WO2023174916A1

WO2023174916A1 - Process and plant for producing a synthetic kerosene from oxygenates

Info

Publication number: WO2023174916A1
Application number: PCT/EP2023/056432
Authority: WO
Inventors: Joachim Engelmann; Mario KUSCHEL; Stephan Schmidt; Norbert VÖLKEL; Petra Seidel; Felix HILDEBRANDT; Karolina WEGRZYN; Valerij TELTEVSKIJ; Sven Kureti; Stefan NOTTELMANN
Original assignee: Cac Engineering Gmbh
Priority date: 2022-03-15
Filing date: 2023-03-14
Publication date: 2023-09-21
Also published as: EP4245828A1; EP4245828B1; EP4245828A8

Abstract

The invention relates to a method for producing a synthetic kerosene, comprising the steps of: I) catalytic production of olefin by reacting oxygenates to form an olefin-containing mixture, II) separation of an olefin-containing fraction from the olefin-containing mixture by absorption using an absorption medium, wherein a remaining circulation gas is returned to step I) and during the absorption in step II) ethylene is returned with the circulation gas, the absorption medium from step II) subsequently being fed directly to an oligomerization in step III) and/or to a separating unit (Va), III) catalytic oligomerization of the olefin-containing fraction obtained in step II) to form an oligomerization product, IV) catalytic hydrogenation of the oligomerization product, V) separation of the absorption medium after step III) in separating unit (Va) and/or after step IV) in a separating unit (Vb). The invention also relates to a corresponding plant and to the use of the plant in the process.

Description

Process and system for producing a synthetic kerosene from oxygenates

The invention relates to a method and an associated system for producing synthetic kerosene from oxygenates, such as methanol.

The production of various, energy-rich hydrocarbon mixtures and fractions from oxygenates, such as methanol, via the steps of olefin production, oligomerization, i.e. H. Chain extension and hydrogenation are known in the art.

An alternative process that can be used to produce high-energy hydrocarbon mixtures is the Fischer-Tropsch process. The synthesis gas must first be produced; this can be done either via gas and coal or via renewable sources such as biomass. However, J. Eilers et al. (1990) that this process primarily produces disadvantageous linear chains over a wide C number range, including a large proportion of long-chain paraffins (waxes). In order to make these usable for synthetic fuels, special hydrocracking processes are necessary.

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. by Bellussi et al. (2012) described. 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 until C4 is olefins, washed out. The ethylene is then discharged as a chemical byproduct of the process. After absorption, the gasoline sorbent stream is led to oligomerization.

EP 0 190 816 A1 describes the production of distillate and gasoline starting from oxygenates.

EP 2 123 736 A1 also describes the production of diesel and aircraft fuels. Production takes place via the dehydration of C1-C5 alcohols.

EP 2 385 093 A1 relates to the production of middle distillate from ethanol. US 10 113 130 B1, on the other hand, describes the production of renewable diesel and renewable kerosene from various natural vegetable oils and ethanol.

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.

The task is solved using the procedure described below and the associated system.

The subject of the invention is a process for producing a synthetic kerosene with the steps:

I) Catalytic olefin production by converting oxygenates into an olefin-containing mixture of substances,

II) Separation of an olefin-containing fraction from the olefin-containing mixture of substances by absorption with an absorbent from step V), with a remaining cycle gas being recycled to step I). III) Catalytic oligomerization of the olefin-containing fraction obtained in step II) to form an oligomer,

IV) catalytic hydrogenation of the oligomer,

V) Separation of the absorbent after step III) in a separation unit (Va) and/or after step IV) in a separation unit (Vb), with ethylene being recycled with the cycle gas during the absorption in step II) (because it is during the absorption Ethylene is converted into the circulating gas), and the absorbent is then passed directly to the oligomerization in step III) and / or to the separation unit (Va) after step II).

The described features of the main claim are shown in Figures 1.1 a-c and Figure 1.2. Reference is made below to these figures and the streams within them.

In the sense of the invention, the above-mentioned process steps also require corresponding system parts and connections between the system parts.

The invention therefore also relates to a system for producing a synthetic kerosene from oxygenates, comprising an olefin production unit (I), an absorption unit (II), an oligomerization unit (III), a hydrogenation unit (IV) and at least one separation unit (Va and/or Vb ), where

• the olefin production unit (I) has a line for supplying the oxygenates,

• the olefin production unit (I) has a line for the removal of an olefin-containing mixture of substances to the absorption unit (II),

• the absorption unit (II) has a line for returning a circulating gas to the olefin production unit (I),

• the absorption unit (II) has a line for the removal of an olefin-containing fraction to the oligomerization unit (III),

• the absorption unit (II) has a line for an absorbent, which leads to the oligomerization unit (III) and/or to the separation unit (Va),

• the separation unit (Va and/or Vb) has a line for the removal of the absorbent to the absorption unit (II). This also includes the fact that further components can be integrated into the lines between these system parts or at other points in the process.

Finally, the subject of the invention is the use of the system according to the invention in the method according to the invention.

In accordance with the invention, in step I) olefins are produced in the olefin production unit (l), 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.

The order of steps I), II), III) and IV) and the order of the system parts (I), (II), (III) and (IV) are fixed. 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).

