EP4245828A1 - Procédé et installation de production d'un kérosène synthétique à partir des composés oxygénés - Google Patents

Procédé et installation de production d'un kérosène synthétique à partir des composés oxygénés Download PDF

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Publication number
EP4245828A1
EP4245828A1 EP22162046.1A EP22162046A EP4245828A1 EP 4245828 A1 EP4245828 A1 EP 4245828A1 EP 22162046 A EP22162046 A EP 22162046A EP 4245828 A1 EP4245828 A1 EP 4245828A1
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EP
European Patent Office
Prior art keywords
unit
iii
oligomerization
olefin
separation unit
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EP22162046.1A
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German (de)
English (en)
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EP4245828A8 (fr
Inventor
Joachim Engelmann
Mario KUSCHEL
Stephan Schmidt
Norbert VÖLKEL
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Cac Engineering GmbH
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Cac Engineering GmbH
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Priority to EP22162046.1A priority Critical patent/EP4245828A1/fr
Priority to PCT/EP2023/056432 priority patent/WO2023174916A1/fr
Publication of EP4245828A1 publication Critical patent/EP4245828A1/fr
Publication of EP4245828A8 publication Critical patent/EP4245828A8/fr
Pending legal-status Critical Current

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    • 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
    • C10G57/00—Treatment of hydrocarbon oils, in the absence of hydrogen, by at least one cracking process or refining process and at least one other conversion process
    • C10G57/02—Treatment of hydrocarbon oils, in the absence of hydrogen, by at least one cracking process or refining process and at least one other conversion process with polymerisation
    • 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
    • C10G3/00—Production of liquid hydrocarbon mixtures from oxygen-containing organic materials, e.g. fatty oils, fatty acids
    • 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
    • C10G67/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only
    • C10G67/02—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only plural serial stages only
    • C10G67/04—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one process for refining in the absence of hydrogen only plural serial stages only including solvent extraction as the refining step in the absence of hydrogen
    • C10G67/0409—Extraction of unsaturated hydrocarbons
    • C10G67/0418—The hydrotreatment being a hydrorefining
    • 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
    • C10G69/00—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process
    • C10G69/02—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only
    • C10G69/12—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only including at least one polymerisation or alkylation step
    • C10G69/126—Treatment of hydrocarbon oils by at least one hydrotreatment process and at least one other conversion process plural serial stages only including at least one polymerisation or alkylation step polymerisation, e.g. oligomerisation
    • 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
    • C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/08—Jet fuel

Definitions

  • the invention relates to a method and an associated system for producing synthetic kerosene from oxygenates, such as methanol.
  • WO2006/076942A1 discloses a process for producing synthetic fuels from oxygenates. Returns are also described here, such as the return of predominantly saturated hydrocarbons after an oligomerization step back to olefin production (MTO step). In addition, unsaturated hydrocarbons are recycled for oligomerization.
  • the oligomerization step i.e. H. chain extension, in the synthesis of diesel fuel, among other things.
  • a fraction of light cracked naphtha was chosen as light olefins, which is composed of C5 to C6 compounds and is characterized by a high proportion of olefins.
  • the catalyst used was the zeolitic catalyst H-ZSM-5.
  • the authors also refer to the Catpoly process, in which a catalyst made from phosphoric acid on kieselguhr converts C3 to C4 olefins into oligomers with a boiling point in the gasoline or aircraft fuel range.
  • WO2018/045397A1 also discloses a process for producing aircraft fuel by oligomerizing light olefins from Fischer-Tropsch synthesis using a ZSM-5 zeolite catalyst. A gasoline fraction is then distilled off and the remaining oligomerization product is hydrogenated using Co/Mo or Pt catalysts. As a product, kerosene can be obtained both with a low aromatic content and with an aromatic content of over 8% by volume.
  • US4,506,106 deals with the entire process, starting from oxygenates, especially methanol, up to the distillate hydrocarbons, with ethylene being produced as an intermediate using a gasoline sorbent stream from a hydrocarbon stream with predominantly light components that are rich in C2 to C4 olefins is washed out. The ethylene is then discharged as a chemical byproduct of the process. After absorption, the gasoline sorbent stream is led to oligomerization.
  • the object of the invention is to develop a process and the associated system that allows the production of synthetic fuels, in particular kerosene as well as gasoline and diesel components.
  • the fuel obtained should be obtained with good yield. In particular, the fuel should contain few aromatics.
  • step I) olefins are produced in the olefin production unit (I), which are then separated off in the absorption II) according to the invention by means of the absorption unit (II).
  • the oligomerization III) is carried out in the oligomerization unit (III), the hydrogenation IV) in the hydrogenation unit (IV) and the separation V) using a separation unit (Va and/or Vb). Explanations for the process steps therefore also directly affect the system components.
  • Step V) and system part (V), ((V) includes (Va) and/or (Vb)) are located at various points in the method according to the invention. That is, (V) is arranged in the form of the separation unit (Va) after (III) and/or in the form of the separation unit (Vb) after (IV).
  • process steps and the associated system parts also require appropriate equipment.
  • the separation process step is carried out by the separation unit in the system.
  • a separation unit in the plant corresponds, for example, to the arrangement of one or more pieces of equipment, such as a distillation column or phase separator.
  • a return in the process requires, for example, a corresponding line from one part of the system to another part of the system to which it is to be returned.
  • the present invention is a process and a plant for producing a synthetic kerosene as the main product, as well as gasoline components and diesel components as co-products.
  • synthetic kerosene is a fuel whose main components are hydrocarbons with chain lengths of C9+, in particular C9 to C17. These are mainly paraffins and naphthenes.
  • the synthetic kerosene produced has an aromatic content of 0% by volume to 20% by volume, in particular it is below 10% by volume, particularly preferably from 0% by volume to 8% by volume.
