WO2022063995A1

WO2022063995A1 - Methanol to olefin (mto) process

Info

Publication number: WO2022063995A1
Application number: PCT/EP2021/076374
Authority: WO
Inventors: Pablo Beato
Original assignee: Haldor Topsøe A/S
Priority date: 2020-09-25
Filing date: 2021-09-24
Publication date: 2022-03-31

Abstract

A process for producing an olefin stream, said process comprising passing a feedstock stream comprising oxygenates over catalyst comprises a zeolite with a framework having a 10-ring pore structure, in which said 10-ring pore structure comprises a three- dimensional (3-D) pore structure, such as MFI, wherein the pressure is 2-20 bar and the temperature is 150-350°C; or the pressure is 2-30 bar and the temperature is 340- 400°C, and wherein an olefin stream comprising C2-C3 olefins is withdrawn from said olefin stream and used as additional feed stream. The olefin stream may be converted to jet fuel, particularly sustainable aviation fuel (SAF) by further oligomerization and hydrogenation.

Description

Title: Methanol to olefin (MTO) process

FIELD OF THE INVENTION

The present invention relates to the conversion of a feedstock comprising oxygenates such as methanol and/or dimethyl ether to an olefin stream having a low content of aromatics and a high content of higher olefins, especially (C4= - C8=), and optionally the subsequent conversion of the olefin stream to the hydrocarbons boiling in the jet fuel range, particularly sustainable aviation fuel (SAF), by oligomerization and hydrogenation.

BACKGROUND OF THE INVENTION

Currently, processes for the conversion of oxygenates such as methanol to olefins (MTO) are used to produce ethylene and propylene as the main olefin products with the purpose of serving as feedstock for plastic production. When higher hydrocarbons are the desired product such as in e.g. methanol to gasoline (MTG) processes, around 30% of aromatics are typically formed. However, when producing hydrocarbons boiling in the jet fuel range, particularly sustainable aviation fuels (SAF), current requirements for do not allow the presence of aromatics in the olefin stream feed.

Due to society concerns about global climate change and the resulting political pressure on the aviation industries, the market for SAFs is expected to increase substantially during the next decades. Currently, a small number of biocatalytic and thermo-catalytic processes have been approved by ASTM to be able to produce SAFs. Hence, a pre-condition for the use of any SAF as aviation turbine fuel is an ASTM certification. So far, only a small number of processes, producing SAF or synthetic paraffinic kerosene (SPK) fuels have been approved by ASTM International (ASTM) Method D7566 for blending into jet fuel at levels up to 50%. One important general requirement is therefore, that the synthetic part of SAF (50 vol%) must be virtually free from aromatics, while the final SAF blend can contain up to 26.5 vol% aromatics.

So far, the only process that is foreseen to be able to produce relevant amounts of SAF is based on biomass derived Fischer-Tropsch (FT) synthesis, followed by oligomerization of the lower olefins and subsequent hydrogenation. However, the selectivity of FT to lower olefins is only moderate. The present invention uses the well- known methanol to olefins (MTO) process as a more attractive route to obtain olefins at a higher selectivity. Currently, the proposed process layouts for the conversion of methanol to jet fuel are multistep processes, consisting of at least MTO, oligomerization and hydrogenation which are all proven technologies but in combination do not appear very efficient, due to high recycle streams and very different process conditions for the individual steps.

Potential feedstocks for producing SAFs are generally classified as (a) oil-based feedstocks, such as vegetable oils, waste oils, algal oils, and pyrolysis oils; (b) solid-based feedstocks, such as lignocellulosic biomass (including wood products, forestry waste, and agricultural residue) and municipal waste (the organic portion); or (c) gas-based feedstocks, such as biogas and synthesis gas (syngas). Syngas, alcohols, sugars, and bio-oils can be further upgraded to jet fuel via a variety of synthesis, either fermentative or catalytic processes.

