WO2021198166A1 - Procédé de conversion de gaz en oléfines avec coproduction d'hydrogène conjointement à un procédé d'intégration thermique - Google Patents

Procédé de conversion de gaz en oléfines avec coproduction d'hydrogène conjointement à un procédé d'intégration thermique Download PDF

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WO2021198166A1
WO2021198166A1 PCT/EP2021/058131 EP2021058131W WO2021198166A1 WO 2021198166 A1 WO2021198166 A1 WO 2021198166A1 EP 2021058131 W EP2021058131 W EP 2021058131W WO 2021198166 A1 WO2021198166 A1 WO 2021198166A1
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stream
process according
hydrogen bromide
methane
mpa
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PCT/EP2021/058131
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Romuald COUPAN
Nikolai Nesterenko
Gleb VERYASOV
Zhongyi John DING
Jingsong Zhou
Miguel A. F. Santos
Mircea G. CRETOIU
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Total Se
Sulzer Management Ag
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Publication of WO2021198166A1 publication Critical patent/WO2021198166A1/fr

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    • C—CHEMISTRY; METALLURGY
    • C25—ELECTROLYTIC OR ELECTROPHORETIC PROCESSES; APPARATUS THEREFOR
    • C25B—ELECTROLYTIC OR ELECTROPHORETIC PROCESSES FOR THE PRODUCTION OF COMPOUNDS OR NON-METALS; APPARATUS THEREFOR
    • C25B1/00—Electrolytic production of inorganic compounds or non-metals
    • C25B1/01—Products
    • C25B1/02—Hydrogen or oxygen
    • C25B1/04—Hydrogen or oxygen by electrolysis of water
    • C—CHEMISTRY; METALLURGY
    • C01—INORGANIC CHEMISTRY
    • C01B—NON-METALLIC ELEMENTS; COMPOUNDS THEREOF; METALLOIDS OR COMPOUNDS THEREOF NOT COVERED BY SUBCLASS C01C
    • C01B32/00—Carbon; Compounds thereof
    • C—CHEMISTRY; METALLURGY
    • C01—INORGANIC CHEMISTRY
    • C01B—NON-METALLIC ELEMENTS; COMPOUNDS THEREOF; METALLOIDS OR COMPOUNDS THEREOF NOT COVERED BY SUBCLASS C01C
    • C01B7/00—Halogens; Halogen acids
    • C01B7/09—Bromine; Hydrogen bromide
    • C01B7/093—Hydrogen bromide
    • C—CHEMISTRY; METALLURGY
    • C01—INORGANIC CHEMISTRY
    • C01B—NON-METALLIC ELEMENTS; COMPOUNDS THEREOF; METALLOIDS OR COMPOUNDS THEREOF NOT COVERED BY SUBCLASS C01C
    • C01B7/00—Halogens; Halogen acids
    • C01B7/09—Bromine; Hydrogen bromide
    • C01B7/096—Bromine
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C1/00—Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon
    • C07C1/26—Preparation of hydrocarbons from one or more compounds, none of them being a hydrocarbon starting from organic compounds containing only halogen atoms as hetero-atoms
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C17/00—Preparation of halogenated hydrocarbons
    • C07C17/093—Preparation of halogenated hydrocarbons by replacement by halogens
    • C07C17/10—Preparation of halogenated hydrocarbons by replacement by halogens of hydrogen atoms
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C17/00—Preparation of halogenated hydrocarbons
    • C07C17/38—Separation; Purification; Stabilisation; Use of additives
    • C07C17/383—Separation; Purification; Stabilisation; Use of additives by distillation
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C4/00—Preparation of hydrocarbons from hydrocarbons containing a larger number of carbon atoms
    • C07C4/02—Preparation of hydrocarbons from hydrocarbons containing a larger number of carbon atoms by cracking a single hydrocarbon or a mixture of individually defined hydrocarbons or a normally gaseous hydrocarbon fraction
    • C07C4/06—Catalytic processes
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C7/00—Purification; Separation; Use of additives
    • C07C7/04—Purification; Separation; Use of additives by distillation
    • C07C7/05—Purification; Separation; Use of additives by distillation with the aid of auxiliary compounds
    • C07C7/08—Purification; Separation; Use of additives by distillation with the aid of auxiliary compounds by extractive distillation
    • C—CHEMISTRY; METALLURGY
    • C09—DYES; PAINTS; POLISHES; NATURAL RESINS; ADHESIVES; COMPOSITIONS NOT OTHERWISE PROVIDED FOR; APPLICATIONS OF MATERIALS NOT OTHERWISE PROVIDED FOR
    • C09C—TREATMENT OF INORGANIC MATERIALS, OTHER THAN FIBROUS FILLERS, TO ENHANCE THEIR PIGMENTING OR FILLING PROPERTIES ; PREPARATION OF CARBON BLACK  ; PREPARATION OF INORGANIC MATERIALS WHICH ARE NO SINGLE CHEMICAL COMPOUNDS AND WHICH ARE MAINLY USED AS PIGMENTS OR FILLERS
    • C09C1/00—Treatment of specific inorganic materials other than fibrous fillers; Preparation of carbon black
    • C09C1/44—Carbon
    • C09C1/46—Graphite
    • C—CHEMISTRY; METALLURGY
    • C09—DYES; PAINTS; POLISHES; NATURAL RESINS; ADHESIVES; COMPOSITIONS NOT OTHERWISE PROVIDED FOR; APPLICATIONS OF MATERIALS NOT OTHERWISE PROVIDED FOR
    • C09C—TREATMENT OF INORGANIC MATERIALS, OTHER THAN FIBROUS FILLERS, TO ENHANCE THEIR PIGMENTING OR FILLING PROPERTIES ; PREPARATION OF CARBON BLACK  ; PREPARATION OF INORGANIC MATERIALS WHICH ARE NO SINGLE CHEMICAL COMPOUNDS AND WHICH ARE MAINLY USED AS PIGMENTS OR FILLERS
    • C09C1/00—Treatment of specific inorganic materials other than fibrous fillers; Preparation of carbon black
    • C09C1/44—Carbon
    • C09C1/48—Carbon black
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G50/00—Production of liquid hydrocarbon mixtures from lower carbon number hydrocarbons, e.g. by oligomerisation
    • C—CHEMISTRY; METALLURGY
    • C07—ORGANIC CHEMISTRY
    • C07C—ACYCLIC OR CARBOCYCLIC COMPOUNDS
    • C07C2529/00—Catalysts comprising molecular sieves
    • C07C2529/04—Catalysts comprising molecular sieves having base-exchange properties, e.g. crystalline zeolites, pillared clays
    • C07C2529/06—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof
    • C07C2529/40—Crystalline aluminosilicate zeolites; Isomorphous compounds thereof of the pentasil type, e.g. types ZSM-5, ZSM-8 or ZSM-11
    • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02E—REDUCTION OF GREENHOUSE GAS [GHG] EMISSIONS, RELATED TO ENERGY GENERATION, TRANSMISSION OR DISTRIBUTION
    • Y02E60/00—Enabling technologies; Technologies with a potential or indirect contribution to GHG emissions mitigation
    • Y02E60/30—Hydrogen technology
    • Y02E60/36—Hydrogen production from non-carbon containing sources, e.g. by water electrolysis
    • Y—GENERAL TAGGING OF NEW TECHNOLOGICAL DEVELOPMENTS; GENERAL TAGGING OF CROSS-SECTIONAL TECHNOLOGIES SPANNING OVER SEVERAL SECTIONS OF THE IPC; TECHNICAL SUBJECTS COVERED BY FORMER USPC CROSS-REFERENCE ART COLLECTIONS [XRACs] AND DIGESTS
    • Y02—TECHNOLOGIES OR APPLICATIONS FOR MITIGATION OR ADAPTATION AGAINST CLIMATE CHANGE
    • Y02P—CLIMATE CHANGE MITIGATION TECHNOLOGIES IN THE PRODUCTION OR PROCESSING OF GOODS
    • Y02P20/00—Technologies relating to chemical industry
    • Y02P20/10—Process efficiency

Definitions

  • the present disclosure relates to gas to olefins processes.
  • Natural gas is an available fossil resource mainly composed of light alkanes.
  • the valorisation of natural gas as feedstock for the petrochemical industry is of interest as natural gas is cheap at source. Accordingly, the conversion of light alkanes into products like syngas, methanol, olefins or aromatics is highly valuable.
  • Oxygen-based processes are the current practices for natural gas conversion. However, processes involving the presence of oxygen leads unavoidably to the formation of carbon dioxide and water in the final product streams. The carbon efficiency of any processes contacting alkanes and oxygen does not exceed 75%.
  • G2A Gas-to-Aromatics
  • GTC Technologies based on bromine mediated activation of methane
  • the G2A process aims at converting methane to liquid hydrocarbons i.e. C6-C9 aromatics, C5 and C10+ products.
  • hydrogen bromide by-product of activation is recycled back to bromine consuming oxygen and coproducing water.
  • This first-generation technology thus offers a low carbon way to transform methane to chemicals and water with high carbon efficiency of about 85%.
  • Current G2A technology is nevertheless not an oxygen-free process as oxygen is used for bromine recovery and produces water.
  • the reforming process is the most practical current commercial process i.e. Steam Methane Reforming (SMR), Auto-Thermal Reforming (ATR), Dry Methane Reforming (DMR), Partial Oxidation Reforming (POX).
  • SMR Steam Methane Reforming
  • ATR Auto-Thermal Reforming
  • DMR Dry Methane Reforming
  • POX Partial Oxidation Reforming
  • the process of reforming methane (SMR, ATR, DMR, POX) is exemplified by the following chemical equations. 4 H 2
  • the methane reforming process produces synthesis gas (i.e. a mixture of carbon monoxide and hydrogen) that can be further converts to liquid hydrocarbons in the Gas-to-Liquid process (GTL) or to methanol in the Gas-to-Methanol process (GTM), as indicated in the following two chemical equations:
  • methanol can be used as feedstock in processes such as Methanol-to-Gasoline (MTG), Methanol-to-Olefins (MTO), and Methanoi-to-Propyiene (MTP).
  • MMG Methanol-to-Gasoline
  • MTO Methanol-to-Olefins
  • MTP Methanoi-to-Propyiene
  • OCM Oxidative Coupling of Methane
  • US2006/0100469 describes a process for converting gaseous alkanes to olefins and liquid hydrocarbons wherein a gaseous feed containing alkanes is reacted with a dry bromine vapour to form alkyl bromides and hydrobromic acid vapour.
  • the mixture of alkyl bromides and hydrobromic acid are then reacted over a synthetic crystalline alumino-silicate catalyst, such as an X or Y type zeolite, at a temperature of from about 250°C to about 500°C to form olefins, higher molecular weight hydrocarbons, hydrobromic acid vapour and coke.
  • a synthetic crystalline alumino-silicate catalyst such as an X or Y type zeolite
  • W02009/152403 describes a process for converting gaseous alkanes to olefins, higher molecular weight hydrocarbons or mixtures thereof wherein a gaseous feed containing alkanes is thermally reacted with a dry bromine vapour to form alkyl bromides and hydrogen bromide. Poly-brominated alkanes present in the alkyl bromides are further reacted with methane over a suitable catalyst to form mono-brominated species. The mixture of alkyl bromides and hydrogen bromide is then reacted over a suitable catalyst at a temperature sufficient to form olefins, higher molecular weight hydrocarbons or mixtures thereof and hydrogen bromide.
  • the catalyst may suffer from coke deactivation, depending on the reaction condition involved.
  • W02008/143940 describes an improved continuous process for converting methane, natural gas, and other hydrocarbon feedstocks into one or more higher hydrocarbons, methanol, amines, or other products which comprises continuously cycling through hydrocarbon halogenation, product formation, product separation, notably from coke, and electrolytic regeneration of halogen combined with hydrogen production.
  • US2009/0308759 describes a process of activation of natural gas by bromination.
  • the brominated products are thus converted by catalysis into C2+ hydrocarbons.
  • it is important to control the temperature to control the rate of deactivation of the catalysis because, during the conversion into C2+ hydrocarbons, the formation of carbonaceous coke is observed.
