Process for obtaining hydrocarbons, and associated installation

Abstract

Process for Obtaining Hydrocarbons, and Associated Installation. This process comprises the following steps: (a) conversion of a C1 to C6 alcohol stream (14) to produce a mixture (16) containing paraffins, olefins, aromatics, and water; (b) separation of the water (40) from the mixture (16) to form a water-depleted mixture (19); The water-depleted mixture (19) is separated and / or treated to recover hydrocarbons. It comprises the addition, in step (a) of converting the C1 to C6 alcohol stream, of a stream (182) containing carbon dioxide, and the simultaneous conversion of carbon dioxide to carbon monoxide in step (a) of converting the C1 to C6 alcohol stream. Figure for the abstract: Figure 1

Description

Title of the invention: Process for obtaining hydrocarbons, and associated installation technical field

[0001] The present invention relates to a process for obtaining hydrocarbons, comprising the following steps:

[0002] (a) conversion of an alcohol stream in Cl to C6 to produce a mixture containing paraffins, olefins, aromatics, and water;

[0003] (b) separation of water from the mixture to form a water-depleted mixture;

[0004] the water-depleted mixture being separated and / or treated to recover hydrocarbons.

[0005] The present invention relates, for example, to the field of the preparation and use of liquid fuels, in particular jet fuel or renewable aviation fuels. The present invention also relates to the production of olefins (in particular ethylene, propylene, butenes and pentenes, hexenes), minimizing the production of paraffins from renewable resources based on the conversion of a C6 alcohol stream to olefins.

[0006] Due to the scarcity of fossil resources and increasing environmental concerns, particularly with the aim of reducing greenhouse gas emissions, the use of alternative molecules with a lower carbon footprint is increasingly sought to replace molecules of fossil origin.

[0007] Renewable fuels derived from biological matter or from carbon dioxide transformed in the presence of decarbonized or electrolytic hydrogen (designated by the English term "E-fuels") are an alternative to conventional fossil fuels.

[0008] Conventional jet fuels can be blended with renewable feedstocks as defined by standard D7566-21, thereby enabling the production of alternative aviation fuels. Examples of renewable feedstocks for aviation fuels that can be blended with fossil jet fuel include:

[0009] - synthetic paraffinic kerosenes [SPK], produced by processes such as the Fischer-Tropsch process;

[0010] - synthetic paraffinic kerosenes produced by the "Alcohol-to-Jet" route (transformation of alcohol into isoparaffini kerosene) [ATJ-SPK];

[0011] - synthetic isoparaffins produced by hydrotreating iso- intermediates olefinics, produced from fermented sugars [SIP-HFS];

[0012] - synthetic aromatic kerosenes obtained by alkylation of light aromatics from non-petroleum source [SPK / A];

[0013] - synthetic kerosenes obtained from the hydrothermal conversion of esters of fatty acid and fatty acid;

[0014] - paraffinic kerosenes [SPK] obtained from hydrocarbons, esters and of hydrotreated fatty acids.

[0015] Currently, most of these renewable aviation fuel bases cannot be used on their own due to their composition being very different from that of fossil fuels. This difference in composition poses, in particular, compatibility problems with the materials of the components with which the fuel comes into contact. In this respect, the absence of aromatic compounds in the majority of available renewable aviation fuel bases can lead to compatibility problems with materials, and especially with certain seals. STATE OF THE ART

[0016] EP2123736 describes a method for producing diesel fuel using a A fuel charge in the form of C5-Cl alcohols, which may be wholly or partly of biogenic origin, wherein a synthetic hydrocarbon is oligomerized or hydrogenated from a mixture of olefinic hydrocarbons obtained at least partially by dehydration of C5-Cl alcohols, with an odd proportion of olefins and iso-olefins. After further hydrogenation and rectification, an aviation fuel is formed with a freezing point of -47 °C or lower.

[0017] US20210078921 describes the conversion of methanol into gasoline that can be carried out using a heavy gasoline treatment, followed by a separation operation.

[0018] US4543435 describes a method for converting an oxygenated feed charge comprising methanol, dimethyl ether or analogue in liquid hydrocarbons, comprising contacting the feed with a zeolite catalyst in a primary catalytic stage at high temperature and moderate pressure to convert the feed into hydrocarbons comprising C2-C4 olefins and C5+ hydrocarbons.

[0019] EP1844125 relates to a process for producing synthetic fuels These methods, according to which, in a first step, a gaseous mixture comprising methanol and / or dimethyl ether and / or another oxygenated molecule, as well as water vapor, is transformed into olefins preferably having between 2 and 8 carbon atoms, at temperatures between 300 and 500 °C, and in a second step, the resulting olefin mixture is oligomerized at higher pressure into olefins. upper olefins consisting essentially of more than 5, preferably between 10 and 20 carbon atoms. According to this process, a) the production of olefins in the first stage is carried out in the presence of a gas stream consisting essentially of saturated hydrocarbons, separated from the product stream of the second stage and returned to the first stage and b) the production of olefins in the second stage is carried out in the presence of a steam stream, which is separated from the product stream of the first stage of the process and returned to the first stage of the process.

[0020] EP2147082 describes a process for producing synthetic fuels from a a mixture, containing hydrogen and oxygenated compounds such as methanol and / or dimethyl ether, in a first step, the mixture is reacted on a catalyst, to obtain a hydrocarbon product containing olefins having, preferably, 2 to 8 carbon atoms, and, in a second step, the hydrocarbon product thus obtained is oligomerized into long-chain olefins, from which it is possible to obtain products which are gasoline and diesel.

[0021] WO2011061198 describes a process for producing hydrocarbons in the form gasoline, by conversion of synthesis gas, to obtain an oxygen-containing compound, such as methanol and / or dimethyl ether, in a first converter, and by further conversion into hydrocarbons in a second converter.

[0022] EP2940103 describes a process for preparing biofuels using ethanol by conversion of ethanol in mixture with hydrocarbons, in a catalytic process on a bed of zeolite-type aluminosilicate, preferably in the presence of a hydrogen form of the zeolitic catalyst.

[0023] EP2720990 describes a process for converting an alcohol into a hydrocarbon, the process comprising bringing said alcohol, as a component of an aqueous solution at a concentration of not more than 20%, into contact with a metal-loaded zeolite catalyst at a temperature of at least 100°C and up to 550°C, wherein said alcohol may be produced by a fermentation process and is selected from ethanol, butanol, isobutanol, or a combination thereof, said metal includes vanadium, and said metal-loaded zeolite catalyst is catalytically active to convert said alcohol into said hydrocarbon.

[0024] EP3795658 describes processes that reduce energy and water consumption in fuel production processes using feedstocks containing renewable alcohol. The alcohol is converted directly into hydrocarbon transport fuels via a catalytic process, with heat transferred between intermediate liquids to reduce thermal energy consumption. Overall water consumption is reduced by recovering water from the catalytic process and by lowering the water temperature, thus enabling reduce losses through evaporation.

[0025] US20160090333 describes methods for producing hydrocarbons in the range Aviation fuels are produced from biorenewable sources, such as the oligomerization of biorenewable C3-C8 olefins, for example, derived from C3-C8 alcohols produced by biomass fermentation. The production of aviation-grade hydrocarbons is increased by using an additional oligomerization zone to oligomerize gasoline separated from the effluent of a primary oligomerization zone in which biorenewable C3-C8 olefins have first undergone oligomerization.

[0026] WO201145535 describes a process for producing distillate from a feed of Heteroatomic organic compounds comprising at least one heteroatom selected from oxygen, sulfur, or halogen, alone or in combination, wherein the feedstock treatment includes at least one step of converting the heteroatomic organic compounds into olefins carried out in a first conversion zone, and, in at least a second oligomerization zone, a step of oligomerizing olefins originating at least in part from the conversion zone, in the presence of at least 0.5% by mass of oxygenated compounds, in order to produce a distillate. This process improves the distillate yield by enabling a higher oligomerization rate compared to the oligomerization of the same feedstock under the same reaction conditions.

[0027] WO2022 / 063994 describes a method for obtaining jet fuel comprising a The conversion step involves a stream of oxygenated compounds being carried out at reduced temperatures (e.g., below 350°C), pressures on the order of 5 to 10 bar, and high hourly spatial velocities (6 h₁ to 10₀'), followed by a combined oligomerization and hydrogenation step in the same reactor. This process tends to keep the levels of ethylene and aromatics produced during the conversion very low.

[0028] These processes are not fully optimized in the conversion of oxygenated compounds to produce a suitable mixture of olefins at the end of the conversion. Summary of the invention

[0029] An object of the invention is therefore to provide a hydrocarbon process derived exclusively or at least partly from renewable feedstocks, which is efficient and productive, particularly in the conversion of feedstocks into olefins.

[0030] To this end, the invention relates to a method for obtaining the aforementioned type, characterized by the addition, at step (a) of converting the alcohol stream into Cl₂ to C₆, of a stream containing carbon dioxide, and the joint conversion of carbon dioxide into carbon monoxide during step (a) of converting the alcohol stream into Cl to C6.

[0031] The method according to the invention may comprise one or more of the following features, taken individually or in any technically feasible combination:

[0032] - the current containing carbon dioxide comprises more than 5% by mass of carbon dioxide;

[0033] - the mass ratio of carbon dioxide to alcohols in Cl to C6 in the charge provided in step (a) of conversion is between 5% and 75%;

[0034] - at least 2% by moles of the carbon dioxide contained in the stream containing carbon dioxide is converted to carbon monoxide during step (a) of converting the alcohol stream to Cl to C6;

[0035] - the conversion step (a) is carried out using at least one catalyst including molecular sieves containing at least 10 oxygen atom (10-MR) pores or larger in their microporous structure

[0036] - the catalyst for implementing step (a) of conversion comprises a phosphorus-modified zeolite, having in particular a P content of at least 0.05% by mass and preferably between 0.3% by mass and 7% by mass, a composite catalyst comprising at least 0.1% by mass of silicate, or a molecular sieve modified by phosphorus and by an alkaline earth or rare earth metal (modified molecular sieve MP), advantageously having an M / P molar ratio in the molecular sieve of less than 1;

[0037] - the catalyst for implementing step (a) of conversion has been modified by addition of one or more metals selected from the metals of group IIB, in particular Zn, of group IIIB, in particular Ga, the transition metals of group VIIIB in particular Fe and / or Ni and / or Pt, of group VIB, in particular Mo, of group IB in particular Cu and / or Ag or of the lanthanide group in particular La;

[0038] - the catalyst for carrying out step (a) of conversion is a zeolite modified by phosphorus having partially an ALPO structure or is a zeolite modified by B;

[0039] - the process includes a step of separating the water-depleted mixture into a fraction of hydrocarbons in C1-C2, and in a fraction of hydrocarbons in C3+, a part of the fraction of hydrocarbons in C1-C2 being advantageously recycled in step (a) of conversion of the alcohol stream in Cl to C6;

[0040] - carbon dioxide, carbon monoxide and hydrogen present in the C1-C2 hydrocarbon fractions are separated from hydrocarbons, particularly by cryogenic distillation, membrane separation or alternating pressure adsorption and their combinations, carbon dioxide, carbon monoxide and hydrogen being then advantageously recycled to a preliminary stage of alcohol synthesis intended to form the C1-C6 alcohol stream, notably by gas fermentation. synthesis and by catalytic conversion of synthesis gas to produce methanol;

[0041] - the process comprises the following steps:

[0042] (c) oligomerization of olefins from the water-depleted mixture;

[0043] (d) alkylation of aromatics from the water-depleted mixture;

[0044] (e) formation of a hydrocarbon stream from at least a portion of the oligomerized olefins in step (c) and at least some of the alkylated aromatics in step (d);

[0045] -The process comprises the following steps:

[0046] (f) hydrogenation of the hydrocarbon stream formed in step (e) to form a hydrogenated hydrocarbon stream;

[0047] (g) recovery of at least a fraction of jet fuel from the stream of hydrogenated hydrocarbons;

[0048] - the jet fuel fraction comprises between 2% by volume and 30% by volume of C8+ aromatics, preferably between 8% by volume and 25% by volume of C8+ aromatics;

[0049] - in the mixture of paraffins, olefins, aromatics and water produced in the step (a) of conversion, the ratio of the mass of C3+ olefins to the total mass of olefins is greater than or equal to 0.8;

[0050] - the C2 to C6 alcohol stream is obtained by fermentation of biomass or gas synthesis, or / and the flow of alcohol in Cl to C6 is obtained by catalytic conversion of carbohydrates, carbon monoxide or carbon dioxide in the presence of hydrogen;

[0051] - step (f) of hydrogenating the hydrocarbon stream to be hydrogenated is carried out separately and downstream of step (c) oligomerization of olefins from the water-depleted mixture and step (d) alkylation of aromatics from the mixture;

[0052] - the catalyst for implementing step (a) of conversion comprises a zeolite with pores of 10 or more oxygen atoms, modified by adding B before, after or simultaneously with the final catalyst formulation step;

[0053] - the catalyst for implementing step (a) of conversion has a ratio atomic Si / Al, measured by chemical analysis in particular by NMR, taking into account only the Al which are part of the network structure of the molecular sieve between 4 and 500, preferably between 5 and 200, or more preferably between 12 and 150;

[0054] - the process includes stripping at least part of the water separated in step (b) separation to produce a stream of extracted hydrocarbons, advantageously recycled in separation step (b), and a stream of treated water.

[0055] The invention also relates to the use of a jet fuel fraction produced at starting from hydrocarbons obtained by implementing the manufacturing process as defined above,

[0056] (i) pure, or

[0057] (ii) in a mixture with jet fuel resulting from the distillation of petroleum and / or another renewable source,

[0058] to power at least one aircraft engine.

[0059] Advantageously, the jet fuel fraction produced comprises between 2% by mass and 30% by mass of aromatics, in particular between 6% by mass and 20% by mass of aromatics.

[0060] In particular, the jet fuel fraction comprises between 2% by mass and 30% by mass of aromatics having at least 8 carbon atoms, in particular between 6% by mass and 20% by mass of aromatics having at least 8 carbon atoms.

[0061] Preferably, more than 50% by mass of the aromatics contained in the jet fuel fraction are monoaromatics having 8 to 14 carbon atoms.

[0062] It advantageously comprises between 5% by mass and 20% by mass of cycloparaffins and at least 50% by mass of isoparaffins.

[0063] The process according to the invention thus makes it possible to obtain a renewable aviation fuel that is totally substitutable (“drop-in”) or compatible with the fuel and engine systems of current aircraft.

[0064] The invention also relates to a hydrocarbon installation, comprising:

[0065] - a stage for converting an alcohol stream in Cl to C6 to produce a mixture containing paraffins, olefins, aromatics, and water;

[0066] - a stage for separating water from the mixture to form a mixture depleted in water ;

[0067] - at least one stage for separating and / or treating the water-depleted mixture to recover hydrocarbons,

[0068] characterized in that the alcohol to Cl to C6 stream conversion stage comprises at least one supply conduit for a stream containing carbon dioxide, the alcohol to Cl to C6 stream conversion stage being configured to convert carbon dioxide from the carbon dioxide-containing stream into carbon monoxide jointly with the alcohol to Cl to C6 stream conversion.

[0069] In one variant, the installation includes at least one recycling conduit at the conversion stage, for at least part of the water from the mixture separated at the water separation stage. DETAILED DESCRIPTION

[0070] The terms "including" and "includes" as used herein are synonymous with "including", "includes" or "contains", "containing", and are inclusive or boundless and do not exclude additional features, elements, or unspecified method steps.

[0071] The expressions % by mass and % by mass have an equivalent meaning and refer to the proportion of the mass of a product relative to 100 g of a composition comprising it.

[0072] Boiling points as mentioned herein are measured at atmospheric pressure, unless otherwise specified. An initial boiling point (hereinafter "IBP") is defined as the temperature at which the first vapor bubble forms. An final boiling point (hereinafter "FBP") is the highest temperature attainable during distillation. At this temperature, no more vapor can be transported to a condenser. The determination of the initial and final boiling points relies on techniques known in the trade, and several methods adapted according to the distillation temperature range are applicable, for example, NF EN 15199-1 (version 2020) or ASTM D2887 for measuring the boiling points of petroleum fractions by gas chromatography, ASTM D7169 for heavy hydrocarbons, and ASTM D7500, D86, or DI 160 for distillates.

[0073] By default, the expression flux, current, fraction, etc. "in Cn to Cm" designates a flux, a current, a fraction, etc. having a majority quantity (for example more than 50% in moles) of compounds having between n and m carbon atoms.

[0074] The expression flux, current, fraction, etc. "in Cn+" designates a flux, a current, a fraction, etc. having a majority quantity (for example more than 50% in moles) of compounds having n carbon atoms or more than n carbon atoms.

[0075] The expression flux, current, fraction, etc. "in Cn-" designates a flux, a current, a fraction, etc. having a majority quantity (for example more than 50% in moles) of compounds having n carbon atoms or less than n carbon atoms.

[0076] Unless otherwise indicated, the percentages used are percentages by mass, and the pressures are absolute pressures. Obtaining the alcohol flow in Cl to C6

[0077] The alcohol stream in Cl to C6 contains predominantly alcohols such as methanol, ethanol, propanols (n-propanol, i-propanol) butanols (n-butanol, i-butanol), pentanols (n-pentanols, i-pentanol) and hexanols.

[0078] It may include minor amounts of C6+ alcohols, and / or oxygenated compounds such as methyl ether; dimethyl ether; diethyl ether; diisopropyl ether; formaldehyde; dimethyl carbonate; dimethyl ketone; acetic acid; furans, tetrahydrofurans and mixtures thereof.

[0079] The C6 Cl alcohol stream forming the process feed advantageously comprises more than 80% by mass of C6 Cl alcohols, preferably more than 90% by mass of C6 Cl alcohols. It advantageously comprises more than 50% by mass of methanol, specifically more than 80% by mass of methanol.

[0080] In one variant, in the case where the conversion is carried out using a plurality of reaction zones having a fixed catalytic bed, a flow of alcohol from C2 to C6 is further advantageously added between two conversion reaction zones of conversion step (a), as will be described below.

[0081] The ratio of the mass flow rate of the alcohol flux from C2 to C6 added between two reaction zones to the mass flow rate of the alcohol flux from Cl to C6 entering the first reaction zone is, for example, less than 0.5, in particular between 0.05 and 0.5±

[0082] Preferably, the C6 Cl alcohol flux and advantageously the additional C2 to C6 alcohol flux are obtained from a renewable source such as biomass (including its constituents and derivatives) or from carbon oxide or carbon dioxide, optionally captured and hydrogen advantageously produced from renewable energy sources such as solar, wind, geothermal, wave or current energy and / or energy whose production does not generate carbon dioxide such as nuclear energy.

[0083] The production routes for the alcohols intended to form the C6 Cl alcohol stream are, for example, but not limited to:

[0084] - anaerobic fermentation of sugars from biomass, in particular to obtain ethanol;

[0085] - catalytic reaction of hydrogen with carbon dioxide or monoxide carbon to obtain, in particular, methanol, ethanol and other alcohols;

[0086] - catalytic reaction of hydrogen with carbohydrates to obtain in particular alcohols Cl to C6;

[0087] - ABE fermentation (ethanol, acetone, butanol), to obtain ethanol and n- butanol;

[0088] - anaerobic fermentation of sugars from biomass, in particular to obtain propanol (iso or n), butanol (iso or n) or isoamyl alcohol;

[0089] - anaerobic fermentation of a mixture containing at least monoxide carbon, carbon dioxide and hydrogen to obtain, in particular, ethanol, propanol (iso or n), butanol (iso or n) or isoamyl alcohol.

[0090] For obtaining alcohols, in particular ethanol of renewable origin by fermentation in an embodiment, the alcohol, in particular ethanol of renewable origin, can be obtained by ethanolic fermentation in a bioreactor containing a culture of one or more microorganisms.