There is currently no viable one-step catalyst/process that would allow to convert a feedstock comprising oxygenates such as methanol directly into a hydrocarbon boiling in the jet fuel range, i.e. jet fuel, at least not at reasonable yields. To produce jet fuel starting from methanol, typically a three-step process is used consisting of: a) Methanol to olefins (MTO), b) Oligomerization of olefins, and c) Hydrogenation of long chain olefins. Conventional approaches to the conversion of methanol to gasoline or diesel hydrocarbon products were envisaged already in the late 1970s and early 1980s. Thus, US 4,021 ,502, US 4,211 ,640, US 4,22,7992, US 4,433,185, US 4,456,779, disclose process layouts based on classical MTO process conditions, i.e. high temperatures e.g. about 500°C and moderate pressures e.g. about 1-3 bar, in order to obtain efficient conversion of methanol to olefins. However, under these conditions a significant amount of aromatic hydrocarbons (aromatics) is produced, e.g. 10-30 wt% in the olefin stream, which needs to be separated and a relatively large volume of MTO product effluent has to be cooled and treated to separate a C2-light gas stream, which is unreac- tive, except for ethene which is reactive to only a small degree. The remaining of the olefin stream has to be pressurized to the substantially higher pressure of the oligomerization (OLI) reactor. Hence, so-called Mobil-Olefin-to-Gasoline-Distillates (MOGD) process patents from the early 1990s such as US 5,177,279 try to solve these problems and improve the overall operation efficiency by further process integration of MTO and oligomerization and reduce the investment costs by splitting the methanol stream between the MTO and the OLI reactor. The splitting of the methanol feed has two advantages: first it reduces the MTO reactor size at the same overall methanol conversion and secondly only half of the methanol is processed at the high temperature conditions of the MTO reactor, thereby reducing both, the aromatics and the C2-content.

Two possible general process layouts are disclosed in US 5,177,279: (1) classical two- step process with MTO and OLI/MOGD reactor in series and (2) a three-step process, including an intermediate “olefin interconversion” (MOI) reactor that converts the lower olefins (C2=-C3=) to higher olefins (C5=-C9=), increasing thereby the amount of higher olefins from 25-35 wt% to 35-70 wt%. The latter process design provides more flexibility and only two reactors (MTO + MOI) are used when gasoline is the desired product, while all three reactors are used when distillates are the preferred product.

US 9,957,449 discloses a process for the producing hydrocarbons in the jet fuel range by oligomerization of renewable olefins having three to eight carbons.

Applicant’s US 20190176136 discloses the use of a ZSM-23 zeolite as catalyst for methanol to olefin conversion in a process step which is conducted at atmospheric pressure (about 1 bar) and 400°C, thereby producing a hydrocarbon stream with, less than 5wt% aromatics.

Yarulina et al, ChemCatChem 8 (2016) 3057-3063, discloses the use of Ca-modified ZSM-5 for methanol to olefins conversion with the purpose of achieving a high propylene selectivity, where the process is conducted at 1 bar and at high temperature of 500°C. The resulting olefin streams shows no formation of aromatics and a high selectivity for light olefins (C2= - C3=), yet low selectivity for higher olefins (C4= - C8=). US 2002/0103406 A1 discloses a process for making olefin dimer and oligomer product using a nickel-based oligomerization catalyst and using as feed an olefin containing stream from an oxygenated to olefin process.

US 2018155637 A1 discloses a process for producing an olefin stream from an oxygenate feedstock over a ZSM-5 catalyst at a pressure of 10-180 psig (0.7-12.4 barg) and a temperature of 440-550°C. A gas phase portion of the olefin stream is separated as light paraffins and light olefins (C4- compounds) and passed to an oligomerization reactor. This citation relates therefore to MTO operation at high temperatures, i.e. about 450°C with C2+C3 olefin yields of at least 50%. Hence, a significant amount of ethylene and which is desirable to minimize efficiently so as to avoid feeding it to the oligomerization, is produced. A portion of said gas phase (C4- compounds) may be recycled to the MTO reactor.

US 8,524,970 discloses a process for producing diesel of better quality, i.e. diesel with a higher cetane number comprising conversion of oxygenates to olefins, oligomerization of olefins and subsequent hydrogenation. More specifically, this citation discloses a similar process in which methanol is first converted to dimethyl ether and which is passed over a ZSM-5 catalyst at a pressure of 2-10 bar and temperature of 300-600°C. The gas phase portion of the resulting product stream is separated as C6-hydrocar- bons and fed to an oligomerization reactor and finally to a hydrogenation reactor. This citation focus therefore also on C2 to C8 olefins, in particular the higher olefins up to C8-olefins by increasing pressure.

