 mol / mol and 20 mol / mol and preferably between 0.1 mol / mol and 10 mol / mol. In a preferred mode of operation, the oligomerization of olefins and the alkylation of aromatics with olefins are carried out on the same catalyst and in the same reactor. The known operating conditions for oligomerization and alkylation are very similar and can easily be adapted to obtain the desired performance in terms of oligomerization and alkylation. When oligomerization and alkylation are carried out simultaneously in the same reactor using the same catalyst, the reactor product contains more than 10% by mass of C8+ aromatics, including more than 6% by mass of C8 to C14 aromatics.When the alkylation in the presence of olefins is carried out separately from the oligomerization, the reactor product contains more than 65% by mass of C8+ aromatics, in particular more than 75% by mass of C8 to C14 aromatics. Step (e) of formation of the hydrocarbon stream to be hydrogenated A hydrocarbon stream to be hydrogenated is formed from at least a portion of the olefins oligomerized in step (c) and at least a portion of the aromatics alkylated in step (d). In the case where a single common stage is used for step (c) of oligomerization of the olefins and for step (d) of alkylation of the aromatics, the product of this stage is used in whole or in part to form the hydrocarbon stream to be hydrogenated.This product comprises for example on a dry basis - less than 15% by mass of paraffins, in particular between 3% and 10% by mass of paraffins, - less than 10% by mass of C4 to C7 olefins, in particular between 0.5% and 5% of C4 to C7 olefins, - more than 50% by mass of C8 to C16 olefins, in particular between 60% and 80% of C4 to C16 olefins, - less than 5% by mass of C17+ olefins, in particular between 0.1% by mass and 1.0% by mass of C17+ olefins; and / or - less than 5% of C6 to C7 aromatics, in particular between 0.5% and 4.0% of C6 to C7 aromatics, - more than 2% of C8+ aromatics, in particular between 6% and 30% of C8+ aromatics. In a variant, an optional step of separating the effluent containing at least a portion of the oligomerized olefins in step (c) and / or at least a portion of the alkylated aromatics in step (d) is carried out in an additional distillation column.The separation produces a C7- hydrocarbon fraction at the top, and a C8+ hydrocarbon fraction at the bottom. The C7- hydrocarbon fraction preferably contains more than 90% by mass, in particular more than 95% by mass of the C7- hydrocarbons contained in the effluent. The C8+ hydrocarbon fraction comprises more than 90% by mass of the C8+ hydrocarbons contained in the effluent. Advantageously, at least a portion, for example less than 50% by mass, of the C7- hydrocarbon fraction is at least partially recycled to step (c) of oligomerization of the olefins and / or to step (d) of alkylation of the aromatics, another portion forming a gasoline stream. The hydrocarbon stream to be hydrogenated is formed by at least a portion, preferably by the entire C8+ hydrocarbon fraction.In the case where step (c) of oligomerization of the olefins is carried out in an oligomerization reactor, and step (d) of alkylation of the aromatics is carried out in an alkylation reactor, separately from step (c) of oligomerization of the olefins, the product of the oligomerization reactor and the product of the alkylation reactor containing the alkylated aromatics are advantageously subjected to separation. In one embodiment, the product of the oligomerization reactor containing the oligomerized olefins and the product of the alkylation reactor containing the alkylated aromatics are separated, at the bottom, into the fraction of C8+ hydrocarbons and at the top, into the fraction of C7- hydrocarbons. The fraction of C7- hydrocarbons is recycled at least in part (for example less than 50% by mass) to step (c) in the oligomerization reactor. At least a portion, preferably all, of the C8+ hydrocarbon fraction forms the hydrocarbon stream to be hydrogenated.In another embodiment, the product of the oligomerization reactor and the product of the alkylation reactor are separated in the distillation column into a C7- hydrocarbon fraction, withdrawn at the top, a C8 to C16 hydrocarbon fraction, taken from an intermediate stage, and a C17+ hydrocarbon fraction, withdrawn at the bottom. At least a portion, preferably all, of the C8 to C16 hydrocarbon fraction forms the hydrocarbon stream to be hydrogenated. The C17+ hydrocarbon fraction is at least partly recycled to step (a) of converting the C1 to C6 alcohol stream. This recycling is carried out in order to crack the C17+ olefins again. Hydrogenation step (f) The hydrocarbon stream to be hydrogenated undergoes hydrogenation to form a hydrogenated hydrocarbon stream. This saturates the olefinic compounds and partially hydrogenates the aromatic compounds.Hydrogenation is for example carried out in one or more fixed bed (descending or ascending) and mixed-phase reactors, the fraction to be hydrogenated being mainly in the liquid phase. Hydrogenation is for example carried out at a temperature between 50°C and 350°C, in particular between 100°C and 300°C. It is carried out under a pressure preferably greater than 10 bara and in particular between 20 bara and 80 bara. A hydrogen stream is fed into the or each reactor in a mixture with the hydrocarbon stream to be treated. The ratio of the volume flow rate of the hydrogen stream to the volume flow rate of the hydrocarbon stream (not including the recycled stream) to be hydrogenated is advantageously between 50 NL / L and 3000 NL / L, in particular between 100 NL / L and 500 NL / L. Hydrogen can be added to the hydrocarbon stream in several stages along the catalytic bed. The hourly space velocity is advantageously between 0.5 and 3 and in particular between 1 and 2 h. -1. Excess hydrogen can be recycled to the reaction zone after separation and compression. The reaction is carried out in the presence of at least one catalyst comprising one or more group VIII metals (typically Pt, Pd, Ni) supported on a support such as silica, alumina or any mixture of these two compounds or carbon. The reaction can also be carried out in the presence of a sulfide catalyst containing a group VIB element (Cr, Mo, W) and a group VIIIB element (Fe, Ru, Co, Os, Co, Rh, Ir, Pd, Ni, Pt) or mixtures of these two groups of metals. The hydrogenated hydrocarbon stream advantageously contains less than 10% by mass of olefins and preferably less than 3% of olefins. It preferably contains more than 50% by mass of paraffins, in particular more than 50% by mass of C7 to C17 paraffins, in particular more than 60% by mass of C7 to C17 paraffins, in particular between 70% by mass and 95% by mass of C7 to C17 paraffins.According to a variant between 10% and 90% by mass of the aromatics contained in the hydrocarbon stream to be hydrogenated, preferably between 30% by mass and 80% by mass of the aromatics contained in the hydrocarbon stream to be hydrogenated are hydrogenated to cycloparaffins. The hydrogenated hydrocarbon stream comprises for example: - more than 20% by mass of C9 hydrocarbons, for example between 20% and 40% by mass of C9 hydrocarbons, - more than 20% by mass, for example between 20% by mass and 40% by mass of C12 hydrocarbons, - more than 6% by mass of C6+ aromatics, in particular more than 6% by mass of C8+ hydrocarbons. Step (g) recovery of the jet fuel fraction To recover the jet fuel fraction, the hydrogenated hydrocarbon stream advantageously undergoes fractionation.This fractionation can be carried out by passing through at least one separation column, for example a distillation column. Preferably, at least one first separation column is used to separate, at the top, a fraction of liquefied petroleum gas from the rest of the hydrogenated hydrocarbon stream which is obtained at the bottom. The fractionation conditions in this column are as follows: pressure advantageously between 2 bara and 15 bara, head condensation temperature adjusted to be able to use an air or cooling water condenser, i.e. preferably between 20°C and 50°C. The residual fraction of hydrogenated hydrocarbon stream is then introduced into a second separation column to produce a naphtha fraction at the top, a diesel fraction at the bottom and at least one intermediate stage the jet fuel fraction.This fractionation can be carried out in a single column comprising a side draw of the jet fuel fraction, or in two separate columns. More than 80% by mass of the hydrogenated hydrocarbon stream introduced into the second column advantageously forms the jet fuel fraction. The liquefied petroleum gas fraction preferably has a final boiling point below 180°C, and even more preferably a final boiling point below 150°C. The initial boiling point can be from 20°C to 60°C and preferably from 25°C to 40°C. The naphtha fraction comprises more than 80% by mass of the C8- paraffins contained in the residual fraction. The jet fuel fraction comprises between 2% by volume and 30% by volume of C8+ aromatics, preferably between 6% by volume and 25% by volume of C8+ aromatics and more preferably between 8% by volume and 25% by volume of C8+ aromatics.It comprises more than 50% by volume of C9 to C16 paraffins, in particular between 60% by volume and 95% by volume of C9 to C16 paraffins. The jet fuel fraction comprises in particular, more than 60% by volume of C9 to C12 paraffins. It preferably has a final boiling point below 400°C, and even more preferably a final boiling point below 350°C. The initial boiling point can be from 130°C to 180°C. The diesel fraction is the heaviest fraction, not all of whose molecules can be used to form jet fuel. This fraction typically has an initial boiling temperature above 300°C, and preferably above 310°C. Advantageously, at least a portion of the diesel fraction is recycled to step (a) of converting the C1 to C6 alcohol stream in the form of a recycle stream in order to generate a new cracking of the compounds present in this fraction.The recycle stream advantageously constitutes between 10% by mass and 50% by mass of the C1 to C6 alcohol stream introduced in conversion step (a). DESCRIPTION OF THE FIGURES The invention will be better understood upon reading the following description, given solely by way of example, and with reference to the appended drawings, in which: - Figure 1 is a schematic view of an installation configured for implementing a first method for manufacturing a jet fuel according to the invention; - Figures 2 to 7 are views similar to that of Figure 1 illustrating installation variants intended for implementing variants of the method of Figure 1: - Figure 8 is a view of a detail illustrating a reactor for implementing the fluidized bed conversion step; - Figure 9 is an NMR spectrum of a ZSM5 parent catalyst and a modified catalyst having an ALPO structure.- Figure 10 illustrates the selectivities obtained during the implementation of an example of a conversion step of the process according to the invention; - Figure 11 illustrates the conversion of certain olefins as a function of time during an example of implementation of the oligomerization step with a first catalyst; - Figure 12 illustrates the conversion of certain olefins as a function of time during an example of implementation of the oligomerization step with a second catalyst; - Figure 13 illustrates the conversion of certain aromatic compounds as a function of time during an example of implementation of the aromatics alkylation step with the second catalyst. DESCRIPTION OF EMBODIMENTS OF THE INVENTION A first manufacturing installation 10, intended for the implementation of a process for manufacturing a jet fuel according to the invention is illustrated schematically in Figure 1.The first installation 10 comprises a stage 12 for converting a stream of alcohol 14 into C1 to C6, intended to produce a mixture 16 containing paraffins, olefins, aromatics, and water, and a stage 18 for separating the water from the mixture 16 to produce a water-depleted mixture 19 comprising a liquid phase 19a and a gas phase 19b. The installation of FIG. 1 also comprises a stage 20 for separating the C1 to C2 hydrocarbons from the water-depleted mixture.In this example, the installation 10 further comprises a joint stage 22 for oligomerization of olefins and alkylation of aromatics originating from the water-depleted mixture, the stage 22 producing a stream 24 of hydrocarbons to be hydrogenated. The installation 10 further comprises a stage 26 for hydrogenation of the stream 24, producing a stream of hydrogenated hydrocarbons 30 and a stage 28 for fractionation of the stream of hydrogenated hydrocarbons 30, intended to fractionate at least one jet fuel fraction 34, and advantageously a diesel fraction 36 and a naphtha fraction 38. The conversion stage 12 is intended to implement the conversion step (a) described above, transforming the C1 to C6 alcohols predominantly into C3 to C7 olefins. As described above, the conversion stage 12 comprises at least one reactor having a fluidized catalytic bed, preferably having a single fluidized catalytic bed reactor.This or these reactors are suitable for implementing the experimental conditions described above. The reactor comprises a reaction zone 111a having a fluidized catalytic bed, and a regeneration zone 111b of the fluidized catalytic bed. A portion of the catalyst present in the reaction zone 111a is continuously withdrawn to be regenerated in the regeneration zone 111b, advantageously by controlled combustion in the presence of oxygen. With reference to the description above, the separation stage 18 is intended for implementing step (b) of water separation. It comprises at least one separator operating by gravity and / or by mechanical drive to separate the water from the mixture 16, making it possible to recover an aqueous fraction concentrated in water (stream 40), a fraction of hydrocarbons in the gas phase 19b and a fraction of hydrocarbons in the liquid phase 19a.The installation 10 optionally comprises at least one conduit 18a for recycling water separated in the separation stage to the conversion stage 12. The recycle conduit 18a opens for example upstream of the catalytic beds. Advantageously, the separation stage 18 comprises a stripping column 41 capable of treating at least a portion of the separated water forming the flow 40 to extract the hydrocarbons it contains and obtain treated water. The installation 10 comprises at least one cooling device (for example a heat exchanger) making it possible to reduce the temperature of the product leaving the reactor and by heating another stream, for example such as the feedstock of the reactor 12, a water or air cooler and / or a combination thereof downstream of the conversion stage 12 to condense the water and produce the flow 40.The separation stage 20 is intended to implement step (c) of separation of hydrocarbons lighter than C3, such as C1-C2, and light compounds CO, CO2, hydrogen, as defined above. It advantageously comprises at least one de-ethanizer. The separation stage 20 operating at a pressure higher than that of the water separation stage 18, the installation comprises at least one pump capable of increasing the pressure of the liquid phase 19a and at least one compressor, capable of increasing the pressure of the gas phase 19b. The stage 22 for oligomerization of olefins and alkylation of aromatics is intended to jointly implement steps (d) and (e). It comprises at least one joint oligomerization and / or alkylation reactor intended to implement the experimental conditions described above. Hydrogenation stage 26 is intended to implement step (f).It comprises at least one fixed-bed hydrogenation reactor intended to carry out the hydrogenation reaction under the conditions described above. The fractionation stage 28 is intended to carry out step (g). In this example, it comprises at least a first upstream column 42 for separating liquefied petroleum gas 44, and a second downstream fractionation column 46, intended to produce fractions 34 to 38. A first example of a process for manufacturing a jet fuel, implemented in the installation of FIG. 1 will now be described. Initially, a stream 14 of C1 to C6 alcohol is brought into the conversion stage 12. The alcohol stream 14 comes for example from a source 50 described above, in which the alcohols of the source 50 are produced for example by fermentation of biomass, by catalytic conversion of carbohydrates or carbon monoxide or carbon dioxide in the presence of hydrogen.Stream 14 comprises the composition described above, for example more than 50% by dry mass of methanol and preferably more than 80% of methanol by dry mass. Stream 14 is introduced into conversion stage 12 where it undergoes a conversion described above comprising a dehydration / aromatization of the C2 to C6 alcohols, and for methanol, a conversion to dimethyl ether followed by a dehydration. The reaction is carried out under the operating conditions of temperatures and pressures described above. One or more catalysts defined above are used. Advantageously, the C1 to C6 alcohol stream is introduced into conversion step (a) at a temperature at least 5°C higher than the bubble point of the C1 to C6 alcohol stream. A mixture 16 containing paraffins (in particular n-paraffins, i-paraffins and cyclo-paraffins), olefins, and aromatics and water is thus obtained.The mixture 16 comprises, for example, the composition described above. Figure 8 illustrates, in a particular embodiment, the flow of the catalyst from the regeneration zone 111b to the reaction zone 111a. For reasons of simplicity, the drawings do not contain details on the internal parts of the tanks forming the zones 111a, 111b. The stream 14 of C1 to C6 alcohol is introduced into the bottom of the reaction zone 111a having a fluidized catalytic bed. At the top of the reaction zone 111a, the products of the conversion reaction are separated from the catalyst in a disengagement zone 203 advantageously provided with cyclones and the mixture 16 produced is transported to