[0091] Ethanol from renewable sources can then advantageously be obtained by:

[0092] - anaerobic fermentation of a sugar-rich substrate derived from biomass, or

[0093] - anaerobic fermentation of a gas comprising CO, which may originate from biomass or not.

[0094] For anaerobic fermentation, ethanol can thus be produced by anaerobic fermentation of a sugar-rich substrate derived from biomass.

[0095] Sugars are composed of chains of 6 or 5 carbons, such as glucose, sucrose (dimer of glucose and fructose), xylose and arabinose.

[0096] This substrate may, for example, comprise or be derived directly from agri-food plants, sugar cane, sugar beet, sweet sorghum, or by depolymerization of starch from maize, wheat, barley, rye, sorghum, triticale, potato, sweet potato, cassava, and / or cellulose and hemicellulose from lignocellulosic biomass.

[0097] The sugar-rich substrate can also be derived from lignocellulose biomass by a treatment comprising (i) a step of separating the lignin, cellulose and hemicellulose contained in the lignocellulose biomass, followed by (ii) a step of converting the cellulose and / or hemicellulose into sugars.

[0098] Obtaining this type of substrate from lignocellulosic biomass is well known to those skilled in the art. Lignocellulosic biomass consists essentially of cellulose, hemicellulose, and lignin. This biomass comes from agricultural and forestry residues or by-products of wood or crop processing, whether woody or herbaceous plants. This lignocellulosic biomass can also include distillers' grains and allow for the production of ethanol as described in document EP2675778.

[0099] The first step (i) is a pretreatment step that allows the lignocellulosic matrix to be separated and the cellulose and hemicellulose to be released from the complex formed with the lignin by means of one or more pretreatments. Known pretreatments include steam pretreatment (or steam blasting), hot water pretreatment (hydrothermal), ammonia blasting (AFEX), acid pretreatment, and alkaline pretreatment. Steam blasting treatment consists of treating the biomass, preferably previously shredded or ground, with high-pressure saturated steam at temperatures of approximately 160 to 240°C and pressures of 0.7 to 4.8 MPa. The efficiency of steam treatment can be improved by the addition of H2SO4, CO2, or SO2 as a catalyst.In AFEX pretreatment, biomass is contacted with anhydrous liquid ammonia in a 1:1 to 2:1 ratio (1 to 2 kg of ammonia / kg of dry biomass) for 10 to 60 min at 60–90°C and pressures above 3 MPa. Hydrothermal pretreatment is similar to steam explosion but uses liquid water at high temperatures instead of steam. In acid pretreatment, typically in the presence of dilute acid, an aqueous suspension of the cellulosic substrate is heated to the desired temperature and pretreated with [the appropriate solution / method]. Preheated sulfuric acid (concentrations <4% by mass) is used in a stainless steel reactor, and the treatment is carried out at a temperature of 140 to 215°C. The residence time varies from a few seconds to a few minutes depending on the treatment temperature. Lime pretreatment is an inexpensive alkaline physicochemical treatment that improves the digestibility of cellulosic biomass. Using 0.1 g of Ca(OH)2 / g of biomass, the treatment can be performed over a wide temperature range from 25 to 130°C. The Organosolv process, which is a process for the delignification and / or saccharification of cellulosic materials and plant cultures, can also be used. In general, the Organosolv process involves the use of a mixture of water and a solvent such as alcohols or ketones, and sometimes other nonpolar solvents, as well as an acidic compound to facilitate hydrolysis.A process of this type is described, for example, in document US4470851A.

[0100] Step (ii) is a step of conversion of cellulose and / or hemicellulose into sugars. It is also well known to those skilled in the art. It is typically a hydrolysis which can be catalyzed by acid or by enzymes such as cellulases, produced for example by the strain Trichoderma reesei, xylanases, xylosidases and arabinofuranosidases.

[0101] The sugar-rich substrate is then subjected to fermentation.

[0102] By way of example, this fermentation can be carried out using specialized microorganisms, and in particular yeasts, which optimize the profitability of the production process, including the following yeasts: Ethanol Red® (Fermentie), Thermosacc® (Lallemand)), Angel Super Alcohol® (Angel®) and Fali® (AB Mauric)), the Saccharomyces cerevisiae yeast strains described in document FR3015985, the Candida Shehatae or Pichia stipitis yeast strains, or any other suitable microorganism.

[0103] For anaerobic fermentation, ethanol can be produced by anaerobic fermentation of a gas containing CO. The substrate is then a gaseous substrate (a gas) containing CO. This gaseous substrate can be a by-product of an industrial process.

[0104] In certain embodiments, the industrial process is selected from the group consisting of the manufacture of ferrous metal products, in particular steel mills, the manufacture of non-ferrous products, petroleum refining processes, coal and / or biomass or biochar gasification, electricity generation, carbon black production, ammonia production, methanol production, coke production, catalytic cracking (in particular during catalyst regeneration, carbon monoxide is produced), and methane reforming. In these embodiments, the gaseous substrate can be captured from the industrial process before it is emitted into the atmosphere, using any appropriate method. appropriate. Depending on the composition of the captured gas, it may also be desirable to treat it to remove any undesirable impurities, such as dust particles, before introducing it into the fermentation process. For example, the gas can be filtered or purified using known methods.

[0105] In other embodiments, the gaseous substrate may be obtained from biomass gasification. The gasification process involves the partial combustion of biomass in a limited supply of air or oxygen. The resulting gas generally comprises mainly CO and H2, with minimal amounts of CO2, methane, ethylene, and ethane. For example, biomass by-products obtained during the extraction and processing of food products, such as sugar from sugarcane or starch from corn or cereals, or non-food biomass waste generated by the forestry industry, can be gasified to produce a CO-containing gas that can be used in the present invention.

[0106] The gaseous substrate used typically contains 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 is within any range defined by two of these limits. Advantageously, the CO-containing gas may comprise 25%, 30%, 35%, 40%, 45%, 50%, 55%, or 60% CO by volume. Gases having lower CO contents, such as 6% by volume, may also be suitable, particularly when H2 and CO2 are also present.

[0107] If the gaseous substrate contains CO, it is not necessary for the gaseous substrate to contain hydrogen, but this is not considered detrimental to ethanol production. The gaseous substrate may also contain CO2, for example, in a proportion of 1% to 80% by volume, or 1% to 30% by volume, or 5% to 10% by volume, or in any range defined by two of these limits.

[0108] Typically, carbon monoxide is added to the fermentation reaction in the gaseous or liquid state. For example, carbon monoxide can be supplied in a liquid by saturation. For instance, a liquid can be saturated with a gas containing carbon monoxide, and then this liquid can be added to a bioreactor. This can be accomplished using a standard methodology. As an example, a microbubble dispersion generator (Hensirisak et al. Scale-up of microbubble dispersion generator for aerobic fermentation; Applied Biochemistry and Biotechnology RV Volume 101, Number 3 / October, 2002) could be used.

[0109] Furthermore, it is often desirable to increase the CO concentration of the gas (or the partial pressure of CO in the gas) and thus increase the efficiency of fermentation reactions using CO as a substrate. Increasing the partial pressure of CO in the gas increases the mass transfer of CO in a fermentation medium. The composition of the gas streams used to feed a fermentation reaction can significantly impact the efficiency and / or cost of that reaction. For example, 1'O2 can reduce the efficiency of an anaerobic fermentation process. Treating unwanted or unnecessary gases in the steps of a fermentation process before or after fermentation can increase the load on those steps (for example, when the gas stream is compressed before entering a bioreactor, unnecessary energy may be used to compress gases that are not required for fermentation). Therefore, it may be desirable to treat substrate streams, especially substrate streams derived from industrial sources, to remove unwanted components and increase the concentration of desirable components.

[0110] Any microorganism capable of fermenting a gaseous substrate comprising CO₂ to produce ethanol may be used in the present invention. By way of example, microorganisms of the genera Moorella, Clostridia, Ruminococcus, Acetobacterium, Eubacterium, Butyribacterium, Oxobacter, Methanosarcina, Methanosarcina, and Desulfotomaculum may be used.

[0111] By way of example, microorganisms of the genus Clostridium may be used, including strains of Clostridium ljungdahlii, Clostridium carboxy-divorans, Clostridium ragsdalei, and Clostridium autoethanogenum; of the genus Moorella, including Moorella sp. HUC22-1; of the genus Carboxydothermus, Moorella thermoacetica, Moorella thermoautotrophica, Ruminococcus productus, Acetobacterium woodii, Eubacterium limosum, Butyribacterium methylotrophicum, Oxobacter pfennigii, Methanosarcina barkeri, Methanosarcina acetivorans, or Desulfotomaculum kuznetsovii. Other specific examples of microorganisms are carboxytrophic anaerobic bacteria. Examples of usable strains are described in document WO201226833.

[0112] It should be noted that the invention can be applied to a mixed culture of two or more microorganisms.

[0113] Regarding the fermentation medium and conditions, regardless of the nature of the substrate (gaseous or non-gaseous) used, for ethanol fermentation to occur through the growth of one or more microorganisms, a suitable nutrient medium must be introduced into the bioreactor in addition to a substrate, under appropriate conditions. A nutrient medium will contain components, such as vitamins and minerals, sufficient to allow the growth of the microorganism used. The reaction conditions to be considered are the temperature, the flow rate of the medium, the pH, the redox potential of the medium, the stirring speed (when using a continuously stirred reactor), the inoculum level, the maximum substrate concentrations, and the rates of substrate introduction into the bioreactor to ensure that the substrate level does not become limiting, and the maximum concentrations of The product is produced in order to avoid product inhibition. Optimal reaction conditions will depend in part on the particular microorganism used. Microorganism culture methods are well known in the art, and those skilled in the art know how to optimize culture conditions for each microorganism, according to its nature. Examples of fermentation conditions suitable for the anaerobic fermentation of a substrate containing CO₂ are detailed in WO2007 / 117157, WO2008 / 115080, WO2009 / 022925, and WO02 / 08438.

[0114] Fermentation reactions can be carried out in any suitable bioreactor. In some embodiments of the invention, the bioreactor may comprise a first growth reactor in which the microorganisms are cultured, and a second fermentation reactor, into which the broth from the growth reactor is introduced and in which most of the fermentation product (ethanol, for example) is produced.

[0115] The fermentation will result in a fermentation broth comprising a desired product (ethanol) and / or one or more by-products (such as acetate and butyrate when the substrate is a gas containing CO) as well as microorganism cells, in a nutrient medium.

[0116] The recovery of ethanol may include the continuous removal of a portion of the broth and the recovery of ethanol from the portion removed from the broth.

[0117] For example, the removed portion of the broth containing ethanol can be passed through a separation unit to separate, for example by filtration, the bacterial cells from the broth and produce a cell-free ethanol-containing permeate, and the return of the microorganism cells to the bioreactor.

[0118] In some embodiments, the recovery of ethanol and / or one or more other products or by-products produced in the fermentation reaction includes the continuous withdrawal of a portion of the broth and the separate recovery of ethanol and one or more other products from the portion withdrawn from the broth.

[0119] By way of example, ethanol can be recovered from the fermentation broth using methods such as filtration, distillation or fractional evaporation, pervaporation, and extractive fermentation. Distilling ethanol from a fermentation broth gives an azeotropic mixture of ethanol and water (i.e., 95% ethanol and 5% water). Anhydrous ethanol can then be obtained using molecular sieve ethanol dehydration technology, which is also well known in the art.

[0120] Ethanol from renewable sources can also be obtained from biomass by conversion of a synthesis gas rich in CO / H2, this synthetic gas being derived from biomass.

[0121] Biomass can, for example, be gasified to produce synthesis gas (or “ Syngas (in English) rich in CO / H2, this synthetic gas is then converted into methanol in the presence of a catalyst. A process of this type is described, for example, in document WO2012003901.

[0122] The biomass used to produce synthesis gas may include, in particular, woody fuels from natural forests and wooded lands (e.g., sawdust), agricultural residues (e.g., rice husks, straw manure), energy crops that are grown exclusively for energy production (e.g., maize and oil palm), urban waste (e.g., wood waste, rice, straw manure), energy crops that are grown exclusively for energy production (e.g., maize and oil palm), urban waste (e.g., municipal solid waste and wastewater) and biomass fuel derived from waste (e.g., wood pellets).

[0123] It is also possible to obtain methanol from renewable sources. By methanol from renewable sources, we mean methanol obtained from biomass.

[0124] Biomass may include in particular woody fuels from natural forests and wooded lands (e.g. sawdust), agricultural residues (e.g. rice husks, straw manure), energy crops that are grown exclusively for energy production (e.g. maize and oil palm), urban waste (e.g. wood waste, rice, straw manure), energy crops that are grown exclusively for energy production (e.g. maize and oil palm), urban waste (e.g. municipal solid waste and wastewater) and biomass fuel derived from waste (e.g. wood pellets).

[0125] Methanol of renewable origin can in particular be obtained by conversion of a synthetic gas rich in CO / H2, this synthetic gas being derived from biomass.

[0126] Biomass can, for example, be gasified to produce a synthetic gas (or "syngas") rich in CO / H2, this synthetic gas then being converted into methanol in the presence of a catalyst. A process of this type is described, for example, in document WO2018134853A1

[0127] A synthetic gas suitable for further conversion to methanol can also be obtained by partial oxidation in the presence of dioxygen of a biogas containing methane and CO2, this biogas resulting, for example, from the anaerobic digestion of biomass in the presence of one or more microorganisms. A process of this type is described, for example, in document WO2019060988A1.

[0128] The alcohols intended to form the C6 Cl alcohol stream and optionally the additional C2 C6 alcohol stream can also be obtained from carbon dioxide, in particular captured.

[0129] Several transformation pathways exist. One example is the catalytic conversion of carbon dioxide into methanol in the presence of hydrogen.

[0130] Another way is to convert carbon dioxide into carbon monoxide by electroconversion or by reverse water gas reaction in the presence of hydrogen.

[0131] Carbon monoxide is then converted by catalytic conversion into methanol, in the presence of hydrogen.

[0132] The hydrogen used for the various operations described above is obtained in particular by steam reforming of methane, by reaction of the gas with water, or is produced by electrolysis from renewable energies such as solar energy, wind, geothermal, waves or currents. Step (a) Conversion of the alcohol stream from Cl to C6

[0133] The conversion of the alcohol stream into Cl to C6 includes, for example, dehydration, carbon-carbon coupling and / or aromatization of at least one alcohol into Cl to C6.

[0134] The dehydration of alcohols in Cl to C6, carbon-carbon coupling and / or their aromatization can be carried out simultaneously.

[0135] For the dehydration of methanol, it is generally converted into dimethyl ether, which is dehydrated to produce olefins having at least two carbon atoms. In this case of methanol and / or dimethyl ether, carbon-carbon coupling occurs during dehydration.

[0136] C2-C6 alcohols can be dehydrated to produce olefins containing the same number of atoms. The dehydration reactions of alcohols to produce alkenes have been known for a long time (J. Catal. 7, p. 163, 1967 and J. Am. Chem. Soc. 83, p. 2847, 1961). Many available solid acid catalysts can be used for the dehydration of alcohols ((Stud. Surf. Sci. Catal. 51, p. 260, 1989), EP0150832, Bulletin of the Chemical Society of Japan, vol. 47(2), 424-429 (1974)). However, alumina-γs are the most commonly used, especially for longer-chain alcohols (with three or more carbon atoms). Indeed, catalysts with higher acidity, such as silica-aluminas, zeolites, heteropoly-acids or "resin catalysts" can promote the displacement of double bonds, the isomerization of the skeleton and other interconversion reactions of olefins.

[0137] In addition, aromatization of alcohols in Cl to C6 occurs by oligomerization of olefinic intermediates, cyclization of a chain having at least 6 carbons and dehydrogenation of cycloparaffins into corresponding aromatics.

[0138] The reaction mechanism during the conversion of Cl to C6 alcohols involves acid catalysis. The catalyst can provide a proton to activate the molecules, alcohols and / or olefins via protonated intermediates. The stability of the intermediates The formation of protonated intermediates or alkylcarbenium 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 the most readily formed, and reactions involving the formation of primary carbenium ions are slow. Primary carbenium ions tend to be converted into secondary or tertiary carbenium ions.

[0139] 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 of the alkylation of aromatics with alkenes.

[0140] Hydride transfer offers a route for converting a neutral molecule into a carbenium ion, or successive hydride transfer from the alkene to the carbenium ion leads to the formation of aromatic compounds. The following reactions demonstrate this mechanism for propylene:

[0141] Protonation: CH3-CH=CH2 + H+ CH3-CH+-CH3

[0142] Dimerization: CH3-CH=CH2 + CH3-CH+-CH3 ' CH3-CH(CH3)-CH2-CH+-CH3

[0143] Deprotonation: CH3-CH(CH3)-CH2-CH+-CH3 “* CH3-CH(CH3)-CH=CH-CH3+ H+

[0144] H transfer: CH3-CH(CH3)-CH=CH-CH3 CH3-CfCH3)-CH=CH-CH3+H

[0145] H + Rp HR]

[0146] Cyclization: CH3-CfCH3)-CH=CH-CH3 CH3-(-C=CH2-CH2-CH-CH2-f + H+

[0147] (methylcyclopentyl cation)

[0148] Isomerization: CH3-(-C=CH2-CH2-CH-CH2-)+ (-CH2-CH2-CH2-CH2-CH2-CH-)+

[0149] (cyclohexyl cation)

[0150] (-CH2-CH2-CH2-CH2-CH2-CH-f (-CH2-CH2-CH2-CH2-CH^CH-) + H+

[0151] (cyclohexene)

[0152] Transfer H: (-CH2-CH2-CH2-CH2-CH-CH-) (-CH2-CH2-CH+-CH2-CH=CH-) + H

[0153] H + Rp H-Rj

[0154] (-CH2-CH2-CHCH2-CHJJH-f (-CH2-CH2-CH=CH-CHJJH-) + H+

[0155] (cyclohexadiene)

[0156] (-CH2-CH2-CH=CH-CHJJH-) (-CH+-CH2-CH=CH-CHXH-) + H

[0157] H + Rp H-Rj

[0158] (-CH+-CH2-CH=CH-ŒCCH-) (-CH=CH2-CH=CH-ŒLCH-) + H+

[0159] (benzene)

[0160] Overall, for the formation of an aromatic ring, three dihydrogen molecules (six atoms) must be removed and these three hydrogen molecules will, according to the hydride transfer mechanism, form three paraffins from the olefins.

[0161] The hydride transfer mechanism typically occurs on catalysts having only an acidic function and often requires harsh conditions which are also conducive to coke formation.

[0162] The production of aromatics from olefins (produced from alcohols) occurs primarily on acid catalysts via hydride transfer from one olefin to another. This produces more unsaturated molecules (and ultimately aromatics) and paraffins. These paraffins are not transformable because they are too inert to be converted under the optimal operating conditions for the conversion of C1-C6 alcohols. It is possible to add a dehydrogenating catalytic function, allowing aromatization by producing molecular hydrogen.

[0163] If the catalyst also possesses a dehydrogenating function, the reactive intermediates can be converted into corresponding aromatics by dehydrogenation. This is particularly the case with bifunctional catalysts, having an acid function and a dehydrogenating function. The dehydrogenating function can be provided by metals of groups VIB, VIIIB, IB, and IIB and mixtures thereof, preferably gallium, zinc, or mixtures thereof.

[0164] In general, olefin dehydrogenation reactions are thermodynamically limited and require high temperatures. To promote the dehydrogenating pathway, another reactant can be added in step (a) of the conversion to shift the thermodynamic equilibrium.

[0165] Surprisingly, according to the invention, this reagent is carbon dioxide, which can be converted into carbon monoxide and water:

[0166] CO2 + H2 CO + H2O

[0167] The CO2, CO and H2 molecules can be separated from other hydrocarbons and recycled towards the synthesis of alcohols, by fermentation of synthesis gas and by catalytic conversion of synthesis gas to make methanol.