  • the kerosene is characterized by a characteristic boiling curve, with the components preferably boiling in the range from 130 to 300 ° C.
  • Kerosene also has a density of 775 to 840 g/cm 3 .
  • the freezing point should generally be in the range of -40 °C to -60 °C, depending on the specification.
  • Steps I), III) and IV) contained in the following description are catalytic, i.e. H. using a catalyst, even if the term “catalytic” is omitted below.
  • Range information for hydrocarbon contents of fractions are, of course, always idealized information, i.e. H. the majority of hydrocarbons fall within this range.
  • Information for chain lengths such as C8- or C9+ means hydrocarbons with a chain length of ⁇ C8 or ⁇ C9.
  • Oxygenates (a)) are used as the feedstock for the olefin production unit (I).
  • Oxygenates are hydrocarbon compounds containing oxygen, such as: B. alcohols and ethers.
  • the number of carbon atoms is preferably C1 to C4. This particularly includes methanol (MeOH) and dimethyl ether (DME), especially MeOH.
  • the starting material contains different proportions of water, particularly preferably 0 to 40% by weight.
  • This feedstock is mixed with the cycle gas (c)) and fed to the olefin production unit (I).
  • the cycle gas preferably contains non-reactive components such as CH 4 , ethane, H 2 , CO and CO 2 as well as reactive components such as ethylene. Furthermore, unreacted oxygenates may be present.
  • the olefin production unit (I) contains at least one fixed bed reactor (an isothermal tubular reactor is particularly preferred) in which the oxygenates (a)) are converted into olefins in the presence of a catalyst. This reaction is exothermic. The main products of the reaction are propene and butenes. In addition, a certain proportion of ethylene and olefins C5+ are formed. Paraffins, naphthenes and aromatics are formed as by-products. The olefin-containing mixture (b) is obtained as a mixture of these components and the non-reactive components.
  • the temperature during olefin production in step I) is preferably 200 ° C to 600 ° C, particularly preferably 300 ° C to 500 ° C, in particular 450 ° C to 500 ° C.
  • the pressure is preferably 1 bar to 6 bar, particularly preferably 2 bar to 5 bar, in particular 2 bar to 4.5 bar.
  • the mass-related space velocity (oxygenate load) of the catalyst used is preferably 1 to 10 kg of oxygenate per h and kg of catalyst (1/h).
  • the olefin-containing mixture of substances is obtained as a product of the olefin production unit (I). This mixture of substances is fed to the absorption unit (II). In the absorption unit (II), the absorbent and the olefin-containing mixture (b)) are contacted and, by means of absorption, predominantly C3+ hydrocarbons are absorbed from the olefin-containing mixture into the absorbent. The remaining cycle gas (c)) is enriched with ethylene and is returned to the olefin production unit (I).
  • the absorbent is, as in the Figures 1.1ac and 1.2 shown, taken from the separation units (Va) and/or (Vb). This is generally a mixture of hydrocarbons.
  • FIG. 1.3 Specific information on the characteristics of the absorbent is given in the description of the separation units (Va) and (Vb).
  • a scheme of the absorption unit (II) is given in Fig. 1.3 shown. This Fig. 1.3 is based on the Fig. 1.1b . However, the absorption unit (II) is in Fig. 1.3 Can also be implemented in all other figures.
  • the main equipment of the absorption unit (II) is usually an absorption column (II Abs.) and a desorption column (II Des.).
  • the absorption column In the absorption column, the unloaded absorbent (e)), coming from separation unit (Va) and/or (Vb), is contacted with the olefin-containing mixture (b)).
  • the olefin-containing mixture of substances is preferably present as a gas phase.
  • components of the olefin-containing mixture (b)) are dissolved in the absorbent. This process is promoted by increased pressures and low temperatures.
  • the temperature is usually 0 °C to 200 °C, particularly preferably 5 °C to 150 °C, particularly preferably 5 °C to 50 °C.
  • the pressure is preferably 1 bar to 50 bar, particularly preferably 3 bar to 30 bar, particularly preferably 5 bar to 15 bar.
  • Predominantly C3+ hydrocarbons from the olefin-containing mixture (b)) pass into the absorbent.
  • olefins especially C3 olefins and C4 olefins
  • the absorbent (s) loaded in this way then leaves the absorption column at the bottom and is led to the desorption column.
  • the remaining cycle gas (c)) leaves the system and is returned to step I) for olefin production.
  • the loaded absorbent (s)) becomes an olefin-containing fraction (d)), which mainly comprises C3/C4 olefins and hydrocarbons C5+ (preferably up to approx. C8).
  • the desorption of these components is promoted by a reduction in pressure and an increase in temperature.
  • the temperature is preferably 0 °C to 400 °C, particularly preferably 20 °C to 300 °C, particularly preferably 30 °C to 200 °C.
  • the pressure is preferably 1 bar to 50 bar, particularly preferably 1 bar to 30 bar, particularly preferably 1 bar to 10 bar.
  • the olefin-containing fraction (d)) leaves the desorption column and is fed to the oligomerization unit (III).
  • a first partial stream is mixed with fresh, unloaded absorbent (e)) within the absorption unit and returned to the absorption column.
  • the second partial stream leaves the absorption unit (II).
  • This partially loaded absorbent (f)) is either converted directly into the oligomerization unit (III) (see Fig. 1.1ac ) or to the separation unit (Va) (see Fig. 1.2 ) guided. In the latter case, the still absorbed species are separated from the absorbent and passed from the separation unit (Va) to the oligomerization unit (III) via the recycling route of C3 to C8 hydrocarbons (m)).
  • olefins are supplied to the oligomerization unit (III).
  • the olefins contained in the absorbent are preferably catalytically oligomerized to form an oligomer.
  • the oligomerization unit (III) contains at least one fixed-bed reactor in which the olefins are converted into long-chain hydrocarbons in the presence of a catalyst.