the separation stage 18. Optionally, the heat of reaction produced by the conversion is extracted from the reaction zone 111a by means of a catalyst cooler 205, which is advantageously a heat exchanger located outside the reaction zone 111a and connected thereto.The reaction zone 111a receives the catalyst regenerated in the regeneration zone 111b via a feed line 207 connecting the regeneration zone 111b to the reaction zone 111a. The deactivated catalyst is removed from the disengagement zone 203 via a discharge line 206 separate from the feed line 207, the discharge line 206 connecting the reaction zone 111a to the regeneration zone 111b. Air is injected via the injection line 221 into the regeneration zone 111b, at the bottom thereof, in a fluidized bed where the coke deposits are burned. The regeneration zone 111b also comprises a disengagement zone 222 advantageously provided with cyclones. In this zone 222, the combustion gases are separated from the regenerated catalyst and are evacuated via the regeneration pipe 223 advantageously at the head of the regeneration zone 111b.Optionally, since the combustion of coke deposits is a very exothermic reaction and the temperature of the regeneration zone 111b must be carefully controlled, a catalyst cooler (not shown in the figure, but similar to the catalyst cooler 205) is connected to the regeneration zone 111b. Hot catalyst extracted from the regeneration zone circulates through this cooler in order to be cooled, which controls the temperature in the regeneration zone 111b. The regenerated catalyst is sent via line 207 to the reaction zone 111a. The mixture 16 is then introduced into the separation stage 18 to produce a water stream 40 at the bottom of the separator, and the water-depleted mixture 19 comprising the gas phase 19b and the liquid phase 19a. A portion 40a of the water stream 40 is optionally recycled to the conversion stage 12, via the conduit 18a, as described above.Another portion 40b of the water stream 40 is introduced into a column 41 to undergo stripping and produce at the top, a stream 41a of extracted hydrocarbons and at the bottom, a treated water stream 40b having a hydrocarbon content lower than that of the water stream 40. The extracted hydrocarbon stream 41a is recycled to the separation stage 18, for example upstream of the separator. The gas phase 19b and the liquid phase 19a are then introduced, after compression, into the deethanizer in the separation stage 20. The deethanizer operates under the conditions defined above and produces at the top, a fraction 60 of C1-C2 hydrocarbons which may contain light compounds such as CO, CO2, hydrogen and at the bottom, a fraction 62 of C3+ hydrocarbons.The fractions 60, 62 obtained have the compositions defined above. Preferably, the fraction 60 of C1-C2 hydrocarbons, optionally after separation, is sent to a steam cracker for recovery of at least a portion of the ethylene it contains. In a variant shown in dotted lines in FIG. 1, at least a portion 64 of the fraction 60 of C1 to C2 hydrocarbons is recycled into the separation stage 12 in the form of a recycle stream 64. Optionally, at least a fraction of the gas phase 19b, for example less than 50% by volume of the gas phase 19b, is also recycled, without passing through the deethanizer. The ratio of the mass flow rate of the recycle stream 64 to the mass flow rate of the fraction 60 from the top of the deethanizer of the separation stage 20 is less than 0.5 as defined above.In the example of Figure 1, the fraction 62 of C3+ hydrocarbons, having the composition defined above, is then introduced into the joint oligomerization and alkylation stage 22. In this stage 22, one or more joint oligomerization and alkylation reactors carry out an oligomerization of the olefins present in the fraction 62, according to the operating conditions defined above in the description, in particular an oligomerization of the C3 to C7 olefins. In addition, jointly, the C6+ aromatics present in the fraction 62 are alkylated to form in particular C8+ aromatics. The reaction is carried out under the operating conditions described above. One or more catalysts defined above are used. At the outlet of the stage 22, a stream 24 of hydrocarbons to be hydrogenated is thus formed, with the composition defined above.Then the stream 24 of hydrocarbons to be hydrogenated is introduced into the hydrogenization stage 26, to induce the hydrogenation of at least a portion of the olefins present in the stream of hydrocarbons to be hydrogenated 24 and the hydrogenation into cycloparaffins of at least a portion of the aromatics present in the stream of hydrocarbons to be hydrogenated 24. The hydrogenation is carried out under the operating conditions described above, with one or more of the catalysts described above. A stream 66 containing hydrogen is introduced into the hydrogenation stage 26, with a ratio of the volume flow rate of hydrogen in the stream 66 to the volume flow rate of the stream of hydrocarbons to be hydrogenated 24 is for example that defined above. A stream of hydrogenated hydrocarbons 30 is formed at the outlet of the hydrogenation stage with the composition described above. The hydrogenated hydrocarbon stream 30 is then fractionated in the fractionation stage 28.In the first column 42, it is separated into a C4- hydrocarbon fraction forming the liquefied petroleum gas fraction 44, and into a C4+ hydrocarbon fraction forming the residual fraction 70 of the hydrogenated hydrocarbon stream. The fractions 44, 70 have the characteristics defined above in terms of cutting points. The residual fraction 70 is introduced into the second column 46 to be fractionated there into the naphtha fraction 38, the jet fuel fraction 34, and the diesel fraction 36, as characterized above. The installation variant 90 illustrated by FIG. 2 is intended for the implementation of a second method according to the invention. It differs from the installation 10 illustrated by FIG. 1 in that the separation stage 20 comprises an additional column 92 for recovering propylene.The fraction 62 of C3+ hydrocarbons from the deethanizer of the separation stage 20 is introduced into the additional column 92 to form at the top of the column, a fraction 80 of C3- hydrocarbons and at the bottom of the column, a fraction 82 of C4+ hydrocarbons intended to be introduced into the joint oligomerization and alkylation stage 22. The fractions 80, 82 have the compositions described above. The fraction 80 contains more than 80% by mass of the propylene contained in the fraction 62 of C3+ hydrocarbons. Such an example makes it possible to recover the propylene formed in the conversion stage 12, when such recovery is economically advantageous. The installation 100 described in FIG. 3 is intended for the implementation of a third method according to the invention.It differs from the installation 10 described in Figure 1, in that at least one tapping 102 for adding C2 to C6 alcohols is provided in the quenching stage 12, for example between two successive catalytic beds of the conversion stage 12. The composition of the C2 to C6 alcohol stream 102 advantageously comprises less than 20% methanol, and more than 80% C2 to C6 alcohol by mass, for example more than 50% ethanol and propanol. The addition of C2 to C6 alcohol in addition to methanol facilitates the conversion reaction of the alcohol stream 14 by making it more isothermal (the conversion of methanol being very exothermic and the conversion of C2 to C6 alcohols being endothermic) and therefore easier to control. The installation 110 described in Figure 4 is similar to that of Figure 1. It differs from the installation 10 visible in Figure 1 in that a stream containing carbon dioxide 182 is added to the conversion stage 12.The carbon dioxide stream 182 comprises more than 5% by mass of carbon dioxide as defined above. The mass ratio in the feed supplied to the conversion stage 12 of carbon dioxide to C1 to C6 alcohols is between 5% and 75%. The carbon dioxide contained in the carbon dioxide stream 182 is converted into carbon monoxide together with the conversion of the C1 to C6 alcohol stream. A fifth installation 120 intended for the implementation of a fifth method according to the invention is illustrated in FIG. 5. The fifth installation 120 differs from the first installation 10 in that it comprises an additional separation stage 122 interposed between the outlet of the joint oligomerization and alkylation stage 22 and the inlet of the hydrogenation stage 26. The additional separation stage 122 comprises at least one distillation column.The product 124 from the joint oligomerization and alkylation stage 22 is separated in the distillation column into a fraction 126 of C7- hydrocarbons and a fraction 128 of C8+ hydrocarbons forming the hydrocarbon stream 24 intended to be hydrogenated. The fraction 126 of C7- hydrocarbons can optionally be recycled to the joint stage 22. A sixth installation 140 according to the invention is illustrated in FIG. 6. This sixth installation 140 is