  • the formation of the coke has an impact on the stability of the catalyst, which must be regenerated when a drop in the conversion is observed.
  • WO2010/009376 describes a continuous process for converting natural gas to liquid hydrocarbons.
  • a stream of natural gas and a stream of bromine are put into contact to perform a bromination reaction in a bromination reactor.
  • the stream exiting the bromination reactor is separated into a polybromides stream and a monobromide stream.
  • the monobromide stream is conducted to a coupling unit in which hydrocarbons are generated.
  • the hydrogen bromide which is form is separated from the liquid hydrocarbons in a liquid-liquid splitter.
  • the aqueous HBr stream is then conducted to a bromine generation unit in which it is going to be oxidized with oxygen in the presence of a suitable catalyst to generate bromine that is going to be recovered and reused.
  • US2010/0087686 describes an integrated process for producing aromatic hydrocarbons and ethylene and/or propylene from low molecular weight hydrocarbons, such as methane.
  • Monobromomethane is generated and is reacted in the presence of a coupling catalyst to produce aromatic hydrocarbons and C2+ alkanes, which are then cracked into ethylene and/or propylene.
  • Hydrogen bromide can be recovered and directed to a bromination generation reactor which may be comprised of several shell/tube exchangers, whose tubes are filled with copper oxide catalyst.
  • the heat released by the exothermic conversion of hydrogen bromide to bromine is removed by the generation of steam which may be subsequently utilized in the process, notably in the fractionation steps.
  • As the olefins market is still continuously growing, there is a need for competitive olefins- producing processes. However, the above-described processes cannot be relied on in term of catalyst stability due to the by-side of coke.
  • the disclosure relates to a process for converting a stream comprising methane into chemicals, said process is remarkable in that it comprises the following steps: a) providing a first stream comprising methane (CH 4 ); b) providing a second stream which is a bromine (Br 2 ) -rich stream; c) putting into contact said first stream with said second stream under bromination reaction conditions to obtain a third stream comprising at least unreacted methane, methyl bromide (CHsBr), dibromomethane (CH 2 Br 2 ), and hydrogen bromide (HBr); d) removing said dibromomethane from said third stream under separation conditions, to produce a dibromomethane stream and a fourth stream comprising unreacted methane, methyl bromide and hydrogen bromide; e) converting the fourth stream into chemicals; f) separating hydrogen bromide from said third stream and/or said fourth stream, to provide a hydrogen bromide-rich stream; g)
  • the disclosure provides for a process allowing to produce hydrogen from an electrolysis reaction of hydrogen bromide and/or water using the electrical energy originated from the conversion of the heat generated by the oxidation of the hydrogen bromide-rich stream.
  • the process of the disclosure is remarkable by the co-production of a hydrogen stream in an energy-efficient way involving a heat integration process.
  • this process allows recovering hydrogen bromide while at the same time generating activated methane under the form of methyl bromide from a stream of methane.
  • the activated methane comprises a low content of dibromomethane, for example, less than 5 mol% based on the total molar content of the sum of methyl bromide and dibromomethane; preferably less than 1 mol%; more preferably the activated methane is devoid of dibromomethane.
  • this limits the formation of coke or carbonaceous components during further processing of methyl bromide into chemicals, which is advantageous in terms of conversion and selectivity of the chemicals.
  • the activation of methane being performed by bromination the recovery of hydrogen bromine allows to obtain a stream of olefins, in particular, a stream of ethylene and/or propylene, devoid of brominated compounds.
  • an electrolysis reaction of water is performed to produce a hydrogen stream and an oxygen stream which is optionally used in said step (g) of oxidation.
  • the disclosure provides for a process in which the hydrogen bromide generated is recovered and recycled by oxidation into bromine without bringing additional reactant except for oxygen.
  • the oxygen stream is produced by water electrolysis conducted thanks to the electricity produced from the thermal energy provided by the exothermic oxidation of hydrogen bromide into bromine.
  • the oxygen stream used in the step (g) of oxidation is one or more selected from air, air enriched in oxygen and an oxygen stream produced by an electrolysis reaction of water.
  • the oxidation of the hydrogen bromide-rich stream is performed on only a part of the said hydrogen bromide-rich stream and the other part of the hydrogen bromide-rich stream is submitted to an electrolysis reaction of hydrogen bromide to produce a stream of a hydrogen stream and a bromine stream being optionally reused in step (c).
  • step (a) of providing a first stream comprising methane comprises providing a natural gas comprising methane at a content of at least 75 mol.% of the total molar content of said natural gas, preferably of at least 85 mol.%, more preferably of at least 90 mol.%, even more preferably of at least 95 mol.%.
  • step (a) of providing a first stream comprising methane comprises providing a natural gas and purifying the natural gas to remove one or more selected from sulphur, nitrogen, water, oxygen and carbon dioxide.
  • the bromine-rich stream comprises a mixture of hydrogen bromide and bromine; with a content of bromine being superior to 80 mol.% based on the total molar content of said stream; with preference, the content of bromine is at least 90 mol.% based on the total molar content of said stream, more preferably at least 95 mol.%, even more preferably at least 99 mol.%.
  • the bromine-rich stream is a stream comprising only bromine.
  • the bromination reaction conditions of step (c) include one or more of the following conditions:
  • molar ratio methane to bromine of at least 7:1, preferably of at least 5:1, more preferably of at least 3:1, even more preferably of at least 2.5:1 ;
  • a pressure ranging from 0.1 MPa to 2.0 MPa, preferably ranging from 0.5 MPa to 1.5 MPa, more preferably ranging from 0.6 MPA to 1.0 MPa.
  • the bromination reaction conditions of step (c) include a temperature ranging from 350°C to 550°C, preferentially from 390°C to 500°C.
  • the bromination reaction conditions of step (c) include a temperature ranging from 300°C to 700°C or from 400°C to 700°C, preferentially from 350°C to 450°C or from 450°C to 650°C.
  • the separation conditions of step (d) comprise temperature conditions of at most 150°C, preferentially of at most 120°C.
  • step (d) comprises a distillation column with temperature conditions of at most 150°C, preferentially of at most 120°C.
  • the separation conditions of step (d) comprise pressure conditions ranging between 0.1 MPa and 2.0 MPa, preferentially between 0.1 MPa and 1.0 MPa or between 0.2 MPa and 1.0 MPa.
  • the dibromomethane stream of step (d) is further converted into carbon by performing carbonization of said dibromomethane stream, said carbonization is carried out using an electrical energy input and/or at a temperature of at least 500°C, or of at least 600°C, preferentially of at least 700°C, more preferentially of at least 1000°C, for example ranging between 500°C and 3500°C, preferentially between 600°C and 3000°C, more preferentially between 700°C and 2500°C.
  • the conversion of the dibromomethane stream of step (d) into carbon is a process of conversion of dibromomethane into carbon black and/or graphite.
  • the temperature which is required to produce carbon black is comprised between at least 500°C and below 3000°C, preferably between 750°C and 2000°C.
  • the temperature is above 2600°C, and up to 3500°C, the production of graphite is favoured.
  • graphite is formed at a temperature comprised between 3000°C and 3400°C.
  • step (f) of separating hydrogen bromide is performed by non-aqueous extraction and/or is a step of separating hydrogen bromide from said fourth stream.
  • said bromine-rich stream is reused in step (c) and is washed and/or dried before being reused in step (c).
  • the step (g) is carried out in the absence of one or more catalysts.
  • Said electrolysis reaction of water is carried out under pressure conditions ranging between 0.1 MPa and 20.0 MPa, preferably between 1.0 MPa and 15.0 MPa.
  • Said electrolysis reaction of water is carried out under temperature conditions ranging between 50°C and 1000°C, preferably between 100°C and 950°C.
  • Said electrolysis reaction of water is carried out upon consumption between 3.0 kWh/m 3 of hydrogen produced and 6.0 kWh/m 3 of hydrogen produced, preferably between 3.5 kWh/m 3 of hydrogen produced and 5.5 kWh/m 3 of hydrogen produced.
  • the electrolysis reaction of hydrogen bromide is carried out under gaseous phase or under liquid phase; with preference, said electrolysis is performed under water-free conditions:
  • said electrolysis reaction when the electrolysis reaction of hydrogen bromide is performed under gaseous phase, said electrolysis reaction comprises temperature conditions ranging between 300°C and 700°C, preferably between 350°C and 650°C, more preferably between 400°C and 600°C.
  • said electrolysis reaction when the electrolysis reaction of hydrogen bromide is performed under liquid phase, said electrolysis reaction comprises temperature conditions ranging between 20°C and 80°C, preferably between 30°C and 60°C, more preferably between 30°C and 40°C.
  • he electrolysis reaction of hydrogen bromide is preferably performed under pressure conditions of at least 0.1 MPa, preferably ranging between 0.1 MPa and 2.0 MPa, more preferably ranging between 0.5 MPa and 1.5 MPa, even more preferably ranging from 0.6 MPa to 1.0 MPa.
  • the process of the present disclosure also, to protect the catalyst from coke deactivation, also provides for enhanced energy management.
  • the step (e) of conversion of the fourth stream into chemicals comprises the following sub-steps: i. providing a first catalytical composition comprising at least one homologation catalyst; ii. putting into contact said fourth stream with the first catalytical composition under first reaction conditions to provide a first product stream comprising C1-C7 hydrocarbons and hydrogen bromide; iii.
  • a second catalytical composition comprising at least one cracking catalyst and putting into contact said first product stream with the second catalytical composition under second reaction conditions, to provide a second product stream comprising C1-C8 hydrocarbons and HBr; iv. separating C2-C4 hydrocarbons from said first product stream and/or from said second product stream when sub-step (iii) is carried out, to form a C2-C4 stream; with preference, said C2-C4 stream is further separated into an ethylene stream and/or into a propylene stream.
  • Said step of conversion of the fourth stream into chemicals further comprises the sub-step (v) of separating an unreacted methane stream from said first product stream and/or from said second product stream when sub-step (iii) is carried out, to form a methane stream; with preference, said methane stream is recycled in the first stream of step (a) and/or said methane stream is purged to form a fuel gas stream.
  • the first catalytical composition is steamed before sub-step (ii) and the one or more homologation catalysts of the first catalytical composition comprise one or more zeolites and a binder, wherein said one or more zeolites comprise at least one 10-membered ring channel.
  • said one or more zeolites of the first catalytical composition contain less than 1000 wt. ppm of alkali metals based on the total weight of the one or more zeolites and/or less than 5000 wt. ppm of transition metals based on the total weight of the one or more zeolites.
  • the content of the alkali metals is below 5000 wt. ppm based on the total weight of the one or more zeolites, preferably below 2500 wt. ppm.
  • the content of the alkaline earth metals is below 5000 wt. ppm based on the total weight of the one or more zeolites, preferably below 2500 wt. ppm.
  • the first catalytical composition may contain a higher content of alkaline earth metals as a component of the binder (e.g. Ca3(PC>4)2). So, additional traces of these metals may be present on the catalyst as impurities from the binder.
  • alkaline earth metals e.g. Ca3(PC>4)2
  • said one or more zeolites of the first catalytical composition contain less than 1000 wt. ppm of alkali metals based on the total weight of the one or more zeolites and/or less than 5000 wt. ppm of transition metals based on the total weight of the one or more zeolites.
  • the content of the alkali metals is below 5000 wt. ppm based on the total weight of the one or more zeolites, preferably below 2500 wt. ppm.
  • the content of the alkaline earth metals is below 5000 wt.
  • the first catalytical composition may contain a higher content of alkaline earth metals as a component of the binder (e.g. Ca3(PC>4)2). So, additional traces of these metals may be present on the catalyst as impurities from the binder.
  • the first catalytical composition is blended with at least one metal-containing material; with preference, the at least one metal-containing material is an alkaline earth metal- containing material which comprises at least one alkaline earth metal is selected from beryllium, magnesium, calcium, strontium, barium and any mixtures thereof, and/or the at least one metal-containing material has an anion selected from the group of oxides, silicates, aluminates, titanates, phosphates, borates and borosilicates.