  • the main reaction that occurs is called oligomerization.
  • the C3 to C5 olefins can be viewed as monomers from which dimers, trimers, tetramers and higher oligomers form through reaction. These reactions are exothermic.
  • the products of the oligomerization reaction are also olefins; side reactions lead to the formation of other hydrocarbons.
  • This product mixture is referred to as an oligomer.
  • the product of the oligomerization unit (III) is passed to the separation unit (Va), then this is referred to as crude oligomer (g)).
  • the crude oligomer (g)) obtained in this way shows a C number distribution of approximately C3 to C25.
  • the light C3 to C5 hydrocarbons are primarily composed of unreacted olefins.
  • the majority of the crude oligomer (g)) consists of C5 to C20 hydrocarbons. It is preferably a mixture of paraffins, olefins, naphthenes and aromatics, in particular they are mostly olefins.
  • the temperature of the oligomerization unit (III) is usually in a range from 50 °C to 500 °C, particularly preferably from 100 °C to 400 °C, in particular from 200 °C to 300 °C.
  • the pressure is preferably in a range from 1 bar to 80 bar, particularly preferably in a range from 10 bar to 70 bar, in particular in a range from 20 bar to 50 bar.
  • the olefin content in the feed mixture for oligomerization III) is preferably adjusted so that the olefins have a proportion of 10 to 90% by mass.
  • a separation unit (Va) is arranged between (III) and (IV).
  • a crude oligomer (g)) is fed to the separation unit (Va).
  • This separation unit takes place, for example, in one embodiment Fig. 1.2 , a separation of C3 to C8 hydrocarbons (m)), which are returned to the oligomerization unit (III).
  • the unloaded absorbent (e)) is obtained.
  • This is a mixture of mainly C5 to C8 hydrocarbons (in particular mainly C5 to C8 olefins). It is characterized by a boiling curve in which most components boil in the range from 30 °C to 130 °C.
  • This fraction is fed to the absorption unit (II) described above as fresh, unloaded absorbent (e)).
  • a heavy oligomer (h)) is obtained in this separation unit from the main product of the oligomerization III). This consists mainly of a mixture of hydrocarbons C9+.
  • the hydrogenation unit (IV) contains at least one fixed bed reactor in which the olefins and aromatics contained in the oligomer, in the preferred embodiment of the separation unit (Va) the heavy oligomer (h)), are converted into paraffins and naphthenes in the presence of a catalyst. These reactions are exothermic.
  • the hydrogenation temperature in step IV) is preferably 20°C to 400°C, particularly preferably 60°C to 300°C, in particular 90°C to 200°C.
  • the pressure is preferably 20 bar to 150 bar, particularly preferably 30 bar to 100 bar, in particular 40 bar to 80 bar.
  • the ratio of supplied hydrogen to hydrocarbons, also referred to as the gas/oil ratio is preferably in a range of 100 to 1000 Nm 3 /m 3 and particularly preferably in a range of 200 to 600 Nm 3 /m 3 .
  • the catalytic hydrogenation in step IV) results in a hydrogenated product which consists mainly of C9+ (C9 to C25) paraffins and C9+ naphthenes.
  • lighter components (C8-) can come from the separation unit (Va) or arise as a result of side reactions (e.g. cracking) during the hydrogenation IV).
  • This hydrogenated product (i)) with even lighter proportions must be processed again and, in a preferred embodiment, is fed to the separation unit (Vb) for this purpose.
  • the hydrogenated product (i)) is divided into the fractions gasoline components (j)), kerosene (k)) and diesel components (l)) in a separation unit (Vb).
  • a low-aromatic stream with gasoline components and diesel components are obtained as by-products.
  • Gasoline components (C5 to C8 hydrocarbons) are characterized by a boiling curve in which the components preferably boil in the range from 30 °C to 130 °C. These components can be further processed into gasoline.
  • Kerosene (C9 to C17 hydrocarbons), as already defined above, is characterized by a boiling curve in which the components preferably boil in the range from 130 ° C to 300 ° C.
  • Diesel components (C18 to C25 hydrocarbons) are characterized by a boiling curve in which the components preferably boil in the range of 300 °C to 400 °C. These components can be further processed into diesel.
  • part of the kerosene (k)) and/or the diesel components (l)) can be used as an absorbent (see Fig. 1.1a , 1.1b and 1.2 ).
  • the absorbent is separated after step IV) in a separation unit (Vb).
  • the absorbent that is separated from the separation unit (Vb) is a hydrogenated fraction of C9+ paraffins and C9+ naphthenes (C9 to C25). It is characterized by a boiling curve in which the components preferably boil in the range from 130 °C to 400 °C, particularly preferably at 200 °C to 400 °C, particularly preferably at 250 °C to 400 °C.
  • this mixture is fed to the absorption unit (II) described above as fresh, unloaded absorbent (e)).
  • an absorbent is thus separated after step III) in the separation unit (Va) and/or after step IV) in the separation unit (Vb), with which the absorption according to the invention is carried out in step II).
  • the absorbent consists mainly of C5 to C8 hydrocarbons.
  • the absorbent is separated off, then the absorbent results as the hydrogenated fraction of the C9+ paraffins and C9+ naphthenes. If separation takes place in both separation units (Va) and (Vb) (see Fig. 1.1a ), the absorbent is a mixture of the C5 to C8 hydrocarbons and the hydrogenated C9+ fraction.
  • the oligomerization unit (III) has a line for the removal of an oligomer to the separation unit (Va) and the separation unit (Va) has a line for the removal of the oligomer to the hydrogenation unit (IV), and / or the hydrogenation unit (IV) has a line for the removal of a hydrogenated product to the separation unit (Vb) and the separation unit (Vb) has at least one line for the removal of the kerosene, the gasoline components and diesel components.