By definition in the sense of the invention, process steps and the associated system parts also require appropriate equipment. For example, 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. Overall, there is a general possibility that some of the streams specified in the process will be removed from the system for technical reasons. Accordingly, a coherent description of the process and system was carried out, since the explanations of the process inevitably result in parts of the system and vice versa.

As stated above, 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.

For the purposes of the invention, 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. In a preferred embodiment, 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.

All pressures are given as absolute pressures. The usual indication of bar(a) or bar(g) is therefore omitted in the following.

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.

Step I) or olefin production unit (I):

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 mass. 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, ethane, H2, CO and CO2 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. Typically, 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).

Step II) or absorption unit (II) of the invention:

According to the invention, the olefin-containing mixture of substances is obtained as the 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 removed from the separation units (Va) and/or (Vb), as shown by way of example in FIGS. 1.1 ac and 1.2. This is generally a mixture of hydrocarbons. Specific information on the characteristics of the absorbent is given in the description of the separation units (Va) and (Vb). A schematic of the absorption unit (II) is shown in Fig. 1.3. This Fig. 1.3 is based on Fig. 1.1b. However, the absorption unit (II) in Fig. 1.3 can also be implemented in all other figures.

Design of the absorption unit with absorption column and desorption column:

In a preferred embodiment, the main equipment of the absorption unit (II) is an absorption column (II Abs.) and a desorption column (II Des.). In the absorption column, the (unloaded) absorbent (e) according to the invention, coming from the 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. During contact with the liquid absorbent, components of the olefin-containing mixture (b)) are dissolved in the absorbent.

In such a preferred embodiment of the process according to the invention, in step II) the absorption is carried out using an absorption column and a desorption column and the absorbent, separated in a separation unit (Va) and/or (Vb), is contacted as an unloaded absorbent in the absorption column with the olefin-containing mixture , and then passed as a loaded absorbent from the absorption column to the desorption column, where the olefin-containing fraction is driven out of the absorbent. In most cases, a partially loaded absorbent remains. This partially loaded absorbent is either fed directly to the oligomerization unit (III) (see Fig. 1.1a-c) and/or to the separation unit (Va) (see Fig. 1.2).

This corresponds to a preferred embodiment of the system according to the invention, in which the absorption unit (II) comprises an absorption column and a desorption column, and wherein the line of the separation unit (Va) and / or (Vb) for the removal of the absorbent to the absorption unit (II) to this Absorption column runs and the line for the removal of an olefin-containing fraction to the oligomerization unit (III) runs from the desorption column to the oligomerization unit.

During the contact of the olefin-containing mixture of substances with the liquid absorbent, as already mentioned above, components of the olefin-containing mixture of substances are dissolved in the absorbent. This 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. As a result, olefins (especially C3 olefins and C4 olefins) accumulate in the absorbent.

The absorbent (s) loaded in this way then leaves the absorption column at the bottom and is led to the desorption column.

At the top of the absorption column, the remaining cycle gas (c)) leaves the system and is returned to step I) for olefin production.

In the desorption column, an olefin-containing fraction (d)) is expelled from the loaded absorbent (n)), which mainly comprises C3/C4 olefins and C5+ hydrocarbons (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 (in this embodiment with an absorption and desorption column) and is fed to the oligomerization unit (III). Depending on the operating conditions of the oligomerization, it may be necessary to adjust the pressure level. This can result in the separation of a liquid phase, which is fed either directly to the oligomerization III) or to the separation unit (Va). However, the components previously absorbed by the absorbent in the absorption column are not completely expelled within the desorption column.

A small proportion of the C3 olefins and C4 olefins as well as C5+ olefins, which are particularly valuable for oligomerization, remains in the absorbent, so that it is called a partially loaded absorbent (f)).

This partially loaded absorbent (f)) particularly preferably leaves the desorption column and is divided into two streams. 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). In this particularly preferred embodiment, it is this second partial stream of the partially loaded absorbent (f)) which, as the absorbent according to the invention, either goes directly to the oligomerization unit (III) (see Fig. 1.1 ac) and/or to the separation unit (Va) (see Fig. 1.2 ) to be led. In the latter case, in a variant of this embodiment, the still absorbed species are separated from the absorbent and passed via a return of C3 to C8 hydrocarbons (m)) from the separation unit (Va) to the oligomerization unit (III).

Step III) or oligomerization unit (III) of the invention:

There are therefore several streams by means of which olefins are supplied to the oligomerization unit (III). In addition, in step III), the olefins contained in the absorbent are preferably catalytically oligomerized to form an oligomer.

According to the invention, 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 05 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. If, as in a preferred embodiment, 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 03 to 025. The light 03 to 05 hydrocarbons are primarily composed of unreacted olefins.

The majority of the crude oligomer (g)) consists of 05 to 020 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. Step V) or separation unit (Va) + definition of absorbent:

In a preferred embodiment of the invention, a separation unit (Va) is arranged between (III) and (IV). A crude oligomer (g)) is fed to the separation unit (Va). In this separation unit, as for example in an embodiment according to FIG. 1.2, C3 to C8 hydrocarbons (m)) are separated off, which are returned to the oligomerization unit (III).

In the preferred embodiments according to FIGS. 1.1a and 1.1c, on the other hand, 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)). Likewise, 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+. Lighter amounts of C8- can result from a non-ideal separation of the separation unit (Va). In order to produce kerosene from this heavy oligomer, it is fed to the hydrogenation unit (IV). This heavy oligomer (h)), like the crude oligomer (g)), can be assigned to the generic term oligomer.