  • the first catalytical composition comprises between 0.1 wt.% and 7.0 wt.% of a phosphorus-containing material as based on the total weight of the first catalytical composition, preferably between 0.3 wt.% and 4.5 wt.%, preferentially between 0.5 wt.% and 4.0 wt.%, more preferentially 2.0 wt.%.
  • the first catalytical composition is modified with 2.3 wt.% of phosphorous.
  • the first catalytical composition modified with phosphorous is blended with at least one metal-containing material; with preference, the at least one metal-containing material is one or more selected from an alkaline earth metal-containing material; magnesium nitrate; and a cerium-containing material.
  • one or more of the following embodiments can be used to better define the one or more zeolites of the first catalytical composition used in the present disclosure:
  • the one or more zeolites of the first catalytical composition have a crystal size below 2000 nm, as determined by scanning electron microscopy (SEM), preferentially below 1750 nm, more preferentially below 1500 nm, even more preferentially below 1250 nm.
  • the one or more zeolites have a Si/AI molar ratio of at least 10 before the step of steaming; and/or a Si/AI molar ratio of at least 80 after the step of steaming; with preference, of at least 150.
  • the one or more zeolites have a Si/AI molar ratio ranging from 80 to 1500 after the step of steaming; preferably ranging from 150 to 1200; more preferably ranging from 400 to 1100 and most preferably from 800 to 1000.
  • the one or more zeolites of the first catalytical composition are dealuminated with an organic acid solution or with an inorganic solution.
  • the one or more zeolites of the first catalytical composition are selected from the group of MFI, MEL, FER, MTT, MWW, TON, EUO and MRE families
  • the one or more zeolites of the first catalytical composition are selected from the group of MFI, MEL, FER, MTT, MWW, TON, EUO and MRE families and said one or more zeolites having a Si/AI molar ratio of at least 10 before the step of steaming; and/or a Si/AI molar ratio of at least 80 after the step of steaming; with preference, of at least 150.
  • the one or more zeolites of the first catalytical composition are selected from the MFI family, with a Si/AI molar ratio of at least 10.
  • the one or more zeolites of the first catalytical composition are selected from the list comprising ZSM-5, silicalites from the MFI family, boralite C, TS-1, ZSM-11, silicalites from the MEL family, boralite D, TS-2, SSZ-46, ferrierite, FU-9, ZSM-35, ZSM-23, MCM-22, PSH-3, ITQ-1, MCM-49, ZSM-22, Theta-1, NU-10, ZSM-50, EU-1 and ZSM- 48.
  • the one or more zeolites of the first catalytical composition are selected from the list comprising ZSM-5, silicalites from the MFI family, boralite C, TS-1, ZSM-11, silicalites from the MEL family, boralite D, TS-2, SSZ-46, ferrierite, FU-9, ZSM-35, ZSM-23, MCM-22, PSH-3, ITQ-1, MCM-49, ZSM-22, Theta-1, NU-10, ZSM-50, EU-1 and ZSM- 48, and said one or more zeolites having a Si/AI molar ratio of at least 10 before the step of steaming; and/or a Si/AI molar ratio of at least 80 after the step of steaming; with preference, of at least 150.
  • the one or more zeolites of the first catalytical composition are ZSM-5, with a Si/AI molar ratio of at least 10.
  • the one or more zeolites have Bronsted acid sites in a concentration inferior to 100 pmol/g-cat as determined by NH3-Temperature Programmed Desorption, preferentially inferior to 90 pmol/g-cat, more preferentially inferior to 80 pmol/g-cat.
  • one or more of the following embodiments can be used to better define the binder of the first catalytical composition used in the present disclosure:
  • the binder is selected from silica, clays, calcium phosphates, magnesium phosphates, and mullite. Most preferentially, the binder is silica.
  • the binder is devoid of aluminium compounds, such as alumina.
  • the binder is present in a content of at least 10 wt.% as based on the total weight of the first catalytical composition; preferably in a content of at least 20 wt.%, most preferably in a content of 30 wt.%, even more preferably in a content of at least 40 wt.%, and most preferably in a content of at least 50 wt.%.
  • the first catalytical composition is calcinated before said sub-step (ii); with preference, the first catalytical composition is calcined at a temperature of at least 400°C.
  • said steaming of the first catalytical composition before sub-step (ii) can be further defined with one or more of the following features:
  • Said steaming before sub-step (ii) is performed at a temperature ranging between 300°C and 800°C, preferentially ranging between 400°C and 750°C
  • the one or more steamed zeolites of the first catalytical composition are leached with an organic or inorganic acid solution, before the sub-step (ii).
  • the steaming and the leaching of the first catalytical composition are performed subsequently, the steaming step being conducted first.
  • Said step of steaming is carried out at a partial pressure of the steam ranging between 0.01 kPa and 20 kPa, preferentially between 0.5 kPa and 1.5 kPa.
  • Said step of steaming is followed by an extraction step, with preference with a monoprotic acid selected from HCI, HNO 3 , HBr, acetic acid or formic acid.
  • a monoprotic acid selected from HCI, HNO 3 , HBr, acetic acid or formic acid.
  • Said step of steaming is followed by an extraction step, with preference with a complexing agent or with an aqueous complexing agent.
  • Said step of steaming is followed by an extraction step and by a calcination step; with preference, said calcination step is carried out in a steam-free atmosphere at a temperature ranging between 550°C and 700°C, preferentially at a temperature ranging between 600°C and 650°C.
  • Said step of steaming is followed by a calcination step; with preference, said calcination step is carried out in a steam-free atmosphere at a temperature ranging between 550°C and 700°C, preferentially at a temperature ranging between 600°C and 650°C.
  • Said step of steaming is followed by a step of modification of the steamed catalyst by phosphorous.
  • said step of steaming of the first catalytical composition before sub-step (ii) is followed by a step of modification of the steamed catalyst by phosphorous under reduced or atmospheric pressure, preferentially at a temperature from 10 to 400°C, more preferentially at a temperature from 50°C to 350°C, even more preferentially at a temperature from 100°C to 300°C.
  • the source of phosphorous in the modification step of the steamed catalyst is mixed in an aqueous or a non-aqueous medium.
  • the source of phosphorous in the modification step of the steamed catalyst is mixed in a non-aqueous medium selected from the group of ethanol, methanol and/or other alcohols.
  • the source of phosphorous is phosphoric acid, preferably a solution of phosphoric acid.
  • the modification step of the steamed catalyst is followed by a calcination step; with preference, said calcination step is carried out in a steam-free atmosphere at a temperature ranging between 550°C and 700°C, preferentially at a temperature ranging between 600°C and 650°C.
  • the modification step of the steamed catalyst is followed by a further step of steaming, preferentially at a steam partial pressure comprised between 0.1 and 1.0 kPa and/or at a temperature comprised between 550 and 750°C and/or for a period of from 0.5 to 10 hours.
  • the modification step of the steamed catalyst is followed by a calcination step and by a further step of steaming, preferentially at a steam partial pressure comprised between 0.1 and 1.0 kPa and/or at a temperature comprised between 550 and 750°C and/or for a period of from 0.5 to 10 hours.
  • the first reaction conditions comprise
  • a pressure ranging between 0.2 MPa and 1.5 MPa, preferably between 0.4 MPa and 1.0 MPa, more preferably between 0.5 MPa and 1.0 MPa;
  • a weight hourly space velocity comprised between 0.1 h 1 and 100 h 1 , preferably comprised between 1.5 h 1 and 10 h 1 .
  • said process is remarkable in that said at least one cracking catalyst comprises one or more zeolites and/or one or more clays.
  • said at least one cracking catalyst comprises one or more zeolites selected from silicalites from the MFI family, crystalline silicate from the MFI family with a Si/AI atomic ratio of at least 180, crystalline silicate from the MEL family with a Si/AI atomic ratio ranging between 150 and 800, and/or phosphorous-modified zeolite from the MFI, MEL, FER, or MOR family.
  • said at least one cracking catalyst comprises one or more zeolites selected from silicalites from the MFI family, optionally with a silica binder.
  • said process is remarkable in that the second reaction conditions comprise
  • a weight hourly space velocity comprised between 0.1 h 1 and 100 h 1 , preferably comprised between 1.5 h 1 and 10 h 1 ;
  • the present disclosure provides for an installation for carrying out the process of conversion of natural gas into chemicals in accordance with the first aspect, said installation being remarkable in that it comprises:
  • the bromination unit, the conversion unit and the physical separation unit are fluidically connected in series, the conversion unit being downstream of said bromination unit and upstream of said physical separation unit; wherein said installation further comprises a first line to conduct the chemicals exiting the physical separation unit into one products-recovery unit; and wherein said installation further comprises a second line to conduct at least a part of the hydrogen bromide-rich stream exiting said physical separation unit to an oxidation unit, said oxidation unit comprising at least one oxidation reactor coupled with at least one electrolysis cell selected from a water electrolysis cell and/or a hydrogen bromide electrolysis cell and a steam-electric power sub-unit is placed downstream of said oxidation reactor and upstream of the one or more electrolysis cells.
  • said installation further comprises a third line to conduct the dibromomethane stream exiting the bromination unit to a fluidized bed reactor.
  • At least one electrolysis cell is a water electrolysis cell.
  • At least one electrolysis cell is a hydrogen bromide electrolysis cell and wherein the installation further comprises a line to conduct a part of the hydrogen bromide-rich stream into the electrolysis cell.
  • said conversion unit comprises one homologation reactor and optionally one cracking reactor.
  • said bromination unit comprises at least one bromination reactor and at least one dibromomethane separator, the one or more dibromomethane separators being downstream of said at least one bromination reactor.
  • said at least one water electrolysis cell is an alkaline electrolysis cell, a proton- exchange membrane electrolysis cell or a solid oxide electrolysis cell.
  • FIG. 1 illustrates the installation in accordance with the present disclosure.
  • Figure 2 shows the NH 3 -TPD profile of the catalyst of the first catalytical composition according to the present disclosure.
  • Figure 3 shows an oxygen-transport membrane reactor (OTMR) having steam as a feed source.
  • OTMR oxygen-transport membrane reactor
  • Figure 4 shows an OTMR having air as a feed source.
  • FIG. 5 shows an example of the settings of the temperature-programmed desorption (TPD) method.
  • Fig 6 is an insert that can be fitted onto figure 1 , in between the hydrogen bromide-rich stream and the bromine-rich stream.
  • the feed gas of the process i.e. the first gaseous stream comprising methane
  • OSBL outside battery limit
  • C# hydrocarbons wherein “#” is a positive integer, is meant to describe all hydrocarbons having # carbon atoms. C# hydrocarbons are sometimes indicated as just C#. Moreover, the term “C#+ hydrocarbons” is meant to describe all hydrocarbon molecules having # or more carbon atoms. Accordingly, the expression “C2+ hydrocarbons” is meant to describe a mixture of hydrocarbons having 2 or more carbon atoms.
  • LPG means “liquefied petroleum gas” and is a mixture of C3-C4 essentially composed of propane and butane.
  • steam is used to refer to water in the gas phase, which is formed when water boils.
  • Zeolite codes are defined according to the “Atlas of Zeolite Framework Types", 6 th revised edition, 2007, Elsevier, to which the present application also refers.
  • the Si/AI atomic ratio corresponds to the content of S1O 2 divided by the content of AI 2 O 3 taking into account the fact there are two atoms of aluminium for one atom of silicon.
  • the silicon to aluminium ratio (also stated as SAR) corresponds to the content of S1O 2 divided by the content of AI 2 O 3 notwithstanding the proportion of the Si atoms over the Al atoms in the chemical formula of the zeolite. Therefore, the value of the SAR always corresponds to twice the value of the Si/AI atomic ratio.
  • the disclosure provides a process and an installation for the conversion of natural gas into chemicals, in particular into ethylene, propylene, LPG, gasoline and hydrogen.