[0168] The process then includes the addition, at step (a) of converting the alcohol stream into C6 Cl, of a stream containing carbon dioxide, and the joint conversion of carbon dioxide into carbon monoxide at step (a) of converting the alcohol stream into C6 Cl.

[0169] Advantageously, the stream containing carbon dioxide comprises more than 5% by mass of carbon dioxide, in particular more than 10% by mass of carbon dioxide.

[0170] The mass ratio in the feed supplied at the conversion step (a) of carbon dioxide supplied in the carbon dioxide stream to the Cl-C6 alcohols supplied in the Cl-C6 alcohol stream is between 5% and 75%.

[0171] The stream containing carbon dioxide is added, for example, mixed with the alcohol stream in Cl to C6 or, if several reaction zones are arranged in series to carry out the conversion step, between two reaction zones or in a reaction zone. tional data.

[0172] Preferably, more than 2 mole percent, in particular more than 6 mole percent of carbon dioxide is converted to carbon monoxide jointly with the conversion of the alcohol stream to Cl to C6 during conversion step (a).

[0173] The main product of acid-catalyzed dehydration of ethanol and / or methanol is ethylene and / or propylene and water.

[0174] More generally, a mixture containing paraffins, olefins, aromatics, and water is produced.

[0175] Paraffins include n-paraffins, i-paraffins and cycloparaffins.

[0176] According to the invention, in the mixture of paraffins, olefins, aromatics and water produced in step (a) of conversion, the ratio of the mass of C3+ olefins to the total mass of olefins is greater than or equal to 0.80, preferably greater than or equal to 0.85, the ratio being calculated on the dry flow, after separation of the water.

[0177] Advantageously, not taking into account the water circulating in a water recycle, the produced mixture containing paraffins, olefins, aromatics, and water contains more than 10% by mass of water, in particular between 10% and 60% by mass of water depending on the alcohol composition in the mixture of alcohols from Cl to C6. In the case where only methanol is present, the water content of the produced mixture (excluding recycle) is between 55% by mass and 60% by mass.

[0178] Advantageously, on a dry basis excluding water and any recycle, the mixture produced containing paraffins, olefins, aromatics, and water contains more than 2% by mass of aromatics, in particular more than 6% by mass of aromatics, in particular between 6% and 30% by mass of aromatics.

[0179] In its dry form, excluding water, it advantageously contains:

[0180] - less than 5% by mass of methane, in particular between 0.1% by mass and 4% by mass mass of methane;

[0181] - less than 5% by mass of dimethyl ether, in particular less than 1% by mass of di- methyl ether;

[0182] - less than 5% by mass of residual Cl to C6 alcohols, in particular less than 1% in mass of residual Cl to C6 alcohols;

[0183] - less than 15% by mass of ethylene, in particular between 5% by mass and 10% by mass mass of ethylene;

[0184] - more than 30% by mass of propylene, in particular between 35% by mass and 60% by mass mass of propylene;

[0185] - less than 15% by mass of paraffins, in particular between 3% and 8% by mass of pa refines;

[0186] - more than 10% by mass of C4 to C7 olefins, in particular between 25% and 40% of C4 to C7 olefins;

[0187] - more than 6% aromatics, in particular between 6% and 10% aromatics, in particular less than 3% aromatics in C9+.

[0188] The production of light olefins (ethylene and propylene) from a feed of mixed alcohols in a process of oxygenated olefin compounds is described, for example, in US patent 7,288,689. The said patent proposes various processes for the production of Cl to C4 alcohols, possibly in a mixed alcohol stream, and possibly for the conversion of alcohols into light olefins.

[0189] Conversion by dehydration and aromatization makes it possible to obtain light olefins, having at least 2 carbon atoms and aromatics from C1-C6 alcohols using a composite catalyst comprising the following steps:

[0190] a) the supply of a catalyst comprising molecular sieves containing at least 10 oxygen atom (10-MR) pores or larger in their microporous structure,

[0191] b) in the case of an implementation of the conversion on a reaction zone having a fluidized bed, provide a reaction zone and a catalyst regeneration zone, said catalyst circulating in both zones, so that at least a part of the regenerated catalyst passes into the reaction zone and at least a part of the catalyst in the reaction zone passes into the regeneration zone;

[0192] or / and in the case of implementation of the conversion on a reaction zone having at least one fixed bed, a step of regeneration of the catalyst in situ, either by directing the flow to react towards a reaction zone which has been previously regenerated, or by stopping the conversion in the reaction zone;

[0193] c) contacting the Cl to C6 alcohols in the reactor with the catalyst under effective conditions to convert at least part of the feed charge to form a reactor effluent comprising a mixture of paraffins, olefins, aromatics, and water.

[0194] The catalyst may be a mixture of two or more catalysts and optionally a binder.

[0195] It is desirable to have a conversion rate of approximately 100% of the alcoholic compound in the reactor. This conversion rate is adjusted by optimizing the contact time, the reaction temperature, and the catalyst regeneration frequency.

[0196] In a specific embodiment, the weight hourly space velocity (WHSV) of the alcohol in the reaction zone is approximately 0.5 h₁ to approximately 10 h₁, advantageously approximately 1 h₁ to approximately 6 h₂*.

[0197] The molecular sieves used in the catalyst composition are selected from the list of molecular sieves with crystalline structure MFI, MOR, MEL, clinop- tilolite, FER, FAU, MWW, BETA, MCM-41, ZSM-21, ZSM-22, ZSM-23, ZSM-42, ZSM-57, LTL, or a mixture thereof. Preferably, the selected molecular sieve is a zeolite, a crystalline aluminosilicate chosen from the MFI, MOR, MEL, clinoptilolite, or FER group, or a mixture thereof. More preferably, in the case of MFI, the molecular sieve is a ZSM-5 zeolite. In another embodiment, the molecular sieve is preferably obtained without the addition of a structuring agent. Other examples are described by the International Zeolite Association (Atlas of Zeolite Structure Types, 1987, Butterworths).

[0198] Crystalline silicates are microporous crystalline inorganic polymers based on a network of tetrahydric XO4 units linked together by oxygen ion sharing, where X can be trivalent (e.g., Al, B, ...) or tetravalent (e.g., Ge, Si, ...). The crystal structure of a crystalline silicate is defined by the specific order in which a network of tetrahedral units are linked together. The size of the pore openings of the crystalline silicate is determined by the number of tetrahedral units, or, alternatively, oxygen atoms, required to form the pores and the nature of the cations present in the pores. They possess a unique combination of the following properties: high internal surface area; uniform pores with one or more discrete sizes; ion exchangeability; good thermal stability; and the ability to adsorb organic compounds.Since the pores of these crystalline aluminosilicates are similar in size to many organic molecules of practical interest, they control the entry and exit of reactants and products, resulting in particular selectivity in catalytic reactions. Crystalline aluminosilicates with the MFI structure possess a bidirectional intersecting pore system with the following pore diameters: a straight channel along

[010] : 0.53–0.56 nm and a sinusoidal channel along

[100] : 0.51–0.55 nm. Crystalline aluminosilicates with the MEL structure possess a bidirectional intersecting straight pore system with straight channels along

[100] having pore diameters of 0.53–0.54 nm and a sinusoidal channel along

[100] : 0.51–0.55 nm.

[0199] The molecular sieves used in the present invention (in the H+ or NH4+ form) have an initial Si / Al ratio advantageously between 4 and 500, preferably from 4 to 100, or more preferably from 4 to 30. The conversion to the H+ or NH4+ form is known per se and described in US3911041 and US5573990. The Si / Al atomic ratio is measured by chemical analysis, for example by NMR. It includes only the Al that are part of the lattice structure of the molecular sieve.

[0200] According to a first embodiment, said zeolite is a phosphorus-modified zeolite manufactured by a process comprising, in this order:

[0201] - Selection of a molecular sieve as described in the list above;

[0202] - Introduction of P under efficient conditions to advantageously introduce to minus 0.05% by mass of P by adding an aqueous solution containing a phosphorus precursor;

[0203] - Possible separation of the solid from the aqueous liquid;

[0204] - Optional washing or optional drying or optional drying followed by washing;

[0205] - Calcination.

[0206] Optionally, the manufacturing process for said phosphorus-modified zeolite includes steam heat treatment and leaching steps. The method consists of steam heating / heat treatment followed by leaching.

[0207] It is generally known to those skilled in the art that steam treatment of zeolites leads to aluminium leaving the zeolite lattice and residing as aluminium oxides both inside and outside the pores of the zeolite. This transformation is known as the dealumination of zeolites, and this term will be used throughout the text.

[0208] In the steam treatment step, the temperature is preferably from 400°C to 870°C, more preferably from 480°C to 760°C. The pressure is preferably atmospheric pressure, and the partial pressure of water can range from 13 kPa to 100 kPa. The steam atmosphere preferably contains from 5% to 100% by volume of steam with from 0% to 95% by volume of an inert gas, preferably nitrogen. The steam treatment is preferably carried out for a duration of 0.01 hours to 200 hours, advantageously from 0.05 hours to 200 hours, more preferably from 0.05 hours to 50 hours. The steam treatment tends to reduce the amount of tetrahedral aluminum in the crystalline silicate lattice by forming alumina.

[0209] Treating steam-treated zeolite with an acidic solution leads to the dissolution of extra-lattice aluminum oxides. This transformation is known as leaching, and this term will be used throughout the text. Leaching can be carried out with an organic acid such as citric acid, formic acid, oxalic acid, tartaric acid, malonic acid, succinic acid, glutahic acid, adipic acid, maleic acid, phthalic acid, isophthalic acid, fumaric acid, nitrilotriacetic acid, hydroxyethylenediaminetetraacetic acid, ethylenediaminetetraacetic acid, trichloroacetic acid, trifluoroacetic acid, or a salt of such an acid (e.g., the sodium salt), or a mixture of two or more of these acids or salts.Other inorganic acids may include an inorganic acid such as nitric acid, hydrochloric acid, hydromethanesulfur acid, phosphoric acid, phosphonic acid, sulfuric acid, or a salt of such an acid (e.g., sodium or ammonium salts), or a mixture of two or more such acids or salts. Leaching with an acidic aqueous solution containing a phosphorus source is advantageously carried out under reflux conditions, i.e., at the boiling point of the solution. The quantity of said... The acidic solution is advantageously between 2 liters and 10 liters per kg of molecular sieve. A typical leaching period is approximately 0.5 to 24 hours. Advantageously, the aqueous acidic solution containing the phosphorus source in the leaching step has a pH of 3, advantageously 2, or lower. Advantageously, said aqueous acidic solution is a solution of phosphoric acids, a mixture of phosphoric acids and organic or inorganic acids, or mixtures of salts of phosphoric acids and organic or inorganic acids. The phosphoric acids or corresponding salts may be phosphate ([PO4]3, being basic), phosphite ([HPO3]2, being dibasic), or hypophosphite ([H2PO2], being monobasic).Unexpectedly, a greater amount of phosphorus than would be expected from the typical pore volume of the molecular sieve, and assuming that the pores of the molecular sieves are filled with the phosphorous acid solution used, remains in the solid molecular sieve material. The two factors—dealumination and phosphorus retention—stabilize the lattice aluminum within the zeolite, thus preventing further dealumination. This leads to higher hydrothermal stability, tuning of the molecular sieve properties, and adjustment of the acid properties, thereby increasing the molecular sieve's selectivity.

[0210] The zeolite is then separated, advantageously by filtration, and optionally washed. A drying step may be considered between the filtration and washing steps. The solution after washing may be either separated, for example, by filtration of the solid, or evaporated. P may be introduced by any means or, for example, according to the recipe described in US 3,911,041, US 5,573,990, and US 6,797,851. The separation of the liquid from the solid is advantageously carried out by filtration at a temperature between 0°C and 90°C, centrifugation at a temperature between 0°C and 90°C, evaporation, or equivalent. Optionally, the zeolite may be dried after separation before washing. Advantageously, said drying is carried out at a temperature between 40°C and 60°C, advantageously for 1h to 100. This drying may be carried out either under static conditions or in a gas flow. Air, nitrogen, or any inert gas can be used.The washing step can be carried out either during filtration (separation step) with a portion of cold water (<40°C) or hot water (>40°C but <90°C), or the solid can be subjected to an aqueous solution and treated under reflux conditions for 0.5 h to 10 h followed by evaporation or filtration. The final calcination step is advantageously carried out at a temperature of 400°C–700°C, either under static conditions or in a gas flow. Air, nitrogen, or any inert gas can be used.

[0211] According to one embodiment of the invention, the phosphorus-modified zeolite is manufactured by a process comprising, in this order:

[0212] - Selection of a molecular sieve as described in the list above;

[0213] - Steam heat treatment at a temperature ranging from 400°C to 870°C for 0.01 -200h;

[0214] - Leaching with an aqueous acid solution under conditions effective for remove a substantial portion of Al from the zeolite;

[0215] - Introduction of P with an aqueous solution containing the source of P into effective conditions for advantageously introducing at least 0.05% by mass of P

[0216] - Separation of solid from liquid;

[0217] - An optional washing step or an optional drying step or a step of optional drying followed by a washing step;

[0218] - A calcination step.

[0219] Possibly between the steam heat treatment step and the leaching step, there is an intermediate step such as, for example, contact with silica powder and drying.

[0220] Advantageously, the final phosphorus (P) content is at least 0.05% by mass and preferably between 0.3% by mass and 7% by mass. Advantageously, at least 10% by mass of aluminum (Al), relative to the parent zeolite, MFI, MEL, FER, MOR, and clinoptilolite, have been extracted and removed from the zeolite by leaching. The zeolite is then either separated from the washing solution or dried without separation from the washing solution. This separation is advantageously carried out by filtration. The zeolite is then calcined, for example, at 400 °C for 2 to 10 hours.

[0221] The residual P content is adjusted by the concentration of P in the aqueous acidic solution containing the P source, the drying conditions and a washing procedure if any. A drying step may be considered between the filtration and washing steps.

[0222] The catalyst consisting of a phosphorus-modified zeolite may be the phosphorus-modified zeolite itself or it may be the phosphorus-modified zeolite formulated in a catalyst by combining with other materials which provide additional hardness or catalytic activity to the finished catalyst product.

[0223] According to a second embodiment, the catalyst of the process is a composite catalyst manufactured by a process comprising the following steps:

[0224] a) Selection of a molecular sieve from the list as defined previously

[0225] b) Contacting the molecular sieve with a metallic silicate comprising at less an alkaline earth metal, so that the composite comprises at least 0.1% by mass of silicate.

[0226] The molecular sieve is preferably brought into contact with the metallic silicate by one of the following two methods:

[0227] - During the catalyst formulation step by mechanical mixing of the sieve mo lecular with metallic silicate forming a precursor to be used in the formulation step;

[0228] - Physical mixing of the previously formulated molecular sieve and the silicate me tallic previously formulated in situ in the reaction medium intended for the implementation of the conversion.

[0229] Said molecular sieve and / or said composite catalyst containing the molecular sieve and the metallic silicate may be post-treated by calcination, reduction, or hydrothermal steam treatment. In the case of using zeolites as components of molecular sieves, phosphorus may be introduced before, simultaneously, or after mixing with the metallic silicate.

[0230] In a particular embodiment of the invention, the molecular sieve can be modified either before or after the introduction of the metal silicate. Preferably, the molecular sieve has undergone some form of modification before the introduction of the metal silicate. By modification, it is understood that the molecular sieve may have undergone steam heat treatment, leaching (for example, acid leaching), washing, drying, calcination, impregnation, or some form of ion exchange. This means that at least some of the cations initially included in the crystal structure can be 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.

[0231] The selected molecular sieve is then formulated into a composite catalyst to comprise at least 10% by mass of a molecular sieve as described herein and at least one metallic silicate comprising at least one alkaline earth metal, so that the composite comprises at least 0.1% by mass of silicate.

[0232] At least one of the metallic silicates included in the composite catalyst comprises at least one alkaline earth metal, preferably Ca. Metallic silicates are insoluble in water, and alkaline earth metal ions, particularly calcium, are versatile and possess a large radius in the hydrated state. Thus, without wishing to be bound by theory, it is thought that the ion exchange reaction with the molecular sieve occurs very slowly, because the alkaline earth metal ion must lose many of its highly coordinated water molecules to penetrate the micropores of the sieve. As a result, the alkaline earth ions expose only the acidic sites located on the outer surface of the molecular sieve, thereby increasing the selectivity of the catalyst.

[0233] Furthermore, without wishing to be bound by theory, it is thought that the presence of silicate anions further enhances the catalytic properties of the composite catalyst. Silicate anions, for example, can provide silicon atoms to repair defects of the molecular sieve. This can thus lead to further stabilization of the catalyst under severe hydrothermal conditions.

[0234] Consequently, the metallic silicate acts as a catalyst promoter. The metallic silicate may comprise more than one alkaline earth metal selected from Ca, Mg, Sr and Ba.

[0235] Metallic silicates may also comprise other metals selected from one or more of the following: Ga, Al, Ce, In, Cs, Sc, Sn, Li, Zn, Co, Mo, Mn, Ni, Fe, Cu, Cr, Ti, and V. Preferably, the other metal is selected from one or more of Al, Mg, Ce, Co, and Zn, or mixtures thereof. These bi-, tri-, or polymetallic silicates may be synthesized by any method known in the art. This may be, for example, by ion exchange in solution or in the solid state (Labhsetwar et al., Reactivity of Solids, vol. 7, no. 3, 1989, 225–233).

[0236] The silicate anion can be present in any form in solid metallic silicate. Examples include SiO32, SiO44, Si2O76, Si3O1O8 and the like

[0237] The preferred catalyst promoter is a calcium silicate with a very open and accessible porous structure. An even more preferred catalyst promoter comprises a synthetic crystalline hydrated calcium silicate having a chemical composition of Ca6Si6Oi7(OH)2 which corresponds to the known mineral xonotlite (having a molecular formula 6CaO.6SiO2.H2O).

[0238] Generally, a synthetic hydrated calcium silicate is synthesized hydrothermally under autogenous pressure. A particularly preferred synthetic hydrated calcium silicate is commercially available from the company Promat of Ratingen, Germany, under the brand name Promaxon.

[0239] Other examples of metallic silicates comprising alkaline earth metals include CaAl2Si2O8, Ca2Al2SiO7, CaMg(Si2O6)x and mixtures thereof.

[0240] Before mixing with the molecular sieve, said metal silicate compounds may be modified by calcination, steam treatment, ion exchange, impregnation, or phosphating. Said metal silicates may be an individual compound or may be part of mixed compounds.

[0241] The metallic silicate can be brought into contact with the molecular sieve by a simultaneous formulation step of a mixture of the metallic silicate with the molecular sieve or an in situ mixture of materials formulated separately in the reaction medium before the implementation of the conversion.

[0242] Said contact can be achieved by mechanically mixing the molecular sieve with the metallic silicate comprising an alkaline earth metal. This can be achieved by any known mixing process. The mixing can last for a period of time ranging from 1 minute up to 24 hours, preferably from 1 minute to 10 hours.

[0243] If not carried out in the in-situ conversion reactor, it can be carried out in a batch mixer or in a continuous process, such as in an extruder, for example a single or twin screw extruder at a temperature of 20°C to 300°C under vacuum or high pressure. This contact can be made in an aqueous or non-aqueous medium. Prior to the formulation step, other compounds facilitating formulation can be added, such as thickening agents or polyelectrolytes that improve the cohesion, dispersion, and flow properties of the precursor. In the case of oil drop drying or spray drying, a rather liquid (high water content) is prepared. In another embodiment, the contact is made in the presence of phosphorus-containing compounds. In a particular embodiment, the contact is made in an aqueous medium at a pH below 5, more preferably below 3.

