Method for producing jet fuel, related jet fuel and plant

Abstract

The method comprises the steps of: (a) converting a C1-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 from a water-depleted mixture (19); (d) alkylating aromatics from the water-depleted mixture (19); (e) forming a hydrocarbon stream (24) to be hydrogenated from the olefins oligomerized in step (c) and the aromatics alkylated in step (d); and (f) hydrogenating the hydrocarbon stream (24) to be hydrogenated. (g) recovering a jet fuel fraction (34) from the hydrogenated hydrocarbon stream (30), wherein in the mixture (16) 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.

Description

[Technical field]

[0001] The present invention relates to a method for producing jet fuel, comprising: (a) converting a C1-C6 alcohol stream to produce a mixture containing paraffins, olefins, aromatics and water; (b) separating water from the mixture to form a water-deficient mixture; (c) oligomerizing olefins from the water-depleted mixture; (d) alkylating aromatic compounds from the water-deficient mixture; (e) forming a hydrocarbon stream to be hydrogenated from at least a portion of the olefins oligomerized in step (c) and at least a portion of the aromatic compounds alkylated in step (d); (f) hydrogenating the hydrocarbon stream to be hydrogenated to form a hydrogenated hydrocarbon stream (30); (g) recovering at least one jet fuel fraction from the hydrogenated hydrocarbon stream; The present invention relates to a method comprising the steps of:

[0002] The field of the invention relates to the preparation and use of liquid fuels, in particular of the type such as jet fuel or renewable aviation fuel. [Background technology]

[0003] Given the scarcity of fossil-based resources and growing environmental concerns, there is an increasing demand to replace fossil-based molecules with alternative molecules that have a lower carbon footprint, particularly with the aim of reducing greenhouse gas emissions.

[0004] Renewable fuels (also called "E-fuels") derived from biological materials or from carbon dioxide converted in the presence of decarbonized hydrogen or electrolytic hydrogen are alternatives to traditional fossil fuels.

[0005] Conventional jet fuels can be blended with base stocks derived from renewable sources, such as those specified in the D7566-21 standard, to allow for the production of alternative aviation fuels. Examples of base stocks for aviation fuels derived from renewable sources that can be incorporated into fossil-based jet fuels include the following: - Synthetic paraffinic kerosene [SPK], derived from processes such as the Fischer-Tropsch process; - Synthetic paraffinic kerosene produced via the "alcohol-to-jet" process (conversion of alcohol to isoparaffinic kerosene) [ATJ-SPK]; - Synthetic isoparaffins produced by hydrotreating iso-olefin intermediates produced from fermented sugars [SIP-HFS]; - Synthetic aromatic kerosene [SPK / A] obtained by alkylation of light aromatic compounds from non-petroleum sources; - synthetic kerosene obtained from hydrothermal conversion of fatty acids and fatty acid esters; - Paraffinic kerosene [SPK] obtained from hydrotreated hydrocarbons, esters and fatty acids.

[0006] Currently, most of these base stocks for renewable aviation fuels cannot be used as is because their composition is too far removed from that of fossil fuels. This difference in composition causes compatibility problems with the materials of the components with which the fuel comes in contact. In this regard, it is believed that the absence of aromatic compounds in most of the available renewable aviation fuel systems may cause compatibility problems with materials, particularly with certain sealants.

[0007] EP 2123736 describes a process for producing diesel fuel using a feedstock in the form of C1-C5 alcohols, which may be entirely or partly of biological origin, in which synthetic hydrocarbons are oligomerized or subjected to hydrogenation, based on a mixture of olefinic hydrocarbons obtained at least partly by dehydration of C1-C5 alcohols with certain proportions of odd-numbered olefins and isoolefins. After subsequent hydrogenation and rectification, an aviation fuel is formed with a freezing point below -47°C.

[0008] U.S. Patent Application Publication No. 2021 / 0078921 describes the conversion of methanol to gasoline, which can be carried out using heavy gasoline processing followed by separation operations.

[0009] U.S. Pat. No. 4,543,435 describes a conversion process for converting oxygenated feedstocks, including methanol, dimethyl ether, and the like, into liquid hydrocarbons, which comprises contacting the feedstock with a zerolite catalyst at elevated temperature and moderate pressure in a first catalytic stage to convert the feedstock into hydrocarbons including C5+ hydrocarbons and C2-C4 olefins.

[0010] EP 1844125 describes a process for producing synthetic fuels, in which in a first step a gas mixture comprising methanol and / or dimethyl ether and / or another oxygenated molecule and steam is converted at a temperature of 300-500° C. into olefins, preferably having 2-8 carbon atoms; in a second step the resulting olefin mixture is oligomerized at higher pressure to higher olefins, substantially containing more than 5, preferably 10-20 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 essentially composed of saturated hydrocarbons, which is 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 steam stream, which is separated from the product stream of the first step of the process and returned to the first step of the process.

[0011] EP 2147082 describes a process for producing synthetic fuels from a mixture containing hydrogen and oxygenated compounds such as, for example, methanol and / or dimethyl ether: in a first step, the mixture is reacted over a catalyst to obtain a hydrocarbon product containing olefins, preferably having 2 to 8 carbon atoms; in a second step, the hydrocarbon product thus obtained is oligomerized to long-chain olefins from which it is possible to obtain products that are gasoline and diesel.

[0012] WO 2011 / 061198 describes a hydrocarbon production process for producing hydrocarbons in the form of gasoline by converting synthesis gas to obtain oxygen-containing compounds such as methanol and / or dimethyl ether in a first converter and by further conversion to hydrocarbons in a second converter.

[0013] EP 2940103 describes a biofuel preparation process for preparing biofuel using ethanol by converting the ethanol into a mixture with hydrocarbons in a catalytic process over a bed of aluminosilicates such as zeolites, preferably in the presence of a zeolite catalyst in hydrogen form.

[0014] EP 2720990 describes an alcohol conversion process for converting alcohol to hydrocarbons, comprising contacting the alcohol, as a component of an aqueous solution at a concentration of 20% or less, with a metal supported zeolite catalyst at a temperature of at least 100°C and up to 550°C, wherein the alcohol may be produced by a fermentation process and is selected from among ethanol, butanol, isobutanol or combinations thereof; the metal comprises vanadium; and the metal supported zeolite catalyst is catalytically active for converting the alcohol to the hydrocarbons.

[0015] EP 3795658 describes a method for reducing energy and water consumption in a fuel production process for producing fuels from renewable alcohol-containing feedstocks. Alcohol is converted directly into 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 by lowering the water temperature, which helps to reduce evaporation losses.

[0016] US Patent Application Publication No. 2016 / 0090333 describes a hydrocarbon production process for producing aviation hydrocarbons from biorenewable sources, such as the oligomerization of C3-C8 biorenewable olefins derived from C3-C8 alcohols produced by fermentation of biomass. The production of aviation hydrocarbons is increased by using an additional oligomerization zone to oligomerize gasoline separated from the effluent of a primary oligomerization zone in which the C3-C8 biorenewable olefins were first subjected to oligomerization.

[0017] WO 2011 / 045535 describes a process for producing a distillate from a feedstock of heteroatomic organic compounds containing at least one heteroatom selected from oxygen, sulfur, halogen, individually or in combination, in which the treatment of the feedstock comprises at least one conversion step, carried out in a first conversion zone, for converting the heteroatomic organic compounds into olefins; and an oligomerization step, in at least one second oligomerization zone, for oligomerizing the olefins, at least partly originating from the conversion zone, in the presence of at least 0.5% by mass of an oxygenated compound, in order to produce a distillate. This process enhances the yield of the distillate and makes it possible to obtain a higher oligomerization rate compared to the oligomerization of the same feedstock under the same reaction conditions.

[0018] WO 2022 / 063994 describes a process that uses low temperatures (e.g. below 350°C), pressure levels of around 5-10 bar and high space velocities (6h -1 ~10h -1 ) describes a process for obtaining jet fuel which includes a conversion step of converting an oxygenated compound stream followed by a joint oligomerization and hydrogenation step in the same reactor. This process tends to ensure that the ethylene and aromatic compounds produced during the conversion are kept at very low content levels.

[0019] None of these methods provide a means to optimize the content and type of compounds, particularly aromatic compounds, in the final jet fuel composition to ensure good compatibility with existing systems while at the same time enhancing combustion quality. [Prior art documents] [Patent documents]

[0020] [Patent Document 1] European Patent No. 2123736 [Patent Document 2] US Patent Application Publication No. 2021 / 0078921 [Patent Document 3] U.S. Pat. No. 4,543,435 [Patent Document 4] European Patent No. 1844125 [Patent Document 5] European Patent No. 2147082 [Patent Document 6] International Publication No. 2011 / 061198 [Patent Document 7] European Patent No. 2940103 [Patent Document 8] European Patent No. 2720990 [Patent Document 9] European Patent No. 3795658 [Patent Document 10] US Patent Application Publication No. 2016 / 0090333 [Patent Document 11] International Publication No. 2011 / 045535 [Patent Document 12] International Publication No. 2022 / 063994 Summary of the Invention [Problem to be solved by the invention]

[0021] It is therefore an object of the present invention to provide an efficient and highly productive fuel production process for producing jet fuel derived exclusively or at least in part from renewable sources, while also effectively meeting the requirements for use in the aviation sector. [Means for solving the problem]

[0022] For this purpose, the subject of the present invention is a method for the production of fuel of the type described above, comprising the steps of: The present invention relates to a method for producing fuels, wherein in the mixture 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 0.8 or more.

[0023] The method according to the invention may comprise one or more of the following features, taken individually or according to any technically feasible combination: - the jet fuel fraction comprises between 2% and 30% by volume of C8+ aromatics, preferably between 8% 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 separating at least a portion of the olefins oligomerized in step (c) and / or at least a portion of the aromatics alkylated in step (d) into a fraction of C7-hydrocarbons and a fraction of C8+hydrocarbons, the C7-hydrocarbon fraction being at least partially recycled to the oligomerization step (c) for oligomerizing the olefins and / or to the alkylation step (d) for alkylating the aromatics, the hydrocarbon stream to be hydrogenated being formed by at least a portion of the C8+hydrocarbon fraction; - the olefin oligomerization step (c) and the aromatics alkylation step (d) are carried out jointly in the same reactor; - the process comprises a separation step for separating the water-depleted mixture into a C1-C2 hydrocarbon fraction and a C3+ hydrocarbon fraction, at least a portion of the C1-C2 hydrocarbon fraction being conveyed to a steam cracker for extracting an ethylene stream therefrom; and at least a portion of the C3+ hydrocarbon fraction being sent to a step (c) for oligomerizing olefins and to a step (d) for alkylating aromatics; optionally a portion of the C1-C2 hydrocarbon fraction being recycled to the conversion step (a) for converting a C1-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 a step (c) for oligomerizing olefins and to a step (d) for alkylating aromatic compounds; - the oligomerization step (c) for oligomerizing olefins is carried out in an oligomerization reactor and the alkylation step (d) for alkylating aromatics is carried out in an alkylation reactor separate from the oligomerization step (c) for oligomerizing olefins; - the method comprises: or 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 at least partially fed to an alkylation step (d) in an alkylation reactor and the C3-C5 hydrocarbon fraction being fed to an oligomerization step (c) in an oligomerization reactor; or 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 fed at least in part to an alkylation step (d) in an alkylation reactor and the C4-C5 hydrocarbon fraction being fed at least in part to an oligomerization step (c) in an oligomerization reactor; - the oligomerization reactor product containing oligomerized olefins and the alkylation reactor product containing alkylated aromatics are separated into a C8+hydrocarbon fraction and a C7-hydrocarbon fraction, the C7-hydrocarbon fraction being at least partially recycled to step (c) in the oligomerization reactor, and at least a portion of the C8+hydrocarbon fraction forming the hydrocarbon stream to be hydrogenated; - the oligomerization reactor product and the alkylation reactor product containing alkylated aromatics are separated into a C7-hydrocarbon fraction, a C8-C16 hydrocarbon fraction and a C17+hydrocarbon fraction, at least a portion of the C8-C16 hydrocarbon fraction forming the hydrocarbon stream to be hydrogenated and the C17+hydrocarbon fraction being at least partially recycled to the conversion step (a) for converting the C1-C6 alcohol stream; - 10% to 90% by weight of the aromatics contained in the hydrocarbon stream to be hydrogenated, preferably 30% to 80% by weight of the aromatics contained in the hydrocarbon stream to be hydrogenated, are hydrogenated to cycloparaffins in step (f); - the C1-C6 alcohol stream contains at least 50% methanol, and the conversion step (a) for converting the C1-C6 alcohol stream comprises the addition of an alcohol stream containing C2-C6 alcohols between two conversion catalyst beds for the conversion step (a); - the method comprises, after a hydrogenation step (f) for hydrogenating the hydrocarbon stream to be hydrogenated, a separation of the hydrogenated hydrocarbon stream into at least a jet fuel fraction and a diesel fraction; - at least a portion of the water from the mixture separated in separation step (b) is recycled to the conversion step (a); - the conversion step (a) is carried out in a series of fixed catalyst beds and a portion of the water from the mixture recycled to the conversion step (a) is introduced upstream of the series of fixed catalyst beds or between two fixed catalyst beds; or the conversion step (a) is carried out in at least one fluidized catalyst bed and a portion of the water from the mixture recycled to the conversion step (a) is introduced into the fluidized catalyst bed, optionally together with a C4-hydrocarbon stream, preferably together with at least a portion of the C1-C2-hydrocarbon fraction obtained by separation of the water-depleted mixture; - 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; - the C2-C6 alcohol stream is obtained by fermentation of synthesis gas or biomass and / or the C1-C6 alcohol stream is obtained by catalytic conversion of carbohydrates, carbon monoxide or carbon dioxide in the presence of hydrogen; - the hydrogenation step (f) for hydrogenating the hydrocarbon stream to be hydrogenated is carried out separately from and downstream of the oligomerization step (c) for oligomerizing olefins derived from the water-depleted mixture and the alkylation step (d) for alkylating aromatics derived from the water-depleted mixture; - the conversion step (a) is carried out using at least one catalyst comprising a molecular sieve containing in its micropore structure pores at least the size of 10 oxygen atoms (10-MR) or more; the catalyst for carrying out the conversion step (a) comprises: a phosphorus-modified zeolite, in particular having a P content of at least 0.05% by weight, preferably between 0.3% and 7% by weight; a composite catalyst comprising at least 0.1% by weight of a silicate; or a molecular sieve modified with phosphorus (P) and an alkaline earth metal or a rare earth metal (M) (MP-modified molecular sieve), advantageously with an M / P molar ratio within the molecular sieve of less than 1; - the catalyst for carrying out the conversion step (a) comprises a phosphorus-modified zeolite having an ALPO structure advantageously characterized by an 27Al NMR signature between 35 ppm and 45 ppm; - the catalyst for carrying out the conversion step (a) is modified by the addition of one or more metals selected from among: group IIB metals, in particular Zn; group IIIB metals, in particular Ga; group VIIIB transition metals, in particular Fe and / or Ni and / or Pt; group VIB metals, in particular Mo; group IB metals, in particular Cu and / or Ag; and also metals from the lanthanide group, in particular La; - the catalyst for carrying out the conversion step (a) comprises a zeolite with pores with a size of 10 oxygen atoms (10-MR) or more, modified by adding B before, after or simultaneously with the preparation step of preparing the final catalyst; the catalyst for carrying out the conversion step (a) has a Si / Al atomic ratio, measured by chemical analysis, in particular by NMR, taking into account only the Al which is part of the framework structure of the molecular sieve, of between 4 and 500, preferably between 5 and 200, or more preferably between 12 and 150; the process comprises stripping at least a portion of the water separated in separation step (b) to produce an extracted hydrocarbon stream and a treated water stream which are advantageously recycled to separation step (b).

[0024] The subject of the present invention is a power source for powering at least one aircraft engine, (i) in pure form; or (ii) in the form of a mixture with jet fuel resulting from the distillation of petroleum and / or another renewable source; It also relates to the use of a jet fuel fraction produced by carrying out the production process defined above.

[0025] Advantageously, the jet fuel fraction produced comprises between 2% and 30% by mass of aromatics, in particular between 6% and 20% by mass of aromatics.

[0026] In particular, the jet fuel fraction comprises from 2% to 30% by weight of aromatic compounds having at least 8 carbon atoms, in particular from 6% to 20% by weight of aromatic compounds having at least 8 carbon atoms.

[0027] Preferably, more than 50 mass % of the aromatic compounds contained in the jet fuel fraction are monocyclic aromatic compounds having 8 to 14 carbon atoms.

[0028] It advantageously contains between 5% and 20% by weight of cycloparaffins and at least 50% by weight of isoparaffins.

[0029] The method according to the invention therefore makes it possible to obtain renewable aviation fuels that are fully replaceable ("drop-in") or compatible with current aircraft engines and fuel systems.

[0030] The subject matter of the present invention is a jet fuel production plant, comprising: - a conversion stage for converting the C1-C6 alcohol stream to form a mixture containing paraffins, olefins, aromatics and water; - a separation stage for separating water from the mixture to form a water-deficient mixture; - an oligomerization stage for oligomerizing olefins from the water-deficient mixture; - an alkylation stage for alkylating aromatic compounds from the water-deficient mixture; - a stream formation stage for forming a hydrocarbon stream to be hydrogenated from at least a portion of the olefins oligomerized in the oligomerization stage and from at least a portion of the aromatic compounds alkylated in the alkylation stage; - a hydrogenation stage for hydrogenating the hydrocarbon stream to be hydrogenated to form a hydrogenated hydrocarbon stream; - a fractionation stage for recovering at least one jet fuel fraction from the hydrogenated hydrocarbon stream; A fuel production plant comprising: The invention also relates to a fuel production plant, wherein the conversion stage is configured to produce a mixture containing paraffins, olefins, aromatics and water, wherein the ratio of the mass of C3+ olefins to the total mass of olefins is 0.8 or greater.

[0031] In one variation, the plant includes at least one recycle conduit for recycling at least a portion of the water from the mixture separated in the water separation stage to the conversion stage.

[0032] [Detailed Description] As used herein, the terms "comprising" and "comprises" are synonymous with the terms "including," "includes," or "contains," "containing," and are inclusive or open-ended and are not intended to exclude additional, unspecified features, elements or method steps.

[0033] The terms "% by mass" and "mass%" have the same meaning and refer to the percentage by mass of a product in 100 g of a composition containing that product.

[0034] The boiling points mentioned herein are measured at atmospheric pressure unless otherwise stated. The initial boiling point (hereinafter referred to as "IBP") is defined as the temperature value at which the first vapor bubble is formed. The final boiling point (hereinafter referred to as "FBP") is the highest temperature that can be reached during distillation. At this temperature, no more vapor can be transported to the condenser. The determination of the initial and final points is based on techniques known to those skilled in the art, and several methods are applicable, adapted depending on the distillation temperature range, such as, for example, NF EN15199-1 (2020 edition) or ASTM D2887 for the measurement of the boiling point of petroleum fractions by gas chromatography; ASTM D7169 for heavy hydrocarbons; ASTM D7500, D86 or D1160 for distillates.

[0035] By default, the terms "Cn-Cm" stream, flow rate, fraction, etc. refer to a stream, flow rate, fraction, etc. having a major proportion (e.g., greater than 50 mol %) of compounds having between n and m carbon atoms.

[0036] The term "Cn+" stream, flow rate, fraction, etc. means a stream, flow rate, fraction, etc. having a major amount (eg, greater than 50 mole %) of compounds having n or more carbon atoms.

[0037] The term "Cn-" stream, flow rate, fraction, etc. refers to a stream, flow rate, fraction, etc. having a majority (eg, greater than 50 mole %) of compounds having n or fewer carbon atoms.

[0038] Unless otherwise indicated, percentages used are percentages by weight and pressures are absolute pressures.

[0039] [Step for obtaining C1-C6 alcohol stream] The C1-C6 alcohol stream contains primarily alcohols such as methanol, ethanol, propanol (n-propanol, i-propanol), butanol (n-butanol, i-butanol), pentanol (n-pentanol, i-pentanol) and hexanol.