  • the process of the present disclosure is a process of conversion of a stream comprising methane into chemicals, said process being remarkable in that it comprises the following steps: a) providing a first stream (1, 5, 15) comprising methane; b) providing a second stream (133, 153, 157), which is a bromine-rich stream; c) putting into contact said first stream 15 with said second stream (133, 153, 157) under bromination reaction conditions to obtain a third stream 21 comprising at least unreacted methane, methyl bromide, dibromomethane, and hydrogen bromide; d) removing said dibromomethane from said third stream 21 under separation conditions, to produce a dibromomethane stream 103 and a fourth stream 27 comprising unreacted methane, methyl bromide and hydrogen bromide; e) converting the fourth stream 27 into
  • the installation for carrying out the process of conversion of natural gas into chemicals is remarkable in that it comprises:
  • the bromination unit 111 , the conversion unit 113 and the physical separation unit 115 are fluidically connected in series, the conversion unit 113 being downstream of said bromination unit 111 and upstream of said physical separation unit 115; wherein said installation further comprises a first line to conduct the chemicals exiting the physical separation unit 115 into one products-recovery unit 119; and wherein said installation further comprises a second line to conduct at least a part of the hydrogen bromide-rich stream 69 exiting said physical separation unit 115 to an oxidation unit 159, said oxidation unit 159 comprising at least one oxidation reactor 73 coupled with at least one electrolysis cell selected from a water electrolysis cell 123 and/or a hydrogen bromide electrolysis cell 171 (visible on figure 6), and a steam-electric power sub-unit 77 is placed downstream of said oxidation reactor 73 and upstream of the one or more electrolysis cells (123, 171).
  • an electrolysis reaction of water is performed to produce a hydrogen stream (127, 131) and an oxygen stream 125 which is optionally used in said step (g) of oxidation.
  • the oxygen stream used in the step (g) of oxidation can be air, air enriched in oxygen or the oxygen stream originating from the electrolysis of water (as shown in figure 1).
  • the oxidation of the hydrogen bromide-rich stream 69 is performed on only a part of the said hydrogen bromide-rich stream 69 and the other part of the hydrogen bromide-rich stream 69 is used in an electrolysis reaction of hydrogen bromide to produce a stream of a hydrogen stream and a bromine stream.
  • the bromine stream can be reused in step (c) for a new cycle.
  • the oxygen stream 125 may be recycled in the step (g) of oxidation, once the reaction has been initiated.
  • a fresh oxygen stream may be provided to initiate the oxidation reaction and/or electrical energy may be provided to initiate the one or more electrolysis reactions.
  • the first stream (1 , 5, 15) is the feed stream of the process.
  • the first stream (1 , 5, 15) is or comprises natural gas.
  • the first stream (1 , 5, 15) is a natural gas comprising methane.
  • the first stream (1, 5, 15) is a natural gas comprising methane at a content of at least 75 mol.% of the total molar content of said natural gas; preferably at least 85 mol.%, more preferably at least 90 mol.%, and even more preferably of at least 95 mol.% of methane.
  • the first stream (1 , 5, 15) comprises methane, and also C2+ hydrocarbons.
  • the C2+ hydrocarbons present in the first stream (1 , 5, 15) may include, for example, lower molecular weight alkanes.
  • lower molecular weight alkanes refers to methane, ethane, propane, butane, pentane, or mixtures thereof.
  • the first stream 1 is advantageously injected into a pre-treatment unit 121 of natural gas.
  • the pre-treatment unit 121 can comprise a demethanizer 13 to separate the methane from the first stream (1 , 5,). At least a part 11 of the first stream (1 ,5) is directed to the demethanizer. A methane stream (14, 15) and a C2-C4 hydrocarbons stream (17, 83) are thus separated.
  • the pre-treatment unit 121 may also include a purification unit 3.
  • the first stream 1 comprising methane can be subjected to an optional preliminary step of purification to remove one or more selected from sulphur, nitrogen, water, oxygen and carbon dioxide.
  • the pre-treatment unit 121 of natural gas may also comprise one or more heat exchangers, such as one or more cold boxes 7 disposed on a line upstream of the demethanizer 13.
  • the purified first stream 5 exiting from the purification unit 3, or the first stream 1 comprising methane, can be conveyed into the cold box 7, to provide a first stream 11 having its temperature adapted to the operating conditions of the demethanizer 13.
  • the demethanizer 13 is a cryogenic distillation column.
  • the operating conditions of the demethanizer 13 comprise a temperature ranging between -120°C and - 80°C, preferably between -110°C and -90°C, and/or a pressure ranging between 2.5 MPa and 3.5 MPa.
  • a stream 14 of methane having a temperature inferior to the first stream (1 , 5) is redirected to the cold box 7 for cooling the first stream (1 , 5).
  • the stream (1 , 5, 14) comprising methane can be purged, to recover a fuel gas stream 9.
  • the methane stream 15 exiting the cold box 7 can be conveyed into an additional heat exchanger 16 before exiting the pre-treatment unit 121.
  • the C2-C4 hydrocarbons stream (17, 83) separated from the first stream (1, 5, 11) can be conveyed into one products-recovery unit 119 of the installation (see below).
  • the methane stream 15 can be put into contact with a second stream (133, 153, 157) which is a bromine-rich stream and then introduced into a bromination reactor 19 of one bromination unit 111 as shown on figure 1.
  • a second stream (133, 153, 157) comprising bromine can be put in contact with the methane stream 15 within the bromination reactor 19 (not shown).
  • the molar ratio of methane to bromine can be advantageously high, for example, the molar ratio of methane to bromine can be of at least 7:1, preferably at least 5: 1 , more preferably at least 3: 1 , even more preferably at least 2.5:1.
  • the molar ratio of methane to bromine is ranging from 2.5:1 to 12:1, preferably of about 3:1 to about 10:1.
  • the reaction selectivity may be above 80 % for methyl bromide and less than 20% dibromomethane, on a molar basis.
  • the bromination reaction conditions of step (c) include a pressure of at least 0.1 MPa; preferably, ranging from 0.1 MPa to 2.0 MPa; more preferably ranging from 0.5 to 1.5 MPa, and more preferably ranging from 0.6 to 1.0 MPa.
  • step (c) can be performed in the absence of a catalyst.
  • the bromination reaction conditions of step (c) when no catalysts are involved include a temperature ranging from 300°C to 550°C or from 350°C and 550°C, or from 300°C to 500°C, or from 390°C and 500°C.
  • step (c) can be performed in the presence of a catalyst.
  • the bromination reaction conditions of step (c) when at least one catalyst is involved include a temperature ranging between 300°C and 700°C or 400°C and 700°C, preferably between 350°C and 450°C or between 450°C and 650°C; more preferably between, 500°C and 600°C.
  • either solid strongly acidic catalysts or (supported Group VIII metal (particularly platinum and palladium) catalysts are capable of catalyzing the gas-phase bromination of methane predominantly to methyl bromide in 85 to 99% selectivity.
  • Subsequent or concurrent catalytic hydrolysis can produce methyl alcohol and/or dimethyl ether
  • a particularly useful class of solid, strongly acidic catalysts are those derived from halides, oxyhalides, oxides, sulfides and oxysulfides of metals, particularly transition metals of Groups IV, V, VI, VIII of the Periodic Table, such as tantalum, niobium, zirconium, tungsten, titanium, chromium, and the like, or mixtures thereof, deposited on suitable chalconite carriers, such as alumina, zirconia or silica-alumina. These catalysts are capable of effecting the ready conversion of methane to methyl halides.
  • a variety of solid oxides and sulfides especially those comprising alumina, silica and mixtures of alumina and silica, either natural or synthetic, in which other oxides such as chromia, magnesia, moiybdena, thoria, tungstic oxide, zirconia, etc., may also be present, as well as sulfides of molybdenum are useful chalcide carriers.
  • Many naturally occurring compositions exist for use as the carriers including bauxite, floridin, Georgia clay, and other natural aluminosilicates.
  • Synthetic chalconites other than those of the silica-alumina type representative of the chalconite carriers are BeO, C ⁇ Os, P 2 O 5 , ThCb, T1O2, AhiSO ⁇ s (which may be regarded as AhC SSOs), AI2O3 ⁇ G2O3, AI2O3, Fe 2 03, AI2O3 ⁇ 0O, AhC nO, A Os V2O3, AI 2 03 Mo 2 03, Cr 2 03-Fe 2 03, M0S2, and M0S3.
  • the acidic chalconite supports are physically and chemically stable. They are generally catalytically active at only higher temperatures, as their acidity is not great enough to lead them to form stable complexes with unsaturated compounds, as do the aluminium halides, for example.
  • the supported Group VI II metal catalysts include the various Group VI II metals supported on suitable chalconite carriers. Particularly useful platinum and palladium supported on alumina, silica, barium sulfate or related carriers.
  • a third stream 21 comprising at least methyl bromide, dibromomethane and hydrogen bromide is then conveyed into a dibromomethane separator 25.
  • the dibromomethane separator 25 can be one or more distillation columns and/or one or more adsorption columns and/or one or more absorption columns, preferably one or more distillation columns.
  • the third stream 21 can be advantageously directed into a heat exchanger 23 to adjust its temperature before entering into the dibromomethane separator 25.
  • the separation conditions that are operated during the step (d) are based on the boiling point (measured under atmospheric pressure, i.e. about 0.1 MPa) of the dibromomethane (96.95°C), of the methyl bromide (3.56°C), methane (-161.5°C) and hydrogen bromide (-66°C). Therefore, when distillation is carried out in one or more distillation columns, it is relatively simple to collect the volatile fraction made of compounds having low boiling points by performing the distillation under temperature conditions of at most 150°C, preferentially of at most 120°C.
  • the pressure conditions operated during step (d) are therefore comprised between 0.1 MPa and 1.0 MPa.
  • a dibromomethane stream 103 which is a non-volatile stream, is then removed from said third stream 21.
  • a fluidized bed reactor preferably an electrothermal fluidized bed reactor 105
  • a carbonization step to produce carbon 109.
  • the carbonization step is carried out using an electrical energy input 107 and/or at a temperature of at least 500°C, or of at least 600°C, preferably of at least 800°C, more preferably of at least 1000°C.
  • the carbonization step is carried out at a temperature ranging between 500°C and 3500°C, preferably between 600°C and 3000°C.
  • Graphite can advantageously be obtained when the temperature is ranging between 2600°C and 3500°C, preferably between 3000°C and 3400°C. Carbon black can advantageously be obtained when the temperature is ranging between 500°C and below 3000°C.
  • the fluidized bed reactor can be fluidized thanks to an inert gas flux, such as a flux of nitrogen, argon and/or helium.
  • the volatile stream exiting the dibromomethane separator 25 is recovered in a fourth stream 27 comprising at least unreacted methane, methyl bromide and hydrogen bromide.
  • the fourth stream 27 is conveyed to a conversion unit 113, comprising at least one homologation reactor 31 and optionally, at least one cracking reactor 37.
  • the fourth stream 27 is then enriched into C2-C4 hydrocarbons to generate valuable chemicals that can be recovered into a C2-C4 stream (17, 83) and further separated into interesting compounds, such as ethylene and/or propylene, but also liquefied petroleum gas.
  • the step (e) of conversion of the fourth stream 27 into chemicals comprises the following sub steps: i. providing a first catalytical composition comprising at least one homologation catalyst, ii. putting into contact said fourth stream 27 with the first catalytical composition under first reaction conditions to provide a first product stream 33 comprising C1-C7 hydrocarbons and hydrogen bromide; iii.
  • a second catalytical composition comprising at least one cracking catalyst and putting into contact said first product stream 33 with the second catalytical composition under second reaction conditions, to provide a second product stream 39 comprising C1-C8 hydrocarbons and HBr; iv. separating C2-C4 hydrocarbons from said first product stream 33 and/or from said second product stream 39 when sub-step (iii) is carried out, to form a C2-C4 stream (17, 83).
  • a homologation reactor 31 loaded with a first catalytical composition comprising one or more homologation catalysts is placed downstream of the dibromomethane separator 25.
  • the homologation reactor 31 can be selected from a fixed bed reactor or a fluidized bed reactor, preferably a fixed bed reactor.
  • the installation may comprise one or more heat exchangers 29 arranged between the dibromomethane separator 25 and the homologation reactor 31 so that one or more steps of adjusting the temperature of the fourth stream 27 before it enters the homologation reactor 31 is included.