Step IV) or hydrogenation unit (IV):

Hydrogen is necessary for hydrogenation IV). 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.

Typically, 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 . According to the invention, the catalytic hydrogenation in step IV) results in a hydrogenated product which consists mainly of C9+ (C9 to C25) paraffins and C9+ naphthenes.

As already described above, 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.

Step V) or separation unit (Vb) + definition of absorbent:

In a preferred embodiment, the hydrogenated product (i)) is divided into the fractions gasoline components (j)), kerosene (k)) and diesel components (I)) in a separation unit (Vb). In a preferred embodiment of the invention, 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.

Separation at least in separation unit (Vb):

In a preferred embodiment of the invention, in which the absorbent is obtained from a separation unit (Vb) at least after step IV), the process has the steps:

II) Separation of an olefin-containing fraction from the olefin-containing mixture of substances

Absorption with an absorbent, a remaining cycle gas being recycled to step I), with ethylene being recycled with the cycle gas during absorption in step II), and wherein the absorbent from step II) is then passed directly to an oligomerization in step III) and/or to a separation unit (Va),

III) Catalytic oligomerization of the olefin-containing fraction obtained in step II) to form an oligomer,

IV) catalytic hydrogenation of the oligomer,

V) Separation of the absorbent after step IV) in a separation unit (Vb), the absorbent being removed from there for absorption in step II).

The advantage of this preferred embodiment is that the absorbent which is separated from the separation unit (Vb) is already a hydrogenated fraction of C9+ paraffins and C9+ naphthenes. Under the conditions of oligomerization, such C9+ paraffins and 09+ naphthenes contained in the absorbent are not reactive. If a stream of the partially loaded absorbent is passed from the absorption to the oligomerization unit (III), then this absorbent allows better control of the temperature during the oligomerization due to the high enthalpy of vaporization of the 09+ paraffins and 09+ naphthenes. This limits undesirable side reactions such as cyclization, hydrogen transfer and aromatic formation to a minimum.

By using this absorbent, advantageously high selectivities for kerosene and thus high kerosene yields are achieved.

This embodiment of the method according to the invention obviously corresponds to a preferred embodiment of the system, where:

• 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, and

• the separation unit (Vb) has a line for the removal of the absorbent to the absorption unit (II).

If the absorbent from step II) is now fed to a separation unit (Va), the following further advantage occurs.

The C3-C8 hydrocarbons contained in this (partially loaded) absorbent are separated in separation unit (Va) and preferably as “a rich in C3 to C8 olefins Hydrocarbon stream “returned to the oligomerization unit (III). The C9+ hydrocarbons originally from the absorbent also enter the product of the oligomerization (the oligomerisate) via the separation unit (Va). This dilution of the oligomer advantageously enables better control of the temperature in the subsequent hydrogenation (step IV)). This minimizes undesirable side reactions such as cracking.

According to the invention, part of the kerosene (k)) and/or the diesel components (I)) can be used as an absorbent (see Fig. 1.1a, 1.1b and 1.2). This corresponds to the variant according to the invention described above, in which in step V) the absorbent is separated after step IV) in a separation unit (Vb). Accordingly, 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. According to the invention, this mixture is fed to the absorption unit (II) described above as fresh, unloaded absorbent (e)).

According to the invention, after step III) in the separation unit (Va) and/or after step IV) in the separation unit (Vb), an absorbent is separated, with which the absorption according to the invention is carried out in step II).

Separation only in separation unit (Va):

If separation only takes place in separation unit (Va) (see Fig. 1.1c) then the absorbent consists mainly of C5 to C8 hydrocarbons.

Separation only in separation unit (Vb):

Alternatively, if the absorbent is only separated in the separation unit (Vb) (see Fig. 1.1 b and 1.2), then the absorbent results as the hydrogenated fraction of the C9+ paraffins and C9+ naphthenes. In such a preferred embodiment, in which (as step V)) the separation of the absorbent only takes place after step IV) in the separation unit (Vb) and not also in the separation unit (Va), the following advantage results. By using the not Reactive long-chain paraffins and naphtenes as absorbents from the separation unit (Vb) prevent reactive C5-C8 olefins from an absorbent (as would be the case when separating the absorbent after step III) in the separation unit (Va)) during absorption (via the cycle gas) can be transferred back to the olefin production in step I). This could lead to the formation of long-chain olefins or even by-products. The corresponding by-product formation would lead to a reduction in selectivity and yield in olefin production (I).

If separation occurs in both separation units (Va) and (Vb) (see Fig. 1.1a), the absorbent is a mixture of the C5 to 08 hydrocarbons and the hydrogenated 09+ fraction.

In a preferred embodiment of the invention, 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). For the olefin production in I), a fixed bed reactor is preferably used, particularly preferably an isothermal tubular reactor with a fixed bed. For the oligomerization in step III), at least one fixed bed reactor is preferably used, particularly preferably at least one adiabatic fixed bed reactor. For the catalytic hydrogenation in step IV), at least one fixed bed reactor is preferably used, particularly preferably at least one adiabatic fixed bed reactor.

In a likewise preferred embodiment of the invention, 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.