intended for the implementation of a sixth method according to the invention. It differs from the first installation 10 in that it comprises a specific stage 22A for oligomerization of olefins from the water-depleted mixture 19, from stage 20, and a specific stage 22B for alkylation of aromatics from the water-depleted mixture 19 from stage 20.Each stage 22A 22B comprises a separate reactor respectively for oligomerization and alkylation in which the operating conditions provided above in the description are implemented. In the separation stage 20, the water-depleted mixture 19 is separated into a fraction 60 of C1-C2 hydrocarbons recovered at the top of the deethanizer, into a fraction 142 of C3 to C5 hydrocarbons recovered at an intermediate stage of the deethanizer and into a fraction 144 of C6+ hydrocarbons recovered at the bottom of the deethanizer. The fraction 142 of C3 to C5 hydrocarbons is sent in its entirety to the oligomerization stage 22A to produce a product 146 from the oligomerization reactor. The fraction 60 of C1-C2 hydrocarbons and the fraction 144 of C6+ hydrocarbons are brought to the alkylation stage 22B to produce a product 152 from the alkylation reactor under the operating conditions defined above.The product 152 from the alkylation reactor is then mixed with the product 146 from the oligomerization reactor. The products 146, 152 are then introduced into the additional separation stage 122 to be separated into the fraction 126 of C7- hydrocarbons and the fraction 128 of C8+ hydrocarbons, described above. At least a portion 150 of the fraction 126 of C7- hydrocarbons is recycled into the oligomerization stage 22A, another portion possibly being recovered in the form of gasoline. The fraction 128 forms the stream of hydrocarbons to be hydrogenated 24. A seventh installation 160 intended for the implementation of a seventh method according to the invention is described in FIG. 7.The seventh method according to the invention differs from the sixth method implemented in the installation 150 in that at the additional separation stage 122, the product 146 of the oligomerization reactor and the product 152 of the alkylation reactor are separated into the fraction 126 of C7- hydrocarbons, taken at the top of the column, into a fraction 162 of C8 to C16 hydrocarbons, taken at an intermediate stage of the column and into a fraction 164 of C17+ hydrocarbons, taken at the bottom of the column. As described previously, at least a portion 150 of the fraction 126 of C7- hydrocarbons is recycled into the oligomerization stage 22A. The fraction 162 of C8 to C16 hydrocarbons is introduced into the hydrogenation stage 26 to be hydrogenated. The fraction 164 of C17+ hydrocarbons is recycled at least partially to the conversion stage 12. Thus, the heavy hydrocarbons present in the fraction of C17+ hydrocarbons are re-cracked in the conversion stage 12.This increases the amount of jet fuel fraction 34 produced. In all the cases described above, the jet fuel fraction 34 produced by the aforementioned methods can be used as such, in pure form, as an aircraft jet fuel intended to propel an aircraft engine, or in a mixture with a jet fuel resulting from the distillation of petroleum. The jet fuel fraction or its mixture is advantageously a renewable aviation fuel (“Sustainable Aviation Fuel”) whose composition is similar to the SAF described according to Standard ASTM D7566. The mixture comprises at least 5% by mass, in particular at least 10% by mass of the jet fuel fraction 34.Thanks to the invention which has just been described, it is possible to have simple and effective methods for manufacturing a jet fuel from a C1 to C6 alcohol stream which is preferably from a renewable source, in particular from fermentation, and / or generated by conversion of carbon monoxide or carbon dioxide captured from the atmosphere in the presence of hydrogen. The jet fuel fraction resulting from the process according to the invention has a very reduced carbon footprint, since it is not derived from petroleum derivatives, but on the contrary from sources contributing to reducing the quantity of carbon dioxide present in the atmosphere. The jet fuel fraction 34 produced by the process according to the invention is furthermore manufactured very economically and can in certain cases be used as such, without additional purification or without mixing, as aircraft engine propulsion fuel.EXAMPLES Specific non-limiting examples of carrying out the conversion step (a), the joint oligomerization and alkylation steps (c) and (d) and the hydrogenation step (f) will now be described. Conversion step (a) PREPARATION OF CATALYSTS A sample of zeolite ZSM-5 (Si / Al=12) in H form (containing 445 ppm Na, less than 25 ppm K, 178 ppm Fe, 17 ppm Ca and synthesized without template) was steamed at 550°C for 6 h in 100% H2O at atmospheric pressure. The sample is hereinafter identified as sample A. The steamed solid A was subjected to contact with a 3.14 M H3PO4 solution for 4 h under reflux conditions (4.2 ml / g of zeolite). Then, solid A was separated from the liquid phase at room temperature by filtration of the solution. The obtained material was dried at 200°C for 16 h. The sample is hereinafter identified as sample B.CATALYST EXAMPLE 1 490 g of sample B was mixed with 490 g of specific binder (P = 15.9 wt%, Si = 13.2%, Mg = 0.27%, Al = 0.15 wt%, K = 230 ppm, Na = 230 ppm, Ca = 19.2 wt%), 588.3 g of low sodium silica sol containing 34 wt% SiO2, 6 g of xonotlite and 2-3 wt% extrusion additives. The mixture was stirred for 30 min and extruded. The specific binder was made by mixing the equivalent mass of NH4H2PO4 and xonotlite in aqueous medium at room temperature (1 g of solid / 4 ml of water). After stirring for 60 minutes, the phosphated xonotlite was separated from the liquid by filtration and dried. The dried product was used as an extrusion component. The extruded solid was dried for 24 h at room temperature, then 16 h at 200°C followed by washing with demineralized water at room temperature and then dried at 110°C overnight.An additional washing step at room temperature is then carried out using demineralized water at pH 3.08. The catalyst is then dried at 110°C overnight and calcined at 700°C for 2 hours. CATALYST EXAMPLE 2 320 g of sample B was mixed with 400 g of specific binder (P = 15.9 wt%, Si = 13.2, Mg = 0.27, Al = 0.15 wt%, K = 230 ppm, Na = 230 ppm, Ca = 19.2 wt%), 165 ml of H2O, 235 g of low sodium silica sol containing 34 wt% SiO2 and 2-3 wt% extrusion additives. The mixture was stirred for 30 min and extruded. The specific binder was made by mixing the equivalent mass of (NH4)H2PO4 (ammonium dihydrogen phosphate) and xonotlite in aqueous medium at room temperature (1g of solid / 4ml of water). After stirring for 60 minutes, the phosphated xonotlite was separated from the liquid by filtration and dried. The dried product was used as an extrusion component.The extruded solid was dried for 24 h at room temperature, then 16 h at elevated temperature followed by washing and heat treatment with steam at 600°C for 2 h. The sample hereinafter identified as sample E. CATALYST EXAMPLE 3 356 g of sample A was extruded with 338.7 g of Nyacol (40% by mass SiO2 sol), 311.3 g of fumed silica (FK500), 480 ml of H2O and 2 to 3% of extrusion additives. The extruded solid was dried for 24 h at room temperature, then 16 h at 110°C followed by calcinations at 500°C for 10 h. The final sample contained 40% by mass of zeolite and 60% by mass of SiO2 binder. The extruded sample was ion-exchanged with 0.5 M NH4Cl under reflux conditions for 18 hours followed by water washing, drying at 110°C for 16 hours and calcinations at 450°C for 6 hours.The shaped and exchanged sample was treated with 3.1 M H3PO4 under reflux conditions for 4 h (1 g / 4.2 ml), followed by cooling, filtration, and drying at 110 °C for 16 h. The phosphated sample was washed at room temperature with 0.1 M calcium acetate solution for 2 h (1 g / 4.2 ml). Then, the washed sample was dried at 110 °C for 16 h and steam-heated in 100 wt% H2O for 2 h at 600 °C. CATALYST EXAMPLE 4 150g of sample B was contacted with 630ml of aqueous solution containing 1.5g of dispersed xonotlite followed by the addition of 450g of low sodium silica sol (34wt% SiO2 in water, 200ppm Na). Then, the solution was stirred for one hour and spray dried.The spray-dried solid was washed with water at room temperature for 2 hours followed by filtration, drying at 110°C for 16 hours and calcinations at 700°C. CATALYST EXAMPLE 5 100 g of Sample A was subjected to contact with 25 g of 85% by mass H3PO4 under reflux conditions for 4 hours, followed by cooling and the addition of 120 ml of aqueous solution containing 7 g of dispersed xonotlite. The resulting slurry was kept stirring for about 1 h, followed by the addition of 300 g of low sodium silica sol (34% by mass SiO2 in water, 200 ppm Na). Then, the solution was stirred for one hour and spray-dried. The spray-dried solid was dried at 200 °C for 16 h and washed with water at room temperature for 2 