  • Reactors or reactor systems are required for the steps of olefin production I), oligomerization III) and hydrogenation IV).
  • a fixed bed reactor is preferably used, particularly preferably an isothermal tubular reactor with a fixed bed.
  • at least one fixed bed reactor is preferably used, particularly preferably at least one adiabatic fixed bed reactor.
  • at least one fixed bed reactor is preferably used, particularly preferably at least one adiabatic fixed bed reactor.
  • the catalytic olefin production in step I), the catalytic oligomerization in step III) and the catalytic hydrogenation in step IV) are each carried out in at least one fixed bed reactor.
  • the reactors are understood to be part of the respective plant part of the plant according to the invention.
  • a zeolitic catalyst is particularly preferably selected from zeolites of the type ZSM-5 (such as H-ZSM-5), ZSM-57, ZSM-22, SAPO, Beta and Y.
  • a non-zeolitic acid catalyst is particularly preferably selected from a solid immobilized acid, such as solid phosphoric acid, and silicate-based catalysts, such as amorphous aluminosilicates or clays.
  • a stream of the partially loaded absorbent (f)) is passed from the absorption unit (II) to the separation unit (Va) (see Fig. 1.2 ), this results in a hydrocarbon stream rich in C3 to C8 olefins, which leads to the oligomerization unit (III) is returned.
  • a stream of the partially loaded absorbent (f)) is passed from the absorption unit (II) to the separation unit (Va) (see Fig. 1.2 ), this results in a hydrocarbon stream rich in C3 to C8 olefins, which leads to the oligomerization unit (III) is returned.
  • an increase in the kerosene yield is achieved analogously to the manner already described above.
  • the C9+ hydrocarbons originally from the absorbent reach the main product of the oligomerization III) (heavy oligomerisate (h)) via the separation unit (Va).
  • kerosene yield 40 to 70% by mass is achieved. This yield is based on the stoichiometric hydrocarbon content of the oxygenate. When focusing on the co-products, lower kerosene yields can also be achieved.
  • Another advantage of the invention is that the kerosene obtained has few aromatic compounds.
  • the proportion of aromatics is 0% by volume to 20% by volume, in particular it is less than 10% by volume, particularly preferably from 0% by volume to 8% by volume. This is achieved by the above-described combination of high olefin yields in step I), minimization of side reactions in step III) and ultimately by the hydrogenation in step IV).
  • the low aromatic content advantageously reduces the formation of particles during combustion of the kerosene.
  • a fraction of low boilers (o)) is additionally separated off and returned to step I).
  • the system has a line for the return of low boilers from the oligomerization unit (III) and/or the separation unit (Va) to the olefin production unit (I).
  • ethylene is advantageously converted into C3+ olefins.
  • step I There is a dilution of the feed stream for olefin production in step I), which serves to better control the heat transfer in the reactor for olefin production I).
  • This and the conversion of the recycled ethylene into C3+ olefins and their subsequent conversion in oligomerization III) increase the yield of kerosene.
  • a C3 to C5 hydrocarbon fraction is additionally separated off in a separation unit (Va) and returned to step III).
  • a C5 to C8 hydrocarbon fraction is additionally separated off in a separation unit (Va) and returned to step III).
  • the olefins are converted to higher olefins in the C number range of the kerosene and this leads to an increase in the yield of the kerosene.
  • an olefin-containing substance mixture (b)) is brought into contact with the absorbent from step V) in the absorption unit (II). It cannot be avoided that a certain proportion of light hydrocarbons C1 to C2 as well as permanent gases (CO, CO 2 , H 2 ) are absorbed in the absorbent and enter the oligomerization unit (III). Under the reaction conditions of oligomerization, these light components can be considered non-reactive or, as in the case of ethylene, unreactive. These light components are therefore also contained in the reaction product of the oligomerization.
  • the oligomerization unit (III) contains a separation stage (III/S1), which in Fig. 2.1a and Fig. 2.1b is shown
  • the crude oligomer 2 (q)) present after the oligomerization reaction is separated in this stage (III/S1), for example by separation, into a fraction of low boilers (o)) and a remaining fraction (crude oligomer (g))).
  • the low boilers (o)) consist mainly of C1 to C2 hydrocarbons, as well as non-reactive components such as CO, CO 2 and H 2 .
  • the crude oligomer (g)) is fed to the separation unit (Va).
  • the separation unit (Va) includes two stages (Va/K1) and (Va/K2). This is exemplified in Fig. 2.1a and Fig. 2.1b shown.
  • the crude oligomer (g)) is in the first stage (Va/K1) by means of a distillation column into a fraction of low boilers (o)), consisting mainly of C1 to C2 hydrocarbons, an olefin-rich C3 to C5 hydrocarbon fraction (p)) and a remaining heavy fraction (crude oligomer 3 (r))) separated.
  • the low boilers (o)) from (III/S1) and (Va/K1) are returned to olefin production I).
  • the olefin-rich C3 to C5 hydrocarbon fraction (p)) is recycled to oligomerization III).
  • This C3 to C5 hydrocarbon fraction is preferably 10% by mass to 99% by mass, particularly preferably 20% by mass to 99% by mass, particularly preferably 40% by mass to 99% by mass. returned.
  • the remaining heavy fraction (crude oligomer 3 (r))) is converted into an olefin-rich C5 to C8 hydrocarbon fraction (v)) in a second stage (Va/K2), for example by means of a distillation column Fig. 2.1b and the heavy oligomer (h)) were separated.
  • the heavy oligomer (h)) is then passed to hydrogenation IV).
  • the olefin-rich C5 to C8 hydrocarbon fraction (v)) is recycled for oligomerization III) (see Fig. 2.1b ).