[0244] According to a third embodiment, the catalyst of the process is a molecular sieve modified by phosphorus (P) and by an alkaline earth or rare earth metal (M) (modified molecular sieve MP) manufactured by a process comprising the following steps:

[0245] a) select at least one molecular sieve chosen from

[0246] - a P-modified molecular sieve containing at least 0.3% by mass of P

[0247] - a molecular sieve modified with P before or during step b) introducing at minus 0.3% by mass of P

[0248] b) contacting said molecular sieve with a compound containing an alkaline earth or rare earth metal (compound containing M) to introduce at least 0.05% by mass of the alkaline earth or rare earth metal M.

[0249] Optionally, the contacting of the molecular sieve with the compound containing P and the compound containing M can be carried out simultaneously.

[0250] The introduction of the alkaline earth metal or rare earth (M) is carried out by contacting the molecular sieve with a solution of one or more compounds containing M. Said solution may contain a higher concentration of alkaline earth metal or rare earth than that found in the final modified molecular sieve MP.

[0251] The molecular sieve is selected from the list described above.

[0252] Before modification P and / or modification by an alkaline earth metal or an rare-earth metal (M modification), the molecular sieve can undergo further treatments, including steam heat treatment, leaching (e.g., acid leaching), washing, drying, calcination, impregnation, or ion exchange. As an addition or alternative, these steps can also be carried out during or after the P-modification. By ion exchange steps, it is understood that at least some of the cations initially included in the crystal structure are replaced by a wide variety of other cations using techniques well known in the art. The replacement cations may be hydrogen, ammonium, or other metallic cations, including mixtures of such cations.

[0253] The modification of molecular sieves with phosphorus is known per se. This modification is carried out by treating molecular sieves with compounds P in aqueous or non-aqueous media, by chemical vapor deposition of organic compounds P, or by impregnation. The catalyst may be pre-formulated with or without a binder. Preferred compounds P typically used for this purpose may be chosen from the phosphoric acid group, NH4H2PO4, or (NH4)2HPO4. The compound containing M may be chosen from organic compounds, salts, hydroxides, and oxides. These compounds may also contain phosphorus. It is essential that these compounds be present in solubilized form before contacting the molecular sieve or by forming a solution upon contact with the molecular sieve.

[0254] The final molar ratio M / P in the MP molecular sieve is preferably less than 1.

[0255] According to a particular embodiment of the invention, the molecular sieve can be modified with phosphorus according to the process comprising the following steps, in the order indicated:

[0256] - steam heat treatment of the molecular sieve at a temperature ranging from 400 °C to 870 °C for 0.01 h - 200 h;

[0257] - leaching with an acidic aqueous solution containing the source of P in effective conditions for removing a substantial portion of Al from the molecular sieve and for introducing at least 0.3% phosphorus by mass from the molecular sieve; further modification can then be carried out according to the following steps, in the order indicated:

[0258] - separation of solid from liquid;

[0259] - an optional washing step or an optional drying step or a step optional drying followed by a washing step;

[0260] - a calcination step.

[0261] Preferably, the separation, optional washing and drying steps, and calcination are carried out after the introduction of the compound containing M 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, Mg, Sr, and Ba can also be used. Possible rare earth metals include La and Ce.

[0262] Advantageously, the final P content of the molecular sieve is at least 0.3% by mass and preferably between 0.3% by mass and 7% by mass. Advantageously, at least 10% by mass of Al has been extracted and removed from the molecular sieve by leaching. The residual P content is adjusted by the P concentration in the solution leaching, separating the conditions during the solid-liquid separation and / or optional washing procedure during which impregnation and / or adsorption may also occur. A drying step may be considered between the separation and / or washing steps.

[0263] The molecular sieve is then either separated from the washing solution or dried without separation from the washing solution. This separation is advantageously carried out by filtration. The molecular sieve is then calcined, for example, at 400 °C for 2 to 10 hours.

[0264] The modification M of the molecular sieve is carried out either on a molecular sieve already modified P, or during / after the modification P process. The modification of P can be carried out as described above in which the sieve is desalinated by steam heat treatment, then leached with an acidic solution containing P. In this case, advantageously, the treatment of the molecular sieve with the solution containing M is carried out after the leaching or washing step, i.e. after the phosphorus compound has been added and the modification P has taken place and before the separation step.

[0265] However, the introduction of M into the molecular sieve can also be considered:

[0266] - during the leaching stage;

[0267] - before the washing stage but after leaching and drying;

[0268] - on calcined molecular sieves brought into contact with P;

[0269] - on a molecular sieve which has not been leached to introduce P but which has been put into contact with P during the washing stage.

[0270] The introduction of M onto the molecular sieves can be carried out either by impregnation or by adsorption from an aqueous solution of compounds containing M.

[0271] The introduction of the compound containing M can be carried out at temperatures ranging from room temperature up to the boiling point of the solution. The concentration of the compound containing M in the solution is at least 0.05-M, preferably between 0.05 and 1.0 M. The amount of alkaline earth metal or rare earth (M) in the molecular sieves MP can vary from at least 0.05% by mass, preferably from 0.05% by mass to 7% by mass, and even better from 0.1% by mass to 4% by mass.

[0272] Before the formulation of the composite catalyst, the molecular sieve may undergo further treatments including steam treatment, leaching (e.g., acid leaching), washing, drying, calcination, impregnation, and ion exchange steps. In addition, or as an alternative, these steps may also be carried out after the formulation of the catalytic composite.

[0273] The alkaline earth or rare earth metal M is preferably selected from one or Several of the following are possible: Mg, Ca, Sr, Ba, La, Ce. More preferably, M is an alkaline earth metal. More preferably, M is Ca. Particularly in the case of a P modification by vaporization and leaching, M can be a rare earth metal such as La and Ce.

[0274] The compound containing M is preferably in the form of an organic compound, a salt, a hydroxide, or an oxide. The compound is preferably in a solubilized form upon contact with the molecular sieve. Alternatively, the solution of the compound containing M can be formed after contacting the molecular sieve with said compound.

[0275] Possible compounds containing M include metal M compounds such as metal M sulfate, formate, nitrate, acetate, halides, oxyhalides, borates, carbonate, hydroxide, oxide, and mixtures thereof. These may be, for example, calcium sulfate, formate, nitrate, acetate, halides, oxyhalides, borates, carbonate, hydroxide, oxide, and mixtures thereof.

[0276] The compound containing M 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 compounds containing M may also further include phosphorus.

[0277] These compounds containing M, which are sparingly soluble in water, can be dissolved to form a well-solubilized solution by heating and / or modifying the pH of the solution by adding phosphoric, acetic, or nitric acid, or the corresponding ammonium acid, or salts thereof. The concentration of the compound containing M is at least 0.05 M.

[0278] Alkaline earth and rare earth metals M, particularly Ca, have a large hydration sphere radius in the hydrated state. Thus, without being bound by theory, it is thought that the ion exchange reaction with the acidic sites located within the microporous structures of the molecular sieve occurs very slowly. Consequently, the chosen metal M exposes only the acidic sites located on the external surface of the molecular sieve, thereby increasing the selectivity of the catalyst.

[0279] In the case of P-modified molecular sieves, the M modification leads to the formation of mixed M-AL phosphates on the outer surface. Given that phosphorus is more strongly bound to the alkali-earth or rare-earth M than to Al, this modification leads to the stabilization of phosphorus on the outer surface of the molecular sieve where phosphorus is most labile. However, it is essential that all M atoms located on the outer surface be saturated with phosphorus. This can be ensured by the presence of excess phosphorus and by the presence of M in solution, which serves, for example, to wash away the excess phosphorus, preventing clogging of the micropore inlets.

[0280] Before, after, or simultaneously with the formulation step to form the composite, other components may optionally be mixed with the modified or unmodified molecular sieve. In a particular embodiment, the modified or unmodified molecular sieve may be combined with other materials that impart additional hardness or catalytic activity to the finished catalytic product. The 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 clays, quartz, alumina or alumina sol, silica or silica sol, and / or metal oxides such as titanium dioxide, zirconia, and mixtures thereof. In one embodiment, certain binder materials may also serve as diluents to control the feed conversion rate to products and thereby improve selectivity.According to one embodiment, the binders also improve the attrition of the catalyst under industrial operating conditions. Natural clays usable as binders include, for example, clays from the kaolin or montmorillonite families. Such clays can be used in their raw state as extracted, or they can be subjected to various pre-use treatments, such as calcination, acid treatment, or chemical modification. In addition to the above, other materials that can be included in the composite catalyst of the invention include various forms of metals, including rare-earth or alkaline earth metals, and phosphates (for example, metallic phosphates, where the metal is selected from one or more of Ca, Ga, Al, Ca, Ce, In, Cs, Sr, Mg, Ba, Sc, Sn, Li, Zn, Co, Mo, Mn, Ni, Fe, Cu, Cr, Ti, and V).Examples of possible phosphates include amorphous calcium phosphate, monocalcium phosphate, dicalcium phosphate, dehydrated dicalcium phosphate, α- or γ-tricalcium phosphate, octacalcium phosphate, hydroxyapatite, zirconia, silica-thorium, silica-beryllium, silica-titanium, and calcium-alumina. Examples of ternary binder compositions include calcium-silica-alumina or silica-alumina-zirconia. These components are effective in increasing catalyst density and strengthening the formulated catalyst. The catalyst usable in fluidized bed reactors has a substantially spherical shape, generally formed by atomization.

[0281] Generally, when using a fluidized bed as a reactor, the catalyst particle size can vary from about 20 pm to 500 pm, more preferably from 30 pm to 100 pm. The size of the molecular sieve crystals contained in the composite catalyst is preferably less than about 10 pm, more preferably less than about 5 pm, and most preferably less than about 2 pm. Generally, when using a fixed bed, the catalyst particle size can vary from about 0.5 to 5 mm in the form of beads or extrudates (length between 1 and 10 mm) cylindrical or trilobed or quadrulobed. The quantity of molecular sieve, which is contained in the final catalytic composite, ranges from 10% by mass to 90% by mass of the total catalytic composite, preferably from 20% by mass to 70% by mass.

[0282] According to another embodiment, unmodified molecular sieves were first formulated with a binder and matrix materials and then modified with phosphorus and alkaline earth metal silicates. According to yet another particular embodiment, the molecular sieves were optionally dealuminized and then modified with phosphorus during the formulation step. The introduction of the alkaline earth metal silicate can be carried out during the formulation step or on the formulated solid.

[0283] According to a preferred embodiment, the molecular sieves were first optionally dealuminized and modified with phosphorus, then formulated. The introduction of the metal is carried out simultaneously with the phosphorus modification step and / or on the already formulated catalyst.

[0284] After formulation, the composite catalyst can undergo further treatments including additional steam treatment, leaching, washing, drying, calcination, impregnation, and ion exchange steps. If the molecular sieve has not been modified with phosphorus prior to the mixture formulation step, i.e., the step of introducing the metallic silicate into the molecular sieve, this modification can be carried out after such a step. According to a particular embodiment of this material, the molecular sieve is a phosphorus-modified (P-modified) zeolite. This phosphorus-modified (P-modified) zeolite has already been described above.

[0285] According to another embodiment, the unmodified molecular sieve was first formulated with a binder and a matrix material and then modified with phosphorus and metals. According to a particular embodiment, the molecular sieves were optionally dealuminized and then modified with phosphorus during the formulation step. The introduction of the metal can be carried out during the formulation step or on the formulated solid. According to a preferred embodiment, the molecular sieve was first optionally dealuminized and modified with phosphorus and then formulated. The introduction of the metal is carried out simultaneously with a modification step using phosphorus and / or on the formulated catalyst.

[0286] The final catalyst containing a phosphorus-modified zeolite advantageously exhibits a 27Al NMR signature between 35 ppm and 45 ppm, characteristic of the presence of an ALPO structure. The mass content of said AlPO4 structure in the catalyst can be up to 99% by mass and is advantageously between 10% and 98% by mass.

[0287] The presence of this ALPO structure is characterized by the following method, illustrated in [Fig. 9]. The measurement is carried out by solid-state NMR by rotation at the angle Magic MAS spectroscopy was performed on the Bruker Avance 500 spectrometer, using a 4 mm zirconia MAS probe at a rotation speed of 15 kHz. To obtain quantitative MAS spectra, a single-pulse excitation was applied using a short pulse length of 0.6 qsec. Each spectrum resulted from 5000 scans separated by a 0.5-second delay. The chemical shifts of the 27Al spectra were referenced to the AlCl3 solution (0.1 M, 0 ppm).

[0288] In the case where there is only a source of zeolitic aluminium in the catalyst, the content of the Al1PO4 phase is estimated directly by a ratio of the signal surface area at 35 ppm -45 ppm (centered at 39 ppm on the [Fig.9]) in the 27Al MAS relative to a total spectrum surface area between -50 ppm and 100 ppm.

[0289] In the case where the binder contains aluminium and phosphorus, the content of the A1PO4 phase in the zeolite is estimated by a ratio of the signal area at 35 ppm-45 ppm in the 27Al MAS to the total area of ​​the spectrum between -50 ppm and 100 ppm after subtracting the signal intensities of the binders.

[0290] To improve selectivity for aromatic formation, catalysts as described above can be further modified by adding one or more metals selected from among the metals of Group IIB (e.g., Zn), Group IIIB (e.g., Ga), the transition metals of Group VIIIB (e.g., Fe and / or Ni and / or Pt), Group VIB (e.g., Mo), Group IB (e.g., Cu and / or Ag), or the lanthanide group (e.g., La). The introduction of the metal(s) can be 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, and chemical vapor deposition (CVD). The introduction of the metal(s) is carried out from one or more solutions containing the metal(s) in the form of salts.The metal salts are previously dissolved in the treatment solution; the metal counter-ions are chosen from among sulfates, nitrates, carbonates, hydroxides, phosphates, carboxylates (formate, acetate, propionate for example), dicarboxylates (oxalate, malonate, succinate for example).

[0291] After introduction of the metal(s), various treatments can be applied, including drying, calcination, steam heat treatment.

[0292] The metal content is generally between 0% by mass and 5% by mass, preferably between 0% by mass and 2.5% by mass. The selected metal can typically be Ga3+ or Zn2+, in the presence / absence of other metals such as Pt.

[0293] The modification of the aforementioned catalyst can also be done by adding B, at a rate of 0.1% by mass to 5% by mass, preferably from 0.1% by mass to 1% by mass, more preferably from 0.1% by mass to 0.5% by mass, in the presence or absence of the other aforementioned metals.

[0294] In another embodiment, the oxygen conversion process is conducted on a catalyst comprising a zeolite with pores of 10 oxygen atoms (10-MR) or more, modified by the addition of B either before, after, or simultaneously with the final catalyst formulation step. Preferably, the catalyst could be, for example, a ZSM-5 modified with B.

[0295] The content of B in the final catalyst is between 0.1% by mass and 5% by mass, preferably between 0.1% by mass and 1% by mass, more preferably between 0.1% by mass and 0.5% by mass.

[0296] Advantageously, the zeolite contained in the final catalyst has an atomic ratio of Si / Al, measured by chemical analysis (for example by NMR), taking into account only the Al that are part of the network structure of the molecular sieve, between 4 and 500, preferably between 5 and 200, or more preferably between 12 and 150.

[0297] With regard to the conversion stage in which the conversion step (a) is implemented, the C1-C6 alcohol stream is brought into contact with the catalyst described above in a reaction zone of at least one reactor under operating conditions to produce the mixture containing paraffins, olefins, aromatics, and water as defined above.

[0298] In this step (a), converting alcohols, the mixture can generally be produced in a temperature range of 300°C to 600°C, in particular between 330°C and 550°C, especially between 350°C and 500°C or between 410°C and 580°C.

[0299] The pressure can also vary over a wide range. Preferred pressures are in the range of about 100 kPa to about 5 MPa, the most preferred range being about 150 kPa to about 1.0 MPa. The preceding pressures refer to the partial pressure of oxygen-containing organic compounds.

[0300] The conversion step (a) can be carried out in a single reaction zone or in several reaction zones arranged in series or in parallel. After a certain operating time, the catalyst must be regenerated.

[0301] In particular, several reactors can be used so that the exothermicity of the reaction is controlled in order to avoid excessive temperatures. Preferably, the maximum temperature difference within the same reactor will not exceed 100°C and preferably 75°C.

[0302] The reactor(s) may be of the isothermal or adiabatic type with fixed bed, moving bed and / or fluidized bed.

[0303] The conversion reaction can be carried out continuously in a configuration comprising a succession of fixed beds mounted in series, in at least one operating reactor in which the raw material passes from one fixed bed to another with cooling between them and at least one analogous reactor mounted in parallel, which undergoes a catalyst regeneration operation.

[0304] Advantageously, as described above, an additional flow of C2 to C6 alcohol is added between two fixed beds to control exothermicity.

[0305] The conversion reaction can also be carried out continuously in a configuration comprising a series of moving beds mounted in series, in which the raw material passes from one moving bed to another with cooling between them and in which the catalyst is mobile and circulates between the reactor and a catalyst regeneration zone.

[0306] The conversion reaction can also be carried out continuously in a configuration comprising a fluidized bed which forms a reaction zone where the reaction takes place and a fluidized bed which forms a regeneration zone where regeneration takes place (for example by controlled combustion in the presence of oxygen) or fluidized beds mounted in series, in which the raw material passes from one bed to another with cooling between them and in which the catalyst is mobile and circulates between the reactor and the catalyst regeneration zone.

[0307] Fluidized beds offer significant advantages, particularly when reactions are 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 fluidized beds exhibit quantitative similarities to those of gas bubbles in liquids.

[0308] The liquid-like behavior of a fluidized bed thus allows solids to be handled like a fluid, making it possible to supply and / or remove solids. Thorough mixing in a fluidized bed ensures a uniform temperature, even for highly exothermic reactions, and therefore allows for more flexible reactor control. Thorough mixing also improves contact between solids and fluids, and enhances heat and mass transfer.

[0309] There are many variations of fluidized beds, which are described, for example, in available technical manuals (e.g., Handbook of Fluidization and Fluid-Particle Systems, Taylor & Francis Group LLC, 2003). The fluidization phenomena of gas-solid systems depend greatly on the types of powders used. Several classifications exist, all based on the original work of Geldart. Many catalysts used in fluidized bed systems are Group A particles, characterized by dense phase expansion after minimal fluidization and before the onset of bubbling. Gas bubbles appear at the minimum bubbling velocity.

[0310] Fluidization regimes can be classified into two main categories: particulate (smooth) fluidization and aggregative (boiling) fluidization. In particulate fluidization, the solid particles generally disperse from The fluidization occurs in a relatively uniform manner within the fluidizing medium, without easily identifiable bubbles. Thus, particulate fluidization is sometimes also called homogeneous fluidization. In heterogeneous or aggregative fluidization, voids (bubbles) not containing solids are generally formed and are observed in a bubbling fluidized bed or in a bed exhibiting slugging or "slugging." For gas-solid systems, there are several distinct fluidization regimes: fixed bed, particulate fluidization, bubble fluidization, slugging fluidization, and turbulent fluidization; criteria are available for each. When the operating velocity is greater than the transport velocity, such that recycling of entrained particles is necessary to maintain a bed, additional fluidization regimes are possible.

[0311] Particulate regime: Umf < U < Umb

[0312] For powders of group A, the fixed bed expands homogeneously (particulate fluidization) above the minimum fluidization velocity (Umf) and no bubbles are observed as long as the velocity remains below the minimum bubbling velocity (Umb).

[0313] Bubble regime: Umb < U < Ums

[0314] Bubbles appear when the gas velocity is increased beyond the minimum bubbling velocity (Umb). The gas bubbles form above the distributor, coalesce, and grow. The bubbling regime is characterized by the coexistence of a bubble phase and a dense / emulsion phase. The majority of the fluidizing gas is present in the form of bubbles, and consequently, the gas velocity through the dense phase is very low.