h followed by filtration, drying and calcinations at 700 °C for 2 h.CATALYST EXAMPLE 6 75 g of sample A were introduced into a solution containing 14.25 g of 85% by mass H3PO4 and 300 ml of deionized water. The suspension was stirred at reflux for 2 h. Then 4.125 g of CaCO3 was added to the suspension. Heating of the solution was stopped, while maintaining stirring of the mixture until a temperature below 30°C. This resulted in suspension A. A solution was then prepared by mixing 450 g of low sodium silica sol (34% by mass SiO2 in water, 200 ppm Na) and 4.5 g of H3PO4 (85% by mass) while stirring at room temperature for 30 minutes. This resulted in suspension B. Suspensions A and B were then mixed together and 120 ml of deionized water was added. Then the solution was stirred for one hour and spray dried.The spray-dried solid was dried at 200°C for 16h and washed with water at room temperature for 2h followed by filtration, drying and calcinations at 700°C for 2h. CONVERSION STEP (a) IMPLEMENTATION Catalyst tests were carried out on 2g (35 mesh to 45 mesh particles) of catalyst with a feed of essentially pure methanol, at Tinjection = 550°C and a pressure of 0.5 barg and a horizontal hourly space velocity (WHSV) = 1.6 h. -1), in a stainless steel fixed-bed reactor with downward flow. Before the catalytic test, all catalysts were activated in a stream of N2 (5 NI / h) up to the reaction temperature. Product analysis was carried out online by a gas chromatograph equipped with a capillary column. The catalytic performances of the catalyst in Table 1 are given on carbon, dry basis and coke-free basis. The results are given for the average catalyst performance during the first 4 hours of operation [Table 1] Performance of step (a) – 550°C -0.5barg Joint steps (c) and (d) of oligomerization and alkylation and step (f) of hydrogenation PROPERTIES OF THE FEEDSTOCK USED FOR THE IMPLEMENTATION OF STEPS (c) AND (d) The properties of the feedstock used are as follows: [Table 2] The detailed composition of the load was determined by the GC method. [Table 3]OLIGOMERIZATION AND ALKYLATION 100 mL of amorphous silica-alumina (ASA) catalyst diluted with 100 mL of inert material (0.21 mm SiC) was loaded into an 18 mm inner diameter fixed-bed tubular reactor. Before testing, the catalyst was activated at 250°C (10°C / h) under 135 NL / h of nitrogen for 8 hours. The temperature was then lowered to 40°C at the beginning of the test program. 100 mL of ZSM-5-based catalyst (80 wt.% MFI with a silica-alumina ratio of 80% at 20 wt.% alumina binder) diluted with 100 mL of inert material (0.21 mm SiC) was loaded into an 18 mm inner diameter fixed-bed tubular reactor. Prior to testing, the catalyst was activated at 400°C (60°C / h) under 160NL / h of nitrogen for 2 hours. The temperature was then lowered to 40°C at the start of the test program.HYDROGENATION Fractionation is carried out on the oligomerization product to recover the oligomerized cuts 145+ and 165°C+ which are hydrotreated on a NiMo catalyst. The following operating conditions were chosen: 80 barg, a liquid hourly space velocity LHSV of 1 h. - 1 , H2 / hydrocarbons volume ratio of 500 NL / L, in one pass without recycle and the temperature was increased from 250°C to 270°C. EXAMPLE 1 - PERFORMANCE OBTAINED WITH A ZEOLITE-BASED CATALYST The feedstock was treated under the following operating conditions: 55 barg, liquid hourly space velocity LHSV of 1h -1 and at temperatures ranging from 240°C to 280°C. [Table 4] The conversion of light olefins (C4-C8 olefins) ranges from 65% by mass at 240°C to 91% by mass at 280°C. C9+ olefins are not included in the conversion calculation as they may result from oligomerization of light olefins (C4 and C5) present in the feed. At 240°C, the conversion of C5-C7 olefins is greater than 88% by mass. Olefins can either react by oligomerization or by alkylation with aromatic compounds. The conversion of aromatics has been observed to range from 10% by mass at 240°C to 26% by mass at 280°C (see Table 5 below). Aromatics are indeed present in the 170-FBP fractions, indicating that alkylation is taking place. [Table 5] EXAMPLE 2 - PERFORMANCE OBTAINED WITH AN AMORPHOUS SILICA-ALUMINA (ASA) CATALYST The feedstock was treated under the following operating conditions: 25 barg, liquid hourly space velocity LHSV of 1h -1and temperatures ranging from 180°C to 220°C. [Table 6] The conversion of light olefins (C4-C8 olefins) ranges from 80% by mass at 180°C to nearly 100°C at 220°C. C9+ olefins are not included in the conversion calculation as they may result from oligomerization of light olefins (C4 and C5) present in the feed. At 180°C, the conversion of C5-C7 olefins is greater than 80% by mass. Olefins can either react by oligomerization or by alkylation with aromatic compounds. The conversion of aromatics has been observed to range from 28% by mass at 180°C to 33% by mass at 220°C (see Table 7 below). Aromatics are indeed present in the FBP fractions at 170°C, indicating that alkylation is taking place. [Table 7] The 145+ cuts were hydrotreated using a NiMo catalyst under the conditions described above. The properties of the hydrogenated cut are described in Table 8 and the detailed composition is summarized in Table 9. [Table 8] [Table 9] Effluent composition as determined by GCxGC ADDITIONAL EXAMPLE OF CONVERSION STEP (a) IMPLEMENTATION CATALYST EXAMPLE 75 g of sample A were introduced into a solution containing 14.25 g of 85% by mass H3PO4 and 225 ml of demineralized water. The suspension was stirred at reflux for 4 h. Then 4.9 g of CaCO3 was added to the suspension. Heating of the solution was stopped, while maintaining stirring of the mixture until a temperature below 30°C. This led to suspension S1. A solution was then prepared by mixing 450 g of low sodium silica sol (34 wt% SiO2 in water, 200 ppm Na) and 4.5 g of H3PO4 (85 wt%) under stirring at room temperature for 30 minutes. This resulted in suspension S2. Suspensions S1 and S2 were then mixed to form a solution. The solution was then stirred for one hour and spray dried to yield catalyst X.IMPLEMENTATION OF CONVERSION STEP (a) Tests 1 to 3 described below were carried out in a fixed bed passivated with ceramic, loaded with 1.31 g of catalyst X described above in a mixture with SiC. The catalytic bed is held in place using quartz wool. The alcohol is introduced at the top of the reactor with a flow rate of 0.05 mol Methanol / h / g catalyst. Nitrogen dilution can be added. In the latter case, the partial pressure of alcohol is different from the total pressure, as visible in the tables below. The total pressure was varied (1.3 bara, 5 bara and 10 bara) as well as the temperature (450°C, 500°C, 550°C). Before the catalytic test, the catalyst is activated under N2 (5 NI / h) up to the reaction temperature.Product analysis was performed by combining on-line gas microchromatography analysis of the reactor outlet stream and off-line gas chromatography-flame ionization detector (GC-FID) analysis of the reaction-generated liquid.The gas micro-chromatography used has four modules: - Module 1 equipped with a molecular sieve capillary column ("Molecular Sieve") for the separation of O2, N2, H2, CO and CH4 - Module 2 equipped with a grafted polystyrene-divinylbenzene capillary column ("Plot Q") to separate MeOH, dimethyl ether (DME), C1 to C3 hydrocarbons and CO2, - Module 3 equipped with an alumina capillary column for the separation of C2 to C5 hydrocarbons - Module 4 equipped with a fused silica or "Stabilwax" column to separate water and C6+ hydrocarbons The results are illustrated by the table below which gives the selectivity in mass percentage for a methanol stream: [Table 10]. Graphs (a) to (e) shown in Figure 10 show the impact of alcohol partial pressure and temperature on the selectivities to ethylene (graph (a)), propylene and butene (graph (b)), C5+ olefins (graph (c)), C5+ paraffins (graph (d)), and aromatics benzene, toluene, xylene (graph (e)). In each graph, the hourly space velocity of methanol relative to the mass of WHSV(MeOH) catalyst is 1.6 / h. The solid triangles correspond to a temperature of 550°C, the open triangles correspond to a temperature of 500°C, and the open diamonds correspond to a temperature of 450°C.These examples illustrate that in the temperature range mentioned above, in particular from 300°C to 600°C, in particular between 330°C and 550°C, in particular between 350°C and 500°C or between 410°C and 580°C, with a partial pressure of methanol in the range mentioned above, in particular the range from 100 kPa to 5 MPa, preferably 100 kPa to approximately 1.0 MPa, with a phosphorus-modified zeolite catalyst, a mixture of paraffins, olefins, aromatics and water is produced according to the invention in conversion step (a), the ratio of the mass of C3+ olefins to the total mass of olefins being greater than or equal to 0.8. The same is true when starting from an ethanol stream, as illustrated by the three tests 1E, 1F and 1G below. The test conditions and selectivity results are as follows:

[0002] [Table 11] Alternatively, a 1H test was also performed in a fluidized bed on 7.2g of catalyst X with a methanol stream to be converted. To ensure adequate fluidization properties, the nitrogen flow rate was set at 23.8mL / min, and the alcohol was co-injected at the bottom of the reactor. The test conditions and selectivity results are as follows: [Table 12] Before the catalytic test, the catalyst was activated under N2 (5 NI / h) up to the reaction temperature. ADDITIONAL EXAMPLE OF OLIGOMERIZATION STEP IMPLEMENTATION On the one hand, 40 mL of ZSM-5-based catalyst (80% by mass of MFI and 20% of alumina binder) cut from 2 mm to 4 mm, diluted with 40 mL of inert material (SiC cut from 1 mm to 1.4 mm) were loaded into a fixed-bed tubular reactor with an inner diameter of 16 mm. Before the test, the catalyst was activated at 400°C (60°C / h) under 160NL / h of nitrogen for 2 hours. The temperature was then lowered to 40°C before introducing the feedstock and raising the temperature to the test conditions. On the other hand, 40 mL of BEA-based catalyst (80% by mass of BEA and 20% of alumina binder) cut 2-4 mm diluted with 40 mL of inert material (SiC 0.21 mm) were loaded into a fixed-bed tubular reactor with an inner diameter of 16 mm.Prior to testing, the catalyst was activated at 400°C (60°C / h) under 160NL / h of nitrogen for 2 hours. The temperature was then lowered to 40°C before introducing the feedstock and raising the temperature to test conditions. The feedstock composition used is as follows for both catalyst tests at 55 barg: [Table 13]. The yield structures are estimated based on the following cut points Light: IBP-80°C Naphtha: 80°C-145°C Jet fuel: 145°C-300°C Diesel: > 300°C EXAMPLE OLIGO-1 - PERFORMANCE OBTAINED WITH ZSM-5 BASED CATALYST The feedstock was treated under the following operating conditions: 55 barg, liquid hourly space velocity LHSV of 1h -1and at temperatures ranging from 180°C to 240°C. The conversion of light olefins (C3 to C6 olefins) is shown in Figure 11. It is above 90% by mass for temperatures above 220°C. At 220°C, the conversion of butenes and hexenes decreases with time. In terms of yield structure, once the contribution of the diluent (n-heptane in this case) is subtracted, the yields at 220°C are as follows: [Table 14] EXAMPLE OLIGO-2 - PERFORMANCE OBTAINED WITH THE BEA-BASED CATALYST The feedstock was treated under the following operating conditions: 55 barg, liquid hourly space velocity LHSV of 1 / h and at temperatures ranging from 150°C to 220°C. The conversion of light olefins to C5 and C6 is shown in Figure 12. It is above 90% by mass for temperatures from 200°C. Olefins can either react by oligomerization or by alkylation with aromatic compounds. The aromatic compounds are partially converted and at iso-temperature, the degree of conversion of these aromatics decreases with time, as illustrated in Figure 13. In terms of yield structure, once the contribution of the diluent (n-heptane in this case) is subtracted, the yields at 200°C are as follows: [Table 15] These results show that for the oligomerization and alkylation conditions defined above, in particular for temperatures from 150°C to 400°C, preferably from 180°C to 350°C, more preferably from 180°C to 290°C, for a weight hourly space velocity (WHSV) of the feedstock of, for example, 0.1 h -1 at 8 p.m. -1 , preferably 0.5 h -1 at 10 a.m. -1 , more preferably 0.8 h -1 at 5 a.m. -1 with zeolite catalysts it is possible to very efficiently convert a feedstock obtained from step (a) into a significant quantity of jet fuel.