According to the invention, the reactors are understood to be part of the respective system section of the system according to the invention. In a preferred embodiment of the invention, the catalyst is

• in step I) or the olefin production unit (I) a zeolitic catalyst, and

• in step III) or the oligomerization unit (III) a zeolitic and/or a non-zeolitic acidic catalyst,

• in step IV) or the hydrogenation unit (IV) preferably a metal catalyst selected from Pt, Pd, Ru, Ni, Rh, Co, Mo and W.

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.

The method or system according to the invention has the following advantages:

The oligomerization step III) of the olefins is carried out in a way that avoids undesirable side reactions, such as. B. Cyclization, hydrogen transfer and aromatic formation, limited to a minimum. Chain growth of olefinic species mainly takes place, with the maximum product formed already being in the C number range and boiling range of kerosene. The formation of long-chain olefins occurs preferably through the reaction of C3+ olefins with each other. The olefins are activated by protonation on the acidic catalyst. In the case of ethylene, a relatively high activation energy is necessary for this. It involves the formation of a primary carbenium ion. At oligomerization temperatures, ethylene shows lower reactivity than the C3 to C5 olefins. The temperature level that is more favorable for activating the ethylene is available in the olefin production unit (I). Here the ethylene is converted into C3+ olefins by means of methylation in the presence of oxygenates. At the same time, new ethylene is created as one of the primary products of the conversion of oxygenates into olefins. As already noted above, the C3+ hydrocarbons are separated from the olefin-containing mixture (b)) obtained in step I) by means of the absorption unit (II) and fed to the oligomerization unit (III). At the same time, light components including ethylene are enriched in the remaining cycle gas (c)), which is returned to olefin production I). • Consequently, the conversion of the recycled ethylene into C3+ olefins in olefin production I) and their subsequent conversion in oligomerization III) advantageously increases the yield of kerosene.

Dilution is also achieved in olefin production I) by the “non-reactive” components of the permanent gases, such as CO2 and CO, contained in the circulating gas (c)). This dilution serves to better control the heat transfer in the olefin production unit (I).

• This advantageously ensures high olefin selectivities and yields, and ultimately high kerosene yields.

The (unloaded) absorbent (e)) necessary for the absorption according to the invention in step II) is, as already described above, separated after step III) in the separation unit (Va) and/or after step IV) in the separation unit (Vb) ( see Fig. 1.1 a-c and 1.2).

When separating the absorbent after step III) in the separation unit (Va)

(see Figures 1.1a and 1.1c), the absorbent consists of an olefin-rich fraction of C5 to C8 hydrocarbons. The absorption taking place in the absorption unit (II) results in a transfer of C3+ hydrocarbons into the absorbent. This results in an enrichment of the absorbent with the main products of olefin production I), i.e. H. C3 olefins and C4 olefins, which after step II) are predominantly present in the olefin-containing fraction (d)). In addition to the olefin-containing fraction, a stream of the partially loaded absorbent (f)) is also passed for oligomerization (III). Both streams are contacted again in oligomerization III). In general, the reactivity of the olefins decreases as the chain length increases. The diffusion inhibition makes the dimerization of long-chain C5+ olefins more difficult. Since C3 olefins and C4 olefins diffuse more quickly in the pore system of the catalyst, the C5+ olefins located at the active centers of the catalyst can be converted more effectively into long-chain olefins in the C number range of kerosene.

• This advantageously increases the yield of kerosene. When separating the absorbent after step IV) in a separation unit (Vb)

(see Fig. 1.1a, 1.1c and 1.2), the absorbent consists of a hydrogenated fraction of C9+ paraffins and C9+ naphthenes. Under the conditions of oligomerization III) these components are not reactive. If a stream of the partially loaded absorbent (f)) is passed from the absorption unit (II) to the oligomerization unit (III) (see Fig. 1.1a and 1.1c), the high enthalpy of vaporization of the C9+ paraffins and C9+ naphthenes enables better control over the temperature in the reaction stages of oligomerization.

• This advantageously ensures high selectivities and ultimately high kerosene yields.

If, in one embodiment, 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 flows to the oligomerization unit (III). is returned. By contacting this mixture of light 0^04 olefins and heavier 05 to 08 olefins in oligomerization III), an increase in the kerosene yield is achieved analogously to the manner already described above. The 09+ hydrocarbons originally from the absorbent reach the main product of the oligomerization III) (heavy oligomerisate (h)) via the separation unit (Va). The resulting proportion of paraffins and naphthenes leads to a dilution of the mainly olefinic oligomer. This dilution enables better control over the temperature in the subsequent hydrogenation IV). Undesirable side reactions that occur here, such as cracking, are minimized. This advantageously ensures high selectivities and ultimately high kerosene yields.

With the measures described, a kerosene yield of 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.

With the invention, a synthetic kerosene is obtained which meets the applicable specifications.

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 oil 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.

Preferred embodiments based on Fig. 2.1a and Fig. 2.1b, according to claims 4/20, 5 and 6

Regarding claim 4:

In a preferred embodiment of the invention, after the catalytic oligomerization, a fraction of low boilers (o)) is additionally separated off and returned to step I).

Regarding claim 20:

This corresponds to a preferred embodiment of the system, in which a fraction of low boilers is separated off in the oligomerization unit (III) by means of a separator and/or in the separation unit (Va) and is returned to the olefin production unit (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).

By returning the low boilers (o)) to the olefin production I), ethylene is advantageously converted into C3+ olefins. 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 03+ olefins and their subsequent implementation in oligomerization III) increase the yield of kerosene.