[0315] Slugging regime: Ums < U < Uc

[0316] With large bed height-to-diameter ratios, the bed provides ample time This allows the bubbles to coalesce into larger ones. When the bubbles reach approximately the size of the bed's cross-section, the bed enters the "slugging" regime, characterized by the periodic passage of large bubbles forming plugs and a large, regular fluctuation in the bed's pressure drop. The velocity Uc corresponds to the bed's operating conditions where the plugs reach their maximum diameter and the amplitude of the pressure fluctuation is highest.

[0317] Transition to turbulent regime: Uc < U < Uk

[0318] When the gas velocity is continuously increased beyond this velocity Uc, large bubbles begin to fragment into smaller bubbles with a lower pressure fluctuation. This velocity is denoted Uk, and characterizes the transition between the bubbling regime and the turbulent regime.

[0319] Turbulent regime: Uk < U < Utr

[0320] Up to the transport velocity (Utr), the bed is in turbulent flow. The bubbles or the Voids are always present, although they are less distinguishable in the dense suspension. In this regime, the interactions between the gas voids and the dense phase / emulsion are vigorous and ensure efficient gas-solid contact.

[0321] Rapid fluidization regime: U > Utr

[0322] Beyond the transport velocity (Utr), particles begin to be entrained, and continuous operation is no longer possible without replacing or recycling the entrained and transported particles. Fast fluidized beds are typically characterized by a dense phase region at the bottom, near the distributor, coexisting with a dilute phase region at the top. The particle velocity increases with height in the bed, and therefore the bed density decreases.

[0323] Pneumatic conveying: U » Utr

[0324] All particles introduced to the bottom of the fluidized bed are transported in dilute phase with a concentration varying along the height of the bed.

[0325] A typical example of a reaction zone is the riser fluidized bed used in fluidic catalytic cracking (FCC) applications. Risers are vertical pipes with a high height-to-diameter ratio (>10), and the ideal riser approximates plug flow conditions, such that the catalyst and the fluid phase pass through the riser with minimal remixing.

[0326] In a transport fluidized bed reactor (fast fluidizing or pneumatically conveyed), core-ring flow can occur in which a high-velocity dilute core is surrounded by a denser, slower-moving ring. When circulating mass fluxes are low, the solids in the ring flow downwards along the wall. When circulating mass fluxes are high, the solids in the ring flow upwards along the wall. This non-uniform flow phenomenon results in inefficient gas-solid contact and suboptimal catalyst performance, and significant remixing of the gas and solids will occur, particularly when there is downward flow in the wall region.For rapid fluidization, internals are used to redistribute the axial and radial gas-solid flow structure, thus improving the uniformity of the gas-solid flow structure in space and promoting radial gas-solid exchange. Fluidized transport reactors require the recirculation of catalyst particles to the reactor bottom. This allows control of the catalyst density in the fluidized bed by recirculating more or less catalyst.

[0327] At the bottom of the fluidized bed, the feed fluid is distributed homogeneously over the cross-section of the reactor vessel. At the top of the reaction zone, the reaction vapors are separated from the entrained catalyst by means of deflectors, a disengagement zone, and cyclones. The catalyst is collected, free of hy remaining hydrocarbons and is advantageously returned to the bottom of the fluidized bed zone by vertical collectors (“standpipe”) and valves.

[0328] For the exothermic reaction such as the conversion carried out in step (a), it is preferable to have a homogeneous temperature throughout the catalytic bed (radially and axially) 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 in the reactor vessel.

[0329] Means of controlling the average reaction temperature consist of introducing the feed into the reaction zone at a temperature lower than the average bed temperature and / or removing heat from the catalyst bed by heat exchange. This heat exchange can be achieved by internal heat exchange tubes through which a cooling medium circulates and removes heat from the reactor vessel, or by external heat exchange by circulating the hot catalyst, collected at the top of the reactor, around the heat exchanger tubes and recirculating the cooled catalyst back into the reactor vessel.

[0330] With regard to catalyst regeneration, the C6 alcohol-to-Cl conversion reactor also includes a regeneration zone (or regenerator) whose main purpose is to remove coke deposits on the catalyst by combustion with oxygen. Regenerators are fast fluidized bed systems. Generally, the regenerator comprises a dense catalyst bed at the bottom of the vessel and a more dilute bed near the top of the vessel.

[0331] There are two types of regenerators, which operate either in partial combustion mode or in total combustion mode. In partial combustion mode, a quantity of air less than the stoichiometric quantity is supplied to the regenerator. Most of the carbon is converted to carbon monoxide and only a portion to carbon dioxide. Ideally, all the oxygen is consumed and no oxygen is present in the flue gases. The CO / CO2 ratio in the flue gases is generally between 0.5 and 2.0. In total combustion mode, an excess of air is supplied to the regenerator. Ideally, all the carbon in the coke is converted to carbon dioxide, and no carbon monoxide is present in the flue gases. The residual oxygen content in the flue gases is between 1.0% by volume and 3.0% by volume on a dry basis.

[0332] Partial combustion regenerators offer several advantages over total combustion regenerators, particularly when the catalyst is sensitive to high temperatures and the vapor environment: (i) it is possible to burn more coke for a given air flow rate, because the required air quantity is less than the stoichiometric quantity, and (ii) the heat of combustion released is lower, allowing for moderate temperature control and better preservation of catalytic activity in the presence of steam produced by the combustion of hydrogen.

[0333] A potential disadvantage of the partial-burn regenerator is the higher amount of coke remaining on the regenerated catalyst. In the case of total-burn regeneration, the carbon remaining on the catalyst is low, and the restoration of catalytic activity is greater. The potential disadvantage of total-burn regenerators is a greater release of heat due to the total combustion reaction and therefore a 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 burn primarily the hydrogen present in the coke, which has a higher reaction rate, as well as some of the carbon.In the second stage, using excess air, the remaining carbon is burned at a higher temperature into carbon dioxide, and thanks to the absence of water vapor in the second-stage regenerator, the deactivation of the catalyst at high temperature can be minimized.

[0334] The use of fluidized beds allows very precise control of the exothermicity of the reaction while offering continuous regeneration of the catalyst, promoting productivity and simplifying operations.

[0335] One or more diluents may be present in the Cl to C6 alcohol stream feeding the reaction zone, for example, in an amount of 1 mol% to 95 mol%, based on the total number of moles of all feed and diluent components introduced into the reaction zone.

[0336] Typical diluents include, but are not limited to, helium, argon, nitrogen, hydrogen, water (possibly recycled), paraffins, alkanes (especially methane, ethane, and propane), aromatic compounds, and mixtures thereof. Preferred diluents are water and nitrogen. Water can be injected in liquid or vapor form. The use of a diluent can offer two advantages. The first advantage is to reduce the partial pressure of the alcohol and thus improve selectivity for light olefins, primarily propylene. Generally, 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.There is an optimum for the yield of light olefins depending on the partial pressure, reaction temperature, hourly spatial velocity of the feed and the properties of the catalyst.

[0337] The second advantage of using a diluent is that it can act as a heat sink for the exothermic conversion of alcohols. Thus, the higher the specific molar heat capacity, the more heat can be absorbed by the Diluents. This second advantage might be less important in the case of fluidized bed reactors, as these are known to be excellent reactors for operating at a nearly homogeneous temperature throughout the catalyst bed. It is preferable that the diluents can be easily separated from the light olefin products, preferably by simple phase separation. Therefore, a preferred diluent is water. Diluents can be added at a rate of 1 mol% to 95 mol% of the combined feed (Cl-6 alcohol stream + diluents), preferably from 10 mol% to 75 mol%.

[0338] Step (b) of separating the water from the mixture

[0339] The water from the mixture containing paraffins, olefins, aromatics, and water produced in step (a) of conversion is separated from the mixture at a water separation stage, to form a water-depleted mixture.

[0340] The water separation step is preceded by a cooling step of the effluent from step (a), which condenses the water as well as some of the hydrocarbons. The operating temperature of this step is generally between 20°C and 100°C.

[0341] The separation is based, for example, on the difference in density and solubility between water and the remaining hydrocarbons. It is generally implemented in a three-phase separator tank allowing the separation of a water-rich aqueous phase (hereinafter referred to as the water separated from the mixture), a phase of liquid hydrocarbons, and a phase of gaseous hydrocarbons.

[0342] The water-depleted mixture contains less than 5% by mass, preferably less than 1% by mass of the water present in the mixture produced in step (a).

[0343] The water separated from the mixture advantageously contains less than 10% by mass of hydrocarbons.

[0344] The water separated from the mixture is optionally at least partially recycled in step (a) of conversion.

[0345] In the case where step (a) of conversion is implemented using several successive fixed catalytic beds, the water separated from the mixture is optionally recycled upstream of the fixed catalytic beds or between two fixed catalytic beds.

[0346] In this case, the mass ratio of recycled water to the alcohol flow in Cl to C6 in the feed supplied in step (a) of conversion is advantageously between 0 and 3, preferably between 0.05 and 2.

[0347] In the case where the conversion step (a) is implemented using at least one fluidized bed, the recycled water is optionally reinjected into the fluidized bed.

[0348] In this case, the mass ratio of recycled water to the alcohol flow in Cl to C6 in the feed supplied in step (a) of conversion is advantageously between 0 and 1, preferably between 0.05 and 0.5.

[0349] In both of the preceding cases, the recycled water forms a flow which controls the exothermicity of the reaction, decreases the partial pressures in hydrocarbons and modifies the acidity of the catalyst, which improves the selectivity in olefins.

[0350] The water separated but not recycled in step (a) of conversion is advantageously treated by stripping in a stripping column to separate the hydrocarbons it contains into an extracted hydrocarbon stream. The extracted hydrocarbon stream is reinjected in step (b) of separation.

[0351] In the case where conversion step (a) is carried out using at least one fluidized bed, a C4- hydrocarbon stream is optionally added to the fluidized bed to supplement or replace the recycled water. This stream is, for example, formed by at least a portion of a C1-C2 hydrocarbon fraction separated from the water-depleted mixture, which will be described below. This stream also controls the exothermicity of the reaction.

[0352] Advantageously, the C6 Cl alcohol stream is introduced at conversion step (a) at a temperature at least 5°C higher than the bubble point of the C6 Cl alcohol stream, and preferably lower than the temperature of the conversion reaction carried out in step (a), for example at least 50°C lower than the conversion reaction temperature, advantageously at least 100°C lower than the conversion reaction temperature.

[0353] Heating this stream absorbs the calories released by the conversion of the alcohols.

[0354] Optional steps for separating C1-C2 hydrocarbons and C3 hydrocarbons

[0355] Advantageously, the water-depleted mixture is introduced into a separation stage comprising at least one distillation column (hereinafter referred to as a de-analyzer) to separate the C1-C2 hydrocarbons (methane, ethane, ethylene) and other gaseous molecules lighter than C2 such as CO, CO2 and hydrogen from the rest of the water-depleted mixture.

[0356] The distillation column operates for example at a pressure greater than 20 barg and preferably greater than 30 barg.

[0357] A fraction of C1-C2 hydrocarbons is extracted from the top of the column. It contains more than 50% by mass of C1-C2 hydrocarbons and other gaseous molecules such as CO, CO2 and hydrogen.

[0358] The C1-C2 hydrocarbon fraction preferably contains more than 90% by mass, in particular more than 95% by mass of the C1-C2 hydrocarbons and other gaseous molecules such as CO, CO2 and hydrogen contained in the water-depleted mixture.

[0359] The C1-C2 hydrocarbon fraction is at least partially conveyed to an ethylene recovery unit, for example within a steam cracker. Thus, ethylene, even if produced in a minor quantity during conversion step (a), can to be valued.

[0360] Advantageously, following the conversion of the carbon dioxide added in step (a) of conversion into the stream containing carbon dioxide, the carbon dioxide, carbon monoxide and hydrogen present in the C1-C2 hydrocarbon fraction are optionally separated from the other hydrocarbons, in particular by cryogenic distillation, membrane separation or alternating pressure adsorption and combinations thereof.

[0361] These compounds are then advantageously recycled to a preliminary stage of alcohol synthesis, in particular by synthesis gas fermentation and by catalytic conversion of synthesis gas to produce in particular methanol.

[0362] In particular, the methanol or ethanol thus produced forms part of the alcohol stream in Cl to C6 which is then converted in conversion step (a).

[0363] In an advantageous embodiment, particularly applicable to the case of a conversion step (a) implemented using at least one fluidized bed, a portion of the C1-C2 hydrocarbon fraction is recycled to the conversion step (a) of the alcohol stream to C6 Cl as a recycle stream. For example, the ratio of the mass flow rate of the portion of the C1-C2 hydrocarbon fraction recycled to the conversion step (a) to the mass flow rate of the C1-C2 hydrocarbon fraction from the distillation column is less than 1 and is in particular between 0.1 and 0.8.

[0364] The C3+ hydrocarbon fraction is recovered at the bottom of the column. It comprises more than 90% by mass of the C3+ hydrocarbons contained in the water-depleted mixture.

[0365] In its dry form, it advantageously contains:

[0366] - less than 5% by mass of C1-C2 hydrocarbons;

[0367] - less than 15% by mass of paraffins, in particular between 3% and 10% by mass of paraffins;

[0368] - more than 40% by mass of C3 to C7 olefins, in particular between 50% and 80% of C3 to C7 olefins;

[0369] - less than 5% by mass of C8+ olefins, in particular between 0.1% by mass and 4.0% in mass of C8+ olefins; and / or

[0370] - more than 6% aromatics, in particular between 6% and 20% aromatics.

[0371] C3 to C7 olefins generally contain propylene. The C3+ hydrocarbon fraction contains, in some cases, more than 30% by mass of propylene.

[0372] In one embodiment, the C3+ hydrocarbon fraction is sent directly to the oligomerization and alkylation step. Alternatively, in a particular embodiment, an additional step for separating C3- hydrocarbons, in particular propylene, is carried out in a second dis- column. tillation. This separation allows for the recovery of propylene.

[0373] The second distillation column operates, for example, at a pressure greater than 5 barg, and preferably greater than 10 barg. It produces at the top a fraction of C3- hydrocarbons, containing more than 50% by mass of propylene, and at the bottom, a fraction of C4+ hydrocarbons.

[0374] The C3- hydrocarbon fraction preferably contains more than 90% by mass, in particular more than 95% by mass of the C3- hydrocarbons contained in the C3+ hydrocarbon fraction from the first distillation column.

[0375] The C4+ hydrocarbon fraction comprises more than 90% by mass of the C4+ hydrocarbons contained in the water-depleted mixture.

[0376] In its dry form, it advantageously contains:

[0377] - less than 5% by mass of C3- hydrocarbons,

[0378] - less than 25% by mass of paraffins, in particular between 10% and 15% by mass of paraffins,

[0379] - more than 15% by mass of C4 to C7 olefins, in particular between 25% and 40% of C4 to C7 olefins,

[0380] - less than 2% by mass of C8+ olefins, in particular between 0.5% by mass and 1.0% in mass of C8+ olefins; and / or

[0381] - more than 6% aromatics, in particular between 7% and 20% aromatics, in particular less than 5% aromatics in C9+.

[0382] Step c) of oligomerization of olefins from the water-depleted mixture and

[0383] step (d) alkylation of aromatics from the water-depleted mixture

[0384] In a first embodiment, step (c) of oligomerization of olefins from the water-depleted mixture and step (d) of alkylation of aromatics from the water-depleted mixture are carried out jointly in the same reactor or the same reactors of the same reaction stage.

[0385] Advantageously, the hydrocarbon charge formed from the C3+ hydrocarbon fraction or the C4+ hydrocarbon fraction described above is oligomerized for its olefins and alkylated for its aromatics by contacting an acid catalyst

[0386] For example, a multi-reactor installation may be used in which the exothermicity of the reaction can be controlled so as to avoid excessive temperatures. Preferably, the maximum temperature difference within the same reactor will not exceed 100°C and preferably 75°C.

[0387] The reactor(s) may be of the isothermal or adiabatic type with a fixed or moving bed. The oligomerization reaction of olefins and alkylation of aromatics may be carried out continuously in a configuration comprising a series of fixed beds mounted in series, in at least one operating reactor in which the The raw material passes from one bed to another with cooling in between, and at least one similar reactor mounted in parallel undergoes a catalyst regeneration operation. The oligomerization of olefins and alkylation of aromatics can be carried out continuously in a configuration comprising a series of moving beds mounted in series, in which the raw material passes from one bed to another with cooling in between, and in which the catalyst is mobile and circulates between the reactor and the catalyst regeneration zone.

[0388] Preferably, the aforementioned steps are carried out jointly using at least two successive reactors. In the case of fixed-bed reactors, alternatively a single reactor may contain several catalytic beds with cooling systems between the beds or be equipped with a quenching flow injection to lower the temperature between the beds.

[0389] The reaction conditions of the first reactor are chosen so as to convert some of the low carbon number olefinic compounds (C3-C8) into intermediate olefins (C8+) and the alkylation of the aromatics by light olefins.

[0390] Advantageously, the first reactor comprises a first catalytic zone and operates at a high temperature, for example greater than or equal to 200°C, and preferably less than 350°C, and a pressure between 25 bar and 60 bar

[0391] The second reactor preferably operates at temperatures and pressures chosen to promote the conversion of some of the low-carbon (C3-C8) olefinic compounds into intermediate (C8+) olefins and the alkylation of aromatics by light olefins. The effluent from the first reactor, comprising unreacted olefins, intermediate olefins, aromatics, water, and possibly other compounds such as paraffins and possibly a reducing gas, then undergoes oligomerization and / or alkylation in this second reactor, which includes a second catalytic zone, resulting in a heavier hydrocarbon effluent rich in distillate.

[0392] A cooling section is advantageously provided between two successive reactors and possibly a flash balloon.

[0393] The mass flow rate through the oligomerization reactor(s) is advantageously sufficient to allow a relatively high conversion, without being too low to avoid undesirable parallel reactions.

[0394] The weight hourly space velocity (WHSV) of the load is, for example, 0.1 h₁ at 20 h₁, preferably 0.5 h₁ at 10 h₁, and even more preferably 0.8 h₁ at 5 h₁.

[0395] The temperature at the inlet of the reactor(s) is advantageously sufficient to allow a relatively high conversion, without being too high to avoid undesirable parallel reactions.

[0396] The temperature at the inlet of the reactor or of each reactor is for example from 150°C to 400°C, preferably from 180°C to 350°C, even more preferably from 200°C to 290°C.

[0397] The pressure through the olefin oligomerization and aromatic alkylation reactor(s) is advantageously sufficient to allow a relatively high conversion, without being too low to avoid undesirable parallel reactions.

[0398] The pressure through the reactor or each reactor is for example from 8 bara to 100 bara, preferably from 10 bara-85 bara, even more preferably from 25 bara to 75 bara (bars, absolute pressure).

[0399] As regards the nature of the catalyst, a first family of catalysts used includes an acid catalyst of either amorphous or crystalline aluminosilicate type, or a silicoaluminophosphate, in the form of H+, chosen from the following list and containing or not alkali elements or rare earths:

[0400] MFI (ZSM-5, silicalite-1, boralite C, TS-1), MEL (ZSM-11, silicalite-2, boralite D, TS-2, SSZ-46), ASA (amorphous silica-alumina), MSA (mesoporous silica-alumina), FER (Ferrierite, FU-9, ZSM-35), MTT (ZSM-23), MWW (MCM-22, PSH-3, ITQ-1, MCM-49), TON (ZSM-22, Theta-1, NU-10), EUO (ZSM-50, EU-1), ZSM-48, MFS (ZSM-57), MTW, MAZ, BEA (Beta zeolite), MOR (mordenite), FAU (faujasite zeolite type), LTL (L zeolite), Omega zeolite and the family of microporous materials composed of silica, aluminum, oxygen and possibly boron.