[0040] It may contain small amounts of C6+ alcohols, and / or oxygenated compounds such as methyl ethyl ether; dimethyl ether; diethyl ether; diisopropyl ether; formaldehyde; dimethyl carbonate; dimethyl ketone; acetic acid; furan; tetrahydrofuran, and mixtures thereof.

[0041] The C1-C6 alcohol stream forming the feedstock for the process advantageously comprises more than 80% by weight of C1-C6 alcohols, preferably more than 90% by weight of C1-C6 alcohols. It advantageously comprises more than 50% by weight of methanol, in particular more than 80% by weight of methanol.

[0042] In one variant, when the conversion is carried out using multiple reaction zones with fixed catalyst beds, as explained below, a C2-C6 alcohol stream is advantageously also added between two conversion reaction zones of the conversion step (a).

[0043] The ratio of the mass flow rate of the C2-C6 alcohol stream added between the two reaction zones to the mass flow rate of the C1-C6 alcohol stream entering the first reaction zone is, for example, less than 0.5, in particular 0.05-0.5.

[0044] Preferably, the C1-C6 alcohol stream and advantageously the additional C2-C6 alcohol stream are obtained from hydrogen produced from renewable sources such as biomass (including its constituents and derivatives) or possibly from captured carbon monoxide or carbon dioxide, and advantageously from renewable energy sources such as solar energy, wind, geothermal energy, ocean waves or tides and / or from energy that does not generate carbon dioxide at the time of its production, such as nuclear energy.

[0045] Production routes for producing alcohols intended to form the C1-C6 alcohol stream include, but are not limited to, for example: - anaerobic fermentation of sugars from biomass, in particular to obtain ethanol; - Catalytic reaction of carbon dioxide or carbon monoxide with hydrogen, in particular to obtain methanol, ethanol and other alcohols; - Catalytic reactions of carbohydrates with hydrogen, specifically to obtain C1-C6 alcohols; - Acetone-butanol-ethanol (ABE) fermentation 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, in particular to obtain ethanol, propanol (iso or n), butanol (iso or n) or isoamyl alcohol.

[0046] In one embodiment, alcohol, particularly ethanol from renewable sources, can be obtained by ethanol fermentation in a bioreactor comprising a culture of one or more microorganisms, with the aim of obtaining alcohol, particularly ethanol from renewable sources, by fermentation.

[0047] Ethanol from renewable sources can then advantageously be obtained by: - Anaerobic fermentation of sugar-rich substrates derived from biomass, or - Anaerobic fermentation of gases containing CO, which may or may not be derived from biomass.

[0048] Thus, in the case of anaerobic fermentation, ethanol may be produced by anaerobic fermentation of a sugar-rich substrate derived from biomass.

[0049] Sugars are made up of chains of six or five carbon atoms, such as glucose, sucrose (a dimer of glucose and fructose), xylose and arabinose.

[0050] The substrate may, for example, comprise or be obtained directly from agro-food plants, sugar cane, sugar beet, sweet sorghum; or may be obtained by depolymerization of cellulose and hemicellulose from starch and / or lignocellulosic biomass from corn, wheat, barley, rye, sorghum, triticale, potato, sweet potato, manioc.

[0051] The sugar-rich substrate may be derived from lignocellulosic biomass by a process that includes: (i) a separation step to separate the lignin, cellulose and hemicellulose contained within the lignocellulosic biomass; followed by (ii) a conversion step to convert the cellulose and / or hemicellulose into sugars.

[0052] The production of this type of substrate from lignocellulosic biomass is well known to those skilled in the art. Lignocellulosic biomass is essentially composed of cellulose, hemicellulose and lignin. This biomass is obtained from agricultural and forestry residues or by-products of wood or crop processing, whether involving woody or herbaceous plants. This lignocellulosic biomass may also include distillers grains and can be used to produce ethanol as described in EP 2675778.

[0053] The first step (i) is a pretreatment step used to decompose the lignocellulosic matrix and liberate cellulose and hemicellulose from the complexes formed with lignin using one or more pretreatment processes. Among the existing pretreatment processes, steam pretreatment (or steam explosion), liquid hot water pretreatment (hot water pretreatment), ammonia fiber explosion (AFEX), acid pretreatment or alkaline pretreatment are well known. Steam explosion treatment consists in treating the biomass, preferably previously chopped or crushed, with high pressure saturated steam at a temperature of around 160-240°C and at a pressure of 0.7-4.8 MPa. The effectiveness of the steam treatment is determined by the use of H as a catalyst. 2 SO 4 , CO 2or SO 2 In AFEX pretreatment, biomass is contacted with anhydrous liquid ammonia feedstock in a ratio of 1:1 to 2:1 (1-2 kg ammonia per kg dry biomass) at 60-90 °C and pressures above 3 MPa for 10-60 minutes. Hydrothermal pretreatment is similar to steam explosion, but instead of steam, water in liquid state at high temperature is used. In acid pretreatment, an aqueous suspension of cellulosic substrate is heated to the desired temperature, typically in the presence of dilute acid, and pretreated with preheated sulfuric acid (concentration level below 4 wt%) in a stainless steel reactor, and the treatment process is carried out at temperatures between 140-215 °C. Residence times vary between a few seconds and a few minutes depending on the treatment temperature. Lime pretreatment is an inexpensive physicochemical alkaline treatment that effectively enhances the digestibility of cellulosic biomass. Using 0.1 g Ca(OH)2 per g biomass, the treatment process can be carried out over a wide temperature range from 25 to 130 °C. Organosolv processes, which are processes for delignification and / or saccharification of cellulosic materials and plant cultures, may also be used. In general, organosolv processes involve the use of mixtures of water and solvents such as alcohols or ketones, and sometimes other non-polar solvents, as well as acidic compounds to promote hydrolysis. This type of process is described, for example, in U.S. Pat. No. 4,470,851.

[0054] Step (ii) is a conversion step for converting cellulose and / or hemicellulose to sugars, which is also well known to those skilled in the art. This typically involves hydrolysis, which can be catalyzed by the use of acids or by enzymes such as cellulases, xylanases, xylosidases and arabinofuranosidases (e.g. produced by strains of Trichoderma reesei).

[0055] The sugar-rich substrate is then subjected to fermentation.

[0056] By way of example, this fermentation can be carried out using microorganisms, and in particular specific yeasts that help to optimize the yield of the production process, such as in particular the following yeasts: Ethanol Red® (Fermentie), Thermosacc® (Lallemand), Angel Super Alcohol® (Angel®) and Fali® (AB Mauric); the yeast strain Saccharomyces cerevisiae described in French Patent No. 3015985; the yeast strains Candida Shehatae or Pichia stipitis; or any other suitable microorganism.

[0057] For anaerobic fermentation, ethanol can be produced by anaerobic fermentation of a gas containing CO. In this case, the substrate is a gaseous substrate (gas) containing CO. The gaseous substrate may be a by-product of an industrial process.

[0058] In some embodiments, the industrial process is selected from the group consisting of: the production of ferrous metal products, particularly steel products; the production of non-ferrous products; petroleum refining processes; the gasification of coal and / or biomass or biochar; the production of electrical energy; the production of carbon black; the production of ammonia; the production of methanol; the production of coke; catalytic cracking (particularly carbon monoxide is produced during catalyst regeneration); and the reforming of methane. In these embodiments, the gaseous substrate can be captured from the industrial process using a suitable method before being released into the atmosphere. Depending on the composition of the gas thus captured, it may also be desirable to treat it to remove any undesirable impurities, such as dust particles, before being introduced into the fermentation process. For example, the gas can be filtered or purified by known methods.

[0059] In other embodiments, the gaseous substrate can be derived from the gasification of biomass. The gasification process involves the partial combustion of biomass in a limited supply of air or oxygen. The resulting gas is typically primarily CO and H 2, and a minimal volume of CO 2 For example, biomass by-products obtained during food extraction and processing, such as sugar from sugar cane or starch from corn or cereals, or non-food biomass waste generated by forestry, can be gasified to produce CO2-containing gases that can be used in the present invention.

[0060] Typically used gaseous substrates have a significant 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 within any range defined by two of these limits. Advantageously, the CO-containing gas may contain 25%, 30%, 35%, 40%, 45%, 50%, 55% or 60% by volume of CO. Gases with a lower CO content, such as 6% by volume, may also be used, particularly H. 2 and CO 2 It may be preferred if also present.

[0061] If the gaseous substrate contains CO, the gaseous substrate need not contain hydrogen, but this is not considered to be detrimental to the production of ethanol. The gaseous substrate may contain, for example, 1% to 80% by volume, or 1% to 30% by volume, or 5% to 10% by volume, or within any range defined by any two of these limits. 2 It may also contain.

[0062] Typically, carbon monoxide is added to the fermentation reaction in gaseous or liquid form. For example, carbon monoxide can be fed into the liquid by saturation. For example, a liquid can be saturated with carbon monoxide-containing gas and then added to the bioreactor. This can be done using any 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) can be used.

[0063] Furthermore, it is often desirable to increase the CO concentration (or partial pressure of CO in a gas) of a gas, thus increasing the effectiveness of fermentation reactions that use CO as a substrate. Increasing the partial pressure of CO in a gas increases the mass transfer of CO in the fermentation medium. The composition of the gas stream used to feed a fermentation reaction can have a significant impact on the effectiveness and / or cost of the reaction. For example, O 2 can reduce the effectiveness of the anaerobic fermentation process. Treating unwanted or unnecessary gases in steps of the fermentation process before or after fermentation can increase the load on those steps (e.g., if the gas stream is compressed before entering the bioreactor, unnecessary energy may be used to compress gases not needed for fermentation). Therefore, it may be desirable to treat substrate streams, especially those derived from industrial sources, to remove undesirable constituents and increase the concentration of desirable constituents.

[0064] Any microorganism capable of fermenting a gaseous substrate, including CO, to produce ethanol can be used in the present invention. By way of example, microorganisms of the genera Moorella, Clostridia, Ruminococcus, Acetobacterium, Eubacterium, Butyribacterium, Oxobacter, Methanosarcina, and Desulfotomaculum can be used.

[0065] As an example, it is possible to use one or more microorganisms of the genera Clostridium, including strains of Clostridium ljungdahlii, Clostridium carboxydivorans, Clostridium ragsdalei and Clostridium autoethanogenum; Moorella, including Moorella sp HUC22-1; Carboxydothermus; Moorella thermoacetica, Moorella thermoautotrophica; Ruminococcus productus; Acetobacterium woodii; Eubacterium limosum; Butyribacterium methylotrophicum; Oxobacter pfennigii; Methanosarcina barkeri; Methanosarcina acetivorans; or Desulfotomaculum kuznetsovii. Another specific example of a microorganism is an anaerobic carboxydotrophic bacterium. Examples of strains that can be used are described in WO 2012 / 026833.

[0066] It should be noted that the present invention is applicable to mixed cultures of two or more microorganisms.

[0067] Regarding fermentation media and conditions, regardless of the nature of the substrate used (gaseous or otherwise), a suitable nutrient medium must be introduced into the bioreactor under suitable conditions in addition to the substrate for ethanol fermentation to occur using the growth of one or more microorganisms. The nutrient medium will contain suitable components such as vitamins and minerals that are sufficient to allow the growth of the microorganisms used. The reaction conditions to be considered are temperature, medium flow rate, pH, redox potential of the medium, agitation rate (if a continuous stirred reactor is used), inoculum level, maximum substrate concentration and rate of introduction of substrate into the bioreactor to ensure that the substrate level is not limiting, and maximum product concentration to avoid product inhibition. Optimal reaction conditions will depend in part on the specific microorganism used. Processes for culturing microorganisms are known in the art, and those skilled in the art know how to optimize the culture conditions for each microorganism based on its properties. Examples of suitable fermentation conditions for anaerobic fermentation of substrates containing CO are detailed in WO 2007 / 117157, WO 2008 / 115080, WO 2009 / 022925 and WO 02 / 08438.

[0068] The fermentation reaction may be carried out in any suitable bioreactor. In some embodiments of the invention, the bioreactor may include a first growth reactor in which the microorganism is cultured and a second fermentation reactor in which the broth from the growth reactor is introduced and the majority of the fermentation product (e.g., ethanol) is produced.

[0069] Fermentation results in a fermentation broth containing the desired product (ethanol) and / or one or more by-products (acetate and butyrate when the substrate is a CO-containing gas) and microbial cells in a nutrient medium.

[0070] Recovering the ethanol can include continuously withdrawing a portion of the culture medium and recovering the ethanol from the removed portion of the culture medium.

[0071] For example, the removed portion of the ethanol-containing broth can be passed through a separation unit to separate the bacterial cells from the broth, e.g., by filtration, and produce a cell-free, ethanol-containing permeate and allow the microbial cells to be returned to the bioreactor.

[0072] In some embodiments, recovering ethanol and / or one or more other products or by-products produced during the fermentation reaction includes continuously removing a portion of the broth and separately recovering the ethanol and one or more other products from the removed portion of the broth.

[0073] As an example, ethanol can be recovered from the fermentation broth using methods such as filtration, distillation or fractional evaporation, pervaporation, and extractive fermentation. Distillation of ethanol from the fermentation broth provides 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 techniques, which are also well known in the art.

[0074] Ethanol from renewable sources also has a CO / H 2 It can also be obtained from biomass by conversion of a fatty acid-rich synthesis gas, said synthesis gas being derived from biomass.

[0075] For example, CO / H 2 Biomass can be gasified to produce a methanol-rich synthesis gas (also known as "syngas"), which is then converted to methanol in the presence of a catalyst. This type of process is described, for example, in WO 2012 / 003901.

[0076] Biomass used to produce syngas may include, in particular, wood fuels from forest and natural woodland sources (e.g., sawdust); agricultural residues (e.g., rice husks, straw compost); energy crops grown exclusively for energy production (e.g., corn and oil palm); urban waste (e.g., wood chips, rice, straw compost); municipal waste (e.g., municipal solid waste and wastewater); and waste-derived biomass fuels (e.g., wood pellets).

[0077] It is also possible to obtain methanol from renewable sources. The term "methanol from renewable sources" is used to mean methanol obtained from biomass.

[0078] Biomass may specifically include wood fuels from forest and natural woodland sources (e.g. sawdust); agricultural residues (e.g. rice husks, straw compost); energy crops grown exclusively for energy production (e.g. corn and oil palm); urban waste (e.g. wood chips, rice, straw compost); municipal waste (e.g. municipal solid waste and wastewater); and waste-derived biomass fuels (e.g. wood pellets).

[0079] Methanol from renewable sources is specifically CO / H 2 The synthesis gas may be obtained by converting a synthesis gas rich in nitric acid, said synthesis gas being derived from biomass.

[0080] For example, CO / H 2 Biomass can be gasified to produce a methanol-rich synthesis gas (also known as "syngas"), which is then converted to methanol in the presence of a catalyst. This type of process is described, for example, in WO 2018 / 134853.

[0081] Synthesis gas, which is suitable for subsequent conversion to methanol, also contains methane and CO 2It can be obtained by partial oxidation in the presence of oxygen of a biogas containing , for example as a result of anaerobic digestion of biomass in the presence of one or more microorganisms. A process of this type is described in WO 2019 / 060988.

[0082] The alcohol intended to form the C1-C6 alcohol stream and optionally the additional C2-C6 alcohol stream may also be obtained from carbon dioxide, in particular captured carbon dioxide.

[0083] There are numerous conversion routes. One example that may be mentioned is the catalytic conversion of carbon dioxide in the presence of hydrogen to methanol.

[0084] Another route consists of converting carbon dioxide to carbon monoxide by electroconversion or by reacting the gas with back-water in the presence of hydrogen.

[0085] The carbon monoxide is then converted to methanol by catalytic conversion in the presence of hydrogen.

[0086] The hydrogen used for the various tasks mentioned above is obtained, in particular, by steam methane reforming, by reaction of gas with water or produced by electrolysis from renewable energy sources such as solar energy, wind power, geothermal energy, waves or tides.

[0087] [Step (a) Conversion of C1-C6 alcohol streams] The conversion of the C1-C6 alcohol stream may include, for example, dehydration, carbon-carbon bonding, and / or aromatization of at least one C1-C6 alcohol.

[0088] The dehydration of the C1-C6 alcohol, the carbon-carbon bond and / or the aromatization thereof can be carried out simultaneously.

[0089] For the dehydration of methanol, it is typically converted to dimethyl ether, which is then dehydrated to produce olefins having at least two carbon atoms. In the case of methanol and / or dimethyl ether, a carbon-carbon bond is generated during the dehydration.

[0090] C2-C6 alcohols may be dehydrated to produce olefins containing the same number of atoms. Dehydration reactions for dehydrating alcohols to produce alkenes have been known for many years (J. Catal. 7, 163, 1967 and J. Am. Chem. Soc. 83, 2847, 1961). For the dehydration of alcohols, many available solid acid catalysts can be used ((Stud. Surf. Sci. Catal. 51, 260, 1989), EP 0150832, Bulletin of the Chemical Society of Japan, vol 47(2), 424-429 (1974)). However, especially for longer chain alcohols (with three or more carbon atoms), γ-alumina is the most commonly used. Indeed, catalysts with higher acidity, such as silica-alumina, zeolites, heteropolyacids or resin catalysts, can promote double bond displacement, backbone isomerization and other olefin interconversion reactions.

[0091] Additionally, aromatization of C1-C6 alcohols occurs via oligomerization of olefin intermediates, cyclization of chains having at least six carbon atoms, and dehydrogenation of cycloparaffins to the corresponding aromatic compounds.

[0092] The mechanism of reaction in the conversion of C1-C6 alcohols involves acid catalysis. The catalyst can provide protons to activate molecules, alcohols and / or olefins via protonated intermediates. The stability of the protonated intermediates or alkyl carbenium ions depends on the inductive effect of the substituents. Tertiary carbenium ions are the most stable, while primary carbenium ions are the least stable. Thus, tertiary carbenium ions are formed most easily, and reactions involving the formation of primary carbenium ions are slow. Primary carbenium ions have a tendency to transform into secondary or tertiary carbenium ions.

[0093] 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 aromatic compounds with alkenes.

[0094] Either hydride transfer provides a pathway for converting a neutral molecule to a carbenium ion, or successive hydride transfers from alkenes to carbenium ions result in the formation of aromatic compounds. The following reaction demonstrates this mechanism for propylene: [ka]

[0095] Overall, three dihydrogen molecules (six atoms) need to be removed to form an aromatic ring, and these three hydrogen molecules will form three paraffins from the olefin via a hydride transfer mechanism.

[0096] The hydride transfer mechanism typically operates on catalysts with only one acid function and often requires harsh conditions that also lead to the formation of coke.

[0097] The production of aromatics from olefins (produced from alcohols) is essentially carried out over acid catalysts via hydride transfer from one olefin to another. This produces more unsaturated molecules (and ultimately aromatics) and paraffins. These paraffins are inconvertible because they are too inert to be converted under optimal operating conditions for the conversion of C1-C6 alcohols. A dehydrogenation catalyst can be added, which allows aromatization by producing molecular hydrogen.

[0098] If the catalyst also has a dehydrogenation function, the reaction intermediates can be converted to the corresponding aromatic compounds by dehydrogenation. This is particularly true for bifunctional catalysts having an acid function and a dehydrogenation function. The dehydrogenation function can be provided by metals from groups VIB, VIIIB, IB and IIB and mixtures thereof; preferably gallium, zinc or mixtures thereof.

[0099] Generally, the dehydrogenation reaction to dehydrogenate olefins is thermodynamically limited and requires high temperatures. To promote the dehydrogenation pathway, additional reagents can be added to the conversion step (a) to shift the thermodynamic balance.

[0100] Advantageously, this reagent is carbon dioxide, which can be converted to carbon monoxide and water: CO 2 +H 2 →CO+H 2 O

[0101] CO 2 , CO and H 2 The molecules can be separated from other hydrocarbons and recycled into the synthesis of alcohols by fermentation of syngas and catalytic conversion of syngas to make methanol therefrom.

[0102] The process then involves adding a carbon dioxide containing stream in a conversion step (a) for converting a C1-C6 alcohol stream, and then simultaneously converting the carbon dioxide to carbon monoxide in a conversion step (a) for converting a C1-C6 alcohol stream.