  • the first catalytical composition comprising one or more zeolites and a binder is steamed before used in sub-step (ii) and said one or more zeolites comprise at least one 10-membered ring channel.
  • the one or more zeolites namely the one or more zeolites before the step of steaming, or the non-steamed one or more zeolites, do not contain any alkali metals since these metals may significantly reduce catalyst activity and neutralize acid sites.
  • said one or more zeolites initially contain less than 1000 wt. ppm of alkali metals, as based on the total weight of the one or more zeolites.
  • the one or more zeolites namely the one or more zeolites before the step of steaming, do not contain any alkaline earth metal since these metals may impact the steam dealumination process and retain halogen after the reaction. The retained halogen will be released during the regeneration and irreversibly deactivate zeolites.
  • the alkaline earth metal is strongly bound with the phosphorous and is less prone to the formation of halides.
  • said one or more zeolites, before the steaming step contain less than 5000 wt. ppm of alkaline earth metals, as based on the total weight of the one or more zeolites.
  • the first catalytical composition does not contain any transition metal since this leads to a completely distinct reactivity resulting in coke formation. This is why the first catalytical composition is devoid of any transition metal.
  • the content of the transition metals is below 5000 wt. ppm in the one or more zeolites, preferably below 2500 wt. ppm in the one or more zeolites. Traces of these metals may be present on the catalyst as impurities from the binder.
  • the one or more zeolites comprise at least one acid 10-membered ring channel; with preference, the one or more zeolites are selected from the list comprising MFI, MEL, FER, MTT, MWW, TON, EUO and MRE families, preferentially from the MFI family or the MEL family.
  • These zeolites or molecular sieves are aluminosilicate catalysts that have a chemical structure that is largely different from the chemical structure of the aluminophosphate and silicoaluminophosphate molecular sieves.
  • the zeolite from the MFI family is selected from ZSM-5, silicalites, boralite C, or TS-1.
  • the zeolites are silicalites from the MFI family or ZSM-5, more preferentially the zeolites are silicalites from the MFI family.
  • the zeolites from the MEL family are, preferentially, selected from ZSM-11 , silicalites, boralite D, TS-2, or SSZ-46.
  • the zeolites are silicalites from the MEL family.
  • the zeolites from the FER family are, preferentially, selected from ferrierite, FU-9 or ZSM-35.
  • the zeolites from the MTT family are, preferentially, ZSM-23.
  • the zeolites from the MWW family are, preferentially, selected from MCM-22, PSH-3, ITQ-1, or MCM-49.
  • the zeolites from the TON family are, preferentially, selected from ZSM-22, Theta-1 , or NU-10.
  • the zeolites from the EUO family are, preferentially, selected from ZSM-50 or EU-1.
  • the zeolites from the MRE family are, preferentially, ZSM-48.
  • the first catalytical composition comprises one or more zeolites with at least one acid 10-membered ring channel.
  • the one or more zeolites have a crystal size below 2000 nm as determined by scanning electron microscopy (SEM), preferentially below 1750 nm, more preferentially below 1500 nm and even more preferentially below 1250 nm.
  • SEM scanning electron microscopy
  • the first catalytical composition comprising one or more zeolites is steamed before step (c) of contacting said feedstream with the said first catalytical composition under reaction conditions to obtain a higher Si/AI molar ratio relative to the non-steamed one or more zeolites.
  • the one or more zeolites are selected from the list comprising ZSM-5, silicalites from the MFI family, boralite C, TS-1, ZSM-11, silicalites from the MEL family, boralite D, TS-2, SSZ-46, ferrierite, FU-9, ZSM-35, ZSM-23, MCM-22, PSH-3, ITQ-1, MCM- 49, ZSM-22, Theta-1, NU-10, ZSM-50, EU-1 and ZSM-48, said one or more zeolites having a Si/AI molar ratio of at least 10 before the step of steaming.
  • the one or more zeolites are selected from the list comprising ZSM-5, silicalites from the MFI family, boralite C, TS-1, ZSM-11, silicalites from the MEL family, boralite D, TS-2, SSZ-46, ferrierite, FU-9, ZSM-35, ZSM-23, MCM-22, PSH-3, ITQ-1, MCM- 49, ZSM-22, Theta-1 , NU-10, ZSM-50, EU-1 and ZSM-48, said one or more zeolites having a Si/AI molar ratio of at least 80 after the step of steaming; with preference, of at least 150.
  • the first catalytical composition comprises 3D zeolites without cages (cavities) and containing at least one acid 10-membered ring channel.
  • the first catalytical composition comprises at least 60 wt.% of one or more zeolites having at least one acid 10-membered ring channel, more preferably at least 70 wt.%, even more preferably at least 80 wt.% and most preferably at least 90 wt.% or 95 wt.%, or 100 wt.%.
  • the one or more zeolites are at least partly in their hydrogen form or at least partly in their ammonia form.
  • Preferably more than 50 wt.% of the total content of the zeolites used are in their hydrogen form or their ammonia form, preferably at least 80 wt.%, more preferably at least 90 wt.%, and even more preferably 100 wt.% of the zeolites are in their hydrogen form or their ammonia form.
  • the one or more zeolites have weak Bronsted acid sites in a concentration inferior to 40 pmol/g-cat and strong Bronsted acid sites in a concentration superior to 40 pmol/g-cat as determined by NH 3 -TPD.
  • the one or more zeolites have Bronsted acid sites in a concentration inferior to 100 pmol/g-cat as determined by NH 3 -TPD, preferentially inferior to 90 pmol/g-cat, more preferentially inferior to 80 pmol/g-cat. This can be obtained by performing a step of steaming the one or more zeolites before the contact of the first catalytical composition with the fourth stream 27.
  • the acidity of the zeolite catalyst was measured by NH 3 -TPD.
  • a temperature at which NH 3 is desorbed is an estimation of the strength of an acid site, i.e. higher the desorption temperature stronger is the acid site.
  • the zeolite catalyst shows two NH 3 -TPD peaks, a first one at 184°C and a second at 363°C ( Figure 2).
  • the one or more zeolites used in the first catalytical composition of the invention have a Si/AI molar ratio of at least 10 before the step of steaming,
  • the Si/AI molar ratio before the step of steaming is typically ranging from 10 to 1500; preferably ranging from 80 to 1200; more preferably ranging from 150 to 1100 and most preferably from 800 to 1000.
  • the steam treatment is conducted at elevated temperature, preferably in the range of from 300 to 800°C, more preferably in the range of from 400 to 750°C and at a partial pressure of steam from 0.01 to 20 kPa, preferentially from 0.5 to 1.5 kPa.
  • the steam treatment is conducted at partial pressure of steam at least 1.5 kPa in the temperature range 300-450°C. If the temperature is above 450°C, the steam treatment is conducted in an atmosphere comprising the steam partial pressure below 1.5 kPa.
  • the concentration of steam in the flow is between 1 to 100%, more preferably from 5 to 20% of steam.
  • the diluent is a gas selected from the group of N2, air, natural gas, CO2 or a mixture of thereof.
  • the steam treatment is preferably carried out for a period of from 0.1 to 200 hours, more preferably from 0.2 hours to 24 hours.
  • the steam treatment tends to reduce the content of tetrahedral aluminium in the crystalline silicate framework, by forming alumina.
  • the particular effect consists of reducing the strong Bronsted external acidity of the zeolites.
  • One or more zeolites used in the first catalytical composition of the invention have a Si/AI molar ratio of at least 80 after the step of steaming,
  • the Si/AI molar ratio after the step of steaming is typically ranging from 80 to 1500; preferably ranging from 150 to 1200; more preferably ranging from 400 to 1100 and most preferably from 800 to 1000.
  • an extraction step is performed to remove the partially dislodged alumina species by leaching.
  • the leaching is performed by a monoprotic acid selected from the HCI, HNO 3 , HBr, acetic or formic or with a complexing agent which tends to form a soluble complex with alumina.
  • the complexing agent is preferably in an aqueous solution thereof.
  • the complexing agent may comprise an organic acid such as citric acid, oxalic acid, tartaric acid, malonic acid, succinic acid, glutaric acid, adipic acid, maleic acid, phthalic acid, isophthalic acid, fumaric acid, nitrilotriacetic acid, hydroxyethylenediaminetriacetic acid, ethylenediaminetetracetic 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 such acids or salts.
  • a particularly preferred complexing agent may comprise an amine, preferably ethylene diamine tetraacetic acid (EDTA) or a salt thereof, in particular, the sodium salt thereof.
  • EDTA ethylene diamine tetraacetic acid
  • the catalyst is advantageously thereafter calcined in absence of steam ( ⁇ 1% of steam) at a temperature of from 550 to 700°C at atmospheric pressure for a period of from 0.5 to 10 hours.
  • the one or more zeolites comprise at least one 10-membered ring channel, with crystal size below 2000 nm, and have a Si/AI molar ratio ranging from 10 to 1500 after the step of steaming; preferably ranging from 80 to 1200; more preferably ranging from 150 to 1100 and most preferably from 800 to 1000.
  • the one or more zeolites are preferably zeolites selected from the silicalites from the MFI family and/or silicalites from the MEL family having a Si/AI molar ratio of at least 10 before the step of steaming; and/or a Si/AI molar ratio of at least 80 after the step of steaming; with preference, of at least 150.
  • the zeolite is ZSM-5.
  • said steamed first catalytical composition is further modified by phosphorous under reduced or atmospheric pressure at a temperature from 10 to 400°C.
  • a non-limiting source of phosphorus can be provided in an aqueous or non-aqueous medium.
  • the non-limiting source of phosphorus is dissolved in a non-aqueous medium selected from the group containing ethanol, methanol or other alcohols.
  • the doping with a phosphorus-containing material consists of a steaming step followed by a leaching step using a solution of phosphoric acid (H 3 PO 4 ) or using any acid solution containing the source of phosphorus.
  • a solution of phosphoric acid H 3 PO 4
  • H 3 PO 4 phosphoric acid
  • the treatment of the steamed zeolite with an acid solution results in the dissolution of the extra-framework aluminiumoxides. This transformation is known as leaching.
  • the zeolite is separated, advantageously by filtration, and optionally washed.
  • a drying step can be envisaged between the filtering and washing steps.
  • the solution after the washing can be either separated, by way of example, by filtering from the solid or evaporated.
  • the residual phosphorus-content is adjusted by the phosphorus concentration in the leaching solution, drying conditions, and washing procedure if any. This procedure leads to dealumination of zeolites and retention of phosphorus.
  • at least 0.1 wt.% and up to 7.0 wt.% of phosphorus is retained after dealumination on zeolite.
  • the first catalytical composition advantageously comprises at least 0.1 wt.% of phosphorous based on the total weight of the first catalytical composition, preferentially with at least 0.5 wt.% of phosphorous, more preferentially with at least 1.0 wt.% of phosphorous, even more preferentially with at least 1.5 wt.% of phosphorous.
  • Both factors dealumination and the retention of phosphorus stabilize the lattice aluminium in the zeolitic lattice, thus avoiding further dealumination.
  • the degree of dealumination can be adjusted by the steaming and leaching conditions.
  • the preferred techniques suitable for the modification by phosphorous are impregnation and chemical vapour deposition.
  • the phosphorus is introduced by a treatment of the catalyst in a solution containing a source of phosphorus at a temperature ranging between 25 and 100°C for 0.1- 96 h followed by filtering or evaporation.
  • the incipient wetness (IW) impregnation techniques are used.
  • IW impregnation techniques the phosphorus is introduced via impregnation using a limited amount of liquid water which is subjected to contact with the catalyst. This method is also known as dry impregnation.
  • IW Incipient wetness
  • IWI incipient wetness impregnation
  • the precursor phosphorus-containing compounds
  • the volume of solution which is used for dissolution of the precursor, is substantially the same as the pore volume of catalyst precursor containing both binder and zeolite.
  • the precursor-containing solution is added to a catalyst precursor. Capillary action draws the solution into the pores.
  • the catalyst can then be dried and calcined to drive off the volatile components within the solution, depositing the phosphorus on the catalyst surface.
  • the sample before impregnation can be dried or calcined.
  • the impregnation could be performed at room or elevated temperature.