[0401] Zeolite can be subjected to various treatments prior to use, which may be: ion exchange, modification with metals, steam treatment, acid treatments or any other dealumination method, surface passivation by silica deposition, or any combination of the above-mentioned treatments.

[0402] The alkali or rare earth content is from 0.05% by mass to 10% by mass, preferably from 0.2% by mass to 5% by mass. Preferably, the metals used are Mg, Ca, Ba, Sr, La, Ce used alone or in mixtures.

[0403] A second family of catalysts used comprises phosphorus-modified zeolites optionally containing an alkali or a rare earth element. In this case, the zeolite may be chosen from the following list:

[0404] 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.

[0405] Zeolite may be subjected to various treatments before use, which may to be: ion exchange, modification with metals, steam treatment, acid treatments or any other dealumination method, mesoporization treatments, surface passivation by silica deposition, or any combination of the aforementioned treatments.

[0406] The alkali or rare earth content is from 0.05% by mass to 10% by mass, preferably from 0.2% by mass to 5% by mass. Preferably, the metals used are Mg, Ca, Ba, Sr, La, Ce used alone or in mixtures.

[0407] A third family of catalysts used includes bifunctional catalysts, comprising:

[0408] - a support, from the following list: MFI (ZSM - 5, silicalite- 1, boralite C, TS-1), MEL (ZSM-11, silicalite-2, boralite D, TS-2, SSZ-46), AS A (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 microporous materials of the ZSM-48 family consisting of silicon, aluminum, oxygen and optionally boron. MFI or MEL (Si / Al> 25), MCM-41, MCM-48, SBA-15, SBA-16, SiO2, A12O3, hydrotalcite, or a mixture of these.

[0409] - a metallic phase (Me) at a level of 0.1% by mass, the metal being selected Among the following elements: Zn, Mn, Co, Ni, Ga, Fe, Ti, Zr, Ge, Sn, and Cr, used alone or in mixtures. These metallic atoms can be incorporated into the tetrahedral structure of the support. The incorporation of this metal can be achieved either by adding it during the synthesis of the support, or by incorporating it after synthesis through ion exchange or impregnation, in which case the metals are incorporated as cations, and not integrated into the structure of the support.

[0410] Zeolite can be subjected to various treatments before use, which may be: ion exchange, modification with metals, steam treatment, acid treatments or any other dealumination method, mesoporization treatments, surface passivation by silica deposition, or any combination of the above-mentioned treatments.

[0411] The alkali or rare earth content is from 0.05% by mass to 10% by mass, preferably from 0.2% by mass to 5% by mass. Preferably, the metals used are Mg, Ca, Ba, Sr, La, Ce used alone or in mixtures.

[0412] A fourth family of catalysts used includes amorphous solids such as silica-alumina, silica-phosphate, silica-borate, silica-titanium, silica-zirconia and / or mixtures.

[0413] The catalyst can be a mixture of the materials described above in the four catalyst families. In addition, the active phases can also be combined with other constituents (binder, matrix) giving the final catalyst a increased mechanical resistance, or an improvement in activity.

[0414] If the hydrocarbon feedstock is oligomerized in an installation comprising several reactors in series, the reactors in series can be loaded with the same catalyst or different catalysts.

[0415] In one variant, step (c) of oligomerization of olefins is carried out in an oligomerization reactor, and step (d) of alkylation of aromatics is carried out in an alkylation reactor, separately from step (c) of oligomerization of olefins.

[0416] Advantageously, the water-depleted mixture is then separated in the first column into the Cl-C2 hydrocarbon fraction, taken, for example, from the top of the column, a C3-C5 hydrocarbon fraction, for example, drawn from an intermediate stage of the column, and a C6+ hydrocarbon fraction, for example, taken from the bottom of the column. The Cl-C2 hydrocarbon fraction and the C6+ hydrocarbon fraction are sent to the alkylation step (d) in the alkylation reactor, while the C3-C5 hydrocarbon fraction is sent to the oligomerization step (c) in the oligomerization reactor.

[0417] Alternatively, the water-depleted mixture is separated into a C3- hydrocarbon fraction, taken, for example, from the top of the column, a C4-C5 hydrocarbon fraction, for example, drawn from an intermediate stage of the column, and a C6+ hydrocarbon fraction, for example, taken from the bottom of the column. The C3- hydrocarbon fraction and the C6+ hydrocarbon fraction are sent, at least in part, to the alkylation step (d) in the alkylation reactor, while the C4-C5 hydrocarbon fraction is sent, at least in part, to the oligomerization step (c) in the oligomerization reactor.

[0418] The oligomerization reactor product contains more than 50% by mass of C7+ olefins, in particular more than 60% by mass of C9 to C12 olefins.

[0419] Alkylation takes place under temperature and pressure conditions effective in maintaining more than 20% by mass of the filler in the liquid phase within the alkylation zone.

[0420] Advantageously, the alkylation of aromatics with alkenes is carried out in the liquid phase, since aromatics are present primarily in the liquid phase. It can be performed using acidic solid catalysts. Zeolites and silica-alumina catalysts exhibiting shape selectivity are generally used.

[0421] In this process, the reactor conditions are chosen so that the alkene introduced into the reactor is predominantly dissolved in the aromatic feedstock. This is generally achieved through an optimal combination of operating conditions, such as pressure, temperature, and the choice of a catalyst with sufficiently high catalytic activity. The presence of alkenes in the gas phase can cause rapid deactivation of the alkylation catalysts typically used in the liquid phase.

[0422] Examples of usable operating conditions are given in US 4,891,458, which describes the liquid-phase synthesis of ethylbenzene with beta zeolite, while US 5,334,795 describes the use of MCM-22 in the liquid-phase synthesis of ethylbenzene; US 7,649,122 describes the use of MCM-22 in the liquid-phase synthesis of ethylbenzene in the presence of a maintained water content. US 4,549,426 describes the liquid-phase synthesis of alkylbenzene with vapor-stabilized Y zeolite. US 8,134,036 describes the liquid-phase aromatic alkylation on at least one catalytic bed containing a first catalyst modified by the inclusion of a rare-earth metal ion.

[0423] The types of products that can preferably be produced correspond to the following generic chemical formulas: monoalkyl benzene, dialkyl benzene, and trialkyl benzene. The alkyl chains (Rx) each have from 2 to 10 carbon atoms, preferably from 2 to 6 carbon atoms. These chains may be of equal length or of different lengths.

[0424] The aromatic compounds produced during step (a) of conversion of Cl to C6 alcohols are typically mono-aromatics, possibly alkylated (benzene, toluene, ethylbenzene and xylenes), the alkylating agent being olefins.

[0425] The alkylation reaction is exothermic, therefore it may be useful to inject some of the aromatic compounds and / or some of the olefins between the different beds of the reactor, if there are several beds. Aromatics with fewer than 8 carbon atoms can be recycled, as well as olefins that are too short, for example, those with fewer than 5 carbon atoms.

[0426] The alkylation catalyst is, for example, in the form of beads, but it is most often in the form of extrudates. It consists of an acidic solid mixed with an amorphous phase. The acidic solid is shaped by means of a matrix, which is an amorphous phase. The acidic solid is preferably at least one zeolite, preferably chosen from among the structurally FAU type zeolites, and more particularly zeolite Y, the structurally MOR type zeolites (mordenite zeolite), the structurally EUO type zeolites (i.e., EU-1, ZSM-50, TPZ-3 zeolites), the structurally NES type NU-87 zeolite, the NU-86 zeolite (described in EP 463 768 A), the NU-85 zeolite (described in EP 462 745 A), the NU-88 zeolite (described in FR 2 752 567), and the IM-5 zeolite (described in FR 2 754 809), the Beta zeolite, the zeolite MCM-22, MCM-36 zeolite, MCM-49 zeolite or MCM-56 zeolite.

[0427] Preferably, the catalyst is a beta zeolite having a silica / alumina molar ratio (expressed as SiO2 / Al2O3) of about 10 to about 200 or about 20 to about 50.

[0428] Beta zeolite may have a low sodium content, for example less than about 0.2% by mass expressed as Na2O, or less than about 0.02% by mass. The content in sodium can be reduced by any method known to man skilled in the art, such as by ion exchange.

[0429] These zeolites are at least partly in acidic form (H+), but may also contain cations other than H+, such as alkaline earth metals or rare earths. The zeolite catalyst may be modified with a rare earth metal ion, such as lanthanum, cerium, neodymium, or praseodymium, for example.

[0430] The BET surface area of ​​the catalyst used is in particular between 50 m² / g and 900 m² / g, preferably between 100 m² / g and 700 m² / g. The Na / Al ratio of the final catalyst is less than 5 atomic % and preferably less than 2%.

[0431] The zeolite content in the catalyst is in particular between 5% by mass and 95% by mass, preferably between 10% by mass and 90% by mass relative to the final catalyst. The overall Si / Al ratio of these zeolites is between 2.6 and 200, preferably between 5 and 100, and even more preferably between 5 and 80.

[0432] The catalyst matrix is ​​a support selected from the group consisting of alumina, silica, silica-alumina, alumina-boron oxide, magnesia, silica-magnesia, zirconia, titanium oxide, and clay, these compounds being used alone or in mixtures. Preferably, an alumina support is used.

[0433] Preferably, the solid acid catalyst has shape selectivity to prevent the formation of excessively large alkylaromatics, such as those with more than 16 carbon atoms. If the molecular size of the alkylaromatics is close to the size of the micropores in the catalyst, formation and diffusion at the pore outlet are still possible; however, the formation of alkylaromatics that are too large to enter, reside, or exit the pores is thus prevented.

[0434] As stated above, the reaction zone is operated at a temperature and pressure such that they maintain phase conditions preferably exhibiting more than 20% by mass of liquid.

[0435] For the production of alkylaromatics, having at least 8 carbon atoms, the reaction temperature is in particular between 140°C and 320°C, and is generally between 160°C and 280°C. In one embodiment, the reaction temperature is between 190°C and 240°C.

[0436] The alkylation pressure is generally maintained at a sufficiently high level to ensure the presence of a liquid phase. In one embodiment, the pressures are between 20 barg and 100 barg, in particular from 30 barg to 50 barg.

[0437] When operating under predominantly liquid phase conditions, an upflow reactor mode is generally used. Flow rates can typically vary from the liquid hourly spatial velocity (LHSV) between about 1 h⁻¹ and 100 h⁻¹ per bed, preferably between about 2 h⁻¹ and 70 h⁻¹ per bed. The aromatic / alkylating agent ratio is, for example, between 0.05 mol / mol and 20 mol / mol and preferably- typically between 0.1 mole / mol and 10 mole / mol.

[0438] In a preferred operating mode, the oligomerization of olefins and the alkylation of aromatics with olefins are carried out on the same catalyst and in the same reactor. The known operating conditions for oligomerization and alkylation are very similar and can easily be adapted to obtain the desired performance in terms of oligomerization and alkylation.

[0439] When oligomerization and alkylation are carried out simultaneously in the same reactor using the same catalyst, the reactor product contains more than 10% by mass of C8+ aromatics, in particular more than 6% by mass of C8 to C14 aromatics.

[0440] When alkylation in the presence of olefins is carried out separately from oligomerization, the reactor product contains more than 65% by mass of C8+ aromatics, in particular more than 75% by mass of C8 to C14 aromatics.

[0441] Step (e) deformation of the hydrocarbon stream to be hydrogenated

[0442] A hydrocarbon stream to be hydrogenated is formed from at least a portion of the oligomerized olefins in step (c) and at least a portion of the alkylated aromatics in step (d).

[0443] In the case where a single common stage is used for step (c) of olefin oligomerization and for step (d) of aromatic alkylation, the product of this stage is used in whole or in part to form the hydrocarbon stream to be hydrogenated.

[0444] This product includes, for example, a dry base

[0445] - less than 15% by mass of paraffins, in particular between 3% and 10% by mass of paraffins,

[0446] - less than 10% by mass of C4 to C7 olefins, in particular between 0.5% and 5% of C4 to C7 olefins,

[0447] - more than 50% by mass of C8 to C16 olefins, in particular between 60% and 80% of C4 to C16 olefins,

[0448] - less than 5% by mass of C17+ olefins, in particular between 0.1% by mass and 1.0 % by mass of olefins as Cl7+; and / or

[0449] - less than 5% aromatics in C6 to C7, in particular between 0.5% and 4.0% aromatics in C6 to C7,

[0450] - more than 2% of C8+ aromatics, in particular between 6% and 30% of aromatics in C8+.

[0451] In one variant, an optional step of separating the effluent containing at least some of the oligomerized olefins in step (c) and / or at least some of the alkylated aromatics in step (d) is carried out in an additional distillation column.

[0452] The separation produces at the top a fraction of C7- hydrocarbons, and at the bottom, a fraction of C8+ hydrocarbons.

[0453] The C7- hydrocarbon fraction preferably contains more than 90% by mass, in particular more than 95% by mass of the C7- hydrocarbons contained in the effluent.

[0454] The C8+ hydrocarbon fraction comprises more than 90% by mass of the C8+ hydrocarbons contained in the effluent.

[0455] Advantageously, at least a part, for example less than 50% by mass, of the C7- hydrocarbon fraction is at least partially recycled in step (c) of olefin oligomerization and / or in step (d) of aromatic alkylation, another part forming a gasoline stream.

[0456] The hydrocarbon stream to be hydrogenated is formed by at least a part, preferably by the whole of the C8+ hydrocarbon fraction.

[0457] In the case where step (c) of oligomerization of olefins is carried out in an oligomerization reactor, and step (d) of alkylation of aromatics is carried out in an alkylation reactor, separately from step (c) of oligomerization of olefins, the product of the oligomerization reactor and the product of the alkylation reactor containing the alkylated aromatics are advantageously subjected to separation.

[0458] In one embodiment, the oligomerization reactor product containing the oligomerized olefins and the alkylation reactor product containing the alkylated aromatics are separated, at the bottom, into the C8+ hydrocarbon fraction and at the top, into the C7- hydrocarbon fraction. The C7- hydrocarbon fraction is recycled at least in part (for example, less than 50% by mass) in step (c) in the oligomerization reactor.

[0459] At least part, preferably all, of the C8+ hydrocarbon fraction forms the hydrocarbon stream to be hydrogenated.

[0460] In another embodiment, the product of the oligomerization reactor and the product of the alkylation reactor are separated in the distillation column into a fraction of C7- hydrocarbons, drawn off at the top, a fraction of C8 to C16 hydrocarbons, taken at an intermediate stage, and a fraction of C17+ hydrocarbons, drawn off at the bottom.

[0461] At least part, preferably all, of the C8 to C16 hydrocarbon fraction forms the hydrocarbon stream to be hydrogenated.

[0462] The C17+ hydrocarbon fraction is at least partly recycled in step (a) of converting the alcohol stream into C6 Cl. This recycling is carried out in order to crack the C17+ olefins again.

[0463] Hydrogenation step (f)

[0464] The hydrocarbon stream to be hydrogenated undergoes hydrogenation to form a hydrogenated hydrocarbon stream. This saturates the olefinic compounds and hydrogen partially binds aromatic compounds.

[0465] Hydrogenation is carried out for example in one or more fixed bed reactors (falling or rising) and in mixed phase, the fraction to be hydrogenated being mainly in liquid phase.

[0466] Hydrogenation is carried out, for example, at a temperature between 50°C and 350°C, in particular between 100°C and 300°C. It is carried out under a pressure preferably greater than 10 bara and in particular between 20 bara and 80 bara.

[0467] A hydrogen stream is fed into the reactor(s) mixed 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 (excluding the recycled stream) to be hydrogenated is advantageously between 50 NL / L and 3000 NL / L, in particular between 100 NL / L and 500 NL / L. The hydrogen can be added to the hydrocarbon stream in several stages along the catalytic bed. The hourly spatial velocity is advantageously between 0.5 and 3, and in particular between 1 and 2 h⁻¹. Excess hydrogen can be recycled to the reaction zone after separation and compression.

[0468] The reaction is carried out in the presence of at least one catalyst comprising one or more Group VIII metals (typically Pt, Pd, Ni) supported on a support such as silica, alumina, or any mixture of these two compounds or carbon. The reaction can also be carried out in the presence of a sulfide-type catalyst containing an element from Group VIB (Cr, Mo, W) and an element from Group VIIIB (Fe, Ru, Co, Os, Co, Rh, Ir, Pd, Ni, Pt) or mixtures of these two groups of metals.

[0469] The hydrogenated hydrocarbon stream advantageously contains less than 10% by mass of olefins and preferably less than 3% of olefins.

[0470] It preferably contains more than 50% by mass of paraffins, in particular more than 50% by mass of C7 to C17 paraffins, in particular more than 60% by mass of C7 to C17 paraffins, especially between 70% by mass and 95% by mass of C7 to C17 paraffins.

[0471] According to a variant, between 10% and 90% by mass of the aromatics contained in the hydrocarbon stream to be hydrogenated, preferably between 30% by mass and 80% by mass of the aromatics contained in the hydrocarbon stream to be hydrogenated are hydrogenated into cycloparaffins.

[0472] The hydrogenated hydrocarbon stream includes, for example:

[0473] - more than 20% by mass of C9 hydrocarbons, for example between 20% and 40% in mass of C9 hydrocarbons,

[0474] - more than 20% by mass, for example between 20% by mass and 40% by mass of hydrocarbons in Cl2,

[0475] - more than 6% by mass of C6+ aromatics, in particular more than 6% by mass of hydrocarbon in C8+.

[0476] Step (g) recovery of the jet fuel fraction

[0477] To recover the jet fuel fraction, the hydrogenated hydrocarbon stream advantageously undergoes fractionation. This fractionation can be carried out by passing through at least one separation column, for example a distillation column

[0478] Preferably at least one first separation column is used to separate, at the top, a fraction of liquefied petroleum gas from the rest of the hydrogenated hydrocarbon stream which is obtained at the bottom.

[0479] The splitting conditions in this column are as follows: pressure advantageously between 2 bara and 15 bara, top condensation temperature adjusted to allow the use of an air or water refrigeration condenser, i.e. preferably between 20°C and 50°C

[0480] The residual fraction of hydrogenated hydrocarbon stream is then introduced into a second separation column to produce a naphtha fraction at the top, a diesel fraction at the bottom, and the jet fuel fraction at at least one intermediate stage. This fractionation can be carried out in a single column including a lateral withdrawal of the jet fuel fraction, or in two separate columns.

[0481] More than 80% by mass of the hydrogenated hydrocarbon stream introduced into the second column advantageously forms the jet fuel fraction.

[0482] The liquefied petroleum gas fraction preferably has a final boiling point below 180°C, and even more preferably a final boiling point below 150°C.

[0483] The initial boiling point can be from 20°C to 60°C and preferably from 25°C to 40°C

[0484] The naphtha fraction comprises more than 80% by mass of the C8- paraffins contained in the residual fraction.

[0485] The jet fuel fraction comprises between 2% by volume and 30% by volume of C8+ aromatics, preferably between 6% by volume and 25% by volume of C8+ aromatics and even more preferably between 8% by volume and 25% by volume of C8+ aromatics.

[0486] It comprises more than 50% by volume of C9 to Cl6 paraffins, in particular between 60% by volume and 95% by volume of C9 to C16 paraffins.

[0487] The jet fuel fraction includes, in particular, more than 60% by volume of C9 to C12 paraffins.

[0488] It preferably has a final boiling point below 400°C, and even more preferably a final boiling point below 350°C.

[0489] The initial boiling point can be from 130°C to 180°C.

[0490] The diesel fraction is the heaviest fraction, whose molecules cannot all can be used to form jet fuel. This fraction typically presents an initial boiling temperature above 300 °C, and preferably above 310 °C.