[0103] Advantageously, the carbon dioxide containing stream comprises more than 5% by weight of carbon dioxide, in particular more than 10% by weight of carbon dioxide.

[0104] In the raw material supplied to the conversion step (a), the mass ratio of carbon dioxide supplied in the carbon dioxide stream to C1-C6 alcohols supplied in the C1-C6 alcohol stream is 5% to 75%.

[0105] The carbon dioxide containing stream is added, for example, in admixture with the C1-C6 alcohol stream; or, when multiple reaction zones are arranged in series to carry out the conversion step, between two reaction zones or within a given reaction zone.

[0106] Preferably, more than 2 mole % of the carbon dioxide, particularly more than 5 mole % is converted to carbon monoxide in conjunction with the conversion of the C1-C6 alcohol stream in the converting step (a).

[0107] The primary products of the acid-catalyzed dehydration of ethanol and / or methanol are ethylene and / or propylene and water.

[0108] More commonly, a mixture containing paraffins, olefins, aromatics and water is produced.

[0109] Paraffins include n-paraffins, i-paraffins and cycloparaffins.

[0110] According to the invention, in the mixture 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 0.80, in particular greater than or equal to 0.82, preferably greater than or equal to 0.85, this ratio being calculated on the dry stream after separation of water.

[0111] Advantageously, the resulting mixture containing paraffins, olefins, aromatics and water, not taking into account the water circulating in the water recycle process, contains more than 10% by weight of water, in particular 10-60% by weight of water, depending on the alcohol composition in the mixture of C1-C6 alcohols. When only methanol is present, the water content of the product mixture (excluding recycle) is 55% to 60% by weight.

[0112] Advantageously, the resulting mixture containing paraffins, olefins, aromatics and water, on an anhydrous basis excluding water and any recycled materials, contains more than 2% by weight of aromatics, particularly more than 6% by weight of aromatics, particularly between 6% and 30% by weight of aromatics.

[0113] On an anhydrous basis, excluding water, it advantageously contains: - less than 5% by weight of methane, in particular between 0.1% and 4% by weight of methane; - less than 5% by weight of dimethyl ether, in particular less than 1% by weight of dimethyl ether; - less than 5% by weight of residual C1-C6 alcohols, in particular less than 1% by weight of residual C1-C6 alcohols; - less than 15% by weight of ethylene, in particular between 5% and 10% by weight of ethylene; - more than 30% by weight of propylene, in particular from 35% to 60% by weight of propylene; - less than 15% by weight of paraffin, in particular between 3% and 8% by weight of paraffin; - more than 10% by weight of C4-C7 olefins, in particular 25% to 40% by weight of C4-C7 olefins; - Greater than 6% aromatics, specifically 6%-10% aromatics, specifically less than 3% C9+ aromatics.

[0114] The production of light olefins (ethylene and propylene) from a mixed alcohol feedstock in an oxygenate-olefin process is described, for example, in U.S. Patent No. 7,288,689, which proposes various methods for optionally producing C1-C4 alcohols in a mixed alcohol stream and optionally converting the alcohols to light olefins.

[0115] The conversion by dehydration and aromatization makes it possible to obtain light olefins having at least two carbon atoms and aromatic compounds from C1-C6 alcohols using a composite catalyst, the process comprising the following steps: a) providing a catalyst comprising a molecular sieve containing pores in its micropore structure that are at least 10 oxygen atom sizes or larger (10-MR); b) when the conversion is carried out in a reaction zone comprising 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 through the reaction zone and at least a portion of the catalyst in the reaction zone passes through the regeneration zone; or / and, when the conversion is carried out in a reaction zone having at least one fixed bed, a catalyst regeneration step for regenerating the catalyst in situ, either by directing the stream to be reacted into a reaction zone which has previously been regenerated or by stopping the conversion in the reaction zone; c) contacting the C1-C6 alcohols in the reactor with a catalyst under conditions effective to convert at least a portion of the feedstock, so as to form a reactor effluent comprising a mixture of paraffins, olefins, aromatics and water.

[0116] The catalyst may be a mixture of two or more catalysts and optionally a binder.

[0117] It is desirable to obtain substantially 100% conversion of the alcohol compounds in the reactor, which is controlled by optimizing contact time, reaction temperature and catalyst regeneration frequency.

[0118] In one specific embodiment, the weight hourly space velocity (hereinafter WHSV) of the alcohol in the reaction zone is about 0.5 h-1 ~about 10 h-1 , advantageously about 1 h-1 ~about 6 h-1 It is.

[0119] The molecular sieve included in the composition of the catalyst is selected from the list of molecular sieves with the crystalline structures MFI, MOR, MEL, clinoptilolite, FER, FAU, MWW, BETA, MCM-41, ZSM-21, ZSM-22, ZSM-23, ZSM-42, ZSM-57, LTL or mixtures thereof. Preferably, the molecular sieve selected is a zeolite, a crystalline aluminosilicate selected from the group comprising MFI, MOR, MEL, clinoptilolite, FER or mixtures thereof. More preferably, in the case of MFI, the molecular sieve is preferably a ZSM-5 zeolite. In a further embodiment, the molecular sieve is preferably obtained without the addition of a structuring agent. Further examples are described by the International Zeolite Institute (Atlas of Zeolite Structure Types, 1987, Butterworths).

[0120] Crystalline silicates (also called zeolites) are made up of a series of atoms linked together by the sharing of oxygen ions, XO, where X can be trivalent (e.g. Al, B, etc.) or tetravalent (e.g. Ge, Si, etc.). 4They are microporous crystalline inorganic polymers based on a tetrahedral framework. The crystal structure of crystalline silicates, determined by X-ray diffraction, is defined by the specific order in which the network of tetrahedral units is linked together. The size of the crystalline silicate pore openings is determined by the number of tetrahedral units, or alternatively oxygen atoms, required to form the pores and the nature of the cations present within the pores. They have a unique combination of properties: high internal surface area; uniform pores with one or more distinct sizes; ion exchange potential; excellent thermal stability; and the ability to adsorb organic compounds. The pores of these crystalline silicates are similar in size to many organic molecules, which is of practical advantage, controlling the ingress and egress of reactants and products, resulting in special selectivity in catalytic reactions. Crystalline silicates with MFI structure have a bidirectional intersecting pore system with pore sizes of linear channels along

[0010] : 0.53-0.56 nm and sinusoidal channels along

[0100] : 0.51-0.55 nm. Crystalline aluminosilicates with the MEL structure have a bidirectional intersecting linear pore system in which the linear channels along

[0100] have pore diameters of 0.53-0.54 nm.

[0121] The molecular sieve (H + or NH 4 + The crystalline form (H) advantageously has an initial Si / Al ratio of 4 to 500, preferably 4 to 100, or more preferably 4 to 30. + or NH 4 + The conversion to the form is known per se and is described in US Pat. Nos. 3,911,041 and 5,573,990. The Si / Al atomic ratio is determined by chemical analysis, for example by NMR. It includes only Al which is part of the framework structure of the molecular sieve.

[0122] According to a first embodiment, the zeolite is a phosphorus-modified zeolite produced by a process comprising, in the indicated order: - selection of a molecular sieve as set out in the list above; - introduction of P under conditions effective to advantageously introduce at least 0.05 wt. % P by adding an aqueous solution containing a phosphorus precursor; - Possible separation of solids from aqueous liquids, if possible; - any washing or any drying or any drying followed by washing; - Baking.

[0123] Optionally, the process for producing said phosphorus modified zeolite includes a steam heat treatment and leaching step. The method consists of steaming / heat treatment followed by leaching.

[0124] Generally, it is known to those skilled in the art that as a result of steam treatment of zeolites, aluminum leaves the zeolite framework and is retained as aluminum oxide inside and outside the zeolite pores, this transformation is known as dealumination of the zeolite, and this term will be used throughout this specification.

[0125] In the steaming step, the temperature is preferably between 400°C and 870°C, more preferably between 480°C and 760°C. The pressure is preferably atmospheric and the partial pressure of water may be in the range of 13 kPa to 100 kPa. The steam atmosphere preferably contains 5% to 100% by volume of steam together with 0% to 95% by volume of an inert gas, preferably nitrogen. The steaming 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. The steaming has a tendency to reduce the amount of tetrahedral aluminum in the crystalline silicate framework by forming alumina.

[0126] Treatment of the steamed zeolite with an acid solution results in the dissolution of the extra-framework aluminum oxide. This transformation is known as leaching, and this term will be used throughout the specification. Leaching can be carried out with organic acids 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, hydroxyethylenediaminetriacetic acid, ethylenediaminetetraacetic acid, trichloroacetic acid, trifluoroacetic acid, or salts of such acids (e.g. sodium salts) or mixtures of two or more such acids or salts. Other inorganic acids may include inorganic acids such as nitric acid, hydrochloric acid, methanesulfonic acid, phosphoric acid, phosphonic acid, sulfuric acid, or salts of such acids (e.g. sodium or ammonium salts) or mixtures of two or more such acids or salts. 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 acid solution is advantageously between 2 and 10 litres per kg of molecular sieve. Typical leaching durations are around 0.5 to 24 hours. Advantageously, the aqueous acid solution containing the P source in the leaching step has a pH of 3, advantageously of 2 or less. Advantageously, said aqueous acid solution is a solution of phosphorous acid, a mixture of phosphorous acid and an organic or inorganic acid or a mixture of a salt of phosphorous acid and an organic or inorganic acid. Phosphorous acid or the corresponding salts are phosphates (basic [PO 4 ] 3- ), phosphites (dibasic [HPO 3 ] 2- ) or hypophosphite (monobasic [H 2 PO 2 ] -). Surprisingly, a larger amount of phosphorus remains in the solid molecular sieve material than would be expected based on the typical pore volume in the molecular sieve and assuming that the molecular sieve pores are filled with the phosphorous acid solution used. Two factors, dealumination and P retention, stabilize the framework aluminum in the zeolite framework, 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.

[0127] The zeolite is then separated, advantageously by filtration, and optionally washed. A drying step may be envisaged between the filtration and washing steps. The solution after washing may be separated, by way of example, by filtration from the solid or evaporated. P may be introduced by any means or, by way of example, according to the recipes described in US Pat. No. 3,911,041, US Pat. No. 5,573,990 and US Pat. No. 6,797,851. The separation of the liquid from the solid is advantageously carried out by filtration at a temperature between 0° C. and 90° C., centrifugation at a temperature between 0° C. and 90° C., evaporation or equivalent. Optionally, the zeolite may be dried after separation and before washing. Advantageously, said drying is carried out at a temperature between 40° C. and 60° C., advantageously for a period between 1 hour and 10 hours. This drying may be carried out either under static conditions or in a gas stream. Air, nitrogen or any inert gas may be used. The washing step may be carried out during filtration (separation step) with a portion of cold water (<40° C.) or hot water (>40° C., but <90° C.), or the solid may 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 between 400° C. and 700° C., under static conditions or in a gas stream. Air, nitrogen or any inert gas can be used.

[0128] According to one embodiment of the present invention, the phosphorus-modified zeolite is produced by a process comprising, in the order listed: - selection of a molecular sieve as set out in the list above; - steam heat treatment at temperatures from 400°C to 870°C for 0.01 to 200 hours; - leaching with an aqueous acid solution under conditions effective to remove a substantial portion of the Al from the zeolite; - introduction of P using an aqueous solution containing a P source under conditions effective to advantageously introduce at least 0.05% by weight of P; - Separation of solids from liquids; - an optional washing step or an optional drying step or an optional drying step followed by a washing step; - Baking step.

[0129] Optionally, there is an intermediate step between the steam heat treatment step and the leaching step, such as contact with silica powder and drying.

[0130] Advantageously, the final P content is at least 0.05% by weight and preferably between 0.3 and 7% by weight. Advantageously, at least 10% of Al, relative to the parent zeolites MFI, MEL, FER, MOR and clinoptilolite, has been extracted and removed from the zeolite by leaching. The zeolite is then separated from the wash solution or dried without separation from the wash solution. Said separation is advantageously achieved by filtration. The zeolite is then calcined, for example at 400° C. for 2 to 10 hours.

[0131] The residual P content is adjusted by the P concentration in the aqueous acid solution containing the P source, the drying conditions and, if any, the washing procedure. A drying step may be envisaged between the filtration and washing steps.

[0132] The catalyst consisting of phosphorus-modified zeolite can be the phosphorus-modified zeolite itself, or it can be a phosphorus-modified zeolite that has been formulated into a catalyst by combining it with other materials that provide additional hardness or catalytic activity to the finished catalyst product.

[0133] According to a second embodiment, the catalyst of the process is a composite catalyst produced by a process comprising the following steps: a) selection of a molecular sieve from the list defined above, b) contacting the molecular sieve with a metal silicate comprising at least one alkaline earth metal such that the composite material contains at least 0.1 wt.% silicate.

[0134] The molecular sieve is preferably contacted with the metal silicate by one of the following two methods: - during the catalyst preparation step, mechanically mixing the metal silicate and the molecular sieve which form the precursor to be used in the preparation step; - Physically mixing a preformulated metal silicate and a preformulated molecular sieve in situ in the reaction medium intended for use in carrying out the conversion.

[0135] The molecular sieve and / or the composite catalyst containing the molecular sieve and metal silicate can be post-treated by calcination, reduction or hydrothermal steam treatment. When zeolite is used as the molecular sieve component, phosphorus can be introduced before, simultaneously with or after mixing with the metal silicate.

[0136] In one particular embodiment of the present invention, the molecular sieve may be modified either before or after the introduction of the metal silicate. Preferably, the molecular sieve is modified in some form prior to the introduction of the metal silicate. The term "modified" is used herein to indicate that the molecular sieve may be steam heat treated, leached (e.g. acid leached), washed, dried, calcined, impregnated or subjected to some form of ion exchange. This means that at least a portion of the cations originally contained in the crystal structure may be replaced by a wide variety of other cations using techniques well known in the art. The replacement cations may be hydrogen, ammonium or other metal cations, including mixtures of such cations.

[0137] The selected molecular sieve is then formulated into a composite catalyst containing at least 10% by weight of the molecular sieve described herein and at least one metal silicate containing at least one alkaline earth metal, such that the composite contains at least 0.1% by weight of the silicate.

[0138] At least one of the metal silicates contained in the composite catalyst contains at least one alkaline earth metal, preferably Ca. Metal silicates are insoluble in water, and alkaline earth metal ions, particularly calcium, are multivalent and have a large radius in the hydrated state. Therefore, without intending to be bound by theory, it is believed that the ion exchange reaction with the molecular sieve occurs very slowly, since the alkaline earth metal ions must lose many of their strongly coordinated water molecules to enter the pores of the sieve. As a result, the alkaline earth ions only expose the acid sites located on the outer surface of the molecular sieve, thus increasing the selectivity of the catalyst.

[0139] Furthermore, without intending to be bound by theory, it is believed that the presence of silicate anions further enhances the catalytic properties of the composite catalyst. For example, silicate anions can provide silicon atoms to repair defects in the molecular sieve. This can therefore lead to additional stabilization of the catalyst under harsh hydrothermal conditions.

[0140] As a result, the metal silicate acts as a catalytic promoter.The metal silicate may contain two or more alkaline earth metals selected from Ca, Mg, Sr and Ba.

[0141] The metal silicates may also contain other metals selected from one or more of Ga, Al, Ce, In, Cs, Sc, Sn, Li, Zn, Co, Mo, Mn, Ni, Fe, Cu, Cr, Ti and V. Preferably, the other metals are selected from one or more of Al, Mg, Ce, Co and Zn or mixtures thereof. These bimetallic, trimetallic or multimetallic silicates can be synthesized by any method known in the art. For example, this may be by ion exchange in solution or in the solid state (Labhsetwar et al., Reactivity of Solids, vol. 7, issue 3, 1989, pp. 225-233).

[0142] The silicate anion may be present in any form in the solid metal silicate. Examples include SiO 3 2- , SiO 4 4- , Si 2 O 7 6- , Si 3 O 10 8- etc.

[0143] A preferred catalyst promoter is calcium silicate with a very open and accessible pore structure. An even more preferred catalyst promoter is calcium silicate (molecular formula 6CaO 6SiO 2 H 2 Ca corresponding to the known mineral xonotlite (containing 1,2-dichlorophenyl) 6 S 6 O 17 (OH) 2 The present invention includes a synthetic crystalline calcium silicate hydrate having a chemical composition of:

[0144] Typically, synthetic calcium silicate hydrates are synthesized by hydrothermal processes under autogenous pressure. A particularly preferred synthetic calcium silicate hydrate is commercially available under the brand name Promaxon from Promat, Ratingen, Germany.

[0145] Other examples of metal silicates containing alkaline earth metals include CaAl 2 S 2 O 8 , Ca 2 Al 2 SiO 7 , CaMg(Si 2 O 6 ) x and mixtures thereof.

[0146] Prior to mixing with the molecular sieve, the metal silicate compound may be modified by calcination, steam treatment, ion exchange, impregnation or phosphating. The metal silicate may be in the form of a separate compound or may be part of a mixed compound.

[0147] The metal silicate may be contacted with the molecular sieve in the reaction medium prior to carrying out the conversion by a blending step to simultaneously prepare a mixture of the molecular sieve and the metal silicate or an in situ mixture of separately prepared materials.

[0148] The contacting can be carried out by mechanically mixing the alkaline earth metal-containing metal silicate with the molecular sieve. This can be carried out by any known mixing process. The mixing can be carried out for a time period ranging from 1 minute to 24 hours, preferably from 1 minute to 10 hours.

[0149] If not carried out in an in situ conversion reactor, it may be carried out in a batch mixer or in a continuous process in an extruder, such as a single or twin screw extruder, under vacuum or high pressure at temperatures between 20°C and 300°C. The contacting may be carried out in an aqueous or non-aqueous medium. Prior to the blending step, other compounds that facilitate blending can be added, such as thickeners or polyelectrolytes that enhance the aggregation, dispersion and flow properties of the precursors. In the case of oil drop drying or spray drying, a rather liquid fluid (high water content) is prepared. In a further embodiment, the contacting is carried out in the presence of a phosphorus-containing compound. 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.

[0150] According to a third embodiment, the catalyst of the process is a phosphorus (P) and alkaline earth metal or rare earth metal (M) modified molecular sieve (MP modified molecular sieve) produced by a process comprising the following steps: a)- P-modified molecular sieves containing at least 0.3% by weight of P; - a P-modified molecular sieve before or during step b) of introducing at least 0.3% by weight of P, selecting at least one molecular sieve selected from b) contacting said molecular sieve with a compound containing an alkaline earth metal or a rare earth metal (M-containing compound) to introduce at least 0.05% by weight of alkaline earth metal or rare earth metal M.

[0151] Optionally, the steps of contacting the P-containing compound and the M-containing compound with the molecular sieve can be performed simultaneously.

[0152] The introduction of alkaline earth or rare earth metals (M) is accomplished by contacting the molecular sieve with a solution of one or more M-containing compounds, which may contain a higher concentration of alkaline earth or rare earth metals than that found in the final MP-modified molecular sieve.

[0153] The molecular sieve is selected from the list previously set forth above.

[0154] Prior to P-modification and / or alkaline earth metal or rare earth metal modification (M-modification), the molecular sieve may undergo other treatments, including steam heat treatment, leaching (e.g. acid leaching), washing, drying, calcination, impregnation or ion exchange. Additionally or as a variant, these steps can be carried out during or after the P-modification. The term "ion exchange step" is understood herein to mean that at least some of the cations originally contained in the crystal structure are replaced by a wide variety of other cations according to techniques well known in the art. The replacement cations may be hydrogen, ammonium or other metal cations, including mixtures of such cations.

[0155] The modification of molecular sieves with phosphorus is known per se. This modification is carried out by treating the molecular sieve with P in aqueous or non-aqueous medium, by chemical vapour deposition in the vapour phase of organic P compounds or by impregnation. The catalyst may be preformulated with or without a binder. Preferred P compounds typically used for this purpose are phosphoric acid, NH 4 H 2 PO 4 or (NH 4 ) 2 HPO 4 The M-containing compounds may be selected from the group consisting of organic compounds, salts, hydroxides and oxides. These compounds may contain phosphorus. It is essential that these compounds are present in a solubilized form before being contacted with the molecular sieve or by forming a solution upon contact with the molecular sieve.