  • the adsorption capacity is typically measured by impregnating the dried extruded zeolite with water until the zeolite was completely wet. Weighing the zeolite before and after impregnation gives the absorption capacity according to formula (1):
  • H 3 PO 4 solution is used for impregnation.
  • a mixture of H 3 PO 4 with their ammonium salts providing a pH of the aqueous solution higher than 2.0 is used for impregnation.
  • the sources of phosphorus are substantially metal-free components, for example, H 3 PO 4 , ammonium phosphates or organic phosphorous-compounds.
  • this proportion can be below 1000 wt. ppm of the total weight of the phosphorous- containing material.
  • the content of phosphorus in the catalyst can be from 0.1 to 30.0 wt.%, preferably from 0.3 to 9.0 wt.%.
  • the content of phosphorous on the catalyst is most preferably 2.0 wt.%.
  • the catalyst is thereafter calcined and/or steamed at a steam partial pressure between 0.1 and 1 kPa at a temperature of from 550 to 750°C for a period of from 0.5 to 10 hours.
  • the crystalline alumino-silicate oxide framework of the one or more zeolite has a portion of the aluminium that is substituted with boron and/or titanium.
  • boron is used to substitute one or more aluminium atoms in the zeolite framework.
  • Boron-substituted zeolite has a very weak acidity.
  • the zeolite catalysts have a Si/(AI+B) molar ratio of at least 80, typically comprised between 100 and 1200, preferentially of 1000.
  • the first catalytical composition modified with a phosphorous containing-material may contain a metal-containing material, which is preferably an alkaline earth metal-containing material.
  • a metal-containing material which is preferably an alkaline earth metal-containing material.
  • the alkaline earth metal-containing material is spatially separated from the zeolite, in which alkaline earth metal is strongly bounded with phosphorous.
  • the said alkaline earth metal is selected from the group of beryllium, magnesium, calcium, strontium, barium and any mixtures thereof.
  • the metal-containing material that can be added to a catalytical composition modified with phosphorous is advantageously in the form of alkaline earth metal salts and comprise at least one inorganic anion selected preferably from the group of oxides, silicates, aluminates, titanates, phosphates, borates and borosilicates.
  • Suitable silicate anions include S1O3 2 , S1O4 4 , S12O7 6 and so on.
  • Suitable borate anions include BO2 , BO3 2 , B2O5 4 , B4O7 2 , Bbqii 4 , B10O19 8' and so on.
  • Suitable aluminate anions include AI2O4 2 , AIO4 5 , AI d Oib 18- and so on.
  • Suitable titanate anions include T1O3 2 , T13O7 2 , TUOg 2 , TiC 4 and so on.
  • Suitable phosphate anions include PO4 3 , HPO4 2 , H2PO4 , PnC>3n+i (n+2)' and so on.
  • Bi-, tri- and poly-metal silicates, borates and borosilicates containing one, two or more alkaline earth metals selected from the list above can be used too.
  • the metal salt may also comprise other anions.
  • alkaline earth metal salts that can be added to a catalytical composition modified with phosphorous include Mg6Al2C03(0H)i6.4(H20) (hydrotalcite), Mg2B2C>5.H20, CaMgB60n.6H20 (hydroboracite), Ca 2 B 6 0n.5H 2 0 (colemanite), Ca 4 BioOig.7H 2 0, Mg(B0 2 ).8H 2 0, Ca(B0 2 ).2H 2 0, BaB 6 O 10 .4H 2 O, CaSi 6 0 17 (0H) 2 (xonotlite), CaMg(Si 2 0 6 ) x , Mg 2 (Si 2 0 6 )x, CaAI 2 Si 2 08, Mg 4 Si60i5(0H)2-6H 2 0 (sepiolite), (Mg,AI)2Si 4 O 10 (OH)-4H 2 O (palygorskite or attap
  • a further example of suitable alkaline earth metals that can be added to a catalytical composition modified with phosphorous is Mg(NC>3)2 (magnesium nitrate).
  • said alkaline earth metal salts Before mixing with the molecular sieve, said alkaline earth metal salts may be modified by calcination, steaming, ion-exchange, impregnation, and/or phosphatation. Said alkaline earth metal salts may be an individual compound or may be a part of mixed compounds, for example, mixed with mineral, natural or chemical fertilizer.
  • the first catalytical composition of the present invention modified with at least one phosphorous-containing material and at least one alkaline earth metal-containing material has for effect to increase the selectivity to olefins (i.e. acyclic C3-C6 olefins) and to decrease subsequently the rate of the alkane formation (i.e. C3-C6 alkanes).
  • olefins i.e. acyclic C3-C6 olefins
  • alkane formation i.e. C3-C6 alkanes
  • the first catalytical composition modified with phosphorous further comprises from 1 to 50 wt.% of hydrotalcite as based on the total weight of the first catalytical composition; with preference from 5 to 25 wt.%.
  • the hydrotalcite is of the formula Mg 6 Al2C03(0H) 16 .4(H20).
  • the one or more zeolites are doped with both at least one phosphorus-containing material and with at least one alkaline earth metal-containing material, preferably at least one magnesium-containing material and/or at least one calcium-containing material.
  • one or more zeolites are shaped with a binder, which is an inorganic material, and preferentially silica.
  • the zeolites shaped with the binder forms a catalytical composition
  • the catalytical composition of the present disclosure preferably comprises at least 10 wt.% of a binder, at most 40 wt.% as based on the total weight of the first catalytical composition and at most 40 wt.%.
  • the first catalytical composition of the present invention comprises between 20 wt.% and 25 wt.% of a binder as based on the total weight of the first catalytical composition.
  • the preferred binder is selected from silica, alpha-alumina, clays, alumina phosphates, calcium phosphates, magnesium phosphates, and mullite. Most preferentially, the binder is silica.
  • the binder preferably does not contain any aluminium compounds, such as alumina. This is because as mentioned above the preferred catalyst for use in the invention is de-aluminated by steaming to increase the Si/AI molar ratio of the crystalline silicate. The presence of alumina in the binder, as well as the presence of hydrogen halides, may lead to the re-alumination of the zeolite. The presence of aluminium in the binder would also tend to reduce the olefins selectivity of the catalyst and to reduce the stability of the catalyst over time.
  • the binder is present in a content of at least 10 wt.% as based on the total weight of the first catalytical composition; preferably in a content of at least 20 wt.%, most preferably in a content of 30 wt.%, even more preferably in a content of at least 40 wt.%, and most preferably in a content of at least 50 wt.%.
  • Non-limiting examples of silicon sources suitable for the binder of the catalytical composition include silicates, precipitated silicas (for example, Zeosil® available from Rhodia), fumed silicas (for example, Aerosil®200 available from Degussa Inc., New York, N.Y.), silicon compounds (such as tetraalkyl orthosilicates, for example, tetramethyl orthosilicate (TMOS) and tetraethylorthosilicate (TEOS)), colloidal silicas or aqueous suspensions thereof (for example, Ludox® HS-40 available from E.l. du Pont de Nemours, Wilmington, Del.), silicic acid, alkali-metal silicate, or any combination thereof.
  • silicates for example, Zeosil® available from Rhodia
  • fumed silicas for example, Aerosil®200 available from Degussa Inc., New York, N.Y.
  • silicon compounds such as tetraalkyl orthosilicate
  • amorphous silica examples include silica powders, such as Ultrasil® VN3 SP (commercially available from Degussa).
  • a suitable solid silica source are special granulated hydrophilic fumed silica, mesoporous silica and high surface area precipitated silica SIPERNAT® from Evonik, Hi-Sil 233 EP (available from PPG Industries) and Tokusil (available from Tokuyama Asia Pacific).
  • suitable amorphous silica sources include silica sols, which are stable colloidal dispersions of amorphous silica particles in an aqueous or organic liquid medium, preferably water.
  • Non-limiting examples of commercially available silica sols include those sold under the tradenames Nyacol® (available from Nyacol Nano Technologies, Inc. or PQ Corp.), Nalco (available from Nalco Chemical Company), Ultra-Sol (available from RESI Inc), Ludox® (available from W.R. Grace Davison), NexSilTM (available from NNTI).
  • Many silica sols are prepared from sodium silicate and inevitably contain sodium. It is, however, found that the presence of sodium ions can cause sintering of the silica body at high temperature and/or affect catalytic performance. Therefore, if silica sols containing sodium are used, a step of ion exchange may be required to reduce or remove sodium.
  • silica sols that contain very little or, ideally, no detectable traces of sodium and have a pH value of less than 7. Most preferably, the silica sol used in the process is slightly acidic with or without polymeric stabilizers.
  • Non-limiting examples of silica sols that contain no detectable traces of sodium include Bindzil® 2034DI, Levasil® 200, Nalco 1034A, Ultra-Sol 7H or NexSilTM 20A.
  • silica dispersion prepared with alkylammonium might be useful.
  • Non-limiting examples of commercially low sodium silica sols stabilized by ammonia or alkylammonium cations include LUDOX® TMA (available from W.R. Grace Davison) or VP WR 8520 from Evonik.
  • the preferred source of silicon is a silica sol or a combination of silica sol with precipitated or fumed silica.
  • the homologation reaction is carried out under first reaction conditions.
  • Said first reaction conditions comprise a temperature ranging between 250°C and 460°C, preferably between 280°C and 420°C, more preferably between 280°C and 400°C.
  • Said first reaction conditions also comprises a pressure ranging between 0.2 MPa and 1.5 MPa, preferably between 0.4 MPa and 1.0 MPa; more preferably between 0.5 MPa and 1.0 MPa.
  • the weight of feed comprising flowing per unit of weight of the catalyst per hour is comprised between 0.1 h 1 and 100 h 1 , preferentially between 1.0 h 1 and 15 h 1 . More preferably, WHSV is comprised between 1.5 h 1 and 10 h 1 . Even more preferably, WHSV is comprised between 2.0 h 1 and 6.0 h 1 . This means that the homologation catalyst of the present disclosure can convert the weight of the fourth stream that is superior to the amount of the catalyst present in the homologation reactor 31.
  • a first product stream 33 of C1-C7 hydrocarbons, including olefins, paraffins and aromatics, can be recovered. Alkyl bromide and hydrogen bromide are also comprised in the first product stream 33.
  • a Sub-step of separation of the C2-C4 hydrocarbons from said first product stream 33 can be carried out to form a C2-C4 stream. Said C2-C4 stream can be further separated into an ethylene stream and/or into a propylene stream.
  • the first product stream 33 can be further processed under a cracking Sub-step (iii).
  • the first product stream 33 is then conveyed into a cracking reactor 37, possibly via one or more heat exchangers 35 arranged between the homologation reactor 31 and the cracking reactor 37, so that one or more steps of adjusting the temperature of the first product stream 33 can be performed.
  • the cracking reactor 37 can be selected from a fixed bed reactor or a fluidized bed reactor, preferably a fixed bed reactor.
  • the cracking reactor 37 is loaded with a second catalytical composition comprising at least one cracking catalyst.
  • said cracking catalyst is a catalyst suitable for an olefin cracking reaction.
  • Preferred catalysts for the olefin cracking reaction can be selected from one or more zeolites and/or one or more clays.
  • said cracking catalyst comprises one or more zeolites selected from silicalites from the MFI family, crystalline silicate from the MFI family with a Si/AI atomic ratio of at least 180, crystalline silicate from the MEL family with a Si/AI atomic ratio ranging between 150 and 800, and/or phosphorous-modified zeolite from the MFI, MEL, FER, MOR family and/or phosphorous-modified clinoptilolite.
  • said cracking catalyst comprises one or more zeolites selected from silicalites from the MFI family, optionally with a silica binder.
  • Examples of crystalline silicate from the MFI family are ZSM-5 and silicalite.
  • An example of crystalline silicate from the MEL family is ZSM-11, which is known in the art.
  • Other suitable non-limiting examples are boralite D and silicalite-2, or any mixtures thereof.
  • the preferred crystalline silicates have pores or channels defined by ten oxygen rings and a high Si/AI atomic ratio.
  • the catalyst having a high Si/AI atomic ratio may be manufactured by removing aluminium from a commercially available catalyst.