[0491] Advantageously, at least part of the diesel fraction is recycled in step (a) of converting the alcohol stream to Cl to C6 as a recycling stream in order to generate a new cracking of the compounds present in this fraction.

[0492] The recycling stream advantageously constitutes between 10% by mass and 50% by mass of the Cl to C6 alcohol stream introduced in conversion step (a). DESCRIPTION OF FIGURES

[0493] The invention will be better understood upon reading the following description, given solely by way of example, and made with reference to the accompanying drawings, in which:

[0494] - [Fig.1] Fig.1 is a schematic view of an installation configured for the implementation of a first method for manufacturing jet fuel according to the invention;

[0495] - [Fig.2][Fig.3][Fig.4][Fig.5][Fig.6][Fig.7] Figures 2 to 7 are analogous views to that of [Fig.1] illustrating installation variants intended for the implementation of variants of the process of [Fig.1]:

[0496] - [Fig.8] [Fig.8] is a detail view illustrating an implementation reactor of the fluidized bed conversion stage;

[0497] - [Fig.9] [Fig.9] is an NMR spectrum of a parent catalyst ZSM5 and a ca modified alyseur having an ALPO structure.

[0498] DESCRIPTION OF WAYS TO EMBODI THE INVENTION

[0499] A first manufacturing installation 10, intended for the implementation of a process for manufacturing jet fuel according to the invention, is schematically illustrated in [Fig. 1].

[0500] The first installation 10 includes a conversion stage 12 of an alcohol stream 14 into C6 Cl, intended to produce a mixture 16 containing paraffins, olefins, aromatics, and water, and a stage 18 for separating the water from the mixture 16 to produce a water-depleted mixture 19 comprising a liquid phase 19a and a gas phase 19b. The installation of [Fig. 1] also includes a stage 20 for separating the hydrocarbons into C2 Cl from the water-depleted mixture.

[0501] In this example, the installation 10 further comprises a joint stage 22 for oligomerizing olefins and alkylating aromatics from the water-depleted mixture, the stage 22 producing a stream 24 of hydrocarbons to be hydrogenated

[0502] The installation 10 further comprises a stream hydrogenation stage 26 for the stream 24, producing a stream of hydrogenated hydrocarbons 30, and a stream fractionation stage 28 for the hydrogenated hydrocarbons 30 stream, intended to fractionate at least a jet fuel fraction 34, and advantageously a diesel fraction 36 and a naphtha fraction 38.

[0503] The conversion stage 12 is intended to implement the conversion step (a) described above, transforming the alcohols in C6 Cl predominantly into C3 to C7 olefins.

[0504] As described above, the conversion stage 12 comprises, for example, at least one fixed-bed reactor, for example several fixed-bed reactors, the fixed-bed reactor(s) advantageously defining several successive catalytic fixed beds, as described above.

[0505] According to the invention, the conversion stage 12 comprises at least one supply conduit for a stream 182 containing carbon dioxide, the conversion stage 12 of the alcohol stream 14 to C6 Cl being configured to convert carbon dioxide from the stream 182 to carbon monoxide jointly with the conversion of the alcohol stream to C6 Cl.

[0506] The supply conduit opens for example upstream of the reactor(s) of the conversion stage 12, so that the stream containing carbon dioxide 182 mixes with the stream 14 of alcohol in Cl to C6, and / or in a reactor of the conversion stage 12 and / or between two reactors of the conversion stage 12.

[0507] With reference to the above description, the separation stage 18 is intended for implementing the water separation step (b). It includes at least one separator operating by gravity and / or mechanical drive to separate the water from the mixture 16, allowing the recovery of a water-concentrated aqueous fraction (flow 40), a hydrocarbon fraction in gas phase 19b and a hydrocarbon fraction in liquid phase 19a.

[0508] The installation 10 optionally includes at least one separate water recycle conduit 18a from the separation stage to the conversion stage 12. The recycle conduit 18a opens, for example, upstream of the fixed catalytic beds or between two successive fixed catalytic beds in the case of a conversion stage 12 comprising at least one fixed bed reactor.

[0509] Advantageously, the separation stage 18 includes a stripping column 41 suitable for treating at least part of the separated water forming the stream 40 to extract the hydrocarbons it contains and obtain treated water.

[0510] The installation 10 includes at least one cooling device (for example a heat exchanger) for reducing the temperature of the product exiting the reactor and heating another stream, for example such as the reactor feed 12, a water or air cooler and / or a combination thereof downstream of the conversion stage 12 to condense the water and produce the stream 40.

[0511] The separation stage 20 is intended to implement step (c) of separating hydrocarbons lighter than C3, such as C1-C2, and light CO compounds, CO2, hydrogen, as defined above. It advantageously includes at least one de-ethane generator. The separation stage 20 operating at a pressure higher than that of the water separation stage 18, the installation includes at least one pump suitable for increasing the pressure of the liquid phase 19a and at least one compressor suitable for increasing the pressure of the gas phase 19b.

[0512] Stage 22 for olefin oligomerization and aromatic alkylation is intended to jointly implement steps (d) and (e). It comprises at least one joint oligomerization and / or alkylation reactor intended to implement the experimental conditions described above.

[0513] The hydrogenation stage 26 is intended to implement step (f). It comprises at least one fixed-bed hydrogenation reactor intended to implement the hydrogenation reaction under the conditions described above.

[0514] The fractionation stage 28 is intended to implement step (g). In this example, it comprises at least a first upstream column 42 for separating liquefied petroleum gas 44, and a second downstream column 46 for fractionation, intended to produce fractions 34 to 38.

[0515] A first example of a process for manufacturing jet fuel, implemented in the installation of [Fig.1], will now be described.

[0516] Initially, a stream 14 of alcohol in Cl to C6 is brought into the conversion stage 12. The alcohol stream 14 is for example from a source 50 described above, in which the alcohols of the source 50 are produced for example by fermentation of biomass, by catalytic conversion of carbohydrates or of carbon monoxide or carbon dioxide in the presence of hydrogen.

[0517] The stream 14 comprises the composition described above, for example more than 50% by dry mass of methanol and preferably more than 80% by dry mass of methanol.

[0518] The stream 14 is introduced into the conversion stage 12 where it undergoes a conversion described above comprising a dehydration / aromatization of the alcohols in C2 to C6, and for methanol, a conversion to dimethyl ether followed by a dehydration.

[0519] The reaction is carried out under the operating conditions of temperature and pressure described above. One or more catalysts defined above are used.

[0520] According to the invention, a current containing carbon dioxide 182 is added to the conversion stage 12.

[0521] The carbon dioxide stream 182 comprises more than 5% by mass of carbon dioxide as defined above.

[0522] The mass ratio in the charge supplied to the conversion stage 12 of carbon dioxide to alcohols in Cl to C6 is between 5% and 75%.

[0523] The carbon dioxide contained in the carbon dioxide stream 182 is converted to carbon monoxide in conjunction with the conversion of the alcohol stream Cl to C6.

[0524] A mixture 16 containing paraffins (in particular n-paraffins, i-paraffins and cycloparaffins), olefins, aromatics and water is thus obtained. The mixture 16 comprises, for example, the composition described above.

[0525] The mixture 16 is then introduced into the separation stage 18 to produce a water flow 40 at the separator foot, and the water-depleted mixture 19 comprising the gas phase 19b and the liquid phase 19a.

[0526] A portion 40a of the water stream 40 is optionally recycled to the conversion stage 12, via the conduit 18a, as described above. Another portion 40b of the water stream 40 is introduced into a column 41 to undergo stripping and produce, at the top, a stream 41a of extracted hydrocarbons and, at the bottom, a treated water stream 40b having a hydrocarbon content lower than that of the water stream 40.

[0527] The extracted hydrocarbon stream 41a is recycled at the separation stage 18, for example upstream of the separator.

[0528] The gas phase 19b and the liquid phase 19a are then introduced, after compression, into the de-anaerobic digester in the separation stage 20. The de-anaerobic digester operates under the conditions defined above and produces at the top, a fraction 60 of C1-C2 hydrocarbons which may contain light compounds such as CO, CO2, hydrogen and at the bottom, a fraction 62 of C3+ hydrocarbons.

[0529] The fractions 60, 62 obtained have the compositions defined above

[0530] Preferably, the C1-C2 hydrocarbon fraction 60, possibly after separation, is sent to a steam cracker for recovery of at least part of the ethylene it contains.

[0531] In a variant shown in dotted lines on [Fig.1], at least part of the 60 fraction of hydrocarbons in Cl to C2 is recycled in the separation stage 12 as 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 deathanizer.

[0532] The ratio of the mass flow rate of the recycle stream 64 to the mass flow rate of the fraction 60 from the head of the deaethanizer of the separation stage 20 is less than 0.5 as defined above.

[0533] In one embodiment, carbon dioxide, carbon monoxide and hydrogen present in the C1-C2 hydrocarbon fraction are separated from other hydrocarbons, in particular by cryogenic distillation, membrane separation or adsorption by alternating pressure and their combinations.

[0534] These compounds are then advantageously recycled to a preliminary stage of alcohol synthesis, in particular by synthesis gas fermentation and by catalytic conversion of synthesis gas to produce in particular methanol.

[0535] The methanol or ethanol thus produced are then advantageously recycled to form part of the 12 alcohol stream in Cl to C6.

[0536] In the example of [Fig.1], the fraction 62 of C3+ hydrocarbons, having the composition defined above, is then introduced into the joint stage 22 of oligomerization and alkylation.

[0537] In this stage 22, one or more joint oligomerization and alkylation reactors carry out oligomerization of the olefins present in the fraction 62, according to the operating conditions defined above in the description, in particular oligomerization of the olefins at C3 to C7.

[0538] In addition, jointly, the C6+ aromatics present in fraction 62 are alkylated to form in particular C8+ aromatics.

[0539] The reaction is carried out under the operating conditions described above. One or more catalysts defined above are used.

[0540] At the exit of stage 22, a stream 24 of hydrocarbons to be hydrogenated is thus formed, with the composition defined above.

[0541] Then the stream 24 of hydrocarbons to be hydrogenated is introduced into the hydrogenation stage 26, to induce the hydrogenation of at least a part of the olefins present in the stream of hydrocarbons to be hydrogenated 24 and the hydrogenation into cycloparaffins of at least a part of the aromatics present in the stream of hydrocarbons to be hydrogenated 24.

[0542] Hydrogenation is carried out under the operating conditions described above, with one or more of the catalysts described above.

[0543] A stream 66 containing hydrogen is introduced into the hydrogenation stage 26, with a ratio of the volumetric flow rate of hydrogen in the stream 66 to the volumetric flow rate of the hydrocarbon stream to be hydrogenated 24 is for example that defined above.

[0544] A stream of hydrogenated hydrocarbons 30 is formed at the outlet of the hydrogenation stage with the composition described above.

[0545] The hydrogenated hydrocarbon stream 30 is then fractionated in the fractionation stage 28.

[0546] In the first column 42, it is separated into a fraction of C4- hydrocarbons forming the liquefied petroleum gas fraction 44, and into a fraction of C4+ hydrocarbons forming the residual fraction 70 of the hydrogenated hydrocarbon stream.

[0547] The fractions 44, 70 have the characteristics defined above in terms of cutting points.

[0548] The residual fraction 70 is introduced into the second column 46 to be fractionated into the naphtha fraction 38, the jet fuel fraction 34, and the diesel fraction 36, as characterized above.

[0549] The installation variant 90 illustrated in [Fig.2] is intended for implementation of a second process according to the invention. It differs from the installation 10 illustrated by [Fig.1] in that the separation stage 20 includes an additional column 92 for recovering propylene.

[0550] The fraction 62 of C3+ hydrocarbons from the deathanizer of the separation stage 20 is introduced into the additional column 92 to form at the top of the column, a fraction 80 of C3- hydrocarbons and at the bottom of the column, a fraction 82 of C4+ hydrocarbons intended to be introduced into the joint oligomerization and alkylation stage 22.

[0551] Fractions 80 and 82 have the compositions described above. Fraction 80 contains more than 80% by mass of the propylene contained in fraction 62 of C3+ hydrocarbons. Such an example allows for the recovery of propylene formed in conversion stage 12, when such recovery is economically advantageous.

[0552] The installation 100 described in [Fig.3] is intended for the implementation of a third process according to the invention. It differs from the installation 10 described in [Fig.1] in that at least one tap 102 for adding alcohols from C2 to C6 is provided in the conversion stage 12, for example between two successive catalytic beds of the conversion stage 12.

[0553] The composition of the C2 to C6 alcohol stream 102 advantageously comprises less than 20% methanol, and more than 80% alcohol by mass in C2 to C6, for example more than 50% ethanol and propanol.

[0554] The addition of C2 to C6 alcohols in addition to methanol facilitates the conversion reaction of the alcohol stream 14 by making it more isothermal (the conversion of methanol being very exothermic and the conversion of C2 to C6 alcohols being endothermic) and therefore easier to control, especially when fixed catalytic beds are used in the conversion stage 12.

[0555] The installation 110 described in [Fig. 4] is similar to that of [Fig. 1]. It comprises a conversion stage 12 having at least one reactor with a fluidized catalytic bed, preferably a single fluidized catalytic bed reactor. This reactor or these reactors are suitable for implementing the experimental conditions described above.

[0556] The reactor comprises a reaction zone 111a having a fluidized catalytic bed, and a regeneration zone 111b for the fluidized catalytic bed. A portion of the catalyst present in the reaction zone 11a is continuously withdrawn to be regenerated in the regeneration zone 111b, advantageously by controlled combustion in the presence of oxygen.

[0557] A portion 40a of the water stream 40 is optionally recycled to the conversion stage 12, via conduit 18a.

[0558] Optionally, part 64 of the C1-C2 hydrocarbon fraction separated from the water-depleted mixture 19 is introduced into the reactor having a fluidized catalytic bed.

[0559] Advantageously, the C6 Cl alcohol stream is introduced at the conversion step (a) at a temperature at least 5°C higher than the bubble point of the C6 Cl alcohol stream.

[0560] Fig. 8 illustrates, in a particular embodiment, the flow of the catalyst from the regeneration zone 111b to the reaction zone 111a.

[0561] For reasons of simplicity, the drawings do not contain details of the internal parts of the tanks forming zones 111a, 111b.

[0562] The flow 14 of alcohol in Cl to C6 is introduced into the foot of the reaction zone 111a having a fluidized catalytic bed.

[0563] At the top of the reaction zone 11la, the products of the conversion reaction are separated from the catalyst in a disengagement zone 203 advantageously equipped with cyclones and the mixture 16 produced is transported to the separation stage 18.

[0564] Optionally, the reaction heat produced by the conversion is extracted from the reaction zone 11la by means of a catalyst cooler 205, which is advantageously a heat exchanger located outside the reaction zone 11la and connected to it.

[0565] The reaction zone 111a receives the regenerated catalyst from the regeneration zone 111b via a supply line 207 connecting the regeneration zone 111b to the reaction zone 111a.

[0566] The deactivated catalyst is removed from the disengagement zone 203 via a discharge line 206 separate from the supply line 207, the discharge line 206 connecting the reaction zone 111a to the regeneration zone 111b.

[0567] Air is injected through the injection line 221 into the regeneration zone 111b, at the foot of it, into a fluidized bed where the coke deposits are burned.

[0568] The regeneration zone 111b also includes a disengagement zone 222 advantageously equipped with cyclones. In this zone 222, the combustion gases are separated from the regenerated catalyst and are discharged through the regeneration line 223 advantageously located at the head of the regeneration zone 111b.

[0569] Optionally, since the combustion of coke deposits is a highly exothermic reaction and the temperature of the regeneration zone 111b must be carefully controlled, a catalyst cooler (not shown in the figure, but similar to catalyst cooler 205) is connected to the regeneration zone 111b. Hot catalyst extracted from the regeneration zone flows through this cooler to be cooled, thus controlling the temperature in the regeneration zone 111b. The regenerated catalyst is sent via line 207 to the reaction zone 111a.

[0570] A fifth installation 120 for implementing a fifth process according to the invention is illustrated in [Fig. 5]. The fifth installation 120 differs from the first installation 10 in that it includes an additional separation stage 122 interposed between the outlet of the joint oligomerization and alkylation stage 22 and the inlet of the hydrogenation stage 26.

[0571] The additional separation stage 122 includes at least one distillation column.

[0572] The product 124 from the joint oligomerization and alkylation stage 22 is separated in the distillation column into a fraction 126 of C7- hydrocarbons and a fraction 128 of C8+ hydrocarbons, forming the hydrocarbon stream 24 intended for hydrogenation. The C7- hydrocarbon fraction 126 can optionally be recycled to the joint stage 22.

[0573] A sixth installation 140 according to the invention is illustrated in [Fig. 6]. This sixth installation 140 is intended for carrying out a sixth process according to the invention. It differs from the first installation 10 in that it comprises a clean stage 22A for the oligomerization of olefins from the water-depleted mixture 19, from stage 20, and a clean stage 22B for the alkylation of aromatics from the water-depleted mixture 19 from stage 20.

[0574] Each stage 22A 22B comprises a separate oligomerization and alkylation reactor respectively, in which the operating conditions described above are implemented.

[0575] At the separation stage 20, the water-depleted mixture 19 is separated into a fraction 60 of C1-C2 hydrocarbons recovered at the top of the de-anaerobic digester, into a fraction 142 of C3 to C5 hydrocarbons recovered at an intermediate stage of the de-anaerobic digester and into a fraction 144 of C6+ hydrocarbons recovered at the bottom of the de-anaerobic digester.

[0576] The fraction 142 of C3 to C5 hydrocarbons is sent in its entirety to the oligomerization stage 22A to produce a product 146 of the oligomerization reactor.

[0577] The C1-C2 hydrocarbon fraction 60 and the C6+ hydrocarbon fraction 144 are brought to the alkylation stage 22B to produce an alkylation reactor product 152 under the operating conditions defined above.

[0578] The product 152 from the alkylation reactor is then mixed with the product 146 from the oligomerization reactor.

[0579] Products 146, 152 are then introduced to the additional separation stage 122 to be separated into fraction 126 of C7- hydrocarbons and fraction 128 of C8+ hydrocarbons, described above.

[0580] At least a part 150 of the fraction 126 of C7- hydrocarbons is recycled in the oligomerization stage 22A, another part possibly being recovered as gasoline.

[0581] Fraction 128 forms the hydrocarbon stream to be hydrogenated 24.

[0582] A seventh installation 160 intended for the implementation of a seventh process according to the invention is described in [Fig.7].

[0583] The seventh process according to the invention differs from the sixth process implemented in the installation 150 in that at the additional separation stage 122, the product 146 of the oligomerization reactor and the product 152 of the alkylation reactor are separated into the fraction 126 of C7- hydrocarbons, taken at the top of the column, into a fraction 162 of C8 to C16 hydrocarbons, taken at an intermediate stage of the column and into a fraction 164 of C17+ hydrocarbons, taken at the bottom of the column.

[0584] As previously described, at least a part 150 of the fraction 126 of C7- hydrocarbons is recycled in the oligomerization stage 22A.

[0585] Fraction 162 of C8 to C16 hydrocarbons is introduced into hydrogenation stage 26 to be hydrogenated.

[0586] Fraction 164 of C17+ hydrocarbons is recycled at least partially at conversion stage 12.

[0587] Thus, the heavy hydrocarbons present in the C17+ hydrocarbon fraction are re-cracked in the conversion stage 12. This increases the amount of jet fuel fraction 34 produced.

[0588] In all the cases described above, the jet fuel fraction 34 produced by the aforementioned processes can be used as such, in its pure form, as an aircraft jet fuel intended to power an aircraft engine, or blended with jet fuel obtained from petroleum distillation. The jet fuel fraction or its blend is advantageously a sustainable aviation fuel (SAF) whose composition is similar to the SAFs described in ASTM D7566.