[0156] The final M / P molar ratio in the MP molecular sieve is preferably less than 1.

[0157] According to one particular embodiment of the invention, the molecular sieve may be modified with phosphorus by a process comprising the following steps, in the order indicated: - Steam heat treatment of the molecular sieve at temperatures from 400℃ to 870℃ for 0.01 to 200 hours; - leaching with an acidic aqueous solution containing a P source under conditions effective to remove a substantial portion of the Al from the molecular sieve and to introduce at least 0.3% phosphorus by weight of the molecular sieve; at this time, additional modification can be carried out according to the following steps in the order presented: - Separation of solids from liquids; - an optional washing step or an optional drying step or an optional drying step followed by a washing step; - Baking step.

[0158] Preferably, the separation, optional washing and drying steps and calcination are carried out after the introduction of the M-containing compound into the molecular sieve. The metal M can be any alkaline earth metal 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.

[0159] Advantageously, the final P content of the molecular sieve is at least 0.3% by weight, preferably between 0.3% and 7% by weight. Advantageously, at least 10% by weight 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, the separation conditions during separation of the solid from the liquid, and / or by any washing procedure, in which impregnation and / or adsorption may also take place. Drying steps can be envisaged between the separation and / or washing steps.

[0160] The molecular sieve is then separated from the washing solution or dried without separation from the washing solution, preferably by filtration, and then calcined, for example at 400° C. for 2 to 10 hours.

[0161] 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, where the sieve is dealuminated by steam heat treatment and then leached with a P-containing acid 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 phosphorus compound has been added and the P-modification has taken place, and before the separation step.

[0162] However, the incorporation of M into the molecular sieve may be contemplated as follows: - during the leaching step; - before the cleaning step, but after leaching and drying; - calcined molecular sieve contacted with P; - molecular sieves that have not been leached with the aim of introducing P but have been brought into contact with P during the washing step.

[0163] The incorporation of M in the molecular sieve may be carried out by impregnation or by adsorption from an aqueous solution of the M-containing compound.

[0164] The M-containing compound may be introduced at a temperature ranging from ambient temperature to the boiling point of the solution. The concentration of the M-containing compound in the solution is at least 0.05 M, preferably 0.05-1.0 M. The amount of alkaline earth metal or rare earth metal (M) in the MP molecular sieve may vary from at least 0.05% by weight, preferably 0.05% to 7% by weight, and better still 0.1% to 4% by weight.

[0165] Prior to preparation of the composite catalyst, the molecular sieve may undergo further treatment processes, including steps such as steaming, leaching (e.g., acid leaching), washing, drying, calcination, impregnation, and ion exchange. Additionally, or as a variation, these steps may be performed after preparation of the catalytic composite.

[0166] The alkaline earth metal or rare earth metal 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. In particular, in the case of P modification by evaporation and leaching, M can be a rare earth metal such as La and Ce.

[0167] 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 when contacted with the molecular sieve. Alternatively, a solution of the M-containing compound may be formed after the molecular sieve is contacted with the compound.

[0168] Possible M-containing compounds include compounds of metal M such as sulfates, formates, nitrates, M metal acetates, halides, oxyhalides, borates, carbonates, hydroxides, oxides, and mixtures thereof, which may be, for example, calcium sulfate, calcium formate, calcium nitrate, calcium acetate, calcium halides, calcium oxyhalides, calcium borates, calcium carbonate, calcium hydroxide, calcium oxide, and mixtures thereof.

[0169] The M-containing compound may also include 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 compound may further include phosphorus.

[0170] These poorly water-soluble M-containing compounds can be dissolved to form a fully solubilized solution by heating and / or modifying the pH of the solution by adding phosphoric, acetic or nitric acids or the corresponding ammonium acids, salts of said acids, the concentration of the M-containing compound being at least 0.05M.

[0171] The alkaline earth and rare earth metals M, especially Ca, have large hydration sphere radii in the hydrated state. Thus, without intending to be bound by theory, it is believed that ion exchange reactions with the acid sites located inside the micropore structure of the molecular sieve occur very slowly. As a result, the selected M metals only expose the acid sites located on the external surface of the molecular sieve, thus increasing the selectivity of the catalyst.

[0172] In the case of P-modified molecular sieves, M modification leads to the formation of mixed M-Al phosphates on the external surface. Given that phosphorus binds more strongly to alkaline or rare earth metals M than to Al, this modification leads to the stabilization of phosphorus on the external surface of the molecular sieve where it is most unstable. However, it is essential that all M atoms on the external surface are saturated with phosphorus. This can be ensured in the presence of excess phosphorus, for example by the presence of M in solution form, which helps to wash away the excess phosphorus and prevent clogging of the micropore entrances.

[0173] Optionally, the molecular sieve, which may or may not be MP-modified, may be mixed with other components before, after, or simultaneously with the compounding step to form the composite. In a particular embodiment, the molecular sieve, either MP-modified or unmodified, may be combined with other materials that impart additional hardness or catalytic activity to the finished catalyst product. Materials that may be mixed with the molecular sieve may be various inert or catalytically active matrix materials and / or various binder materials. Such materials include clay, quartz, alumina or alumina sol, silica or silica sol, and / or metal oxides, such as titanium oxide, zirconia, and mixtures thereof. In an embodiment, some binder materials may also be used as diluents to control the conversion of product feed, thus increasing selectivity. According to an embodiment, the binder also improves the attrition of the catalyst under industrial operating conditions. Natural clays that can be used as binders are, for example, clays from the kaolin group or the montmorillonite family. Such clays may be used in their raw state as mined or may be subjected to various treatments such as calcination, acid treatment or chemical modification prior to use. In addition to those mentioned above, other materials that may be incorporated into the composite catalyst of the present 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, anhydrous dicalcium phosphate, alpha- or beta-tricalcium phosphate, octacalcium phosphate, hydroxyapatite, etc.: zirconia, silica-thoria, silica-beryllium, silica-titanium, calcium-alumina. Examples of ternary binders include, for example, calcium-silica-alumina or silica-alumina-zirconia. These components are effective in increasing the catalyst density and increasing the strength of the catalyst formulation.The catalysts usable in fluidized bed reactors generally have a substantially spherical shape formed by atomization.

[0174] Generally, when using a fluidized bed as the reactor, the size of the catalyst particles may vary from about 20 μm to 500 μm, more preferably from 30 μm to 100 μm. The size of the molecular sieve crystals contained within 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 altogether. Generally, when using a fixed bed, the size of the catalyst particles may vary from approximately 0.5 to 5 mm in the form of beads, cylinders, or extrudates (length 1 to 10 mm) in either trilobe or quadrilobate shapes. The amount of molecular sieve contained in the final catalyst composite is in the range of 10% to 90% by weight, preferably 20% to 70% by weight of the total catalyst composite.

[0175] According to another embodiment, the unmodified molecular sieve is first formulated with a binder and a matrix material and then modified with phosphosilicate and alkaline earth metal silicate. According to another particular embodiment, the molecular sieve is optionally dealuminated and then modified with phosphorus during the formulation step. The alkaline earth metal silicate may be introduced during the formulation step or onto the formulated solid.

[0176] According to a preferred embodiment, the molecular sieve is first optionally dealuminated and modified with phosphorus and then prepared. The introduction of the metal is carried out simultaneously with the phosphorus modification step and / or with the already prepared catalyst.

[0177] After preparation, the composite catalyst may undergo further processing, including further steps of steaming, leaching, washing, drying, calcination, impregnation and ion exchange. If the molecular sieve is not modified with phosphorus prior to the preparation step of preparing the mixture, i.e. the step of introducing the metal silicate into the molecular sieve, this may be performed after said step. According to a particular 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.

[0178] According to another embodiment, the unmodified molecular sieve is first prepared with a binder and a matrix material and then modified with phosphorus and a metal. According to a particular embodiment, the molecular sieve is optionally dealuminated and then modified with phosphorus during the preparation step. The introduction of the metal may be carried out during the preparation step or on the prepared solid. According to a preferred embodiment, the molecular sieve is optionally first dealuminated and modified with phosphorus and then prepared. The introduction of the metal may be carried out simultaneously with the phosphorus modification step and / or with the prepared catalyst.

[0179] The final catalyst containing the phosphorus-modified zeolite advantageously has an 27Al NMR signature between 35 ppm and 45 ppm, which is characteristic for the presence of the ALPO structure. The mass content of said AlPO4 structure in the catalyst can be up to 99% by mass, advantageously between 10% and 98% by mass.

[0180] The presence of this ALPO structure is characterized using the following method illustrated in FIG. 9. Measurements are performed by magic angle spinning (MAS) in solid-state nuclear magnetic resonance (NMR) spectroscopy performed on a Bruker Avance 500 spectrometer using a 4 mm zirconia MAS probe at a spinning speed of 15 kHz. To obtain quantitative MAS spectra, a single pulse excitation was applied using a short pulse length of 0.6 μsec. Each spectrum is the result of 5000 scans separated by a delay of 0.5 seconds. The chemical shifts of the 27Al spectra were calculated using the formula: AlCl 3The solution was checked against (0.1M, (0 ppm)).

[0181] When only one zeolitic aluminum source is present in the catalyst, AlPO 4 The phase content is directly estimated by the ratio of the signal area between 35 ppm and 45 ppm in the 27Al MAS (centered at 39 ppm in Figure 9) to the total spectral area between -50 ppm and 100 ppm.

[0182] If the binder contains aluminum and phosphorus, AlPO in the zeolite 4 The phase content is estimated by the ratio of the signal area at 35 ppm to 45 ppm in 27Al MAS to the total spectral area from -50 ppm to 100 ppm after subtraction of the binder signal intensity.

[0183] In order to enhance the selectivity towards the formation of aromatic compounds, the above-mentioned catalysts may be further modified by the addition of one or more metals selected from among metals in group IIB (e.g. Zn), metals in group IIIB (e.g. Ga), transition metals in group VIIIB (e.g. Fe and / or Ni and / or Pt), metals in group VIB (e.g. Mo), metals in group IB (e.g. Cu and / or Ag), and even metals from the lanthanide group (e.g. La). The introduction of the metals can be carried out 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 metals is carried out using one or more solutions containing the metals in the form of a salt. The metal salts are first pre-dissolved in the treatment solution; the metal counterions are selected from among sulfates, nitrates, carbonates, hydroxides, phosphates, carboxylates (e.g. formates, acetates, propionates), and dicarboxylates (e.g. oxalates, malonates, succinates).

[0184] After introduction of the metal, a variety of different treatments may be applied including drying, calcining, and steam heat treatment.

[0185] The metal content is generally between 0% and 5% by weight, preferably between 0% and 2.5% by weight. The selected metal is typically Ga in the presence / absence of other metals such as Pt. 3+ or Zn 2+ It could be.

[0186] The above mentioned catalysts may be modified by the addition of B in an amount of 0.1% to 5% by weight, preferably 0.1% to 1% by weight, more preferably 0.1% to 0.5% by weight, in the presence or absence of the other above mentioned metals.

[0187] In another embodiment, the conversion process for converting oxygenates is carried out with a catalyst comprising a zeolite having a pore size of 10 oxygen atoms (10-MR) or greater that has been modified by adding B either before, after or simultaneously with the preparation step of preparing the final catalyst. Suitably, for example, the catalyst may be B-modified ZSM-5.

[0188] The B content in the final catalyst is 0.1% by mass to 5% by mass, preferably 0.1% by mass to 1% by mass, and more preferably 0.1% by mass to 0.5% by mass.

[0189] Advantageously, the zeolite contained in the final catalyst has a Si / Al atomic ratio, determined by chemical analysis (e.g. NMR) taking into account only the Al that is part of the framework structure of the molecular sieve, of between 4 and 500, preferably between 5 and 200 or more preferably between 12 and 150.

[0190] As regards the conversion stage in which the conversion step (a) is carried out, the C1-C6 alcohol stream is contacted with the above-mentioned catalyst in the reaction zone of at least one reactor under operating conditions leading to the production of a mixture containing paraffins, olefins, aromatics and water, as defined above.

[0191] In this step (a) of converting the alcohol, the mixture can generally be produced within a temperature range of 300°C to 600°C, in particular 330°C to 550°C, particularly 350°C to 500°C, or 410°C to 580°C.

[0192] The pressure may also vary over a wide range. The preferred pressure is within the range of about 100 kPa to about 5 MPa, with the most preferred range being about 150 kPa to about 1.0 MPa. The pressures referred to above refer to the partial pressure of the oxygen-containing organic compound.

[0193] The conversion step (a) can be carried out in a single reaction zone or in multiple reaction zones arranged in series or parallel.After a period of operation the catalyst should be regenerated.

[0194] In particular, multiple reactors can be used to ensure that the exothermicity of the reaction is controlled in such a way as to avoid excessive temperatures. Preferably, the maximum temperature difference within the same reactor should not exceed 100°C, preferably 75°C.

[0195] Regarding the type of reactor, either isothermal or adiabatic fixed bed, moving bed and / or fluidized bed reactors may be used.

[0196] The conversion reaction may be carried out in a continuous mode in a configuration comprising a series of fixed beds connected in series, in at least one operating reactor through which raw feed is passed from one fixed bed to the other with cooling in between; and in at least one similar reactor connected in parallel undergoing a catalyst regeneration operation.

[0197] Advantageously, an additional C2-C6 alcohol stream is added between the two fixed beds to control exothermicity, as described above.

[0198] The conversion reaction may be carried out in a continuous mode in a configuration including a series of moving beds connected in series, in which the raw feed passes from one moving bed to the other with cooling in between, and the catalyst is movable and circulates between the reactor and a catalyst regeneration zone.

[0199] The conversion reaction may be carried out in a continuous mode in a configuration comprising one fluidized bed forming a reaction zone where reaction occurs, and one fluidized bed forming a regeneration zone where regeneration occurs (e.g., by controlled combustion in the presence of oxygen); or, alternatively, serially connected fluidized beds, with raw feed passing from one bed to the other with cooling in between, and catalyst movingly circulating between the reactor and the catalyst regeneration zone.

[0200] Fluidized beds offer significant advantages, especially when the reaction is highly exothermic. Once the bed solids are fluidized, the solids within the bed behave like a liquid. The size, shape, formation, rise rate and coalescence of gas bubbles in a fluidized bed are quantitatively similar to those of gas bubbles in a liquid.

[0201] Thus, the fluidized bed behaves like a liquid, providing the ability to handle solids like a fluid, thus allowing the solids to be fed and / or extracted. The severe mixing in the fluidized bed also allows for uniform temperature to be achieved, even for highly exothermic reactions, thus allowing for more versatile control of the reactor. The severe mixing also improves the contact between the solids and the fluid, thus increasing heat and mass transfer.

[0202] There are numerous fluidized bed variations described in available technical manuals (e.g. Handbook of fluidization and fluid-particle systems, Taylor & Francis Group LLC, 2003). The fluidization phenomenon of gas-solid systems depends greatly on the type of powder used. There are several classifications, all based on Geldart's own research work. Most catalysts used in fluidized bed systems are group A particles, characterized by dense phase expansion after minimal fluidization and before the onset of bubbling. Bubbles appear at the lowest bubbling rates.

[0203] Fluidization regimes can be broadly divided into two categories: fluidization in particulate mode (smooth) and agglomerated mode (bubbly). In particulate fluidization, the solid particles are generally dispersed in a relatively uniform manner in the fluidization medium, without easily identifiable bubbles. Particulate fluidization is therefore sometimes also called homogeneous fluidization. In heterogeneous or agglomerated fluidization, interstices (bubbles) that are completely free of solids are generally formed and are found in bubbling fluidized beds or in beds that exhibit "slugging". For gas-solid systems, there are many distinct fluidization regimes: fixed bed, particulate fluidization, bubbling fluidization, slugging fluidization and turbulent fluidization; for each of these, relevant criteria are available. Additional fluidization regimes are possible when the operating velocity is higher than the transport velocity, and thus recycling of entrained particles is necessary to maintain the bed.

[0204] Fine particle regime: Umf≦U <Umb For powders in Group A, the fixed bed expands homogeneously (fine particle fluidization) above the minimum fluidization velocity (Umf) and no foaming is observed as long as the velocity remains below the minimum foaming velocity (Umb).

[0205] Foaming regime: Umb≦U <Ums Foam appears when the gas velocity increases above the minimum bubbling velocity (Umb). Bubbles form, coalesce and grow above the distributor. The bubbling regime is characterized by the coexistence of a foam phase and a dense / emulsion phase. Most of the fluidizing gas is present in the form of bubbles and, as a result, the gas velocity through the dense phase is very low.

[0206] Slugging regime: Ums≦U <Uc Considering the high bed height to diameter ratio, the bed provides ample time for the bubbles to coalesce into larger bubbles. When the bubbles reach approximately the size of the cross section of the bed, the bed enters a "slugging" regime with periodic passage of larger bubbles forming plugs and large regular fluctuations in pressure drop across the bed. The velocity Uc corresponds to the operating condition of the bed where the plug reaches its maximum diameter and the amplitude of the pressure fluctuations is maximum.

[0207] Transition to the turbulent regime: Uc≦U <Uk When the gas velocity continues to increase beyond this velocity Uc, the large bubbles begin to break up into smaller bubbles with smaller pressure fluctuations. This velocity is denoted as Uk and characterizes the transition between the bubbling and turbulent regimes.

[0208] Turbulence regime: Uk≦U <Utr Up to the transport velocity (Utr), the bed is in the turbulent regime. Bubbles or voids continue to exist, but these are not as distinct as in dense suspensions. In this regime, the interactions between the gas voids and the dense / emulsion phase are vigorous, ensuring effective gas-solid contact.

[0209] Fast fluidization regime: U>Utr Above the transport velocity (Utr), particles start to become entrained and it is no longer possible to operate continuously without replacing or recycling the entrained and transported particles. Fast fluidized beds are typically characterized by the coexistence of a dense phase at the bottom near the distributor with a dilute phase region at the top. The velocity of the particles increases with height within the bed, thus decreasing the density of the bed.

[0210] Air transport: U>>Utr All particles introduced at the bottom of the fluidized bed are transported into the dilute phase with the concentration varying along the height of the bed.

[0211] A typical example of a reaction zone is a fluidized bed with a riser used in fluid catalytic cracking (FCC) applications. A riser is a vertical pipe with a high height-to-diameter ratio (greater than 10), and an ideal riser would approximate plug flow conditions, such that the catalyst and fluid phases pass through it with minimal backmixing.

[0212] In transport fluidized bed reactors (either fast fluidization or pneumatic transport), a core annular flow can occur, where a lean fast core is surrounded by a denser slow ring. When the circulating mass flow rate is low, the solids in the ring flow downward across the wall. When the circulating mass flow rate is high, the solids in the ring flow upward along the wall. This non-uniform flow phenomenon leads to inefficient gas-solid contact and suboptimal catalyst performance, and significant backmixing of gas and solids occurs, especially when there is a downward flow down in the wall region. For fast fluidization, the interior is used to redistribute the axial and radial gas-solid flow structures, i.e. to enhance the uniformity of the gas-solid flow structures in the space and thus promote radial gas-solid exchange. In fluidized transport reactors, it is necessary to recirculate the catalyst particles to the bottom of the reactor. This offers the possibility to control the density of the catalyst in the fluidized bed by recirculating more or less catalyst.

[0213] At the bottom of the fluidized bed, the make-up fluid is distributed homogeneously across the cross section of the reactor vessel. At the top of the reaction zone, the reaction vapors are separated from the entrained catalyst using baffles, disengagement zones and cyclones. The catalyst is collected, stripped of residual hydrocarbons and returned to the bottom of the fluidized bed zone, advantageously via a vertical manifold ("standpipe") and valves.

[0214] For exothermic reactions such as the conversion carried out in step (a), it is preferable to have a homogeneous temperature (radially and axially) across the catalyst bed in order to avoid hot spots and to properly control the catalytic reaction. This can be achieved by rapid recirculation and possibly remixing of the catalyst within the reaction vessel.