  • This C5 to C8 hydrocarbon fraction (v)) is preferably recycled at 10% by mass to 99% by mass, particularly preferably at 20% by mass to 99% by mass, particularly preferably at 40% by mass to 99% by mass .
  • the same fraction is used as unloaded absorbent (e)) in the absorption unit (II) (see Fig. 2.1a ).
  • a portion is returned to the oligomerization unit (III) as a C5 to C8 hydrocarbon fraction (v)) and to the absorption unit (II) as an unloaded absorbent (e)).
  • the low boilers (o)), the C3 to C5 hydrocarbon fraction (p)) and the C5 to C8 hydrocarbon fraction (v)) or the unloaded absorbent (e)) can also be obtained using fewer steps. In general, it is possible to remove some of the fractions from the process after separation.
  • ethylene is advantageously converted into C3+ olefins.
  • the "non-reactive" components of the permanent gases such as CO 2 and CO contained in the low boilers (o)) advantageously achieve a dilution of the reactive feed in olefin production I). This dilution serves to better control the heat transfer in the olefin production reactor I). This and the conversion of the recycled ethylene into C3+ olefins and their subsequent conversion in oligomerization III) increase the yield of kerosene.
  • C3 to C8 olefins are converted to higher olefins in the C number range of kerosene and thus the yield of the kerosene is increased Kerosene.
  • the hydrogenation unit (IV) contains a separation stage (IV/S1). This is exemplified in Fig. 2.1a shown.
  • the hydrogenated product 2 (s)) present after the reactor is separated in this stage (IV/S1), for example by separation, into a gaseous hydrogen-rich fraction (t)) and a heavy fraction (hydrogenated product (i))).
  • the hydrogen-rich fraction (t)) is returned to the hydrogenation unit (IV).
  • This fraction can also be removed from the process or added with fresh hydrogen (u)) before recycling.
  • the recycling advantageously reduces the need for hydrogen and improves the economic efficiency of the process.
  • the heavier fraction (hydrogenated product (i)))) remaining after the separation (IV/S1) is fed to the separation unit (Vb).
  • a separation takes place into a fraction of gasoline components (j)), kerosene (k)) and diesel components (l)), with the synthetic kerosene typically accounting for the majority of the product.
  • a light fraction of C3/C4 hydrocarbons can also be separated off in this column.
  • the separation unit (Vb) it is also possible for the separation unit (Vb) to carry out the complex separation task in a two-stage distillation.
  • Oligomerization is a highly exothermic reaction.
  • the reaction temperature For the synthesis of kerosene hydrocarbons, it is necessary to limit the reaction temperature to a certain range. For this, the reaction mixture must be cooled.
  • the reaction mixture is mixed directly with a non-reactive liquid fluid.
  • a non-reactive liquid fluid for this purpose, an absorbent consisting of C9+ paraffins and C9+ naphthenes is used.
  • the high enthalpy of vaporization of these components allows for better control over the temperature within a catalyst bed. If direct cooling is not used within the bed If the measures taken for this alone are not sufficient, then the temperature increase caused by the exothermicity of the reaction must be limited by adjusting the progress of the reaction. This is achieved by dividing the catalyst bed.
  • the reaction mixture leaves the catalyst bed at an elevated temperature and is cooled before entering another catalyst bed. This so-called intermediate cooling can be carried out using a corresponding heat exchanger and/or by directly injecting a cooler fluid into the reaction mixture.
  • the oligomerization unit (III) is divided into several stages. This is exemplified in Fig. 2.2a and in Fig. 2.2b shown.
  • a stage represents either a catalyst bed within a reactor or a reactor with one or more catalyst beds.
  • the stages (III.1) and (III.2) were shown for illustration purposes, but further stages are also possible.
  • the Figures 2.2a and 2.2b contain all possible feeds from II) and returns from the separation unit (Va), although not all feeds and returns have to be combined when using this embodiment.
  • the olefin-containing fraction (d)) coming from the absorption unit (II) is divided into the corresponding stages. This is also possible for the partially loaded absorbent (f)).
  • the returns of the C3 to C5 hydrocarbon fraction (p)) and the C5 to C8 hydrocarbon fraction (v)) from the separation unit (Va) are also divided into the different stages.
  • the corresponding streams (d)), (f)), (p)) and (v)) can each be distributed to the stages in different proportions. These streams are generally at a lower temperature level than that of the oligomerization reaction. (please refer Fig. 2.2a )
  • the intermediate product of the oligomerization (w)) is cooled by partial streams from (d)), (f)), (p)) and (v)).
  • control over the reaction temperature is advantageously achieved. This ensures high selectivities and high kerosene yields.
  • olefin-rich fractions of C5 to C8 hydrocarbons are advantageously mixed with C3 olefins and C4 olefins before their implementation in the oligomerization. This results in an increase in kerosene yield.
  • the returns of the C3 to C5 hydrocarbon fraction (p)) and the C5 to C8 hydrocarbon fraction (v)) contain a certain proportion of paraffins, which were mainly originally created as by-products in olefin production I). This Paraffins are not reactive in oligomerization III).
  • a dilution of the reaction mixture is achieved at the entrance to a stage. This advantageously achieves additional control over the reaction temperature and consequently ensures high kerosene yields.
  • the intermediate product of the oligomerization (w)) is passed after stage (III.1) to a separation stage (e.g. by means of separation) (III/S2) and subsequently to the separation unit (Va).
  • the intermediate product of the oligomerization (w)) is separated into a fraction of low boilers (o)) and a remaining heavy fraction (intermediate product oligomerization 2 (x))).
  • the latter is fed to the separation unit (Va).
  • the C3 to C8 olefins are transferred from the separation unit (Va) to a subsequent stage (III.2) of the oligomerization III) or to the first Stage (III.1) returned.
  • a partial conversion of light olefins with intermediate separation of the C9+ fraction is achieved via one or more successive stages of oligomerization, as in Fig. 2.2b will be shown.