Regarding claim 5:

In a likewise preferred embodiment of the invention, after step III), an additional 03 to 05 hydrocarbon fraction is separated off in a separation unit (Va) and returned to step III). Regarding claim 6:

In a likewise preferred embodiment of the invention, after step III), a C5 to C8 hydrocarbon fraction is additionally separated off in a separation unit (Va) and returned to step III).

Advantageously, with the last two embodiments, 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.

Below are more details about these embodiments:

According to the invention, 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, CO2, H2) are absorbed in the absorbent and reach 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. In a preferred embodiment, the oligomerization unit (III) contains a separation stage (III/S1), which is shown in Fig. 2.1a and Fig. 2.1b.

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, CO2 and H2.

The crude oligomer (g)) is fed to the separation unit (Va). In a preferred embodiment, the separation unit (Va) includes two stages (Va/K1) and (Va/K2). This is shown as an example in Fig. 2.1a and Fig. 2.1b.

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)). (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 in a second stage (Va/K2), for example by means of a distillation column, into an olefin-rich C5 to C8 hydrocarbon fraction (v)) in Fig. 2.1 b and the heavy oligomer (h )) separated. The heavy oligomer (h)) is then passed to hydrogenation IV). In a preferred embodiment, the olefin-rich C5 to C8 hydrocarbon fraction (v)) is returned to oligomerization III) (see Fig. 2.1b). This C5 to C8 hydrocarbon fraction (v)) is preferably 10% to 99% by mass, particularly preferably 20% to 99% by mass, particularly preferably 40% to 99% by mass Mass % returned.

In a further preferred embodiment, the same fraction is used as unloaded absorbent (e)) in the absorption unit (II) (see Fig. 2.1a). Likewise, in a preferred embodiment, 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 03 to 05 hydrocarbon fraction (p)) and the 05 to 08 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.

By returning the low boilers (o)) to the olefin production I), ethylene is advantageously converted into 03+ olefins. Due to the “non-reactive” components of the permanent gases such as CO2 and CO contained in the low boilers (o)), a dilution of the reactive feed is advantageously achieved 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.

By recycling the olefin-rich hydrocarbon fractions C3 to C5 (p)) and C5 to C8 (v)) for oligomerization III), 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. In a preferred embodiment, the hydrogenation unit (IV) contains a separation stage (IV/S1). This is shown as an example in Fig. 2.1a. Here, 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 passed to the separation unit (Vb). Here, for example by means of a distillation column, a separation takes place into a fraction of gasoline components (j)), kerosene (k)) and diesel components (I)), with the synthetic kerosene typically accounting for the majority of the product. In addition, a light fraction of C3/C4 hydrocarbons can also be separated off in this column. In principle, it is also possible for the separation unit (Vb) to carry out the complex separation task in a two-stage distillation.

Preferred embodiment based on Fig. 2.2a and Fig. 2.2b, according to claim 9

In a preferred embodiment of the invention, step III) is carried out in several stages of oligomerization, wherein, i. an intermediate product of the oligomerization is separated between these stages and fed to the separation unit (Va), and/or ii. an intermediate product of the oligomerization when entering the subsequent stage of oligomerization by recycling the C3 to 05 hydrocarbon fraction according to claim 5 and / or by recycling the 05 to 08 hydrocarbon fraction according to claim 6 and / or by the olefin-containing fraction according to claim 1 and / or the absorbent from step II) is cooled. The latter concerns the case in which the absorbent from step II) is passed directly to the oligomerization in step III).

Advantageously allowed i. the separation of an intermediate product of the oligomerization, the securing of components in the kerosene range with a partial conversion and corresponding to the associated kerosene selectivity. This avoids conversion of these components to higher C numbers. Advantageously allowed ii. control over the reaction temperature and consequently high selectivities and high kerosene yields are ensured. By controlling the reaction temperature, a long service life of the catalyst is also achieved.

Oligomerization is a highly exothermic reaction. 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.

In a preferred embodiment, the reaction mixture is mixed directly with 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 within the bed is not used or the measures taken for this purpose 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.

In a preferred embodiment, the oligomerization unit (III) is divided into several stages. This is shown as an example in Fig. 2.2a and in Fig. 2.2b. Here, a stage represents either a catalyst bed within a reactor or a reactor with one or more catalyst beds. The stages (111.1) and (III.2) were shown for illustration purposes, but other stages are also possible. 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. (see Fig. 2.2a) In this way, the intermediate product of the oligomerization (w)) is cooled by partial streams from (d)), (f)), (p)) and (v)). By dividing the streams and using them for intermediate cooling of the reaction mixture, control over the reaction temperature is advantageously achieved. This ensures high selectivities and high kerosene yields.

Due to the described division of the streams, olefin-rich fractions of CS 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). These paraffins are not reactive in oligomerization III). By dividing the partial streams of (p)) and (v)), 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.