[0215] Means for controlling the average temperature of the reaction consist of introducing a feed into the reaction zone at a temperature lower than the average bed temperature, and / or removing heat from the catalyst bed via heat exchange, which can be accomplished by internal heat exchange tubes through which a cooling medium is circulated to remove heat from the reaction vessel, or external heat exchange by circulating hot catalyst collected at the top of the reactor around the heat exchange tubes and recirculating the cooled catalyst inside the reaction vessel.

[0216] Regarding catalyst regeneration, conversion reactors for converting C1-C6 alcohols also contain a regeneration zone (regenerator) whose main purpose is to remove coke deposits on the catalyst by combustion with oxygen. The regenerator is a fast fluidized bed system. Typically, the regenerator contains a dense catalyst bed at the bottom of the vessel and a leaner bed near the top of the vessel.

[0217] There are two types of regenerators that operate in either partial or complete burn mode. In partial burn mode, less than the stoichiometric amount of air is fed to the regenerator. The carbon is mostly converted to carbon monoxide and only a portion is converted to carbon dioxide. Ideally, all the oxygen is consumed so that there is no oxygen in the flue gas. The CO / CO ratio in the flue gas 2 The ratio is generally between 0.5 and 2.0. In the complete combustion mode, excess air is fed to the regenerator. Ideally, all the carbon contained in the coke is converted to carbon dioxide and no carbon monoxide is present in the flue gas. The residual oxygen content in the flue gas is between 1.0% and 3.0% by volume on an anhydrous basis.

[0218] Partial combustion regenerators offer several advantages over full combustion regenerators, especially when the catalyst is sensitive to high temperature and steam environments: (i) less air is required than stoichiometric, allowing more coke to be burned for a given air flow rate, and (ii) less heat of combustion is released, thus allowing for better temperature control and better maintenance of catalyst activity in the presence of steam generated by hydrogen combustion.

[0219] A potential drawback of the partial combustion regenerator is that a larger amount of coke remains on the regenerated catalyst. In the case of regeneration by complete combustion, the amount of carbon remaining on the catalyst is smaller and the degree of recovery of catalytic activity is higher. A potential drawback of the complete combustion regenerator is that due to the complete combustion reaction, more heat is released, thus resulting in more irreversible loss of catalytic activity. The use of two-stage regeneration can reduce catalyst deactivation. In two-stage regeneration, the first stage operates at a moderate temperature to mainly burn off the hydrogen present in the coke, as well as a portion of the carbon, which has a higher reaction rate. In the second stage, with the use of excess air, the remaining carbon is burned at a higher temperature to produce carbon dioxide, and the absence of water vapor in the second stage regenerator can minimize catalyst deactivation at high temperatures.

[0220] The use of a fluidized bed offers the ability to very precisely control the exotherm of the reaction while providing continuous catalyst regeneration which facilitates manufacturability and simplifies operation.

[0221] One or more diluents may be present in the C1-C6 alcohol stream that supplies the reaction zone in an amount of, for example, 1 mole % to 95 mole %, based on the total moles of all supply and diluent components introduced into the reaction zone.

[0222] Typical diluents include, but are not limited to, helium, argon, nitrogen, hydrogen, water (optionally recycled), paraffins, alkanes (specifically methane, ethane and propane), aromatics and mixtures thereof. Preferred diluents are water and nitrogen. Water may be injected in liquid or vapor form. The use of a diluent can provide two advantages. The first advantage is to reduce the partial pressure of the alcohol, increasing the selectivity for light olefins, primarily propylene. In general, the lower the partial pressure of the alcohol, the higher the selectivity for light olefins, and conversely, the higher the partial pressure, the higher the selectivity for heavy olefins such as butenes and pentenes. With regard to the yield of light olefins, there is an optimum yield depending on the partial pressure, reaction temperature, make-up space velocity and catalyst properties.

[0223] A second advantage of using a diluent is that it can act as a heat sink for the conversion of exothermic alcohols. Thus, the higher the specific molar heat capacity, the more heat the diluent can absorb. This second advantage may be less significant in the case of fluidized bed reactors, as these have proven to be excellent reactors operating at a nearly homogeneous temperature throughout the catalyst bed. The diluent is preferably easily separated from the light olefin product, preferably by simple phase separation. Thus, water is the preferred diluent. The diluent may be added in a proportion of 1 mol % to 95 mol %, preferably 10 mol % to 75 mol %, of the combined feed (C1-C6 alcohol stream + diluent).

[0224] Step (b) Separating water from the mixture Water from the mixture containing paraffins, olefins, aromatics and water produced in the conversion step (a) is separated from the mixture in a water separation stage to form a water-depleted mixture.

[0225] The water separation step is preceded by a cooling step for cooling the effluent from step (a) which condenses the water as well as a portion of the hydrocarbons. The temperature at which this step is carried out is generally between 20°C and 100°C.

[0226] The separation is based, for example, on the density difference and solubility between the water and the remainder of the hydrocarbons. It is generally carried out in a three-phase separation flask, which serves to separate the water-rich aqueous phase (hereafter referred to as water separated from the mixture), the liquid hydrocarbon phase and the gaseous hydrocarbon phase.

[0227] The water-deficient mixture contains less than 5% by weight, preferably less than 1% by weight, of the water present in the mixture produced in step (a).

[0228] The water separated from the mixture advantageously contains less than 10% by weight of hydrocarbons.

[0229] Water separated from the mixture is optionally at least partially recycled to the conversion step (a).

[0230] When the conversion step (a) is carried out using several successive fixed catalyst beds, the water separated from the mixture is optionally recycled upstream of a fixed catalyst bed or between two fixed catalyst beds.

[0231] In this case, the mass ratio of recycled water to the C1-C6 alcohol stream in the feed supplied to the conversion step (a) is advantageously between 0 and 3, preferably between 0.05 and 2.

[0232] When the converting step (a) is carried out using at least one fluidized bed, the recycled water is optionally reinjected into the fluidized bed.

[0233] In this case, the mass ratio of recycled water to the C1-C6 alcohol stream in the feed supplied to the conversion step (a) is advantageously between 0 and 1, preferably between 0.05 and 0.5.

[0234] In both the above cases, the recycle water forms steam which controls the exothermicity of the reaction, reduces the hydrocarbon partial pressure, and modifies the acidity of the catalyst, thereby enhancing the olefin selectivity.

[0235] The separated water which is not recycled to the conversion step (a) is advantageously treated by stripping in a stripping column in order to separate the hydrocarbons contained therein into an extracted hydrocarbon stream which is reinjected into the separation step (b).

[0236] When the conversion step (a) is carried out using at least one fluidized bed, a C4-hydrocarbon stream can be optionally added to the fluidized bed to supplement or replace the recycled water, for example formed by at least a portion of the C1-C2 hydrocarbon fraction separated from the water-depleted mixture, as discussed below, and which also controls the exothermicity of the reaction.

[0237] Advantageously, the C1-C6 alcohol stream is introduced in the conversion step (a) at a temperature at least 5° C. higher than the bubble point of the C1-C6 alcohol stream and preferably lower than the temperature of the conversion reaction carried out in step (a), such as 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.

[0238] Heating this stream absorbs the calories released by the conversion of alcohol.

[0239] Optional step for separation of C1-C2 and C3 hydrocarbons

[0240] Advantageously, the water-depleted mixture is separated from the remainder of the water-depleted mixture by the addition of C1-C2 hydrocarbons (methane, ethane, ethylene) and CO, CO 2 and other gas molecules lighter than C2, such as hydrogen, into a separation stage that includes at least one distillation column (hereafter referred to as a deethanizer).

[0241] The distillation column may for example operate at a pressure in excess of 20 barg, preferably in excess of 30 barg.

[0242] A C1-C2 hydrocarbon fraction is extracted from the top of the column. This fraction is composed of more than 50% by mass of C1-C2 hydrocarbons and CO, CO 2 and other gas molecules such as hydrogen.

[0243] The C1-C2 hydrocarbon fraction preferably comprises more than 90% by mass, in particular more than 95% by mass, of the C1-C2 hydrocarbons contained in the water-deficient mixture, and CO, CO 2 and other gas molecules such as hydrogen.

[0244] The C1-C2 hydrocarbon fraction is at least partially conveyed to an ethylene recovery unit, for example in a steam cracker, etc. Thus, ethylene can be recovered even if only small amounts are produced during the conversion step (a).

[0245] In the variants in which a carbon dioxide-containing stream is added to the 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.

[0246] These compounds are then advantageously recycled to a preliminary alcohol synthesis step, in particular by fermentation of syngas and catalytic conversion of syngas, in particular to produce methanol.

[0247] In particular, the methanol or ethanol thus produced is then advantageously recycled to form part of the C1-C6 alcohol stream.

[0248] In one advantageous variant, which is particularly applicable in the case of a conversion step (a) carried out using at least one fluidized bed, a part of the C1-C2 hydrocarbon fraction is recycled in the form of a recycle stream to the conversion step (a) for converting the C1-C6 alcohol 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 withdrawn from the distillation column is less than 1, in particular between 0.1 and 0.8.

[0249] A C3+ hydrocarbon fraction is recovered at the bottom of the column, this fraction comprising greater than 90% by weight of the C3+ hydrocarbons contained in the water-depleted mixture.

[0250] On an anhydrous basis, this fraction advantageously contains: - less than 5% by weight of C1-C2 hydrocarbons; - less than 15% by weight of paraffin, in particular between 3% and 10% by weight of paraffin; - more than 40% by mass of C3-C7 olefins, in particular 50% to 80% by mass of C3-C7 olefins; - less than 5% by weight of C8+ olefins, particularly between 0.1% and 4.0% by weight of C8+ olefins; and / or - More than 6% aromatics, specifically 6%-20% aromatics.

[0251] The C3-C7 olefins generally include propylene. In particular, the C3+ hydrocarbon fraction contains, in some cases, more than 30% by weight of propylene.

[0252] In one embodiment, the C3+ hydrocarbon fraction is sent directly to the oligomerization and alkylation steps. As a variant, in one particular embodiment, an additional separation step is carried out in a second distillation column to separate the C3- hydrocarbons, in particular propylene. Said separation allows the recovery of propylene.

[0253] The second distillation column operates, for example, at a pressure above 5 barg, preferably above 10 barg. This column produces at its top a C3-hydrocarbon fraction containing 50% by mass of propylene and at the bottom a C4+hydrocarbon fraction.

[0254] The C3-hydrocarbon fraction preferably contains more than 90% by mass, particularly more than 95% by mass, of the C3-hydrocarbons contained in the C3+hydrocarbon fraction from the first distillation column.

[0255] The C4+ hydrocarbon fraction comprises greater than 90% by mass of the C4+ hydrocarbons contained in the water-depleted mixture.

[0256] On an anhydrous basis, this fraction advantageously contains: - less than 5% by weight of C3-hydrocarbons; - less than 25% by weight of paraffin, in particular between 10% and 15% by weight of paraffin; - more than 15% by weight of C4-C7 olefins, in particular 25% to 40% by weight of C4-C7 olefins; - less than 2% by weight of C8+ olefins, particularly 0.5% to 1.0% by weight of C8+ olefins; or - More than 6% aromatics, specifically 7%-20% aromatics, specifically less than 5% C9+ aromatics.

[0257] Step (c) oligomerization of olefins from the water-depleted mixture; and step (d) alkylation of aromatic compounds from the water-depleted mixture.

[0258] In a first embodiment, the oligomerization step (c) for oligomerizing olefins from the water-depleted mixture and the alkylation step (d) for alkylating aromatics from the water-depleted mixture are carried out jointly in the same one or more reactors of the same given reaction stage.

[0259] Advantageously, the hydrocarbon feedload formed by the above-mentioned C3+ or C4+ hydrocarbon fraction is oligomerized with respect to its olefins and alkylated with respect to its aromatic compounds by contact with an acid catalyst.

[0260] For example, multiple reactor plants can be used, which allow the exothermicity of the reaction to be controlled in a way that avoids excessive temperatures. Preferably, the maximum temperature difference within the same given reactor should not exceed 100°C, preferably 75°C.

[0261] The one or more reactors may be of the following types: isothermal or adiabatic fixed or moving bed reactors. The olefin oligomerization and aromatic alkylation reactions may be carried out in a continuous mode in a configuration with a series of fixed beds connected in series, in at least one reactor in operation, where the raw feed passes from one bed to the other with cooling between the beds; and in at least one similar reactor connected in parallel, where a catalyst regeneration operation is being performed. The olefin oligomerization and aromatic alkylation reactions may be carried out in a continuous mode in a configuration with a series of moving beds connected in series, where the raw feed passes from one bed to the other with cooling between the beds, and where the catalyst is mobile and circulates between the reactor and the catalyst regeneration zone.

[0262] Preferably, the above steps are carried out jointly using at least two reactors in series. In the case of fixed bed reactors, alternatively, a single reactor may contain multiple catalyst beds with a cooling system between the beds or be equipped with quench injection for the purpose of lowering the temperature between the beds.

[0263] The reaction conditions in the first reactor are selected to allow the conversion of a portion of the olefinic compounds with low carbon numbers (C3-C8) to intermediate olefins (C8+) and the alkylation of aromatic compounds with light olefins.

[0264] Advantageously, the first reactor contains a first catalytic zone and operates at an elevated temperature, for example above 200° C., preferably below 350° C., and at a pressure between 25 bar and 60 bar.

[0265] The second reactor is preferably operated at a temperature and pressure selected to promote the conversion of a portion of the low carbon (C3-C8) olefinic compounds to intermediate olefins (C8+) and the alkylation of aromatic compounds with light olefins. The effluent from the first reactor, containing unreacted olefins, intermediate olefins, aromatic compounds, water and possibly other compounds such as paraffins, and possibly reducing gases, can then be oligomerized and / or alkylated in this second reactor, which now contains a second catalytic zone, to obtain a heavier hydrocarbon effluent rich in distillates.

[0266] Advantageously, between two successive reactors, a cooling section and, optionally, a flash drum is provided.

[0267] The mass flow rate through the oligomerization reactor is advantageously sufficient to permit relatively high conversion without being too low to avoid undesirable concurrent reactions.

[0268] The weight space velocity (WHSV) of a supply is, for example, 0.1h -1 ~20h -1 , preferably 0.5h -1 ~10h -1 , and even more preferably 0.8 h -1 ~5h -1 It is.

[0269] The temperature at the inlet of the one or more reactors is advantageously sufficient to allow relatively high conversion without being too high so as to avoid undesirable concurrent reactions.

[0270] The temperature at the inlet of the or each reactor is, for example, 150°C to 400°C, preferably 180°C to 350°C, and even more preferably 200°C to 290°C.

[0271] The pressure through the reactor or reactors for olefin oligomerization and aromatics alkylation is advantageously sufficient to permit relatively high conversion without being too low to avoid undesirable concurrent reactions.

[0272] The pressure through the or each reactor is between 8 bara and 100 bara, preferably between 10 bara and 85 bara, more preferably between 25 bara and 75 bara (bar, absolute pressure).

[0273] As regards the nature of the catalyst, the first group of catalysts used are selected from the following list, whether or not they contain alkali elements or rare earth elements: + The acid catalysts include any of the following types: amorphous or crystalline aluminosilicates, or silicoaluminophosphates in the form of: 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), T ON (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 (L zeolite), zeolite omega, and a group of microporous materials composed of silica, aluminum, oxygen and optionally boron.

[0274] Prior to its use, the zeolite may be subjected to various treatments which may include ion exchange; modification with metals; steam treatment; acid treatment or any other suitable dealumination method; silicon surface passivation by silicon deposition; or any combination of the foregoing treatments.

[0275] The content of alkali or rare earth is 0.05% to 10% by weight, preferably 0.2% to 5% by weight.Preferably, the metals used are Mg, Ca, Ba, Sr, La, Ce, used individually or in the form of a mixture thereof.

[0276] A second group of catalysts used comprises phosphorus-modified zeolites, optionally containing alkali or rare earths, in which case the zeolites may be selected 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.

[0277] Prior to its use, the zeolite may be subjected to various treatments which may include ion exchange; modification with metals; steam treatment; acid treatment or any other suitable dealumination method; mesoporosification treatment, silicon surface passivation by silicon deposition; or any combination of the aforementioned treatments.

[0278] The content of alkali or rare earth is 0.05% to 10% by weight, preferably 0.2% to 5% by weight.Preferably, the metals used are Mg, Ca, Ba, Sr, La, Ce, used individually or in the form of a mixture thereof.

[0279] A third group of catalysts that can be used includes bifunctional catalysts, including: - substrates 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 the ZSM-48 group of microporous materials consisting of silicon, aluminum, oxygen and optionally boron. MFI or MEL (Si / Al>25), MCM-41, MCM-48, SBA-15, SBA-16, SiO 2 , Al 2 O 3 , hydrotalcite, or mixtures thereof. - a metallic phase (Me) corresponding to 0.1% by weight, in which the metals are selected from the following elements: Zn, Mn, Co, Ni, Ga, Fe, Ti, Zr, Ge, Sn and Cr, used individually or as a mixture. These metal atoms may be inserted into the tetrahedral structure of the substrate. The incorporation of the metal may be carried out by adding it during the synthesis of the substrate or by post-synthesis incorporation by ion exchange or impregnation, so that the metal is then not integrated inside the structure of the substrate but in the form of a cation.

[0280] Prior to its use, the zeolite may be subjected to various treatments which may include ion exchange; modification with metals; steam treatment; acid treatment or any other suitable dealumination method; mesoporosification treatment, silicon surface passivation by silicon deposition; or any combination of the above mentioned treatments.

[0281] The content of alkali or rare earth is 0.05% to 10% by weight, preferably 0.2% to 5% by weight.Preferably, the metals used are Mg, Ca, Ba, Sr, La, Ce, used individually or in the form of a mixture thereof.

[0282] A fourth group of catalysts used comprises amorphous solids such as silica-alumina, silica-phosphate, silica-borate, silica-titanium, silica-zirconia and / or mixtures thereof.

[0283] The catalysts can be mixtures of the above mentioned materials within the four catalyst groups. Additionally, the active phase can be combined with other components (binders, matrices) that impart greater mechanical strength or enhanced activity to the final catalyst.

[0284] When the hydrocarbon feed is oligomerized in a plant containing multiple reactors in series, the reactors in series may be fed with the same or different catalysts.

[0285] In one variant, the oligomerization step (c) for oligomerizing olefins is carried out in an oligomerization reactor and the alkylation step (d) for alkylating aromatic compounds is carried out in an alkylation reactor separate from the oligomerization step (c) for oligomerizing olefins.

[0286] Advantageously, the water-depleted mixture is then separated in a first column into a C1-C2 hydrocarbon fraction, for example taken from the top of the column; a C3-C5 hydrocarbon fraction, for example taken from an intermediate stage of the column; and a C6+ hydrocarbon fraction, for example taken from the bottom of the column. The C1-C2 and C6+ hydrocarbon fractions are sent to an alkylation step (d) in an alkylation reactor, while the C3-C5 hydrocarbon fraction is sent to an oligomerization step (c) in an oligomerization reactor.

[0287] In one variant, the water-depleted mixture is separated into a C3-hydrocarbon fraction, for example taken from the top of the column, a C4-C5 hydrocarbon fraction, for example taken from an intermediate stage of the column, and a C6+hydrocarbon fraction, for example taken from the bottom of the column. The C3- and C6+hydrocarbon fractions are sent at least partially to an alkylation step (d) in an alkylation reactor, while the C4-C5 hydrocarbon fraction is sent at least partially to an oligomerization step (c) in an oligomerization reactor.

[0288] The product from the oligomerization reactor contains greater than 50% by weight C7+ olefins, specifically greater than 60% by weight C9-C12 olefins.

[0289] The alkylation occurs under temperature and pressure conditions effective to maintain greater than 20 weight percent of the feed in the alkylation zone in the liquid phase.

[0290] Advantageously, the alkylation of aromatic compounds with alkenes is carried out in the liquid phase, since aromatic compounds are essentially present in the liquid phase. The alkylation may be carried out using solid acid catalysts. Shape-selective zeolite and silica-alumina catalysts are generally used.