  • a proportion of components in the kerosene range is advantageously secured.
  • Each additional stage with partial sales increases the overall level of sales.
  • a high yield of kerosene is advantageously achieved with a high selectivity in the oligomerization III) with an increase in the overall degree of conversion.
  • the oligomer from step III) is diluted upon entering step IV) by recycling a portion of the hydrogenated product to step IV).
  • the oligomer preferably comes from the separation unit (Va) and the hydrogenated product for dilution is taken between (IV/S1) and (Vb) (see Fig. 2.3 ).
  • step IV) is carried out in several stages, wherein hydrogen and/or a portion of a heavy oligomer and/or the product hydrogenated in IV) is used for intermediate cooling (between the hydrogenation stages).
  • the dilution or intermediate cooling advantageously leads to better control over the reaction temperature. This also ensures that the catalyst has a long service life.
  • Hydrogenation is also a highly exothermic reaction. At elevated temperatures, cracking can occur as a side reaction and the catalyst can deactivate more quickly.
  • the heavy oligomer (h) originating from the separation unit (Va) serves as the feed mixture. It is a mixture of C9+ hydrocarbons, which consists mainly of olefins. A certain proportion of aromatics is also included.
  • the reaction temperature In order to achieve almost complete hydrogenation of the product, it is necessary to set the reaction temperature within a certain range. The temperature increase must be limited by adjusting the progress of the reaction. This is done analogously to oligomerization by cooling the reaction mixture and dividing the catalyst bed.
  • the hydrogenation unit (IV) is therefore divided into several stages. This is exemplified in Fig. 2.3 shown. Here a stage represents either a Catalyst bed within a reactor or a reactor with one or more catalyst beds.
  • the corresponding intermediate product of the hydrogenation (y)) is passed from one stage to the subsequent stage.
  • part of the hydrogenated product (i)) is also removed between (IV/S1) and (Vb) and returned to the hydrogenation unit (IV).
  • the hydrogenated product (i)) consists mainly of C9+ paraffins and C9+ naphthenes. These components do not behave reactively within the hydrogenation IV). This fraction of the hydrogenated product (i)) is preferably recycled to 10 to 90% by weight.
  • the heavy oligomer (h)) coming from the separation unit (Va) is, like the returns of the hydrogenated product (i)) and the hydrogen-rich fraction (t)), including the supply of fresh hydrogen (u)), divided into the various stages .
  • the corresponding streams (h)), (i)) and (t)) can each be distributed to the stages in different proportions. These streams are generally at a lower temperature level than the hydrogenation reaction.
  • the hydrogenation intermediate product (y)) is cooled at the entrance to stage (IV.2).
  • Contacting the heavy oligomer (h)) with a partial stream of the hydrogenated product (i)) leads to a dilution of the reactive feed and enables the high enthalpy of vaporization of the C9+ paraffins and C9+ naphthenes to be used for cooling within the bed.
  • control over the reaction temperature and the suppression of side reactions are advantageously achieved.
  • the hydrogenation is optimized accordingly and a high kerosene yield is ensured.
  • By controlling the reaction temperature a long service life of the catalyst is also achieved.
  • water is separated from the olefin-containing mixture of substances and returned to the catalytic olefin production in step I).
  • This advantageously leads to a dilution of the feed from step I), which serves to better control the heat transfer and thus better control over the temperature, with the advantages already mentioned above, among others. also high olefin selectivities and yields, and ultimately high kerosene yields.
  • the return of the water also advantageously slows down the formation of coke.
  • the formation of coke on the catalyst blocks active centers and thereby reduces the activity of the catalyst. Consequently, by recycling the water, a long cycle time and service life of the catalyst is achieved.
  • a fraction of C8+ hydrocarbons is separated from the olefin-containing mixture of substances and passed to step IV) and/or step V).
  • the oligomerization is protected from the coke formation potential emanating from these components . Due to the lower deactivation, long cycle times and a long service life of the catalyst are achieved in the oligomerization. Furthermore, a high kerosene yield is advantageously achieved.
  • the product of olefin production I) is an olefin-containing mixture of substances. Olefins contained therein are removed using absorption II). During the contact of the olefin-containing substance mixture (b)) with the liquid absorbent, predominantly hydrocarbons C3+ are dissolved in the absorbent. As noted above, this process is facilitated by elevated pressures and low temperatures. Consequently, the olefin-containing mixture of substances must be cooled and/or its pressure level raised before entering the absorption unit (II). In both cases this is associated with the formation of a liquid product portion.
  • the product range from olefin production I) contains two separation stages (I/S1) and (I/S2). This is in Fig. 2.4 shown. Here, the reaction product of olefin production (olefin-containing mixture 2 (z)) is separated into a gaseous phase, a liquid organic phase and a liquid-aqueous phase using a three-phase separator (I/S1).
  • the liquid-aqueous phase consists mainly of water (aa)). Part of this water (aa)) is added to the feedstock for olefin production I). This is advantageous in the Olefin production I) achieves a dilution of the reactive feed. This dilution serves to better control the heat transfer and thus better control over the temperature in the olefin production reactor I). This advantageously ensures high olefin selectivities and yields, and ultimately high kerosene yields. By recycling the water (aa)), coke formation is also advantageously slowed down and thus a long cycle time and service life of the catalyst is achieved.
  • the liquid organic phase contains C8+ hydrocarbons (ab)), which are already formed to a small extent during the olefin production reaction.
  • the liquid organic phase (ab)) is fed to the separation unit (Va) and/or the hydrogenation unit (IV).
  • Light hydrocarbons, in particular light olefins, which have dissolved in the liquid organic phase during phase deposition, are separated off in the separation unit (Va) and fed to the oligomerization III) via the corresponding returns.