Claims

CLAIMS 1. A process for manufacturing jet fuel comprising the following steps: (a) conversion of a C1 to C6 alcohol stream (14) to produce a mixture (16) containing paraffins, olefins, aromatics,and water; (b) separation of water (40) from the mixture (16) to form a water-depleted mixture (19); (c) oligomerization of olefins from the water-depleted mixture (19); (d) alkylation of aromatics from the water-depleted mixture (19); (e) formation of a stream (24) of hydrocarbons to be hydrogenated from at least a portion of the olefins oligomerized in step (c) and at least a portion of the aromatics alkylated in step (d); (f) hydrogenation of the stream (24) of hydrocarbons to be hydrogenated to form a stream (30) of hydrogenated hydrocarbons; (g) recovery of at least one jet fuel fraction (34) from the stream (30) of hydrogenated hydrocarbons; in which the conversion step (a) is carried out in a reaction zone (111a) comprising at least one fluidized catalytic bed, and in which, in the mixture (16) of paraffins, olefins, aromatics and water produced in the conversion step (a),the ratio of the mass of C3+ olefins to the total mass of olefins is greater than or equal to 0.

8.

2. Process according to claim 1, wherein the jet fuel fraction (34) comprises between 2% by volume and 30% by volume of C8+ aromatics, preferably between 8% by volume and 25% by volume of C8+ aromatics, the mass content of aromatics in the water-depleted mixture (19) preferably being greater than 6% by mass, in particular between 6% and 10% by mass.

3. A process according to any one of the preceding claims, comprising separating at least a portion of the oligomerized olefins in step (c) and / or at least a portion of the alkylated aromatics in step (d) into a fraction (126) of C7- hydrocarbons, and into a fraction (128) of C8+ hydrocarbons, the fraction (126) of C7- hydrocarbons being at least partially recycled to step (c) of olefin oligomerization and / or to step (d) of aromatic alkylation, the stream (24), of hydrocarbons to be hydrogenated being formed by at least a part of the fraction (128) of C8+ hydrocarbons.