[0291] In this process, the reactor conditions are selected to ensure that the alkene introduced into the reactor is mostly dissolved in the aromatic feedstock. This is usually achieved by an optimum combination of operating conditions such as pressure, temperature and catalyst selection, together with sufficiently high catalyst activity. The vapor phase alkenes present can cause rapid deactivation of the alkylation catalysts that are typically used in the liquid phase.

[0292] Examples of usable operating conditions are provided in U.S. Patent No. 4,891,458, which describes the liquid phase synthesis of ethylbenzene using zeolite beta, while U.S. Patent No. 5,334,795 describes the use of MCM-22 in the liquid phase synthesis of ethylbenzene; U.S. Patent No. 7,649,122 describes the use of MCM-22 in the liquid phase synthesis of ethylbenzene, where a given water content level is maintained. U.S. Patent No. 4,549,426 describes the liquid phase synthesis of alkylbenzenes using steam-stabilized zeolite Y; U.S. Patent No. 8,134,036 describes liquid phase aromatic alkylation in at least one catalyst bed containing a first catalyst modified with inclusions of rare earth metal ions.

[0293] The types of products which can be produced preferably correspond to the generic formulas monoalkylbenzenes, dialkylbenzenes and trialkylbenzenes. 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 length.

[0294] The aromatic compounds produced in the conversion step (a) to convert C1-C6 alcohols are typically mono-aromatics, optionally alkylated mono-aromatics (benzene, toluene, ethylbenzene and xylenes), and the alkylating agent is an olefin.

[0295] The alkylation reaction is exothermic, and as a result, in the case of multiple beds, it may be useful to inject a portion of the aromatics and / or a portion of the olefins between the different reactor beds. Aromatics having less than 8 carbon atoms may be recycled, as may overly short olefins having, for example, less than 5 carbon atoms.

[0296] The alkylation catalyst may, for example, take the form of a bead, but more often an extruded form. It is composed of an acidic solid mixed with an amorphous phase. The formation of the acidic solid is carried out with the aid of a matrix which is an amorphous phase. The acidic solid is preferably at least one zeolite selected from among zeolites of the FAU structural type and more particularly zeolites Y, zeolites of the MOR structural type (i.e. mordenite zeolite), zeolites of the EUO structural type (i.e. zeolites EU-1, ZSM-50, TPZ-3), zeolites of the NES structural type NU-87, zeolites NU-86 (described in EP-A-463768), zeolites NU-85 (described in EP-A-462745), zeolites NU-88 (described in FR-A-2752567) and zeolites IM-5 (described in FR-A-2754809), zeolites beta, zeolites MCM-22, zeolites MCM-36, zeolites MCM-49 or zeolites MCM-56.

[0297] Preferably, the catalyst has a silica / alumina molar ratio (SiO 2 / Al 2 O 3 It is a zeolite beta having a molecular weight of 1.01 to 1.01 g / mol.

[0298] Zeolite beta, for example, contains Na 2 The sodium content may be low, less than about 0.2% or less than about 0.02% by weight, expressed as O. The sodium content can be reduced by any method known to those skilled in the art, such as ion exchange.

[0299] These zeolites are at least partially in the acid form (H + ), but alkaline earth or rare earth H + The zeolite catalyst may contain cations other than those listed above. The zeolite catalyst may be modified with rare earth metal ions such as lanthanum, cerium, neodymium or praseodymium.

[0300] The Brunauer-Emmett-Teller (BET) surface area of ​​the catalyst used is 50 m 2 / g~900m 2 / g, preferably 100m 2 / g~700m 2 The Na / Al ratio of the final catalyst is less than 5 atomic %, preferably less than 2%.

[0301] The content of zeolites in the catalyst is in particular between 5% and 95% by weight, preferably between 10% and 90% by weight, relative to the final catalyst. The overall Si / Al ratio of these zeolites is between 2.6 and 200, preferably between 5 and 100, even more preferably between 5 and 80.

[0302] The matrix of the catalyst is a substrate selected from the group formed by alumina, silica, silica-alumina, alumina-boron oxide, magnesia, silica-magnesia, zirconia, titanium oxide and clay, these compounds being used individually or in their mixtures. It is preferred to use an alumina substrate.

[0303] Preferably, the solid acid catalyst has shape selectivity to avoid the formation of oversized alkylaromatic compounds, such as those having more than 16 carbon atoms, such that if the molecular size of the alkylaromatic compounds is close to the size of the pores in the catalyst, formation and diffusion at the output from the pores is still feasible, while thus not resulting in the formation of alkylaromatic compounds that are too large to enter, reside within, or exit the pores.

[0304] As stated above, the reaction zone is preferably operated at a temperature and pressure to maintain phase conditions presenting greater than 20% by weight liquid.

[0305] For the preparation of alkylaromatic compounds having at least 8 carbon atoms, the reaction temperature is in particular between 140° C. and 320° C., typically between 160° C. and 280° C. In one embodiment, the reaction temperature is between 190° C. and 240° C.

[0306] The alkylation pressure is generally maintained at a level high enough to ensure the presence of a liquid phase, in one embodiment the pressure is between 20 barg and 100 barg, particularly between 30 barg and 50 barg.

[0307] When operating under predominantly liquid phase conditions, the upflow reactor mode is generally used. Flow rates are typically around 1 h per bed. -1 ~100h -1 , preferably about 2 h per bed -1 ~7h -1 The liquid hourly space velocity (LHSV) of the aromatic compound / alkylating agent may vary from, for example, 0.05 mol / mol to 20 mol / mol, preferably 0.1 mol / mol to 10 mol / mol.

[0308] In one preferred mode of operation, the oligomerization of olefins and the alkylation of aromatics with olefins are carried out over the same catalyst and in the same reactor. The known operating conditions for oligomerization and alkylation are very similar and can be easily adapted to achieve the desired oligomerization and alkylation performance.

[0309] When oligomerization and alkylation are carried out simultaneously in the same reactor using the same catalyst, the reactor product contains more than 10% by weight of C8+ aromatics, specifically more than 6% by weight of C8-C14 aromatics.

[0310] When alkylation in the presence of olefins is carried out separately from oligomerization, the reactor product contains more than 65 mass % C8+ aromatics, specifically more than 75 mass % C8-C14 aromatics.

[0311] Step (e) Forming a Hydrocarbon Stream to be Hydrogenated The 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 aromatic compounds alkylated in step (d).

[0312] When a single common stage is used for the olefin oligomerization step (c) and the aromatics alkylation step (d), the product from this stage is used partially or completely to form the hydrocarbon stream to be hydrogenated.

[0313] The product may, for example, comprise, on an anhydrous basis: - less than 15% by weight of paraffin, in particular between 3% and 10% by weight of paraffin, - less than 10% by weight of C4 to C7 olefins, in particular 0.5% to 5% by weight of C4 to C7 olefins, - more than 50% by mass of C8 to C16 olefins, in particular 60% to 80% of C4 to C16 olefins, - less than 5% by weight of C17+ olefins, particularly between 0.1% and 1.0% by weight of C17+ olefins; and / or - less than 5% C6-C7 aromatic compounds, specifically between 0.5% and 4.0% C6-C7 aromatic compounds; - Greater than 2% C8+ aromatics, specifically between 6% and 30% C8+ aromatics.

[0314] In one variation, an optional separation step for separating the effluent containing at least a portion of the oligomerized olefins from step (c) and / or at least a portion of the alkylated aromatic compounds from step (d) is carried out in an additional distillation column.

[0315] The separation produces a fraction of C7-hydrocarbons at the top of the column and a fraction of C8+hydrocarbons at the bottom.

[0316] The C7-hydrocarbon fraction preferably contains more than 90% by weight, in particular more than 95% by weight, of the C7-hydrocarbons contained in the effluent.

[0317] The C8+ hydrocarbon fraction comprises greater than 90% by mass of the C8+ hydrocarbons contained in the effluent.

[0318] Advantageously, at least a portion of the C7-hydrocarbon fraction, for example less than 50% by mass, is at least partially recycled in the olefin oligomerization step (c) and / or the aromatics alkylation step (d), and another portion forms a gasoline stream.

[0319] The hydrocarbon stream to be hydrogenated is formed at least in part, preferably entirely, of the C8+ hydrocarbon fraction.

[0320] When the oligomerization step (c) for oligomerizing olefins is carried out in an oligomerization reactor and the alkylation step (d) for alkylating aromatics is carried out in an alkylation reactor separate from the oligomerization step (c) for oligomerizing olefins, the oligomerization reactor product and the alkylation reactor product containing the alkylated aromatics are advantageously subjected to separation.

[0321] In one embodiment, the oligomerization reactor product containing oligomerized olefins and the alkylation reactor product containing alkylated aromatics are separated into a C8+ hydrocarbon fraction at the bottom and a C7- hydrocarbon fraction at the top, which is at least partially (e.g., less than 50% by weight) recycled in step (c) in the oligomerization reactor.

[0322] At least a portion, preferably all, of the C8+ hydrocarbon fraction forms the hydrocarbon stream to be hydrogenated.

[0323] In another embodiment, the products from the oligomerization reactor and the alkylation reactor are separated in a distillation column into a C7- hydrocarbon fraction withdrawn at the top; a C8 to C16 hydrocarbon fraction withdrawn from the middle stage; and a C17+ hydrocarbon fraction withdrawn at the bottom.

[0324] At least a portion, preferably the entire C8-C16 hydrocarbon fraction forms the hydrocarbon stream to be hydrogenated.

[0325] The C17+ hydrocarbon fraction is at least partially recycled in the conversion step (a) for converting the C1-C6 alcohol stream, with the aim of cracking the C17+ olefins again.

[0326] Step (f) Hydrogenation The hydrocarbon stream to be hydrogenated is subjected to hydrogenation to form a hydrogenated hydrocarbon stream, whereby the olefinic compounds are saturated and the aromatic compounds are partially hydrogenated.

[0327] The hydrogenation is carried out, for example, in one or more mixed-phase fixed-bed (downward or upward) reactors, the fraction to be hydrogenated being mainly in the liquid phase.

[0328] The hydrogenation is carried out at a temperature between 50° C. and 350° C., in particular between 100° C. and 300° C. It is preferably carried out using a pressure of more than 10 bara, in particular between 20 bara and 80 bara.

[0329] The hydrogen stream is fed into the or each reactor in a mixed state with the hydrocarbon stream to be treated. The ratio of the volumetric flow rate of the hydrogen stream to the volumetric flow rate of the hydrocarbon stream to be hydrogenated (excluding the recycle stream) is advantageously between 50NL / L and 3000NL / L, in particular between 100NL / L and 500NL / L. Hydrogen may be added to the hydrocarbon stream in several stages along the catalyst bed. The space velocity is advantageously between 0.5 and 3, in particular between 1 and 2h -1 The excess hydrogen, after separation and compression, may be recycled back into the reaction zone.

[0330] 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 substrate such as silica, alumina or any suitable 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 a mixture of metals from these two groups.

[0331] The hydrogenated hydrocarbon stream advantageously contains less than 10% by weight of olefins, and preferably less than 3% by weight of olefins.

[0332] The hydrogenated hydrocarbon stream preferably comprises more than 50% by mass of paraffins, particularly more than 50% by mass of C7-C17 paraffins, particularly more than 60% by mass of C7-C17 paraffins, particularly between 70% and 95% by mass of C7-C17 paraffins.

[0333] According to one variant, between 10% and 90% by weight of the aromatics contained in the hydrocarbon stream to be hydrogenated, preferably between 30% and 80% by weight of the aromatics contained in the hydrocarbon stream to be hydrogenated, are hydrogenated to cycloparaffins.

[0334] Hydrogenated hydrocarbon streams include, for example: - more than 20% by weight of C9 hydrocarbons, for example between 20% by weight and 40% by weight of C9 hydrocarbons; - more than 20% by weight, for example between 20% and 40% by weight, of C12 hydrocarbons; - More than 6% by mass of C6+ aromatics, specifically more than 6% by mass of C8+ hydrocarbons.

[0335] Step (g) Recovery of the jet fuel fraction For the purpose of recovering the jet fuel fraction, the hydrogenated hydrocarbon stream is advantageously subjected to fractionation, which may be carried out by passing the stream through at least one separation column, for example a distillation column.

[0336] Preferably, at least one first separation column is used to separate the liquefied petroleum gas fraction at the top from the remainder of the hydrogenated hydrocarbon stream obtained at the bottom.

[0337] The fractionation conditions in this column are: a pressure advantageously between 2 bara and 15 bara, a condensation temperature at the top of the column adjusted to allow the use of an air or cooling water condenser, ie a temperature preferably between 20° C. and 50° C.

[0338] The remaining fraction of the hydrogenated hydrocarbon stream is then introduced into a second separation column for the purpose of producing a naphtha fraction at the top, a diesel fraction at the bottom, and a jet fuel fraction at at least one intermediate stage. This fractionation may be carried out in a single column with a side draw of the jet fuel fraction, or in two separate columns.

[0339] More than 80% by mass of the hydrogenated hydrocarbon stream introduced into the second column advantageously forms the jet fuel fraction.

[0340] The liquefied petroleum gas fraction preferably has a final boiling point of less than 180°C, even more preferably less than 150°C.

[0341] The initial boiling point may be in the range of 20°C to 60°C, preferably 25°C to 40°C.

[0342] The naphtha fraction contains greater than 80% by weight of the C8-paraffins contained within the residual fraction.

[0343] The jet fuel fraction contains between 2% and 30% by volume of C8+ aromatics, preferably between 6% and 25% by volume of C8+ aromatics, and even more preferably between 8% and 25% by volume of C8+ aromatics.

[0344] This fraction contains more than 50% by volume of C9-C16 paraffins, specifically 60% to 95% by volume of C9-C16 paraffins.

[0345] Specifically, the jet fuel fraction contains greater than 60% by volume of C9-C12 paraffins.

[0346] This fraction preferably has a final boiling point of less than 400°C, even more preferably less than 350°C.

[0347] The initial boiling point is in the range of 130°C to 180°C.

[0348] The diesel fraction is the heaviest fraction and not all of its molecules are used to form jet fuel. The initial boiling temperature of this fraction is typically above 300°C, preferably 310°C.

[0349] Advantageously, at least a portion of the diesel fraction is recycled to the conversion step (a) for converting a C1-C6 alcohol stream in the form of a recycle stream with a view to producing further cracking of the compounds present in this fraction.

[0350] The recycle stream advantageously constitutes between 10% and 50% by weight of the C1-C6 alcohol stream introduced into the conversion step (a).

[0351] The invention can be better understood from reading the following description, given purely by way of example, and with reference to the accompanying drawings, in which: [Brief description of the drawings]

[0352] [Figure 1] FIG. 1 is a schematic diagram of a plant configured for carrying out a first fuel production process for producing jet fuel according to the present invention. [Diagram 2] FIG. 2 is a diagram similar to that of FIG. 1 illustrating a variant of a plant designed for carrying out a variant of the process shown in FIG. 1; [Diagram 3] FIG. 2 is a diagram similar to that of FIG. 1 illustrating a variant of a plant designed for carrying out a variant of the process shown in FIG. 1; [Figure 4]FIG. 2 is a diagram similar to that of FIG. 1 illustrating a variant of a plant designed for carrying out a variant of the process shown in FIG. 1; [Diagram 5] FIG. 2 is a diagram similar to that of FIG. 1 illustrating a variant of a plant designed for carrying out a variant of the process shown in FIG. 1; [Figure 6] FIG. 2 is a diagram similar to that of FIG. 1 illustrating a variant of a plant designed for carrying out a variant of the process shown in FIG. 1; [Figure 7] FIG. 2 is a diagram similar to that of FIG. 1 illustrating a variant of a plant designed for carrying out a variant of the process shown in FIG. 1; [Figure 8] FIG. 2 is a detailed diagram illustrating a reactor for carrying out a fluidized bed-based conversion step. [Figure 9] NMR spectra of ZSM5 parent catalyst and modified catalyst with ALPO structure. [Figure 10] 1 illustrates the selectivity obtained when carrying out an exemplary conversion step of the process according to the present invention. [Figure 11] 1 illustrates the conversion of certain olefins as a function of time during an exemplary run of a first catalytic oligomerization step. [Figure 12] 1 illustrates the conversion of certain olefins as a function of time during an exemplary run of the second catalytic oligomerization step. [Figure 13] 1 illustrates the conversion of certain aromatic compounds as a function of time during an exemplary run of an aromatic compound alkylation step using a second catalyst. DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS

[0353] A first production plant 10 designed to carry out the jet fuel production process according to the present invention is illustrated generally in FIG.

[0354] The first plant 10 comprises a conversion stage 12 for converting a C1-C6 alcohol stream 14 intended to produce a mixture 16 containing paraffins, olefins, aromatics and water; and a separation stage 18 for separating water from the mixture 16 in order to produce a water-depleted mixture 19 comprising a liquid phase 19a and a gas phase 19b. The plant shown in Figure 1 also comprises a separation stage 20 for separating C1-C2 hydrocarbons from the water-depleted mixture.

[0355] In this embodiment, the plant 10 further includes a combined oligomerization and alkylation stage 22 for oligomerizing olefins and alkylating aromatics from the water-depleted mixture, which produces a hydrocarbon stream 24 to be hydrogenated.

[0356] The plant 10 further includes a hydrogenation stage 26 for hydrogenating the stream 24 to produce a hydrogenated hydrocarbon stream 30; and a fractionation stage 28 for fractionating the hydrogenated hydrocarbon stream 30, which is designed to fractionate at least one jet fuel fraction 34 and advantageously one diesel fraction 36 and one naphtha fraction 38.

[0357] Conversion stage 12 is designed to carry out conversion step (a) described above, which converts C1-C6 alcohols to primarily C3-C7 olefins.

[0358] As discussed above, conversion stage 12 includes at least one fixed bed reactor, e.g., a plurality of fixed bed reactors, with, e.g., one or more fixed bed reactors advantageously defining a plurality of successive fixed beds of catalyst, as discussed above.

[0359] With reference to the above description, separation stage 18 is designed to perform a separation step (b) for separating water. This separation stage comprises at least one separator operating by gravity and / or mechanical drive for the purpose of separating water from mixture 16 to recover a water-enriched aqueous fraction (stream 40), a gaseous hydrocarbon fraction 19b and a liquid hydrocarbon fraction 19a.

[0360] Optionally, plant 10 includes at least one recycle conduit 18a for recycling water separated in the separation stage to conversion stage 12. If conversion stage 12 includes at least one fixed bed reactor, recycle conduit 18a opens, for example, upstream of a fixed catalyst bed or between two successive fixed catalyst beds.

[0361] Advantageously, separation stage 18 includes a stripping column 41 capable of treating at least a portion of the separated water forming stream 40 to extract the contained hydrocarbons and obtain a treated water.

[0362] Plant 10 includes at least one cooling device (e.g., a heat exchanger) that serves to reduce the temperature of the product output from the reactor that heats another stream, such as the make-up feed to reactor 12; a water-cooled or air-cooled cooler; and / or a combination of the foregoing downstream of conversion stage 12 to condense water and produce stream 40.

[0363] As defined above, the separation stage 20 separates hydrocarbons lighter than C3 hydrocarbons, such as C1-C2 hydrocarbons, CO, CO 2 and lighter compounds such as hydrogen. Advantageously, this stage comprises at least one deethanizer. Since the separation stage 20 operates at a higher pressure than the water separation stage 18, the plant 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.

[0364] The olefin oligomerization and aromatics alkylation stage 22 is designed to jointly carry out steps (d) and (e). This stage includes at least one joint oligomerization and / or alkylation reactor designed to carry out the experimental conditions described above.

[0365] Hydrogenation stage 26 is designed to carry out step (f). This stage comprises at least one fixed bed hydrogenation reactor designed to carry out the hydrogenation reaction under the conditions described above.

[0366] Fractionation stage 28 is designed to carry out step (g). In this embodiment, the stage includes at least a first upstream separation column 42 for separating liquefied petroleum gas 44, and a second downstream fractionation column 46 designed to produce fractions 34-38.

[0367] A first embodiment of a jet fuel production process carried out in the plant shown in FIG. 1 is described below.