  • the higher hydrocarbons of the liquid organic phase advantageously enter the hydrogenation IV) together with the heavy oligomer (h)).
  • the hydrogenation of the higher hydrocarbons contained in the liquid organic phase advantageously increases the kerosene yield.
  • the C8+ hydrocarbons of the liquid organic phase (ab)) consist largely of aromatics.
  • hydrocarbons When hydrocarbons are reacted over acidic catalysts, such components act as precursors to coke formation. By removing these components for hydrogenation IV), the oligomerization III) is protected from the coke formation potential emanating from these components. Due to the lower deactivation, long cycle times and a long service life of the catalyst can be achieved in oligomerization. By hydrogenating the aromatics, a kerosene low in aromatics is advantageously obtained.
  • the gaseous phase (olefin-containing mixture (b))) is fed to the absorption unit (II).
  • the liquid organic phase (olefin-containing mixture 4 (ad))) consists mainly of C3+ olefins and is fed to the oligomerization unit (III) and/or the separation unit (Va).
  • the olefins contained in the olefin-containing mixture 4 (ad)) are passed from the separation unit (Va) to the oligomerization III) by means of the corresponding returns of the C3 to C8 hydrocarbons (m)).
  • the yield of kerosene is advantageously increased.
  • the conversion of methanol into olefins was carried out in a pilot plant.
  • the reactor of this plant consisted of a tube.
  • the reaction tube was filled with H-ZSM-5 catalyst.
  • the reaction temperature within the catalyst bed was monitored using thermocouples. With this reaction tube, an almost isothermal operation of the reaction was achieved.
  • the feed mixture contained methanol/water/ethylene in the proportions 10 to 30% by volume / 0 to 30% by volume / 0 to 25% by volume.
  • the remaining portions of the feed mixture consisted of non-reactive gaseous components.
  • the liquid mixture of methanol and water was conveyed to the reactor using a pump. The flow was controlled using a flow meter. Ethylene and the non-reactive gaseous components were dosed using a mass flow controller. The mixture of these gases was added to the evaporator along with the liquid feed. The feed mixture produced in this way was then fed to the reactor. After the reactor, the product stream was cooled using ice water. This caused part of the product to condense. The liquid reaction product was separated from the gaseous reaction product in a separator.
  • the amount of condensate was recorded and the corresponding mass flow was determined based on the duration of the experiment.
  • the mass flow of the gaseous product resulted from the difference in the mass flows of the feed mixture and condensate.
  • the compositions of gaseous and liquid product fractions were analyzed by gas chromatography. On this basis, the degree of conversion of methanol and the yield of olefins were calculated.
  • Table 1 Composition of the hydrocarbon product: experiment C1 to C4 paraffins (mass%) Ethylene (mass%) Propene (% by mass) Butenes (mass%) C5+ olefins (mass%) Remaining C5+ hydrocarbons (mass%) Total olefins (mass%) 1 10.5 13.0 33.3 23.7 13.3 6.2 83.3
  • Table 2 Standardized mass proportions of methanol, ethylene and C3+ olefins: experiment Feed methanol Feed ethylene Product C3+Olefins ⁇ Product C3+ Olefins (%) 1 1 - 0.30 2 0.65 0.35 0.37 + 23
  • a tube was used as the reactor, into which an H-ZSM-5 catalyst was filled.
  • the reaction temperature was recorded using thermocouples.
  • a C3/C4 olefin mixture was used as feed because these components are the main reactive products of the olefin production unit (I) upstream in the overall process.
  • the respective proportion of C3 olefins/C4 olefins metered was 5 to 40% by volume/5 to 40% by volume.
  • the reaction gas consisted of a proportion of non-reactive components.
  • the product mixture was cooled, relaxed and separated into a gaseous and a liquid phase using a separator.
  • the product fractions obtained were examined by gas chromatography.
  • the conditions for experiment 2 were as follows: temperature: 250 °C; Pressure: 50 bar; Proportion of non-reactive components: 67% by volume and mass-related space velocity: 0.5 kg olefins per h and kg of catalyst.
  • Table 3 Sales, composition of the hydrocarbon mixture and product selectivities for selected examples: Exp. Sales of C3 olefins (mass%) Sales of C4 olefins (mass%) Composition (% by mass) Product selectivity (mass%) Intermediate product (C2 to C8) product Kerosene (C9 to C17) Diesel components C18+) Kerosene (C9 to C17) Diesel components (C18+) number 1 16 13 95.4 4.6 0.01 99.8 0.2 No. 2 94 80 35.8 42.5 21.7 66.2 33.8
  • the amount of intermediate product (C2 to C8) remaining as a result of a partial conversion can be converted into the target product kerosene with high selectivity in the overall process through appropriate further reaction stages and recycling of the C3 to C8 hydrocarbon fraction. This is based on the preferred embodiments Fig. 2.2a and Fig. 2.2b and the associated ones Figures 2.2a and 2.2b shown.
  • the hydrogenation tests IV were carried out on a test facility.
  • a tube was used as a reactor, into which a metal catalyst with nickel was filled.
  • Thermocouples are installed throughout the reactor to control the reaction temperature.
  • the liquid feed was metered into the reactor using pumps.
  • the hydrogen was introduced into the reactor using a mass flow controller.
  • a mixture of olefins including dienes in the range from C8 to C24 was fed to the reactor together with a proportion of approximately 9% by weight of aromatics in the range from C7 to C9.
  • This reaction feed was diluted with a liquid non-reactive hydrocarbon mixture.
  • the product was cooled and separated into a gaseous and a liquid phase using a separator. Samples of both phases were analyzed by gas chromatography and wet chemistry.
  • a temperature of 150 °C and a pressure of 60 bar were set.
  • the hydrogen/hydrocarbon ratio (gas/oil ratio) was set to 400 Nm 3 /m 3 .
  • the system is based schematically on the Fig. 2.5 .