4. Process according to any one of the preceding claims, wherein step (c) of oligomerization of olefins and step (d) of alkylation of aromatics are carried out jointly in the same reactor.

5. A process according to claim 4, comprising a step of separating the water-depleted mixture (19) into a fraction (60) of C1-C2 hydrocarbons, and into a fraction (62) of C3+ hydrocarbons, at least a portion of the C1-C2 hydrocarbon fraction (60) being conveyed to a steam cracker to extract an ethylene stream, at least a portion of the C3+ hydrocarbon fraction (60) being sent to the olefin oligomerization step (c) and the aromatics alkylation step (d), a portion (64) of the C1-C2 hydrocarbon fraction being optionally recycled to the C1 to C6 alcohol stream conversion step (a). 6.A process according to claim 5, comprising separating the C3+ hydrocarbon fraction (62) to form a C3- hydrocarbon fraction (80) and a C4+ hydrocarbon fraction (82), the C4+ hydrocarbon fraction (82) being sent to the olefin oligomerization step (c) and the aromatic alkylation step (d).

7. A process according to any one of claims 1 to 3, wherein the olefin oligomerization step (c) is carried out in an oligomerization reactor, and the aromatic alkylation step (d) is carried out in an alkylation reactor, separately from the olefin oligomerization step (c). 8.A process according to claim 7, comprising: separating the water-depleted mixture (19) into a fraction (60) of C1-C2 hydrocarbons, a fraction (142) of C3-C5 hydrocarbons, and a fraction (144) of C6+ hydrocarbons, the fraction (60) of C1-C2 hydrocarbons and the fraction (144) of C6+ hydrocarbons being sent at least in part to the alkylation step (d) in the alkylation reactor, the fraction (142) of C3-C5 hydrocarbons being sent to the oligomerization step (c) in the oligomerization reactor or comprising: separating the water-depleted mixture into a fraction of C3- hydrocarbons, a fraction of C4-C5 hydrocarbons, and a fraction of C6+ hydrocarbons, the. fraction of C3- hydrocarbons and fraction of C6+ hydrocarbons being sent at least in part to step (d) of alkylation in the alkylation reactor, fraction of C4 to C5 hydrocarbons being sent at least in part to step (c) of oligomerization in the oligomerization reactor.

9. A process according to claim 8, taken in combination with claim 3, wherein a product of the oligomerization reactor containing the oligomerized olefins and a product (152) of the alkylation reactor containing the alkylated aromatics are separated into the fraction (128) of C8+ hydrocarbons and the fraction (126) of C7- hydrocarbons, the fraction (126) of C7- hydrocarbons being recycled at least in part in step (c) to the oligomerization reactor, at least a portion of the fraction (128) of C8+ hydrocarbons forming the hydrocarbon stream to be hydrogenated (24). 10.A process according to claim 8, wherein the oligomerization reactor product (146) and the alkylation reactor product (152) containing the alkylated aromatics are separated into a fraction (126) of C7- hydrocarbons, a fraction (162) of C8-C16 hydrocarbons and a fraction (164) of C17+ hydrocarbons, at least a part of the fraction (162) of C8-C16 hydrocarbons forming the hydrocarbon stream to be hydrogenated (24), the fraction (164) of C17+ hydrocarbons being at least partly recycled in step (a) of converting the alcohol stream to C1-C6.

11. A process according to any one of the preceding claims, wherein between 10% and 90% by mass of the aromatics contained in the hydrocarbon stream to be hydrogenated (24), preferably between 30% by mass and 80% by mass of the aromatics contained in the hydrocarbon stream to be hydrogenated, are hydrogenated to cycloparaffins in step (f). 12.A process according to any one of the preceding claims, wherein the C1-C6 alcohol stream (14) is introduced at the conversion step (a) at a temperature at least 5°C higher than the bubble point of the C1-C6 alcohol stream, and lower than the temperature of the conversion reaction carried out in step (a), in particular at least 50°C lower than the temperature of the conversion reaction carried out in step (a), advantageously at least 100°C lower than the temperature of the conversion reaction carried out in step (a).

13. A process according to any one of the preceding claims, comprising stripping at least a portion (40b) of the water separated in the separation step (b) to produce a stream (41a) of extracted hydrocarbons, advantageously recycled in the separation step (b), and a treated water stream (41b).

14. A process according to any one of the preceding claims, wherein the fluidized catalytic bed receives, in addition to the C1-C6 alcohol stream (14), an exothermicity control stream comprising recycled water (40) separated from the mixture (16) in the separation step (b) and / or C4- hydrocarbons obtained from the water-depleted mixture (19), the mass ratio of recycled water (40a) in the exothermicity control stream to the C1-C6 alcohol stream (14) in the feed supplied in the conversion step (a) advantageously being between 0 and 1, preferably between 0.05 and 0.

5. 15.A process according to any one of the preceding claims, wherein the conversion step (a) is carried out in the presence of a conversion catalyst comprising a phosphorus-modified zeolite partially having an ALPO structure or in the presence of a conversion catalyst comprising a B-modified zeolite.

16. Use of the jet fuel fraction (34) produced by carrying out the process according to any one of the preceding claims, (i) pure, or (ii) blended with jet fuel resulting from the distillation of petroleum, to power at least one aircraft engine. 17.Installation (10; 90; 100; 110; 120; 140; 160) for manufacturing a jet fuel comprising: - a stage (12) for converting a stream (14) of C1 alcohol to C6 to form a mixture (16) containing paraffins, olefins, aromatics, and water; - a stage (18) for separating water (40) from the mixture (16) to form a water-depleted mixture (19); - a stage (22; 22A) for oligomerizing olefins from the water-depleted mixture (19); - a stage (22; 22B) for alkylating aromatics from the water-depleted mixture (19); - a stage for forming a hydrocarbon stream to be hydrogenated (24) from at least a portion of the oligomerized olefins in the oligomerization stage and from at least a portion of the alkylated aromatics in the alkylation stage; - a stage (26) for hydrogenating the hydrocarbon stream to be hydrogenated (24) to form a stream (30) of hydrogenated hydrocarbons; - a fractionation stage (28) for recovering at least one jet fuel fraction (34) from the hydrogenated hydrocarbon stream (30); wherein the conversion stage (12) comprises a reaction zone having at least one fluidized catalytic bed, the conversion stage (12) being configured to produce a mixture containing paraffins, olefins, aromatics, and water in which the ratio of the mass of C3+ olefins to the total mass of olefins is greater than or equal to 0.8.

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