[0368] Initially, a C1-C6 alcohol stream 14 is delivered to the conversion stage 12. The alcohol stream 14 originates, for example, from a source 50 as described above, where the alcohol from the source 50 is produced, for example, by fermentation of biomass, catalytic conversion of carbohydrates or carbon monoxide or carbon dioxide in the presence of hydrogen, etc.

[0369] Stream 14 may be comprised of the composition described above, for example with greater than 50% methanol by dry weight, and preferably greater than 80% methanol by dry weight.

[0370] Stream 14 is introduced into conversion stage 12 where it undergoes the conversions described above, including dehydration / aromatization of C2-C6 alcohols, and, in the case of methanol, conversion to dimethyl ether followed by dehydration.

[0371] The reaction is carried out under operating conditions in terms of temperature and pressure as described above. One or more of the catalysts defined above are used.

[0372] A mixture 16 is thus obtained that contains paraffins (specifically n-paraffins, i-paraffins and cycloparaffins), olefins, aromatics and water. For example, the mixture 16 may have the composition described above.

[0373] The mixture 16 is then introduced into a separation stage 18 to produce a separator bottom water stream 40 and a water-depleted mixture 19 comprising a gas phase 19b and a liquid phase 19a.

[0374] A portion 40a of water stream 40 is optionally recycled to conversion stage 12 via conduit 18a, as described above. Another portion 40b of water stream 40 is introduced into column 41 for stripping to produce an extracted hydrocarbon stream 41a at the top and a treated water stream 40b at the bottom, which has a lower hydrocarbon content than water stream 40.

[0375] The extracted hydrocarbon stream 41a is recycled to the separation stage 18, for example upstream of the separator.

[0376] 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 its top CO, CO 2 at the bottom, a fraction 60 of C1-C2 hydrocarbons containing light compounds such as hydrogen; and at the bottom, a fraction 62 of C3+ hydrocarbons.

[0377] The resulting fractions 60 and 62 have the composition defined above.

[0378] Preferably, the C1-C2 hydrocarbon fraction 60, optionally after separation, is sent into a steam cracker for recovery of at least a portion of the ethylene contained therein.

[0379] In one variant represented by a dotted line in Figure 1, at least a portion 64 of the C1-C2 hydrocarbon fraction 60 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.

[0380] The ratio of the mass flow rate of recycle stream 64 to the mass flow rate of fraction 60 output from the top of the deethanizer of separation stage 20, as defined above, is less than 0.5.

[0381] In the embodiment shown in FIG. 1, the C3+ hydrocarbon fraction 62 having the composition defined above is then introduced into a joint oligomerization and alkylation stage 22.

[0382] In this stage 22, one or more combined oligomerization and alkylation reactors carry out the oligomerization of the olefins present in fraction 62, in particular the oligomerization of the C3 to C7 olefins, according to the operating conditions defined herein above.

[0383] Additionally, jointly, the C6+ aromatic compounds present in fraction 62 are alkylated to form, inter alia, C8+ aromatic compounds.

[0384] The reaction is carried out under the operating conditions described above, using one or more of the catalysts defined above.

[0385] At the outlet of the stage 22, a hydrocarbon stream 24 to be hydrogenated is formed with a composition as defined above.

[0386] The hydrocarbon stream 24 to be hydrogenated is then introduced into a hydrogenation stage 26 for the purpose of inducing hydrogenation of at least a portion of the olefins present in the hydrocarbon stream 24 to be hydrogenated, as well as hydrogenation of at least a portion of the aromatic compounds present in the hydrocarbon stream 24 to be hydrogenated to cycloparaffins.

[0387] The hydrogenation is carried out using one or more of the catalysts described above and under the operating conditions described above.

[0388] A hydrogen-containing stream 66 is introduced into the hydrogenation stage 26, the ratio of the volumetric flow rate of hydrogen in stream 66 to the volumetric flow rate of the hydrocarbon stream 24 to be hydrogenated being, for example, as defined above.

[0389] A hydrogenated hydrocarbon stream 30 is formed at the outlet of the hydrogenation stage with the composition described above.

[0390] The hydrogenated hydrocarbon stream 30 is then fractionated in fractionation stage 28 .

[0391] In the first column 42, the hydrogenated hydrocarbon stream is separated into a C4- hydrocarbon fraction forming a liquefied petroleum gas fraction 44 and a C4+ hydrocarbon fraction forming a residual fraction 70 of the hydrogenated hydrocarbon stream.

[0392] Fractions 44, 70 have the characteristics defined above in terms of cut points.

[0393] The residue fraction 70 is introduced into the second column 46 and fractionated therein into the naphtha fraction 38, the jet fuel fraction 34 and the diesel fraction 36 as characterized above.

[0394] The plant variant 90 illustrated in Figure 2 is designed to carry out the second process according to the invention. It differs from the plant 10 illustrated in Figure 1 in that the fractionation stage 20 includes an additional recovery column 92 for recovering propylene.

[0395] The C3+ hydrocarbon fraction 62 obtained from the deethanizer of the separation stage 20 is introduced into an additional column 92 to form a C3- hydrocarbon fraction 80 at the top of the column and a C4+ hydrocarbon fraction 82 at the bottom of the column, which are intended to be introduced into the joint oligomerization and alkylation stage 22.

[0396] Fractions 80, 82 have the compositions previously described above. Fraction 80 contains greater than 80% by weight of the propylene contained in C3+ hydrocarbon fraction 62. Such an embodiment provides a means for recovering the propylene formed in conversion stage 12, when such recovery is economically attractive.

[0397] The plant 100 depicted in Figure 3 is designed for carrying out the third process according to the invention. It differs from the plant 10 depicted in Figure 1 in that it is equipped with at least one tap 102 for the addition of C2-C6 alcohols for quenching in the conversion stage 12, for example between two successive catalyst beds of the conversion stage 12.

[0398] The composition of the C2-C6 alcohol stream 102 advantageously comprises less than 20% methanol and more than 80% by weight C2-C6 alcohols, such as more than 50% ethanol and propanol.

[0399] The addition of C2-C6 alcohols in addition to methanol facilitates the conversion reaction for converting the alcohol stream 14 by making it more isothermal (the conversion of methanol is highly exothermic and the conversion of C2-C6 alcohols is endothermic) and therefore easier to control, particularly when a fixed catalyst bed is used in the conversion stage 12.

[0400] The plant 110 depicted in Figure 4 is similar to the plant shown in Figure 1. The plant includes a conversion stage 12 having at least one reactor with a fluidized catalyst bed, preferably having a single fluidized catalyst bed reactor, said one or more reactors being suitable for carrying out the experimental conditions described above.

[0401] The reactor comprises a reaction zone 111a having a fluidized catalyst bed and a regeneration zone 111b for regenerating the fluidized catalyst bed. A portion of the catalyst present in the reaction zone 111a is continuously withdrawn for regeneration in the regeneration zone 111b, advantageously by controlled combustion in the presence of oxygen.

[0402] A portion 40a of water stream 40 is optionally recycled to conversion stage 12 via conduit 18a.

[0403] Finally, as an option, a portion 64 of the C1-C2 hydrocarbon fraction separated from the water-depleted mixture 19 is introduced into a reactor having a fluidized catalyst bed.

[0404] Advantageously, the C1-C6 alcohol stream is introduced into the conversion step (a) at a temperature at least 5° C. above the bubble point of the C1-C6 alcohol stream.

[0405] FIG. 8 illustrates the flow of catalyst from regeneration zone 111b to reaction zone 111a in one particular embodiment.

[0406] For reasons of simplicity, the figure does not include details of the internal parts of the vessel that form the zones 111a, 111b.

[0407] The C1-C6 alcohol stream 14 is introduced into the bottom of reaction zone 111a, which is characterized by a fluidized catalyst bed.

[0408] At the top of the reaction zone 111a, the products from the shift reaction are separated from the catalyst in a disengagement zone 203, advantageously equipped with a cyclone, and the resulting mixture 16 is conveyed to a separation stage 18.

[0409] Optionally, the heat of reaction produced by the conversion is extracted from reaction zone 111a using catalyst cooler 205, which is a heat exchanger advantageously located outside and connected to reaction zone 111a.

[0410] Reaction zone 111a receives catalyst regenerated in regeneration zone 111b via a supply line 207 that connects regeneration zone 111b to reaction zone 111a.

[0411] Deactivated catalyst is removed from the disengagement zone 203 via a discharge line 206, separate from the make-up line 207, which connects the reaction zone 111a to the regeneration zone 111b.

[0412] Air is injected into the regeneration zone 111b through an injection conduit 221 at the bottom of this zone, in a fluidized bed where the coke deposits are burned off.

[0413] The regeneration zone 111b also includes a disengagement zone 222, advantageously equipped with a cyclone, in which the flue gas is separated from the regenerated catalyst and discharged via a regeneration conduit 223, advantageously located at the top of the regeneration zone 111b.

[0414] Optionally, a catalyst cooler (not depicted in the figure but similar to catalyst cooler 205) is connected to the regeneration zone 111b because the combustion of coke deposits is a highly exothermic reaction and the temperature in the regeneration zone 111b needs to be carefully controlled. Hot catalyst extracted from the regeneration zone is circulated through this cooler for cooling, which thus controls the temperature in the regeneration zone 111b. The regenerated catalyst is sent via line 207 to the reaction zone 111a.

[0415] A fifth plant 120 designed to carry out a fifth process according to the invention is illustrated in Figure 5. The fifth plant 120 differs from the first plant 10 in that it includes an additional separation stage 122 located between the outlet of the joint oligomerization and alkylation stage 22 and the inlet of the hydrogenation stage 26.

[0416] The additional separation stage 122 includes at least one distillation column.

[0417] The product 124 from the joint oligomerization and alkylation stage 22 is separated in a distillation column into a C7-hydrocarbon fraction 126 and a C8+hydrocarbon fraction 128, which form the hydrocarbon stream 24 intended to be hydrogenated. The C7-hydrocarbon fraction 126 may optionally be recycled to the joint oligomerization and alkylation stage 22.

[0418] A sixth plant 140 according to the invention is illustrated in Figure 6. This sixth plant 140 is designed to carry out the sixth process according to the invention. It differs from the first plant 10 in that it has a dedicated oligomerization stage 22A for oligomerizing olefins derived from the water-depleted mixture 19 coming from stage 20, and a dedicated alkylation stage 22B for alkylating aromatics derived from the water-depleted mixture 19 coming from stage 20.

[0419] Each stage 22A, 22B respectively includes separate oligomerization and alkylation reactors in which the operating conditions provided herein above are carried out.

[0420] In separation stage 20, the water-depleted mixture 19 is separated into a fraction 60 of C1-C2 hydrocarbons that is recovered at the top of the deethanizer; a fraction 142 of C3-C5 hydrocarbons that is recovered at an intermediate stage of the deethanizer; and a fraction 144 of C6+ hydrocarbons that is recovered at the bottom of the deethanizer.

[0421] The C3-C5 hydrocarbon fraction 142 is sent in its entirety to oligomerization stage 22A for producing oligomerization reactor product 146.

[0422] The C1-C2 hydrocarbon fraction 60 and the C6+ hydrocarbon fraction 144 are conveyed to the alkylation stage 22B for producing an alkylation reactor product 152 under the operating conditions defined above.

[0423] The product 152 from the alkylation reactor is then mixed with the product 146 from the oligomerization reactor.

[0424] The products 146, 152 are then introduced into the additional separation stage 122 and separated into a C7- hydrocarbon fraction 126 and a C8+ hydrocarbon fraction 128, as described above.

[0425] At least a portion 150 of the C7-hydrocarbon fraction 126 is recycled into the oligomerization stage 22A, and another portion may be recovered in the form of gasoline.

[0426] Fraction 128 forms the hydrocarbon stream 24 to be hydrogenated.

[0427] A seventh plant 160 designed to carry out a seventh process according to the invention is shown in FIG.

[0428] The seventh process according to the invention differs from the sixth process carried out in the plant 150 in that in an additional separation stage 122, the product 146 from the oligomerization reactor and the product 152 from the alkylation reactor are separated into a C7- hydrocarbon fraction 126 withdrawn from the top of the column; a C8 to C16 hydrocarbon fraction 162 withdrawn from an intermediate stage of the column; and a C17+ hydrocarbon fraction 164 withdrawn from the bottom of the column.

[0429] As previously noted above, at least a portion 150 of the C7-hydrocarbon fraction 126 is recycled back to the oligomerization stage 22A.

[0430] The C8-C16 hydrocarbon fraction 162 is introduced into the hydrogenation stage 26 for hydrogenation.

[0431] The C17+ hydrocarbon fraction 164 is at least partially recycled to the conversion stage 12.

[0432] Thus, the heavy hydrocarbons present in the C17+ hydrocarbon fraction are re-cracked in the conversion stage 12. This results in an increased amount of jet fuel fraction 34 being produced.

[0433] In all the above mentioned cases, the jet fuel fraction 34 produced by the above process can be used in pure form as an aviation jet fuel intended for use in propelling aircraft engines, for example, or in the form of a blended mixture with jet fuel derived from the distillation of petroleum. The jet fuel fraction or its blend is advantageously a sustainable aviation fuel (SAF), the composition of which is similar to the SAF described according to the ASTM D7566 standard.

[0434] The blend comprises at least 5% by mass of the jet fuel fraction 34, particularly at least 10% by mass.

[0435] Thanks to the above mentioned invention it is possible to provide a simple and efficient fuel production method for producing jet fuel from a C1-C6 alcohol stream, i.e. preferably from a renewable source, in particular produced by converting carbon monoxide or carbon dioxide derived from fermentation or / and captured from the atmosphere in the presence of hydrogen.

[0436] The jet fuel fraction produced by the method according to the invention has a very low carbon footprint, since it does not come from petroleum derivatives, but on the contrary comes from sources that contribute to reducing the amount of carbon dioxide present in the atmosphere.

[0437] The jet fuel fraction 34 produced by the method of the present invention is, moreover, very economical to produce and, in some cases, can be used as an aircraft engine propulsion fuel directly without the need for further refining or blending.

[0438] It will be appreciated that the plants shown in FIGS. 2, 3, 5-7 may also feature a conversion stage 12 equipped with a fluidized catalyst bed, as shown in FIG. EXAMPLES

[0439] Some specific, non-limiting implementations of the conversion step (a), the joint oligomerization and alkylation steps (c) and (d), and the hydrogenation step (f) are described below.

[0440] Step (a) Conversion [Catalyst preparation] A sample of zeolite ZSM-5 (Si / Al=12) in the H form (containing 445 ppm Na, less than 25 ppm K, 178 ppm Fe, 17 ppm Ca, and synthesized without a matrix) was heated at atmospheric pressure in 100% H 2 The sample was steam treated at 550° C. for 6 hours in O. In the following, this sample is identified as Sample A.

[0441] The steamed solid A was dissolved in 3.14 M H under reflux conditions (4.2 ml per gram of zeolite) for 4 hours. 3 PO 4 The solid was then separated from the liquid phase at ambient temperature by filtration of the solution. The resulting material was dried at 200° C. for 16 hours. Hereinafter, this sample is identified as Sample B.

[0442] Catalyst Example 1 490 g of sample B was mixed with 490 g of a specific binder (P=15.9 mass%, Si=13.2 mass%, Mg=0.27 mass%, Al=0.15 mass%, K=230 ppm, Na=230 ppm, Ca=19.2 mass%), 34 mass% SiO 2 The mixture was mixed with 588.3 g of low sodium silica sol containing 6 g of xonotlite and 2-3 wt. % of an extrusion additive. The mixture was stirred for 30 minutes and then extruded.

[0443] A particular binder is dissolved in an aqueous medium (1 g solid / 4 ml water) at ambient temperature with an equivalent mass of NH 4 H 2 PO 4 and xonotlite. After stirring for 60 minutes, the phosphorylated xonotlite was separated from the liquid by filtration and dried. The dried product was used as the extrusion component.

[0444] The extruded solid was dried at ambient temperature for 24 hours and then at 200° C. for 16 hours, then washed with demineralized water at ambient temperature and then dried overnight at 110° C. An additional washing step was then performed at ambient temperature with demineralized water at pH 3.08. The catalyst was then dried overnight at 110° C. and calcined at 700° C. for 2 hours.

[0445] Catalyst Example 2 320 g of sample B was mixed with 400 g of a specific binder (P = 15.9 mass%, Si = 13.2, Mg = 0.27, Al = 0.15 mass%, K = 230 ppm, Na = 230 ppm, Ca = 19.2 mass%), 165 ml of H 2 O, 34% by mass SiO 2 and 235 g of low sodium silica sol containing 2-3 wt % extrusion additive. The mixture was stirred for 30 minutes and then extruded.

[0446] A particular binder has an equivalent mass of (NH 4 )H 2 PO 4 (ammonium dihydrogen phosphate) and xonotlite. After stirring for 60 minutes, the phosphorylated xonotlite was separated from the liquid by filtration and dried. The dried product was used as the extrusion component.

[0447] The extruded solid was dried at ambient temperature for 24 hours and then at elevated temperature for 16 hours, then washed and steam heat treated at 600° C. for 2 hours. This sample is hereinafter identified as Sample E.

[0448] Catalyst Example 3 356 g of sample A was mixed with 338.7 g of Nyacol (40% by weight SiO 2 sol), 311.3g of fumed silica (FK500), 480ml of H 2The extruded solid was dried at ambient temperature for 24 h, then at 110 °C for 16 h, and then calcined at 500 °C for 10 h. The final sample contained 40 wt. % zeolite and 60 wt. % SiO 2 The extruded samples were incubated in 0.5 M NH 4 The sample was subjected to ion exchange with Cl, then washed with water, dried at 110 °C for 16 h, and calcined at 450 °C for 6 h. The molded and ion-exchanged sample was dissolved in 3.1 M H 2 O 4 for 4 h under reflux conditions (1 g / 4.2 mL). 3 PO 4 After that it was cooled, filtered and dried at 110° C. for 16 hours.

[0449] The phosphate samples were washed with 0.1 M calcium acetate solution (1 g / 4.2 ml) for 2 h at ambient temperature. The washed samples were then dried at 110 °C for 16 h and soaked in 100 wt. % H at 600 °C for 2 h. 2 The mixture was heat-treated with steam in O.

[0450] Catalyst Example 4 150 g of sample B was contacted with 630 ml of an aqueous solution containing 1.5 g of dispersed xonotlite, followed by 450 g of low sodium silica sol (34% by weight SiO in water). 2 , 200 ppm Na) was added. The solution was then stirred for 1 hour and spray dried. The spray dried solid was washed with water at ambient temperature for 2 hours, then filtered, dried at 110°C for 16 hours, and calcined at 700°C.

[0451] Catalyst Example 5 100 g of sample A was refluxed for 4 hours and dissolved in 25 g of 85% H 3 PO 4 The resulting slurry was contacted with 120 ml of an aqueous solution containing 7 g of dispersed xonotlite, then cooled and added with stirring for a period of time of around 1 hour, after which 300 g of low sodium silica sol (34% by weight SiO in water) was added. 2, 200 ppm Na) was added. The solution was then stirred for 1 hour and spray dried. The spray dried solid was dried at 200° C. for 16 hours, washed with water at ambient temperature for 2 hours, then filtered, dried and calcined at 700° C. for 2 hours.

[0452] Catalyst Example 6 75 g of sample A was mixed with 14.25 g of 85% H 3 PO 4 and 300 ml of demineralized water. The suspension was stirred under reflux for 2 hours. Then, 4.125 g of CaCO 3 was added to the suspension. Heating of the solution was stopped while stirring the mixture was maintained until it reached a temperature below 30° C. As a result, suspension A was obtained.

[0453] Then, 450 g of low-sodium silica sol (34% by weight SiO in water) was added under stirring at ambient temperature for 30 min. 2 , 200 ppm Na) and 4.5 g H 3 PO 4 (85% by mass) was mixed to prepare a solution, as a result of which suspension B was obtained.

[0454] Suspensions A and B were then mixed with the addition of 120 ml of demineralized water. The solution was then stirred for 1 hour and spray-dried. The spray-dried solid was dried at 200° C. for 16 hours, washed with water at ambient temperature for 2 hours, then filtered, dried and calcined at 700° C. for 2 hours.