  • This figure has the origin Fig. 1.1c and the system contains features from the Fig. 2.1.a , 2.1b and 2.4.
  • the plant therefore contains the corresponding plant parts of the olefin production (I), the absorption unit (II), the Oligomerization unit (III), the hydrogenation unit (IV) and the separation unit (V).
  • This system therefore represents a design of the overall process.
  • the feed for olefin production I) consists of the oxygenate methanol (a)) (100% methanol), of which 300 kg/h were used in the present example, a dilution with water (aa)) from the separator (I/S1) and the recycled cycle gas (c)) from step II).
  • Olefin production I takes place using the methods described.
  • the temperature is in the range of 450 to 500 °C and the pressure is in the range of 3 to 4.5 bar.
  • the reaction product of the olefin production is cooled, whereby liquid is condensed, which is predominantly water.
  • This liquid-aqueous phase is separated off in the separator (I/S1) and a portion is returned to the olefin production reaction (aa)).
  • This is followed by compression stages, each with downstream liquid separation (separation stage (I/S2)) for the necessary pressure increase for the absorption unit (II).
  • a liquid aqueous phase is removed from the process and an organic liquid phase (olefin-containing substance mixture 4 (ad))) is collected and this is fed to the oligomerization (III).
  • the gas phase now obtained is fed to the absorption unit (II) as an olefin-containing mixture (b)).
  • the C5 to C8 hydrocarbon fraction (e)) is used as an absorbent.
  • the cycle gas (c)) is obtained during absorption.
  • This cycle gas (c)) contains the components not absorbed by the absorbent, such as. Ethylene. These are returned to the olefin production unit (I).
  • an olefin-containing fraction (d)) enriched predominantly with C3 to C4 hydrocarbons and a part of the absorbent (f)) are transferred to step III).
  • All of the described streams (ad), f) and d)) are fed to the oligomerization III).
  • the oligomerization takes place in a temperature range of 200 to 300 °C and in a pressure range of 10 to 70 bar.
  • the product of oligomerization III), the so-called crude oligomer (g)), is separated in the separation unit (Va).
  • the C5 to C8 hydrocarbon fraction is separated off as unloaded absorbent (e)) and returned to the absorption unit (II). Overall, non-condensable components resulting from the separations are removed from the process. In addition, the C3 to C8 hydrocarbons (m)) are recycled. Thus, part of the C5 to C8 hydrocarbon fraction (v)) becomes Oligomerization III) and used as an absorbent in order to be able to use as many olefins as possible in the oligomerization III).
  • the heavy oligomer (h)) is fed to hydrogenation IV). To the hydrogenation IV), excess hydrogen (u)) is added and the excess hydrogen is released.
  • the hydrogenation takes place in a temperature range of 90 to 200°C and in a pressure range of 40 to 80 bar.
  • Table 4 Overview table of the streams included according to Fig. 2.5: No Physical parameters of material flows Composition of the streams (Ma.-%) T (°C) p (bar) Mass flow (kg/h) MeOH H2O Permanent gases C2 to C4 C5 to C8 C9 to C17 C18+ a) 20.0 4.4 300.0 100 b) 10.0 10.5 298.6 0.3 48.1 38.1 13.4 c) 11.0 4.4 183.3 77.3 10.9 11.7 0.1 d) 78.0 3.0 205.4 0.1 1.6 46.7 50.9 0.7 e) 85.4 13.0 100.0 0.1 2 94.4 3.5 f) 40.0 21.5 10.5 0.1 0.4 75.5 22.6 1.4 G) 31.7 15.8 544.0 0.2 28.1 57 12.1 2.6 H) 127.5 63.5 80.8 5.5 77.3 17.2 i) 50.1 66.0 81.1 0.1 1.3 8.6 73.3 16.7 j) 115.0 2.0 7.9 0.6 12.7 60.5 2

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EP22162046.1A 2022-03-15 2022-03-15 Procédé et installation de production d'un kérosène synthétique à partir des composés oxygénés Pending EP4245828A1 (fr)

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WO2006076942A1 (fr) 2005-01-22 2006-07-27 Lurgi Ag Procede de production de combustibles synthetiques a base d'oxygenats
EP2123736A1 (fr) * 2008-05-19 2009-11-25 C.E.-Technology Limited Procédé de fabrication de carburants diesel et carburants d'avion à partir d'alcools C1-C5
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WO2018045397A1 (fr) 2016-09-01 2018-03-08 The Petroleum Oil & Gas Corporation Of South Africa (Pty) Ltd Procédé de production d'un carburéacteur différent obtenu par synthèse - kérosène paraffinique synthétique (spk)
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US4506106A (en) 1984-01-04 1985-03-19 Mobil Oil Corporation Multistage process for converting oxygenates to distillate hydrocarbons with interstage ethene recovery
EP0190816A1 (fr) * 1985-01-17 1986-08-13 Mobil Oil Corporation Procédé de conversion de produits oxygénés en hydrocarbures liquides
WO2006076942A1 (fr) 2005-01-22 2006-07-27 Lurgi Ag Procede de production de combustibles synthetiques a base d'oxygenats
EP2123736A1 (fr) * 2008-05-19 2009-11-25 C.E.-Technology Limited Procédé de fabrication de carburants diesel et carburants d'avion à partir d'alcools C1-C5
EP2385093A1 (fr) * 2010-05-06 2011-11-09 IFP Energies nouvelles Procédé flexible de transformation de l'éthanol en distillats moyens mettant en oeuvre un système catalytique homogène et un système catalytique héterogène
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WO2018045397A1 (fr) 2016-09-01 2018-03-08 The Petroleum Oil & Gas Corporation Of South Africa (Pty) Ltd Procédé de production d'un carburéacteur différent obtenu par synthèse - kérosène paraffinique synthétique (spk)

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