In the oligomerization of olefins, increasing sales lead to an increase in the proportion of C9+ hydrocarbons in the product. The proportion of 09 to 017 hydrocarbons in this product later results in the end product kerosene (k)) after hydrogenation IV) and the separation unit (Vb). The proportion of hydrocarbons 018+ later results in the byproduct of the diesel components (I) after the hydrogenation IV) and separation unit (Vb). As the conversion rate of the light olefins increases, the ratio of the 09 to 017 hydrocarbons to the 018+ hydrocarbons in the 09+ hydrocarbon product decreases. This corresponds to a reduction in kerosene selectivity. In order to benefit from high selectivities, only a partial conversion of the olefins is permitted in the process via one or more stages of oligomerization III). The proportion of hydrocarbons 09+ is removed from the corresponding product streams and fed to hydrogenation IV). The 03 to 08 olefins are fed back into the oligomerization III) in streams (p)) and (v)). This is shown as an example in Fig. 2.2b. Accordingly, in a preferred embodiment, the intermediate product of the oligomerization (w)) is passed after stage (II 1.1) to a separation stage (e.g. by means of separation) (III/S2) and subsequently to the separation unit (Va). As a result, 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 03 to 08 olefins are added from the separation unit (Va) via the returns of the 03 to 05 hydrocarbon fraction (p)) and the 05 to 08 hydrocarbon fraction (v)). a subsequent stage (III.2) of the oligomerization III) or to the first stage (111.1).

A partial conversion of light olefins with intermediate separation of the C9+ fraction is achieved via one or more successive stages of oligomerization, as shown in Fig. 2.2b. By separating the intermediate product of the oligomerization (w)), a proportion of components in the kerosene range is advantageously secured. Each additional stage with partial sales increases the overall level of sales. Thus, 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.

Preferred embodiments based on Fig. 2.3, according to claims 11 and 12

Regarding claim 11:

In a preferred embodiment of the invention, the oligomer from step III) is diluted upon entering step IV) by recycling a portion of the hydrogenated product to step IV).

The oligomer is preferably taken from the separation unit (Va) and the hydrogenated product for dilution is taken between (IV/S1) and (Vb) (see Fig. 2.3).

Regarding claim 12:

In a likewise preferred embodiment, step IV) is carried out in several stages, with hydrogen and/or a portion of a heavy oligomer and/or the product hydrogenated in IV) being 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.

With the measures taken, the hydrogenation is optimized accordingly and a high kerosene yield is ensured.

More details about both embodiments below:

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 feed mixture from the separation unit (Va) serves as the feed mixture Oligomerisate (h)). It is a mixture of C9+ hydrocarbons, which consists mainly of olefins. A certain proportion of aromatics is also included. 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.

In a preferred embodiment, the hydrogenation unit (IV) is therefore divided into several stages. This is shown as an example in Fig. 2.3. Here, a stage represents either a catalyst bed within a reactor or a reactor with one or more catalyst beds.

For illustrative purposes, a two-stage system was shown in Fig. 2.3, but further stages are possible. All possible feed and return routes were also shown, although they do not all have to be combined. The corresponding intermediate product of the hydrogenation (y)) is passed from one stage to the subsequent stage. In this preferred embodiment, 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.

By contacting the partial streams (i)) and (t)), 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. By dividing the streams and using them to cool the reaction mixture, control over the reaction temperature and the suppression of side reactions are advantageously achieved. With the measures taken, the hydrogenation is optimized accordingly and consequently a high kerosene yield ensured. By controlling the reaction temperature, a long service life of the catalyst is also achieved.

Preferred embodiments based on Figure 2.4, corresponding to claims 7 and 8

Regarding claim 7:

In a preferred embodiment of the invention, after the catalytic olefin production in step I), 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.

Regarding claim 8:

In a likewise preferred embodiment, after the catalytic olefin production in step I), a fraction of C8+ hydrocarbons is separated from the olefin-containing mixture of substances and passed to step IV) and/or step V). Advantageously, 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.

More details about both embodiments below:

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 03+ 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. In a preferred embodiment, the product range from olefin production I) contains two Isolation stages (I/S1) and (I/S2). This is shown in Fig. 2.4. 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 advantageously results in a dilution of the reactive feed in olefin production I). 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 08+ hydrocarbons of the liquid organic phase (ab)) consist largely of aromatics. 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 3 (ac)) remaining after the three-phase separator (I/S1) is raised by compression to a pressure level favorable for absorption II) and then cooled. This can lead to the formation of a liquid organic phase again. This is separated from the gaseous phase using a separator in (I/S2), for example. 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). In the latter case, 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)). By reacting the olefins contained in the olefin-containing mixture 4 (ad)) in the oligomerization III), the yield of kerosene is advantageously increased.

All embodiments of the invention, including those in the subclaims, can expediently be combined with one another.

The invention is illustrated by the following exemplary embodiments, without being limited to them.

Embodiment 1: Olefin production unit (I)

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 based on the duration of the experiment corresponding mass flow is determined. 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.

Two different experiments were carried out at 4 bar and 475°C. In experiment 1, no ethylene was added. The degree of conversion of the methanol was >99% by mass. Based on the CFh content contained in the methanol, an oil yield of 83% by mass could be achieved.

A composition of the hydrocarbon product of this Experiment 1 is as follows

Table 1 shown:

Table 1: Composition of the hydrocarbon product:

In experiment 2, ethylene was added. The procedure was carried out analogously to Experiment 1. In both experiments, the same mass flow of methanol was supplied. Experiment 2 was used to investigate the conversion of the ethylene, which in the process according to the invention is returned to the olefin production unit (I) via the cycle gas (c)). The recycled ethylene is converted into C3+ olefins by the oxygenates, such as methanol, through methylation. To illustrate this aspect, the mass fractions of methanol and ethylene supplied as well as the C3+ olefins formed, normalized to the feed, are shown in Table 2 below for both experiments.