  • the commercially available catalysts may be modified by steaming to remove at least part of inter-framework aluminium followed by a leaching step to remove external aluminium.
  • the cracking catalyst can be formulated with a binder, preferably an inorganic binder, and shaped to the desired shape, e.g. extruded pellets.
  • the binder is an inorganic material selected from clays, silica, metal oxides.
  • the binder content ranges from 5 to 50% by weight, more typically from 15 to 35% by weight, based on the weight of the cracking catalyst. More preferably, the binder is a silica binder.
  • the cracking catalyst can be subjected to a steaming step before the sub step (iii).
  • the olefin cracking reaction is known per se. It has been described in EP1035915, EP1036133, EP1036134, EP1036135, EP1036136, EP1036137, EP1036138, EP1036139, EP1190015, EP1194500, EP1194502, and EP1363983; the content of which is incorporated in the present description.
  • the second reaction conditions comprise a temperature ranging from 400°C to 600°C, preferably ranging from 450°C to 550°C; and/or a weight hourly space velocity comprised between 0.1 h 1 and 100 h 1 , preferably comprised between 1.5 h 1 and 10 h 1 ; and/or a pressure ranging between 0.5 MPa and 1.5 MPa, preferentially between 0.6 MPa and 1.0 MPa.
  • the stream exiting the cracking reactor 37 is a second product stream 39 comprising C1-C8 hydrocarbons and hydrogen bromide, in particular comprising unreacted methane, C2-C4 olefins, C2-C8 paraffins, C6-C8 aromatics, hydrogen bromide, alkyl bromides including monobromomethane.
  • a sub-step of separation of the C2-C4 hydrocarbons from said second product stream 39 can be carried out to form a C2-C4 stream (17, 83).
  • Said C2-C4 stream (17, 83) can be further separated into an ethylene stream 91 and/or into a propylene stream 99.
  • the second product stream 39 is conveyed into a physical separation unit 115.
  • the conversion unit 113 does not comprise the cracking reactor, this is the first product stream 33 that is conveyed into the physical separation unit 115 (not shown).
  • the second product stream 39 can then be directed into one or more separation columns (43, 59, 65).
  • the second product stream 39 is cooled down in one or more heat exchangers 41 before being processed in a first separation column, being a debutanizer 43.
  • the debutanizer 43 is a distillation column and/or is working at a temperature ranging between 50°C and 70°C, and/or is working at a pressure ranging between 0.3 MPa to 1.0 MPa. Passing the product stream 39 in the debutanizer 43 separates the product stream 39 into a fifth stream (51, 55) and a sixth stream (45), the fifth stream (51 , 55) comprising C1-C4 hydrocarbons and hydrogen bromide and the sixth stream (45) comprising C5-C8 hydrocarbons, aromatics (/.e., at least one of benzene, toluene and/or xylene) and hydrogen bromide.
  • An optional separation step can be provided on the sixth stream 45, through for example a separator 47, such as an adsorption column, to remove hydrogen bromide and recover a stream of a C5-C8 hydrocarbons stream and aromatics (/.e., at least one of benzene, toluene and/or xylene), /.e. a gasoline stream 49.
  • a stream 167 of hydrogen bromide can also be recovered.
  • the gasoline stream 49 can be further conveyed into one or more distillation columns 161 to separate a C5-C6 hydrocarbons stream 163 that can be further recycled back to the homologation reactor 31.
  • the distillation column 161 is a dehexanizer.
  • the pressure conditions implemented in the one or more distillation columns 161 are comprised between 0.1 MPa and 0.4 MPa, preferably between 0.2 MPa and 0.3 MPa.
  • the temperature conditions implemented in the one or more distillation columns 161 are comprised between 50°C and 150°C, preferably between 55°C and 145°C.
  • the remaining part of the gasoline stream 49, namely a stream 165 comprising aromatics (/.e., at least one of benzene, toluene and/or xylene) and C7-C8 hydrocarbons can be optionally further separated into a C7-C8 hydrocarbons stream and an aromatics stream (not shown).
  • the C5-C6 hydrocarbons stream 163 can be purged, to recover a fuel gas stream 169 comprising C5-C6 hydrocarbons.
  • the separation and recycling of the C5-C6 hydrocarbons stream 163 back to the homologation reactor 31 allows the first catalytical composition to convert the fourth stream 27 comprising unreacted methane, methyl bromide and hydrogen bromide into a first product stream 33 comprising C1-C7 hydrocarbons and hydrogen bromide.
  • the continuous feeding of C5-C6 hydrocarbons to the homologation reactor 31 allows their reaction with methyl bromide to generate C6-C7 hydrocarbons and therefore form the first product stream 33 comprising C1- C7 hydrocarbons.
  • the fifth stream 51 can be submitted to one or more further separation steps such as a step of separation of hydrogen bromide from the C1-C4 hydrocarbons.
  • this step can be performed in an extractive distillation system, comprising one or more separation columns (59, 65).
  • the removal of hydrogen bromide is performed by non-aqueous extraction.
  • the fifth stream 51 comprising C1-C4 hydrocarbons and hydrogen bromide is passed through one extractive distillation column 59, which is loaded with one specific solvent, such as solvent comprising alcohol, carboxylic acid, ketone, organobromine compounds, ionic liquid, organic acid anhydride and/or nitrile.
  • one suitable solvent is acetic acid.
  • the hydrogen bromide which is contained in the fifth stream 51 is thus absorbed on said specific solvent, which allows the separation of the C1-C4 hydrocarbons and then the production of a stream 61 comprising C1- C4 hydrocarbons.
  • Said stream 61 comprising C1-C4 hydrocarbons can be recycled into the first stream (1 , 5) comprising natural gas and/or into the one or more of the reactors 19 of the bromination unit 111.
  • the stream 61 comprising C1-C4 hydrocarbons is conveyed into the products- recovery unit 115 (see below).
  • the installation 1 may also comprise one or more water adsorbers, such as the water adsorption column 53, arranged between the debutanizer 43 and the extractive distillation system so that one or more steps of drying the fifth stream 51 before it enters the extractive distillation system is included.
  • one or more heat exchangers 57 can be potentially placed between the water adsorption column 53 and the extractive distillation system.
  • the remaining stream 63 exiting the extractive distillation column 59 is directed to a second separation column 65 of the extractive distillation system to recover the solvent which was used in the extractive distillation column 59 and to produce a hydrogen bromide-rich stream 69.
  • Said second separation column 65 is preferably a distillation column.
  • the extractive distillation system as described provides the advantage that the hydrogen bromide-rich stream 69 does not comprise water.
  • the recovered solvent 67 is preferably redirected to the extractive distillation column 59 so that the generation of the hydrogen bromide-rich stream 69 can be continuously achieved.
  • the hydrogen bromide-rich stream 69 comprises only hydrogen bromide.
  • the hydrogen bromide-rich stream 69 exiting the physical separation unit 115, is conveyed to an oxidation unit 159 to produce a stream (127, 131) comprising hydrogen and a stream (133, 153, 157) which is a bromine-rich stream. More particularly, the hydrogen bromide-rich stream 69 can be directed in an oxidation reactor 73, optionally after one or more steps of adjusting its temperature in one or more heat exchangers 71.
  • the oxidation reactor 73 can be a thermal oxidizer, multi-tubular fixed bed reactor or a shell and tube heat exchanger reactor, preferably a shell and tube heat exchanger reactor, optionally a catalyst bed to complete the oxidation.
  • the oxidation reactor 73 can be an oxygen-transport membrane reactor (OTMR) as shown in figure 3 with steam as a feed source.
  • OTMR oxygen-transport membrane reactor
  • such membranes are in ceramics, preferably from the perovskite-type (see study entitled “Perovskite-based proton conducting membranes for hydrogen separation: a review” of Hashim S. S. et al., Int. J. of Hydrogen Energy, 2018, 43, 15281-15305).
  • the oxidation reactor 73 can be an oxygen-transport membrane reactor (OTMR) as shown in figure 4 with air as a feed source.
  • OTMR oxygen-transport membrane reactor
  • the electrical potential which is applied allows to generate oxygen (O 2 ) from steam (boiling water) or to extract oxygen (O 2 ) from the air, such source of oxygen being used as an oxidant in the oxidation of hydrogen bromide into bromine.
  • the membrane reactor can be composed of a porous cermet (i.e.
  • the oxidation of at least a part of the hydrogen bromide-rich stream 69 is advantageously carried out in the absence of one or more catalysts.
  • hydrogen bromide is let to react with oxygen (O2) in the gas phase.
  • the pressure conditions of the oxidation reactor are comprised between 0.5 MPa and 1.5 MPa, preferably between 0.7 MPa and 1.2 MPa.
  • the temperature conditions of the oxidation reaction are comprised between 500°C and 1000°C, preferably between 600°C and 1000°C.
  • the oxidation of at least a part of the hydrogen bromide-rich stream 69 is carried out in the presence of one or more catalysts, it is advantageous to use a shell and tube heat exchanger reactor, so that the one or more catalysts can be installed in the tube side.
  • one or more catalysts are metal-oxides supported on alumina. Further details about the one or more catalysts can be found in US 2011/0015458, the content of which is incorporated by reference.
  • the oxidation of at least a part of hydrogen bromide into bromine is an exothermic reaction (i.e. the heat reaction is about -69 kJ/mol of hydrogen bromide).
  • the shell and tube heat exchanger reactor allow recovering the heat by vaporizing at the shell side of the shell and tube heat exchanger reactor water into a high-pressure stream.
  • the shell side of the shell and tube heat exchanger can act as a steam generator which can be included in a thermodynamic steam cycle for power production.
  • the steam cycle is composed of a feed water pump for compression, preferably an isentropic compression; a steam generator for heat addition, preferably for isobaric heat addition; a steam turbine for expansion, preferably for isentropic expansion; and a condenser for heat rejection, preferably for isobaric heat rejection.
  • thermodynamic efficiency It is expected that about 30% to 50% thermodynamic efficiency leads to 0.006 (69/3600*30%) to 0.01 (69/3600*50%) kWh of produced work per mole of HBr oxidized.
  • the thermodynamic efficiency has been determined by simulation via ASPEN PLUS V9 software.
  • the thermal energy 75 recovered from the oxidation reactor that has been transferred to water is further converted into electrical energy 79 thanks to its passage within a steam-electric power sub-unit 77.
  • the steam-electric power sub-unit 77 can be a steam turbine.
  • the steam-electric power sub-unit 77 is placed downstream of the oxidation reactor 73 and upstream of one or more water electrolysis cells 123.
  • the oxidation reactor 73 is thus coupled with at least one water electrolysis cell 123.
  • the electrical energy 79 serves then to feed at least one water electrolysis cell 123, for example, an alkaline electrolysis cell, a proton-exchange membrane electrolysis cell or a solid oxide electrolysis cell.
  • This configuration is interesting in the sense that it prevents the use of an external electricity supply for the water electrolysis cell 123.
  • Said electrolysis reaction of water is advantageously carried out under pressure conditions ranging between 0.1 MPa and 20.0 MPa, preferably between 0.5 MPa and 17.0 MPa, more preferably between 1.0 MPa and 15.0 MPa.
  • Said electrolysis reaction of water can also be carried out under temperature conditions ranging between 50°C and 1000°C, preferably between 75°C and 975°C, more preferably between 100°C and 900°C.
  • Said electrolysis reaction of water is advantageously carried out upon consumption between 3 kWh/m 3 of hydrogen produced and 6 kWh/m 3 of hydrogen produced, preferably between 3.5 kWh/m 3 of hydrogen produced and 5.5 kWh/m 3 of hydrogen produced.
  • the hydrogen stream 127 can be further dried over the drier system 129.
  • the drier system 129 is a desiccant.
  • the desiccant can be a molecular sieve, such as one or more zeolite from the LTA family. Among the LTA family, zeolites from LTA-3A, LTA- 4A and/or LTA-5A can be selected.
  • the drier system 129 is a hydrogen drying system as described in US2016/0129390.
  • the drier system 129 works under room temperature, such as between 15°C and 25°C, and/or under pressure conditions of at least 0.1 MPa, preferably ranging between 0.1 MPa and 2.0 MPa, more preferably ranging between 0.5 MPa and 1.5 MPa, even more preferably ranging from 0.6 MPa to 1.0 MPa.