[0455] Step (a) Carrying out the conversion In a downflow stainless steel fixed bed reactor, T injection was performed at 550°C, pressure of 0.5 barg, and hourly space velocity (WHSV) of 1.6 h -1 ) catalytic tests were carried out on 2 g (35 mesh to 45 mesh particles) of catalyst using an essentially pure methanol feedstock.

[0456] Prior to the catalytic test, N was added to the reaction mixture until the reaction temperature was reached. 2All catalysts were activated in a stream (5 Nl / h). On-line product analysis was performed using a gas chromatograph equipped with a capillary column. The catalytic performance results for the catalysts in Table 1 are reported on a carbon, dry and coke-free basis. Results are provided for the average performance of the catalysts during the first 4 hours of operation.

[0457] [Table 1]

[0458] Combined steps (c) oligomerization and (d) alkylation, and hydrogenation step (f). The characteristics of the ingredients used to carry out steps (c) and (d). The characteristics of the raw materials used are as follows:

[0459] [Table 2]

[0460] The detailed composition of the raw material was determined by GC method.

[0461] [Table 3]

[0462] 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 a fixed-bed tubular reactor with an internal diameter of 18 mm. Prior to testing, the catalyst was activated at 250° C. (10° C. / h) for 8 hours under 135 NL / h nitrogen. The temperature was then reduced to 40° C. at the start of the test program.

[0463] 100 mL of ZSM-5 based catalyst (80 wt. % MFI and 20 wt. % alumina binder) diluted with 100 mL of inert material (0.21 mm SiC) was loaded into a fixed bed tubular reactor with an internal diameter of 18 mm. Prior to testing, the catalyst was activated at 400° C. (60° C. / h) for 2 hours under 160 NL / h nitrogen. The temperature was then reduced to 40° C. at the start of the test program.

[0464] Hydrogenation Fractionation was carried out on the oligomerization product to recover the 145+ and 165°C+ oligomerized cuts hydrotreated over NiMo catalyst. Single pass with no recycle, 80 barg, 1 h -1 Liquid Hourly Space Velocity (LHSV), 500NL / L 2 The operating conditions were selected at a 1000 MPa / hydrocarbon volume ratio and the temperature was increased from 250°C to 270°C.

[0465] Example 1 - Performance results obtained using zeolite-based catalysts The feedstock was processed under the following operating conditions: 55 barg, 1 h -1 Liquid Hour Velocity (LHSV) and temperature of 240℃~280℃.

[0466] [Table 4]

[0467] The conversion of light olefins (C4-C8 olefins) varies from 65% by mass at 240° C. to 91% by mass at 280° C. C9+ olefins are not taken into account in the conversion calculations since they may result from oligomerization of the light olefins (C4 and C5) present in the feed. At 240° C., the conversion of C5-C7 olefins is greater than 88% by mass.

[0468] Olefins can react with aromatics either by alkylation or oligomerization. It was observed that aromatics conversion varied 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 fraction, which means that alkylation is indeed occurring.

[0469] [Table 5]

[0470] Example 2 - Performance results obtained using amorphous silica-alumina (ASA) catalyst The feedstock was processed under the following operating conditions: 25 barg, 1 h -1 Liquid HSV and temperature of 180℃~220℃.

[0471] [Table 6]

[0472] The conversion of light olefins (C4-C8 olefins) varies from 80% by mass at 180° C. to almost 100% by mass at 220° C. C9+ olefins are not taken into account in the conversion calculations since they may result from oligomerization of the light olefins (C4 and C5) present in the feed. At 180° C., the conversion of C5-C7 olefins is greater than 80% by mass.

[0473] Olefins can react with aromatics either by alkylation or oligomerization. It was observed that aromatics conversion varied from 28% by mass at 180° C. to 33% by mass at 220° C. (see Table 7 below). Aromatics were indeed present in the FBP fraction at 170° C., which means that alkylation is indeed occurring.

[0474] [Table 7]

[0475] The 145+ cut was hydrotreated using NiMo catalyst under the conditions described above. The properties of the hydrotreated cut are given in Table 8 and the detailed composition is summarized in Table 9.

[0476] [Table 8]

[0477] [Table 9]

[0478] Additional examples of step (a) conversion Catalyst Examples 14.25g of 85% by mass H 3 PO 4 and 225 ml of demineralized water, 75 g of sample A were introduced. The suspension was stirred under reflux for 4 hours.

[0479] Next, 4.9 g of CaCO was added to the suspension. 3 Heating of the solution was stopped while maintaining stirring of the mixture until a temperature below 30° C. was reached. As a result, suspension S1 was obtained.

[0480] Then, 450 g of low-sodium silica sol (34% by weight SiO in water) was added under stirring at ambient temperature for 30 min. 2 , 200 ppm Na) and 4.5 g H 3 PO 4 (85% by weight) was mixed to prepare a solution, resulting in suspension S2.

[0481] The suspensions S1 and S2 were then mixed to form a solution, which was then stirred for 1 hour and spray-dried, resulting in catalyst X.

[0482] Step (a) Carrying out the conversion Tests 1 to 3 described below were carried out in a ceramic passivated fixed bed fed with 1.31 g of the above catalyst X mixed with SiC. Quartz wool was used to keep the catalyst bed in place. At the top of the reactor, alcohol was introduced at a flow rate of 0.05 moles of methanol per hour per gram of catalyst. Nitrogen diluent may be added. In that case, the alcohol partial pressure is different from the total pressure, as shown in the table below.

[0483] The total pressure was varied (1.3 bara, 5 bara, and 10 bara), as was the temperature (450° C., 500° C., 550° C.).

[0484] Prior to catalytic testing, the catalyst was heated in N until the reaction temperature was reached. 2 (5Nl / h).

[0485] The analysis of the products was carried out using a combination of on-line analysis of the reactor output stream via gas phase microchromatography and off-line analysis of the liquid produced by the reaction by gas phase chromatography coupled with a flame ionization detector (GC-FID). The micro gas chromatograph used has four modules: -O 2 , N 2 , H 2 , CO and CH 4 Module 1 equipped with a molecular sieve capillary column for the separation of; - MeOH, dimethyl ether (DME), C1-C3 hydrocarbons and CO 2 module 2 ("Plot Q") equipped with a polystyrene-divinylbenzene grafted capillary column for the separation of; - module 3 equipped with an alumina capillary column for separating C2-C5 hydrocarbons; - Module 4 with a fused silica type column or a "Stabilwax" column for separating water and C6+ hydrocarbons.

[0486] The results are illustrated in the table below which provides the mass percentage selectivities for the methanol stream.

[0487] [Table 10]

[0488] Graphs (a)-(e) shown in FIG. 10 demonstrate the effect of alcohol partial pressure and temperature on the selectivity for ethylene (graph (a)), propylene and butenes (graph (b)), C5+ olefins (graph (c)), C5+ paraffins (graph (d)), and the aromatics benzene, toluene, and xylenes (graph (e)).

[0489] In each graph, the hourly space velocity of methanol relative to the mass of the catalyst, WHSV(MeOH), is 1.6 / h. The black triangles correspond to a temperature of 550° C., the white triangles correspond to a temperature of 500° C., and the white diamonds correspond to a temperature of 450° C.

[0490] These examples illustrate that, in the above-mentioned temperature range, specifically 300°C-600°C, specifically 330°C-550°C, specifically 350°C-500°C or 410°C-580°C, at a methanol partial pressure in the above-mentioned range, specifically 100 kPa-5 MPa, preferably around 100 kPa-1.0 MPa, a mixture of paraffins, olefins, aromatics and water is produced in the conversion step (a) according to the present invention using a phosphorus-modified zeolite catalyst, and the ratio of the mass of C3+ olefins to the total mass of olefins is 0.8 or more.

[0491] The same is true when starting with an ethanol flow, as illustrated by the following three tests 1E, 1F and 1G.

[0492] The test conditions and selectivity results are as follows:

[0493] [Table 11]

[0494] In one variant, a 1H test was also carried out in a fluidized bed on 7.2 g of catalyst X with a methanol stream to be converted. To ensure adequate fluidization properties, the nitrogen flow rate was set at 23.8 mL / min and alcohol was co-injected at the bottom of the reactor.

[0495] The test conditions and selectivity results are as follows:

[0496] [Table 12]

[0497] Prior to catalytic testing, the catalyst was heated in N until it reached reaction temperature. 2 (5Nl / h).

[0498] Example of additional oligomerization step On the one hand, 40 mL of ZSM-5 catalyst (80 wt.% MFI and 20% alumina binder) with a size of 2 mm–4 mm, diluted with 40 mL of inert material (SiC with a size of 1 mm–1.4 mm) was fed into a fixed-bed tubular reactor with an inner diameter of 16 mm.

[0499] Prior to testing, the catalyst was activated at 400°C (60°C / h) for 2 hours under 160NL / h nitrogen, after which the temperature was reduced to 40°C before the feed was introduced and the temperature was increased back to the test conditions.

[0500] On the other hand, 40 mL of BEA-based catalyst (80 wt.% BEA and 20% alumina binder) with a size of 2–4 mm, diluted with 40 mL of inert material (0.21 mm SiC) was fed into a fixed-bed tubular reactor with an inner diameter of 16 mm.

[0501] Prior to testing, the catalyst was activated at 400°C (60°C / h) for 2 hours under 160NL / h nitrogen, after which the temperature was reduced to 40°C before the feed was introduced and the temperature was increased back to the test conditions.

[0502] The feed composition used to test the two catalysts at 55 barg was as follows:

[0503] [Table 13]

[0504] Estimate the yield profile based on the following cut points: Light olefin: IBP~80℃ Naphtha: 80℃~145℃ Jet fuel: 145℃~300℃ Diesel:>300℃

[0505] Example OLIGO-1 - Performance results obtained with ZSM-5 catalyst

[0506] The feedstock was treated and processed under the following operating conditions: 55 barg, 1 h -1 Liquid Hourly Space Velocity (LHSV), and temperature of 180℃~240℃.

[0507] The conversion of light olefins (C3-C6 olefins) is shown in Figure 11. It is greater than 90% by weight for temperatures above 220°C. At 220°C, the conversion of butenes and hexenes decreases over time.

[0508] From a yield structure perspective, once the contribution of the diluent (in this case n-heptane) is subtracted, the yield at 220° C. is as follows:

[0509] [Table 14]

[0510] Example OLIGO-2 - Performance results obtained with BEA-based catalyst The feedstock was treated and processed under the following operating conditions: 55 barg, liquid hourly space velocity LHSV of 1 / h, and temperature of 150°C to 220°C.

[0511] The conversion of C5 and C6 light olefins is shown in Figure 12. It is greater than 90 wt% for temperatures above 200°C.

[0512] Olefins can react with aromatics either by alkylation or oligomerization. The aromatics are partially converted at isotherms, and the degree of conversion of these aromatics decreases over time, as illustrated in FIG.

[0513] From a yield structure perspective, once the contribution of the diluent (in this case n-heptane) is subtracted, the yield at 200° C. is as follows:

[0514] [Table 15]

[0515] These results are based on the oligomerization and alkylation conditions defined above, in particular for temperatures between 150° C. and 400° C., preferably between 180° C. and 350° C., and even more preferably between 180° C. and 290° C.; -1 ~20h -1 , preferably 0.5h -1 ~10h -1 , and even more preferably 0.8 h -1 ~5h -1 of the feed weight hourly space velocity (WHSV); using a zeolite type catalyst, it is possible to very efficiently convert the feedstock obtained at the end of step (a) into significant amounts of jet fuel. [Explanation of symbols]

[0516] 10, 90, 100, 110, 120, 140, 160 Fuel Production Plant 12 Conversion stage 14 C1-C6 alcohol stream 16 Mixtures containing paraffins, olefins, aromatics and water 18 separation stage 19 Water Deficient Mixture 22, 22A Oligomerization stage 22, 22B Alkylation stage 24 Hydrocarbon stream to be hydrogenated 26 Hydrogenation Stage 28 fractionation stages 30 Hydrogenated Hydrocarbon Streams 34 Jet Fuel Fraction 36 Diesel Fraction 40 water 60 C1-C2 Hydrocarbon Fraction 62 C3+ Hydrocarbon Fraction 80 C3-Hydrocarbon Fraction 82 C4+ Hydrocarbon Fraction 126 C7-Hydrocarbon Fraction 128 C8+ Hydrocarbon Fraction 142 C3-C5 Hydrocarbon Fraction 144 C6+ Hydrocarbon Fraction 146 Oligomerization Reactor Products 152 Alkylation reactor product 162 C8~C16 Hydrocarbon Fraction 164 C17+ Hydrocarbon Fraction

Claims

1. 1. A method for producing jet fuel, comprising: (a) converting a C1-C6 alcohol stream (14) to produce a mixture (16) containing paraffins, olefins, aromatics and water; (b) separating water (40) from the mixture (16) to form a water-deficient mixture (19); (c) oligomerizing olefins from the water-depleted mixture (19); (d) alkylating aromatic compounds from the water-depleted mixture (19); (e) forming a hydrocarbon stream (24) to be hydrogenated from at least a portion of the olefins oligomerized in step (c) and at least a portion of the aromatic compounds alkylated in step (d); (f) hydrogenating the hydrocarbon stream (24) to be hydrogenated to form a hydrogenated hydrocarbon stream (30); (g) recovering at least one jet fuel fraction (34) from the hydrogenated hydrocarbon stream (30); A method comprising: 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 0.8 or more; method.

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

3. 3. The process according to claim 1 or 2, comprising separating at least a portion of the olefins oligomerized in step (c) and / or at least a portion of the aromatics alkylated in step (d) into a C7- hydrocarbon fraction (126) and a C8+ hydrocarbon fraction (128), the C7- hydrocarbon fraction (126) being at least partially recycled to the oligomerization step (c) for oligomerizing the olefins and / or the alkylation step (d) for alkylating the aromatics, and the hydrocarbon stream to be hydrogenated (24) being formed by at least a portion of the C8+ hydrocarbon fraction (128).

4. 4. The process according to any one of claims 1 to 3, wherein the olefin oligomerization step (c) and the aromatic compound alkylation step (d) are carried out jointly in the same reactor.

5. 5. The process according to claim 4, comprising a separation step for separating the water-depleted mixture (19) into a C1-C2 hydrocarbon fraction (60) and a C3+ hydrocarbon fraction (62), wherein at least a portion of the C1-C2 hydrocarbon fraction (60) is conveyed to a steam cracker for extracting an ethylene stream therefrom; and at least a portion of the C3+ hydrocarbon fraction (60) is sent to step (c) for oligomerizing olefins and step (d) for alkylating aromatics; and a portion of the C1-C2 hydrocarbon fraction (64) is optionally recycled to the conversion step (a) for converting a C1-C6 alcohol stream.

6. 6. The process of 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 step (c) for oligomerizing olefins and to step (d) for alkylating aromatics.

7. 4. The process according to claim 1, wherein the oligomerization step (c) for oligomerizing olefins is carried out in an oligomerization reactor and the alkylation step (d) for alkylating aromatic compounds is carried out in an alkylation reactor separate from the oligomerization step (c) for oligomerizing olefins.

8. or - separating the water-depleted mixture (19) into a C1-C2 hydrocarbon fraction (60), a C3-C5 hydrocarbon fraction (142) and a C6+ hydrocarbon fraction (144), the C1-C2 hydrocarbon fraction (60) and the C6+ hydrocarbon fraction (144) being sent at least in part to an alkylation step (d) in an alkylation reactor and the C3-C5 hydrocarbon fraction (142) being sent to an oligomerization step (c) in an oligomerization reactor, or 8. The method of claim 7, 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 fed at least partially to an alkylation step (d) in an alkylation reactor and the C4-C5 hydrocarbon fraction being fed at least partially to an oligomerization step (c) in an oligomerization reactor.

9. 9. The process according to claim 8, with reference to claim 3, wherein the oligomerization reactor product containing oligomerized olefins and the alkylation reactor product containing alkylated aromatics (152) are separated into a C8+ hydrocarbon fraction (128) and a C7- hydrocarbon fraction (126), the C7- hydrocarbon fraction (126) being at least partially recycled to step (c) in the oligomerization reactor, and at least a portion of the C8+ hydrocarbon fraction (128) forming the hydrocarbon stream (24) to be hydrogenated.

10. 9. The process according to claim 8, wherein the oligomerization reactor product (146) and the alkylation reactor product (152) containing alkylated aromatics are separated into a C7- hydrocarbon fraction (126), a C8 to C16 hydrocarbon fraction (162) and a C17+ hydrocarbon fraction (164), at least a portion of the C8 to C16 hydrocarbon fraction (162) forming the hydrocarbon stream to be hydrogenated (24) and the C17+ hydrocarbon fraction (164) being at least partially recycled to the conversion step (a) for converting the C1 to C6 alcohol stream.

11. 11. The method according to any one of claims 1 to 10, wherein from 10% by weight to 90% by weight of the aromatics contained in the hydrocarbon stream to be hydrogenated (24), preferably from 30% by weight to 80% by weight of the aromatics contained in the hydrocarbon stream to be hydrogenated, are hydrogenated to cycloparaffins in step (f).

12. 12. The process according to any one of claims 1 to 11, wherein the C1-C6 alcohol stream (14) contains at least 50% methanol and the converting step (a) for converting the C1-C6 alcohol stream comprises the addition of an alcohol stream containing C2-C6 alcohols between two conversion catalyst beds of the converting step (a).

13. 13. The method according to any one of claims 1 to 12, comprising, after the hydrogenation step (f) for hydrogenating the hydrocarbon stream (24) to be hydrogenated, separation of the hydrogenated hydrocarbon stream (30) into at least a jet fuel fraction (34) and a diesel fraction (36).

14. 14. The method of any one of claims 1 to 13, wherein at least a portion (40a) of the water (40) from the mixture (16) separated in the separation step (b) is recycled to the conversion step (a).

15. 15. The process according to claim 14, wherein the conversion step (a) is carried out in a series of fixed catalyst beds and a portion (40a) of the water (40) from the mixture (16) recycled to the conversion step (a) is introduced upstream of the series of fixed catalyst beds or between two fixed catalyst beds; or the conversion step (a) is carried out in at least one fluidized catalyst bed and a portion (40a) of the water (40) from the mixture (16) recycled to the conversion step (a) is introduced into the fluidized catalyst bed, optionally together with a C4-hydrocarbon stream, preferably together with at least a portion of the C1-C2 hydrocarbon fraction (60) obtained by separation of the water-depleted mixture (19).

16. 16. The process according to any one of claims 1 to 15, wherein the conversion step (a) is carried out in the presence of a conversion catalyst comprising a phosphorus-modified zeolite having a partial ALPO structure, or in the presence of a conversion catalyst comprising a B-modified zeolite.

17. for powering at least one aircraft engine, (i) in pure form, or (ii) in the form of a mixture with jet fuel obtained as a result of the distillation of petroleum; 17. Use of a jet fuel fraction (34) produced by carrying out the process according to any one of claims 1 to 16 in

18. In a jet fuel production plant (10, 90, 100, 110, 120, 140, 160), a conversion stage (12) for converting a C1-C6 alcohol stream (14) to form a mixture (16) containing paraffins, olefins, aromatics and water; a separation stage (18) for separating water (40) from the mixture (16) to form a water-depleted mixture (19); - an oligomerization stage (22; 22A) for oligomerizing olefins from the water-depleted mixture (19); an alkylation stage (22; 22B) for alkylating aromatic compounds from the water-depleted mixture (19); a formation stage for forming a hydrocarbon stream (24) 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 hydrogenation stage (26) for hydrogenating the hydrocarbon stream (24) to be hydrogenated to form a hydrogenated hydrocarbon stream (30); a fractionation stage (28) for recovering at least one jet fuel fraction (34) from the hydrogenated hydrocarbon stream (30); 1. A jet fuel production plant comprising:

1. A jet fuel production plant, wherein the conversion stage (12) is configured to produce a mixture containing paraffins, olefins, aromatics and water, wherein the ratio of the mass of C3+ olefins to the total mass of olefins is 0.8 or greater.

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