Table 2: Standardized mass proportions of methanol, ethylene and 03+ olefins:

When comparing the experiments it can be seen that the co-reaction with ethylene significantly increases the product formation of C3+ olefins (Am = + 23%).

In the process according to the invention, a larger mass flow of C3+ olefins is consequently achieved by recycling ethylene to produce olefins (I). Its conversion in the oligomerization unit (III) ultimately leads to an increase in the amount and yield of synthetic kerosene produced.

Embodiment 2: Oligomerization unit (III)

An experimental system was used for the experiments on the oligomerization of C3/C4 olefins. The system could be used to set the pressure in the range 0-65 bar and the temperature up to 350 °C.

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. In addition to the olefins, the reaction gas consisted of a proportion of non-reactive components.

All (liquid) gases were metered using mass flow controllers, heated to the reaction temperature and mixed before entering the reactor.

After leaving the reactor, 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 corresponding conversion, the composition of the hydrocarbon mixture and the product selectivity are listed below in Table 3 for two selected experiments.

The operating conditions for Experiment 1 were as follows: temperature: 200 °C; Pressure: 20 bar; Proportion of non-reactive components: 46% by volume and mass-related space velocity: 1.0 kg olefins per hour and kg of catalyst.

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:

It can be seen from Table 3 that in experiment 1 a relatively low conversion of the C3 olefins and C4 olefins was achieved under moderate conditions. Increasing the reaction temperature and pressure causes an increase in conversion in Experiment 2. In experiment 1, a small proportion of kerosene (4.6% by mass) is obtained in the hydrocarbon mixture with a low conversion. However, based on the C9+ hydrocarbon product (kerosene & diesel components), the result is a very high kerosene selectivity. In contrast, in Experiment 2, although a high proportion of kerosene is obtained in the entire hydrocarbon mixture, a lower kerosene selectivity is achieved in the C9+ hydrocarbon product. Therefore, a larger proportion of diesel components is obtained here as an undesirable by-product. Both experiments show the opposite relationship between the degree of conversion and selectivity. With the help of the experiments listed it is shown that in the process according to the invention there is an optimum of C3/C4 conversion and kerosene selectivity for the implementation at the respective stages

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 illustrated in the preferred embodiments using FIGS. 2.2a and 2.2b and the associated figures 2.2a and 2.2b.

Embodiment 3: Hydrogenation unit (IV)

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. To simulate the Reaction feeds from the hydrogenation unit (IV) were fed to the reactor with a mixture of olefins including dienes in the range from C8 to C24 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. After the reactor, 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.

In the specific example, 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 .

As a result, 98% by mass of the olefins used, including the dienes, could be hydrogenated. The conversion of the C7 aromatics was 99% by weight and the conversion of the C9 aromatics was 89% by weight.

The results show that olefins including dienes are almost completely hydrogenated over the C number range from C8 to C24. It can also be shown that the majority of the aromatics present are hydrogenated and thus a fuel with a low aromatic content is obtained.

Embodiment 4: Plant for producing synthetic kerosene

The appendix described in this example describes a concrete form of appendix to describe the invention without limiting it.

The system is based schematically on Fig. 2.5. This figure has as its origin Figure 1.1c and the system contains features from Figures 2.1.a, 2.1b and 2.4. The plant therefore contains the corresponding plant parts for 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). During these liquid separations, 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)).

In the absorption unit (II), 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). From the absorption unit (II) according to the invention, 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).

Here, 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 08 hydrocarbons (m)) are recycled. Thus, part of the 05 to 08 hydrocarbon fraction (v)) is returned to the 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. The hydrogenated product (i)) is separated in a separation unit (Vb) using distillation into gasoline components (j)), the main product kerosene (k)) and diesel components (I)). The following result is achieved:

Yield of gasoline components (C5 to C8) = 6% by mass

Yield of kerosene (C9 to C17) = 44% by mass

Yield of diesel components (C18+) = 12% by mass

Aromatics < 5% by mass

The currents described in this description are shown in Table 4 below.

Table 4: Overview table of the streams included according to Fig. 2.5:

Embodiment 5 - Absorbent after hydrogenation from separation unit (Vb)

The appendix described in this example describes a concrete form of appendix to describe the invention without limiting it. In this system, the absorbent is separated after hydrogenation in a separation unit (Vb) and from there taken to absorption in step II).

The organic liquid phase (olefin-containing mixture 4 (ad))) obtained in the separation stage (I/S2) is fed to the separation unit (Va). The partially loaded absorbent (f)) is fed to the separation unit (Va). The C3 to C8 hydrocarbons (m)) are recycled from the separation unit (Va) for oligomerization (III).

The remaining material flows proceed analogously to exemplary embodiment 4.

With this arrangement, the yield of kerosene (k) is 47% by mass.

Non-patent literature

J. Eilers, S.A: Posthuma and S.T. She

The Shell Middle Distillate Synthesis Process (SMDS)

Calatysis Letters 1990, 7, 253-270.

Bellussi, G.; Mizia, F.; Calemma, V.; Pollesel, P.; Millini, R.

Oligomerization of olefins from Light Cracking Naphtha over zeolite-based catalyst for the production of high quality diesel fuel

Microporous and Mesoporous Materials 2012, 164, 127-134.