  • the oxygen stream 125 is advantageously mixed with the hydrogen bromide-rich stream 69 before entering the oxidation reactor 73, as shown in figure 1.
  • the oxygen stream 125 can also be directed directly into the oxidation reactor 73 (not shown).
  • the oxygen stream 125 is thus the oxidant used in the oxidation of hydrogen bromide.
  • the thermal energy 75 recovered from the oxidation reactor 73 during the oxidation step is at least partially recycled, which therefore improved the energy management of the described installation.
  • the thermal energy 75, or the corresponding electrical energy 79 could be re directed in another installation. In that case, a cheap and easy-to-handle oxidant for working the oxidation reaction could be the use of air.
  • the bromine-rich stream 133 exiting the oxidation reactor 73 comprises traces of water and subsequently forms hydrogen bromide by hydrolysis.
  • the bromine-rich stream 133 comprising thus a mixture of bromine and hydrogen bromide is advantageously directed into a washing tower 135, which allows separating a first part of hydrogen bromide from said mixture.
  • the aqueous hydrogen bromine stream 137 exiting the washing tower 135 can be optionally redirected upstream of the oxidation reactor 73, for additional recovery by oxidation into bromine.
  • the stream 139 exiting the washing tower comprises the remaining part of said mixture and can be directed, optionally after one or more steps of adjusting its temperature in one or more heat exchangers 141, to a decantation sub-unit 143.
  • the optional decantation sub-unit 143 allows the separation on one hand of the bromine, to recover a bromine-rich stream 153, and on the other hand of the hydrogen bromide in gaseous form and aqueous form (respectively, a gaseous stream (147, 151) and an aqueous stream 144). Finally, water can be recovered, forming a water stream 145, and can be advantageously conveyed into the water electrolysis cell 123.
  • the bromine-rich stream 153 exiting the optional decantation sub-unit 143 is preferably dried in a dryer system 155 to produce the bromine-rich stream 157 that exits the oxidation unit 159 and that can be optionally put into contact with the methane stream 15 before being redirected into the bromination reactor 19 or reused in the bromination unit 111.
  • the dryer system 155 is a desiccant.
  • the desiccant can be a molecular sieve, such as one or more zeolite from the LTA family. Among the LTA family, zeolites from LTA-3A, LTA-4A and/or LTA-5A can be selected.
  • the gaseous stream 147 can be optionally passed through a purification column 149 to eliminate any residual traces of bromine.
  • the removal of the gaseous stream 147 allows the optional decantation sub-unit 143 to diminish the internal pressure and to make it work efficiently.
  • the aqueous stream 144 can be optionally redirected into the washing tower 135, so that the aqueous hydrogen bromide is either redirected on a line upstream to the oxidation reactor 73 or redirected into the optional decantation sub-unit 143.
  • the oxidation of the hydrogen bromide-rich stream 69 is performed on only a part of the said hydrogen bromide-rich stream 69 and the other part of the hydrogen bromide- rich stream 69 is submitted to an electrolysis reaction of hydrogen bromide to produce a stream of the hydrogen stream 173 and a bromine stream 133 being optionally reused in step (c).
  • the installation further comprises a line to conduct a part of the hydrogen bromide-rich stream into electrolysis cell 171.
  • the hydrogen stream 173 exiting the hydrogen bromide electrolysis cell 171 can be recovered as such and does not need to be directed into a drier system.
  • a part of the hydrogen bromide-rich stream, exiting the physical separation unit, is conveyed to an electrolysis cell to produce a stream comprising hydrogen and a stream which is a bromine-rich stream.
  • the hydrogen bromide-rich stream can be directed in an electrolysis cell, being preferentially a gas-phase electrolysis cell.
  • the gas-phase electrolysis cell can be a proton exchange membrane reactor and is supplied with an energy stream.
  • the energy stream is an electrical energy stream produced from the thermal energy recovered in the oxidation step of the other part of the hydrogen bromide-rich stream which is converted into electrical energy.
  • the hydrogen bromide-rich stream may comprise only hydrogen bromide
  • its electrolysis in the electrolysis unit under gaseous phase provide a bromine-rich stream that comprises only bromine and can be thus put in contact without any further treatment with the methane stream.
  • the electrolysis can be achieved in a proton-exchange membrane electrolysis cell.
  • it may comprise an anode side and a cathode side, the anode side and the cathode side is separated by a proton-conducting membrane, for example, a proton-conducting membrane made in ceramic materials.
  • the electrolysis when the electrolysis of a part of the hydrogen bromide-rich stream is carried out in gas phase, the electrolysis is performed under temperature conditions ranging between 300°C and 700°C, preferably between 325°C and 675°C, more preferably between 375°C and 625°C.
  • the electrolysis when the electrolysis is carried out in liquid phase including the aqueous phase, the electrolysis is performed under temperature conditions ranging between 20°C and 80°C, preferably between 25°C and 70°C, more preferably between 30°C and 40°C.
  • the electrolysis is performed under pressure conditions of at least 0.1 MPa, preferably ranging between 0.1 MPa and 2.0 MPa, more preferably ranging between 0.5 MPa and 1.5 MPa, even more preferably ranging from 0.6 MPa to 1.0 MPa.
  • the bromine-rich stream exiting from the electrolysis cell can be preferentially used in the bromination unit, either by being mixed with the methane stream before entering the bromination unit or by being directly conveyed into the bromination reactor.
  • the process comprises this recycling step, there is no net consumption of bromine, because it is fully regenerated in the process.
  • the most preferred way to carry out the electrolysis step is to do it under gaseous phase, with preference in the absence of water, since the bromine-rich stream only comprises bromine and the installation does not need to have a bromine-recovery unit disposed downstream of the electrolysis cell.
  • the electrolysis of hydrogen bromine-rich stream allows the formation of hydrogen as a pure stream product and avoids the use of oxidative bromine recovery routes.
  • the electrolysis When the electrolysis is performed in liquid phase, or more particularly in aqueous phase, the electrolysis does not fully convert hydrogen bromide into bromine.
  • the electrolysis when the electrolysis is performed in aqueous phase, stable performances of the electrolysis cell are achieved to obtain up to about 60% of hydrogen bromide conversion along with an energy efficiency of about 85%.
  • the presence of at least one bromine-recovery unit may allow recovering a bromine-rich stream comprising only bromine.
  • the installation for carrying out the present process further comprises one products-recovery unit 119 downstream of the demethanizer 13 of the pre-treatment unit 121.
  • a demethanizer other than the demethanizer 13 of the pre-treatment unit 121 can be installed upstream of the products-recovery unit 119.
  • the product-recovery unit 119 is supplied with the C2-C4 hydrocarbons stream 17 separated from the first stream (1, 5, 11), but also, by the stream 61 comprising C1-C4 hydrocarbons coming from the dibromomethane and hydrogen bromide separation unit 111 which has been processed to a demethanizer, preferably the demethanizer 13 of the pre-treatment unit 117.
  • a demethanizer preferably the demethanizer 13 of the pre-treatment unit 117.
  • the use of the demethanizer 13 of the pre-treatment unit 117 allows maximizing the potential of the component of the installation, rendering the installation cost-effective.
  • the C2-C4 hydrocarbons stream 17 can be optionally conveyed into one or more heat exchangers 81 to adjust the temperature of the C2-C4 hydrocarbons stream 83 before it enters into a de-ethanizer 85.
  • the de-ethanizer 85 can be a cryogenic distillation column.
  • the de-ethanizer 85 works under a temperature ranging between - 5°C and 15°C, more preferably between 0°C and 10°C, and/or under a pressure ranging between 1.5 MPa and 2.5 MPa.
  • Said C2 hydrocarbons stream 87 is conveyed to at least one C2 splitter 89.
  • one or more C2 splitters are one or more cryogenic distillation columns.
  • the one or more C2 splitters work under a temperature ranging between -40°C and -20°C, more preferably between -35°C and -25°C, and/or under a pressure ranging between 1.5 MPa and 2.5 MPa, preferably between 1.7 MPa and 2.3 MPa. This allows to remove the ethane from the C2 hydrocarbons stream 87 to recover an ethylene stream 91 and optionally an ethane stream 93.
  • Said C3-C4 hydrocarbons stream 95 is conveyed to at least one C3 splitter 97.
  • one or more C3 splitters are one or more cryogenic distillation columns.
  • the one or more C3 splitters work under a temperature ranging between 20°C and 60°C, more preferably between 30°C and 50°C, and/or under a pressure ranging between 1.5 MPa and 2.5 MPa, preferably between 1.7 MPa and 2.3 MPa. This allows recovering a propylene stream 99 and a liquid petroleum gas (LPG) stream 101.
  • LPG liquid petroleum gas
  • Temperature Programmed Desorption is the method of observing desorbed molecules from a surface when the surface temperature is increased. It has been performed by following the heating sequences I, II and III shows in figure 5, respectively corresponding to activation, saturation and analysis.
  • the temperature has been gradually increased to 600°C at a rate of 20°C/min.
  • the zeolite sample is considered as being activated and the temperature is then gradually decreased to 100°C at a rate of 10°C/min.
  • the temperature is maintained at 100°C and in the first 1 hour, 10% of ammonia (NH 3 ) is added to the helium flow (which is decreased to 30 cc/min).
  • NH 3 ammonia
  • the surface of the zeolite is thus saturated with the molecules of ammonia that are going to be adsorbed onto the surface.
  • the initial flow of helium is reinstated.
  • the temperature is increased again to 600°C at a rate of 10°C/min to desorb the ammonia.
  • the sample is maintained at 600°C for an additional one hour.
  • the homologation reactor 31 and cracking reactor 37 were charged with a silicalite catalyst from the MFI family having a Si/AI molar ratio of >800 (as defined by TPD) and shaped with a S1O2 binder. For the catalytic test, the catalyst was pressed, then crushed and seized between 35-45 mesh screens and preactivated in an N2 flow at 525°C for 6 h. Said homologation reactor 31 and cracking reactor 37 are fixed-bed tubular reactors with inner diameter of 11 mm made of INCONEL alloy. The homologation reactor 31 had a catalyst loading ⁇ 10 mL (6 g) and the cracking reactor 37 was ⁇ 3.5 mL (2 g).
  • the sampling was done from a 2-way split valve installed in between the homologation reactor and the cracking reactor and from the effluent of the cracking reactor.
  • Table 1 shows the molar composition of the first product stream 33 of C1-C7 hydrocarbons exiting the homologation reactor 31 and of the second product stream 39 comprising C1-C8 hydrocarbons exiting the cracking reactor 37.
  • Table 1 Molar composition of first product stream 33 and second product stream 39.
  • Table 2 shows an ASPEN simulation of the production of the installation.
  • the fluxes of the products displayed in table 2 show the potential of the installation of the present disclosure.

Abstract

La présente invention concerne, selon un premier aspect, un procédé de conversion d'un flux contenant du méthane en produits chimiques, ledit procédé étant remarquable en ce qu'il comprend les étapes consistant à : fournir un premier flux (1, 5, 11) contenant du méthane, fournir un deuxième flux (79) qui est un flux riche en brome, mettre en contact le premier flux (15) avec le deuxième flux (79) pour obtenir un troisième flux (21) contenant au moins du méthane n'ayant pas réagi, du bromure de méthyle, du dibromométhane et du bromure d'hydrogène, éliminer le dibromométhane du troisième flux (21) pour obtenir un flux de dibromométhane (103) et un quatrième flux (27) contenant du méthane n'ayant pas réagi, du bromure de méthyle et du bromure d'hydrogène qui est converti en produits chimiques, et séparer le bromure d'hydrogène du quatrième flux (27) pour fournir un flux riche en bromure d'hydrogène. Selon son second aspect, la présente invention concerne une installation destinée à la mise en œuvre du procédé selon le premier aspect.
PCT/EP2021/058131 2020-03-30 2021-03-29 Procédé de conversion de gaz en oléfines avec coproduction d'hydrogène conjointement à un procédé d'intégration thermique WO2021198166A1 (fr)

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