[0589] The mixture comprises at least 5% by mass, in particular at least 10% by mass of the jet fuel fraction 34.

[0590] Thanks to the invention which has just been described, it is possible to have simple and efficient processes for manufacturing jet fuel from a Cl to C6 alcohol stream which is preferably from a renewable source, in particular from fermentation, and / or generated by conversion of carbon oxide or carbon dioxide captured from the atmosphere in the presence of hydrogen.

[0591] The jet fuel fraction from the process according to the invention has a very low carbon footprint, since it is not derived from petroleum products, but on the contrary from sources which contribute to reducing the amount of carbon dioxide present in the atmosphere.

[0592] The jet fuel fraction 34 produced by the process according to the invention is further manufactured very economically and can in some cases be used as such, without further purification or mixing, as aircraft engine propulsion fuel.

[0593] Naturally, the installations in Figures 2, 3 and 5 to 7 may also feature a conversion stage 12 equipped with a fluidized catalytic bed as in [Fig.4].

[0594] In one variant, the process is implemented to produce olefins (in particular ethylene, propylene, butenes and pentenes, hexenes), minimizing the production of paraffins from renewable resources on the basis of a conversion of a Cl alcohol stream to C6 into olefins.

[0595] EXAMPLES

[0596] Specific, non-limiting examples of implementation of step (a) of conversion, joint steps (c) and (d) of oligomerization and alkylation and step (f) of hydrogenation will now be described.

[0597] Conversion step (a) CATALYST PREPARATION

[0598] A sample of ZSM-5 zeolite (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 matrix) was steam-baked at 550°C for 6 h in 100% H2O at atmospheric pressure. The sample is hereafter identified as Sample A.

[0599] The steam-cured solid A was subjected to contact with a 3.14 M H3PO4 solution for 4 h under reflux conditions (4.2 ml / g zeolite). Solid A was then separated from the liquid phase at room temperature by filtration of the solution. The resulting material was dried at 200°C for 16 h. This sample is hereafter referred to as sample B.

[0600] EXAMPLE 1 OF A CATALYST

[0601] 490 g of sample B were mixed with 490 g of specific binder (P = 15.9% in (mass, Si = 13.2%, Mg = 0.27%, Al = 0.15% by mass, K = 230 ppm, Na = 230 ppm, Ca = 19.2% by mass), 588.3 g of low-sodium silica sol containing 34% by mass of SiO2, 6 g of xonotlite, and 2 to 3% by mass of extrusion additives. The mixture was stirred for 30 min and extruded.

[0602] The specific binder was prepared by mixing an equivalent mass of NH4H2PO4 and xonotlite in an aqueous medium at room temperature (1 g of solid / 4 ml of water). After stirring for 60 minutes, the phosphated xonotlite was separated from the liquid by filtration and dried. The dried product was used as an extrusion component.

[0603] The extruded solid was dried for 24 h at room temperature, then for 16 h at 200°C followed by washing with demineralized water at room temperature and then dried at 110°C overnight. A further washing step at room temperature was then carried out using demineralized water at pH 3.08. The catalyst was then dried at 110°C overnight and calcined at 700°C for 2 hours.

[0604] EXAMPLE 2 OF A CATALYST

[0605] 320 g of sample B were mixed with 400 g of specific binder (P = 15.9% in mass, Si = 13.2, Mg = 0.27, Al = 0.15% by mass, K = 230 ppm, Na = 230 ppm, Ca = 19.2% by mass), 165 ml of H2O, 235 g of low-sodium silica sol containing 34% by mass of SiO2 and 2 to 3% by mass of extrusion additives. The mixture was stirred for 30 min and extruded.

[0606] The specific binder was prepared by mixing an equivalent mass of (NH4)H2PO4 (ammonium dihydrogen phosphate) and xonotlite in an aqueous medium at room temperature (1 g of solid / 4 ml of water). After stirring for 60 minutes, the phosphated xonotlite was separated from the liquid by filtration and dried. The dried product was used as an extrusion component.

[0607] The extruded solid was dried for 24 h at room temperature, then for 16 h at elevated temperature followed by washing and steam heat treatment at 600°C for 2 h. The sample hereafter identified as sample E.

[0608] EXAMPLE 3 OF A CATALYST

[0609] 356 g of sample A were extruded with 338.7 g of Nyacol (40% SiO2 sol) (by mass), 311.3 g of fumed silica (FK500), 480 mL of H2O, and 2–3% extrusion additives. The extruded solid was dried for 24 h at room temperature, then for 16 h at 110°C followed by calcination at 500°C for 10 hours. The final sample contained 40 wt% zeolite and 60 wt% SiO2 binder. The extruded sample was subjected to ion exchange with 0.5 M NH4Cl under reflux conditions for 18 hours, followed by washing with water, drying at 110°C for 16 hours, and calcination at 450°C for 6 hours. The shaped and exchanged sample was treated with 3.1 M H3PO4 under reflux conditions for 4 h (1 g / 4.2 ml), followed by cooling, filtration and drying at 110 °C for 16 h.

[0610] The phosphate sample was washed at room temperature with a 0.1 M calcium acetate solution for 2 h (1 g / 4.2 ml). Then, the washed sample was dried at 110°C for 16 h and heat-treated with steam in 100% by mass of H2O for 2 h at 600°C.

[0611] EXAMPLE 4 OF A CATALYST

[0612] 150g of sample B were subjected to contact with 630 ml of aqueous solution containing 1.5 g of dispersed xonotlite, followed by the addition of 450 g of low-sodium silica sol (34 wt. SiO2 in water, 200 ppm Na). The solution was then stirred for one hour and spray-dried. The spray-dried solid was washed with room-temperature water for two hours, followed by filtration, drying at 110°C for 16 hours, and calcination at 700°C.

[0613] EXAMPLE 5 OF A CATALYST

[0614] 100 g of sample A were subjected to contact with 25 g of 85% H3PO4 in The mass was refluxed for 4 hours, followed by cooling and the addition of 120 mL of aqueous solution containing 7 g of dispersed xonotlite. The resulting slurry was stirred for approximately 1 h, followed by the addition of 300 g of low-sodium silica sol (34 wt. SiO2 in water, 200 ppm Na). The solution was then stirred for 1 hour and spray-dried. The spray-dried solid was dried at 200 °C for 16 h and washed with room-temperature water for 2 h, followed by filtration, drying, and calcination at 700 °C for 2 h.

[0615] EXAMPLE 6 OF A CATALYST

[0616] 75 g of sample A were introduced into a solution containing 14.25 g of 85% H3PO4 by mass and 300 mL of demineralized water were used. The suspension was stirred under reflux for 2 hours. Then, 4.125 g of CaCO3 were added to the suspension. Heating of the solution was stopped, while stirring continued until the temperature dropped below 30°C. This resulted in suspension A.

[0617] A solution is then prepared by mixing 450 g of low sodium silica sol (34% by mass of SiO2 in water, 200 ppm Na) and 4.5 g of H3PO4 (85% by mass) under stirring at room temperature for 30 minutes. This led to suspension B.

[0618] Suspensions A and B are then mixed together and 120 ml of demineralized water is added. The solution was then stirred for one hour and spray-dried. The spray-dried solid was dried at 200°C for 16 h and washed with water at room temperature for 2 h, followed by filtration, drying, and calcination at 700°C for 2 h.

[0619] IMPLEMENTATION OF CONVERSION STEP (a)

[0620] Catalyst tests were carried out on 2 g (35 mesh to 45 mesh particles) of catalyst with a charge of essentially pure methanol, at Tinjection = 550 °C and at a pressure of 0.5 barg and a space velocity hour (WHSV = 1.6 h *), in a downflow stainless steel fixed bed reactor.

[0621] Before the catalytic test, all catalysts were activated in a stream of N2 (5 Nl / h) up to the reaction temperature. Product analysis was performed online using a gas chromatograph equipped with a capillary column. The catalytic performance of the catalyst in Table 1 is given on the carbon, dry, and coke-free basis. The results are given for the average performance of the catalyst during the first 4 hours of operation.

[0622] [Tables 1] WHSV h-1 1.6 1.6 4 1.6 1.6 1.6 1.6 MeOH Conversion % mass C 100 100 100 98 100 100 100 CH4 % mass C 2.3 1.6 1.6 3.4 1.3 2.1 1.5 Paraffins % mass C 6.2 6.7 5.7 8.8 8.2 8.2 8.8 Olefins % mass C 86 85.4 87.5 79.1 83.6 82.8 82.3 Dienes % mass C 1 0.5 0.7 1.2 0.8 0.9 0.6 Aromatics % mass C 6.6 7.4 6.1 10.7 6.9 8 7.6 Ethylene % mass C 9.8 13.9 10 6.1 15.4 12.1 15.7 Propylene % mass C 41.4 41.8 43.3 35.4 39.6 39.2 38.4 C4+ Olefins % mass C 34.8 29.7 34.2 37.6 28.6 31.5 28.3 Catalyst of Example 1 2 2 3 4 5 6

[0623] Performance of stage (a) - 550°C -0.5 bar

[0624] Joint steps (c) and (d) of oligomerization and alkylation and step (f) of hydrogenation

[0625] AI PROPERTIES LOAD USED FOR AI IMPLEMENTATION OF STEPS (c) AND (d)

[0626] The properties of the charge used are as follows:

[0627] [Tables2] Charge Properties: Density at 15°C (g / mL) 0.7251; Bromine Number (g Br / 100 g) 75; Distillation Range: PI (°C) 38.8; T50 (°C) 84.5; T95 (°C) 160.7; FBP (°C) 164.1

[0628] The detailed composition of the charge was determined by the GC method.

[0629] [Tables3] (% mass) Hydrocarbon distribution (% mass) Olefin composition n-paraffins 7.1 C4= 0.3 i-paraffins 25.7 C5= 15.7 naphthenes 11.3 C6= 10.1 n-olefins 10.6 C7= 7.6 i-olefins 19.6 C8= 3.5 c-olefins 7.9 C9= 0.7 Aromatics 17.7 C10= 0.2 C11= 0.1 Total olefins 38.1

[0630] OL1GOMER1SAT1ON AND ALKYLATION

[0631] 100 mL of amorphous silica-alumina (ASA) catalyst diluted with 100 mL of inert material (SiC 0.21 mm) were loaded into a fixed-bed tubular reactor with an internal diameter of 18 mm. Before the test, the catalyst was activated at 250°C (10°C / h) under 135 NL / h of nitrogen for 8 hours. The temperature was then lowered to 40°C at the start of the test program.

[0632] 100 mL of ZSM-5 based catalyst (80% by mass of MFI having a ratio 80% silica-alumina (20% by mass of alumina binder) diluted with 100 mL of inert material (0.21 mm SiC) were loaded into a fixed-bed tubular reactor of Inner diameter 18 mm. Before the test, the catalyst was activated at 400°C (60°C / h) under 160NL / h of nitrogen for 2 hours. The temperature was then lowered to 40°C at the start of the test program.

[0633] HYDROGENATION

[0634] Fractionation is performed on the oligomerization product to recover the 145+ and 165°C+ oligomerized cuts, which are hydrotreated over a NiMo catalyst. The following operating conditions were chosen: 80 barg, an hourly space velocity in LHSV liquid of 1 h*, H2 / hydrocarbon volume ratio of 500 NL / L, in a single pass without recycle, and the temperature was increased from 250°C to 270°C.

[0635] Example 1 - Performance obtained with a Zeolite-based catalyst

[0636] The charge was processed under the following operating conditions: 55 barg, velocity hourly spatial LHSV liquid of Ih 1 and at temperatures ranging from 240°C to 280°C.

[0637] [Tables4] Yield structure (% mass) 240°C 260°C 280°C Yield 145-245°C 23 27 25 Yield 245+ 5 7 13

[0638] 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 calculation because they may result from the oligomerization of the light olefins (C4 and C5) present in the feedstock. At 240°C, the conversion of C5-C7 olefins is greater than 88% by mass.

[0639] Olefins can react either by oligomerization or by alkylation with aromatic compounds. The conversion of aromatics has been observed to vary from 10 wt% at 240°C to 26 wt% at 280°C (see Table 5 below). Aromatics are indeed present in the 170-FBP fractions, indicating that alkylation does occur.

[0640] [Tables5] Loading temperature 240°C 260°C 280°C Conversion of aromatics IBP-170°C (% mass) - 10 14 26 Aromatic concentration in 170-FBP (% mass) - 3.7 5.5 11.4

[0641] Example 2 - Performance obtained with an amorphous silica-alumina catalyst (ASA)

[0642] The charge was processed under the following operating conditions: 25 barg, liquid LHSV hourly space velocity of Ih 1 and temperatures ranging from 180°C to 220°C.

[0643] [Tableauxô] Yield structure (% mass) 180°C 200°C 220°C Yield 145-245°C 19 20 20 245+ yield 12 18 18

[0644] The conversion of light olefins (C4-C8 olefins) varies from 80% by mass at 180°C to nearly 100% at 220°C. C9+ olefins are not taken into account in the conversion calculation because they may result from the oligomerization of the light olefins (C4 and C5) present in the feedstock. At 180°C, the conversion of C5-C7 olefins is greater than 80% by mass.

[0645] Olefins can react either by oligomerization or by alkylation with aromatic compounds. It has been observed that the conversion of aromatics varies from 28% by mass at 180°C to 33% by mass at 220°C (see Table 7 below). Aromatics are indeed present in the FBP fractions at 170°C, indicating that alkylation does occur.

[0646] [Tables7] Loading temperature 180°C 200°C 220°C Conversion of aromatics IBP-170°C (% mass) - 28 31 33 Aromatic concentration in 170-FBP (% mass) - 13.3 18.5 23.3

[0647] The 145+ sections were hydrotreated using a NiMo catalyst under the conditions described above. The properties of the hydrogenated section are described in Table 8 and the detailed composition is given in Table 9.

[0648] [Tables8] ASA (200°C) ZEOLITH (240°C) ZEOLITH (280°C) 145-245 section after hydrotreatment 145-245 section after hydrotreatment 145-245 section after hydrotreatment Unit Method Density @ 15°C kg / m³ NF EN ISO 12185 809.23 775.29 797.06 Solidification Point °C ASTM D7153 <-100 <-100 <-100 Flash Point ABEL °C IP170 47.0 Calculated Cetane - ISO4264 39.4 53.6 43.2 Measured Cetane 23.1 38.6 31.4 Distillation ISO ISO3405 IBP °C 166 165 167.5 5% °C 174.5 174.3 174 50% °c 193 188.2 189 95% °c 226 222.8 2242 FBP °c 231.5 230.6 2305

[0649] [Tables9] 145-245% Cut (% mass) Zeolite 280°C ASA 200°C Paraffins 50.18 42.29 Naphthenes 27.11 28.85 Dinaphenes 6.58 5.01 Alkylaromatics 14.49 22.38 Monoaromatic naphthenes 1.52 1.35

[0650] Composition of effluents as determined by GCxGC

Claims

Demands

1. A process for obtaining hydrocarbons, comprising the following steps: (a) conversion of a stream (14) of C6 Cl alcohol to produce a mixture (16) containing paraffins, olefins, aromatics, and water; (b) separation of the water (40) from the mixture (16) to form a water-depleted mixture (19); the water-depleted mixture (19) being separated and / or treated to recover hydrocarbons, characterized by the addition, at step (a) of converting the C6 Cl alcohol stream, of a stream (182) containing carbon dioxide, and the co-conversion of carbon dioxide to carbon monoxide at step (a) of converting the C6 Cl alcohol stream.

2. A method according to claim 1, wherein the stream containing carbon dioxide (182) comprises more than 5% by mass of carbon dioxide.

3. A process according to claim 1 or 2 wherein the mass ratio of carbon dioxide to Cl- to C6 alcohols in the feed supplied in step (a) of conversion is between 5% and 75%.

4. A process according to any one of the preceding claims, wherein at least 2 mole percent of the carbon dioxide contained in the carbon dioxide-containing stream (182) is converted to carbon monoxide during step (a) of converting the alcohol stream to Cl to C6.

5. A method according to any one of the preceding claims, wherein the conversion step (a) is carried out using at least one catalyst comprising molecular sieves containing at least 10 oxygen atom (10-MR) or larger pores in their microporous structure.

6. A process according to claim 5, wherein the catalyst for carrying out conversion step (a) comprises a phosphorus-modified zeolite, having in particular a P content of at least 0.05 wt% and preferably between 0.3 wt% and 7 wt%, a composite catalyst comprising at least 0.1 wt% of silicate, or a molecular sieve modified with phosphorus (P) and an alkaline earth or rare earth metal (M) (modified molecular sieve MP), advantageously having an M / P molar ratio in the mo- ecular less than 1.

7. A process according to any one of claims 5 or 6 wherein the catalyst for carrying out the conversion step (a) has been modified by adding one or more metals selected from the metals of group IIB, in particular Zn, of group IIIB, in particular Ga, the transition metals of group VIIIB in particular Fe and / or Ni and / or Pt, of group VIB, in particular Mo, of group IB in particular Cu and / or Ag or from the lanthanide group in particular La.

8. A method according to any one of claims 5 to 7, wherein the catalyst for carrying out conversion step (a) is a phosphorus-modified zeolite partially having an ALPO structure or is a B-modified zeolite.

9. A process according to any one of the preceding claims, comprising a step of separating the water-depleted mixture (19) into a fraction (60) of C1-C2 hydrocarbons, and into a fraction (62) of C3+ hydrocarbons, a portion (64) of the C1-C2 hydrocarbon fraction being advantageously recycled in step (a) of converting the alcohol stream into C6 Cl.

10. A process according to claim 9, wherein the carbon dioxide, carbon monoxide and hydrogen present in the C1-C2 hydrocarbon fraction are separated from the hydrocarbons, in particular by cryogenic distillation, membrane separation or alternating pressure adsorption and combinations thereof, the carbon dioxide, carbon monoxide and hydrogen then being advantageously recycled to a preliminary step of synthesis of alcohols intended to form the C6 Cl alcohol stream, in particular by synthesis gas fermentation and by catalytic conversion of the synthesis gas to produce methanol.

11. A process according to any one of the preceding claims, comprising the following steps: (c) oligomerization of olefins from the water-depleted mixture (19); (d) alkylation of aromatics from the water-depleted mixture (19); (e) formation of a hydrocarbon stream (24) from at least a portion of the olefins oligomerized in step (c) and at least a portion of the aromatics alkylated in step (d).

12. A method according to claim 11, comprising the following steps: (f) hydrogenation of the hydrocarbon stream (24) formed in step (e) to form a hydrogenated hydrocarbon stream (30); (g) recovery of at least one jet fuel fraction (34) from the hydrogenated hydrocarbon stream (30).

13. A process according to claim 12, wherein the jet fuel fraction (34) comprises between 2% by volume and 30% by volume of C8+ aromatics, preferably between 8% by volume and 25% by volume of C8+ aromatics.

14. A process according to claim 12 or 13, wherein, in the mixture (16) of paraffins, olefins, aromatics and water produced in step (a) of conversion, the ratio of the mass of C3+ olefins to the total mass of olefins is greater than or equal to 0.

8.

15. Hydrocarbon production plant, comprising: - a stage (12) for converting a stream (14) of alcohol in Cl to C6 to produce a mixture (16) containing paraffins, olefins, aromatics, and water; - a stage (18) for separating the water (40) from the mixture (16) to form a water-depleted mixture (19); - at least one stage (20, 22, 24, 26) for separating and / or treating the water-depleted mixture (19) to recover hydrocarbons, characterized in that the conversion stage (12) of the C6 alcohol to Cl stream (14) comprises at least one supply conduit for a stream (182) containing carbon dioxide, the conversion stage (12) of the C6 alcohol to Cl stream (14) being configured to convert carbon dioxide from the carbon dioxide-containing stream (182) into carbon monoxide jointly with the conversion of the C6 alcohol to Cl stream.

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