Reference symbol list

(I) Olefin production unit

(II) Absorption unit

(III) Oligomerization unit

(IV) Hydrogenation unit

(V) Separation unit a) oxygenates b) olefin-containing substance mixture c) cycle gas d) olefin-containing fraction e) unloaded absorbent f) partially loaded absorbent g) crude oligomer h) heavy oligomer i) hydrogenated product j) gasoline components k) kerosene l) diesel components m) C3 up to C8 hydrocarbons n) loaded absorbent o) low boilers p) C3 to C5 hydrocarbon fraction q) crude oligomerisate 2 r) crude oligomerisate 3 s) hydrogenated product 2 t) hydrogen-rich fraction u) hydrogen v) C5 to C8 hydrocarbon fraction w) intermediate product oligomerization x) intermediate product oligomerization 2 y) hydrogenation intermediate product z) olefin-containing mixture of substances 2 aa) water ab) C8+ hydrocarbons ac) olefin-containing mixture of substances 3 ad) olefin-containing mixture of substances 4

Claims

Patent claims

1. Process for producing a synthetic kerosene with the steps:

II) Separation of an olefin-containing fraction from the olefin-containing mixture of substances by absorption with an absorbent, a remaining cycle gas being recycled to step I), with ethylene being recycled with the cycle gas during the absorption in step II), and then the absorbent from step II ) is led directly to an oligomerization in step III) and / or to a separation unit (Va),

IV) catalytic hydrogenation of the oligomer,

V) Separation of the absorbent after step III) in a separation unit (Va) and/or after step IV) in a separation unit (Vb).

2. The method according to claim 1, with the step:

3. The method according to any one of claims 1 or 2, wherein in step II) the absorption is carried out by means of an absorption column and a desorption column and the absorbent, separated in a separation unit (Va) and / or (Vb), as an unloaded absorbent in the absorption column is contacted with the olefin-containing mixture of substances, and is then passed as a loaded absorbent from the absorption column to the desorption column, where the olefin-containing fraction is expelled from the absorbent.

4. The method according to any one of claims 1 to 3, wherein after the catalytic oligomerization, a fraction of low boilers is additionally separated off and returned to step I). Method according to one of claims 1 to 4, wherein after step III) a C3 to C5 hydrocarbon fraction is additionally separated off in a separation unit (Va) and returned to step III). Method according to one of claims 1 to 5, wherein after step III) a C5 to C8 hydrocarbon fraction is additionally separated off in a separation unit (Va) and returned to step III). Process according to one of claims 1 to 6, wherein after the catalytic olefin production in step I), water is separated from the olefin-containing mixture of substances and recycled to the catalytic olefin production in step I). Process according to one of claims 1 to 7, wherein after the catalytic olefin production in step I), a fraction of C8+ hydrocarbons is separated from the olefin-containing mixture of substances and passed to step IV) and/or step V). Method according to one of claims 1 to 8, wherein step III) is carried out in several stages of oligomerization,

• an intermediate product of the oligomerization is separated between these stages and fed to the separation unit (Va), and/or

• an intermediate product of the oligomerization when entering the subsequent stage of oligomerization by recycling the 03 to 05 hydrocarbon fraction according to claim 5 and / or by recycling the 05 to 08 hydrocarbon fraction according to claim 6 and / or through the olefin-containing fraction according to claim 1 and / or that Absorbent from step II) is cooled. Process according to one of claims 1 to 9, wherein a low-aromatic stream with gasoline components and diesel components is obtained as by-products. Process according to one of claims 1 to 10, wherein the oligomer from step III) is diluted upon entering step IV) by recycling a portion of the hydrogenated product to step IV). Method according to one of claims 1 to 11, wherein step IV) is carried out in several stages, with hydrogen and / or part of an oligomer and / or the product hydrogenated in IV) being used for intermediate cooling. Process according to one of claims 1 to 12, wherein 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. Process according to one of claims 1 to 13, wherein as a catalyst

• in step I) a zeolitic catalyst, and

• in step III) a zeolitic and/or a non-zeolitic acidic catalyst, and

• in step IV) a metal catalyst selected from Pt, Pd, Ru, Ni, Rh, Co, Mo and W is used. Process according to one of claims 1 to 14, wherein in step III) the olefins contained in the absorbent are also catalytically oligomerized to form an oligomer. Plant for producing a synthetic kerosene from oxygenates, comprising an olefin production unit (I), an absorption unit (II), an oligomerization unit (III), a hydrogenation unit (IV) and at least one separation unit (Va and/or Vb), wherein

• the olefin production unit (I) has a line for supplying the oxygenates,

• the separation unit (Va and/or Vb) has a line for the removal of the absorbent to the absorption unit (II). Plant according to claim 16, wherein

• the separation unit (Vb) has a line for the removal of the absorbent to the absorption unit (II). Plant according to one of claims 16 or 17, wherein the absorption unit (II) comprises an absorption column and a desorption column, and wherein the line of the separation unit (Va and / or Vb) for the removal of the absorbent to the absorption unit (II) runs to this absorption column and the line for the removal of an olefin-containing fraction to the oligomerization unit (III) runs from the desorption column to the oligomerization unit. Plant according to one of claims 16 to 18, wherein 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 the diesel components. Plant according to one of claims 16 to 19, wherein a line for the return of low boilers leads from the oligomerization unit (III) and / or the separation unit (Va) to the olefin production unit (I). Use of a system according to one of claims 16 to 20 for the method according to one of claims 1 to 15.