SUMMARY OF THE INVENTION

As used herein, “MTO” (methanol to olefins) means the conversion of an oxygenate such as methanol to olefins.

As used herein, “OU” means oligomerization.

As used herein, “Hydro” means hydrogenation.

As used herein, “Hydro/OLI” means a single combined step comprising hydrogenation and oligomerization.

As used herein, “MTJ” means methanol to jet fuel and is interchangeable with the term “overall process” or “overall process and plant”, which means a process/plant combining MTO, OLI and Hydro, whereby a feedstock comprising oxygenates such as methanol is converted into jet fuel.

As used herein, the terms “jet fuel” and “hydrocarbons boiling in the jet fuel range” are used interchangeably and have the meaning of a mixture of C8-C16 hydrocarbons boiling in the range of about 130-300° at atmospheric pressure.

As used herein, “SAF” means sustainable aviation fuel or aviation turbine fuel, in compliance with ASTM D7566 and ASTM D4054.

As used herein, “olefin stream” means a hydrocarbon stream rich in olefins comprising higher and lower olefins, and optionally also aromatics, paraffins, iso-paraffins and naphthenes, and in which the combined content of higher and lower olefins is at least 25 wt%, such as 30 wt% or 50 wt%.

As used herein, the term “higher olefins” means olefins having four (4) or more carbons (C4+ olefins), in particular C4-C8 olefins (C4= - C8=)

As used herein, the term “lower olefins” means olefins having three or fewer carbons, in particular C2-C3 olefins (C2= - C3=)

As used herein, the term “high content of higher olefins” means that the weight ratio in the olefin stream of higher olefins to lower olefins is above 1 , for instance 2-4. Conversely, the term “low content of higher olefins” means that the weight ratio in the olefin stream of higher olefins to lower olefins is 1 or below.

As used herein, the term “selectivity to higher olefins” means the weight ratio of higher to lower olefins. “High selectivity to higher olefins” or “higher selectivity to higher olefins” means a weight ratio of higher to lower olefins of above 1.

As used herein, the terms “C2-light fraction” means C2= and C1-2 hydrocarbons.

As used herein, the term “lower hydrocarbons” means C1-2 (e.g. methane, ethane) and optionally also C2= and C3=. The term is also used interchangeably with the term “light paraffins”.

As used herein, the term “substantially free of aromatics”, “aromatic-free” or “low aromatics” means less than 10 wt% aromatics in an olefin stream, in particular less than 5 wt%, or even less than 1 wt%.

As used herein, the term “partial conversion of the oxygenates” or “partly converting the oxygenates” means a conversion of the oxygenates of 20-80%, for instance 40-80%, or 50-70%.

As used herein, the term “full conversion of the oxygenates” or “fully converting the oxygenates” means above 80% conversion of the oxygenates, for instance 90% or 100%. As used herein, the term “substantial methanol conversion” is used interchangeably with the term “full conversion of the oxygenates”, where the oxygenate is methanol.

As used herein, the terms “catalyst comprising a zeolite” and “zeolite catalyst” are used interchangeably.

It is an object of the present invention to provide a process for the conversion of oxygenates such as methanol to olefins (MTO), that is capable of producing a more efficient olefin stream as feed for oligomerization, in particular an olefin stream (olefin feed) having low aromatics and high content of higher olefins.

It is another object of the present invention to provide such an olefin feed while still maintaining full conversion of the oxygenates.

It is yet another object of the present invention to provide a process for the conversion of oxygenates with a low content of C2-light fraction, in particular C2= (ethylene), while at the same time being able to reduce the temperature of the MTO and increase the catalyst lifetime.

These and other objects are solved by the present invention.

In a broad aspect there is provided a process for producing an olefin stream, said process comprising passing a feedstock stream comprising oxygenates over a catalyst active in the conversion of oxygenates, in which the catalyst comprises a zeolite with a framework having a 10-ring pore structure, in which said 10-ring pore structure comprises a three-dimensional (3-D) pore structure, such as MFI, at a pressure of 1-50 bar and a temperature of 150-480°C.

More specifically, the present invention provides a process for producing an olefin stream, said process comprising passing a feedstock stream comprising oxygenates over a catalyst active in the conversion of oxygenates, in which the catalyst comprises a zeolite with a framework having a 10-ring pore structure, in which said 10-ring pore structure comprises a three-dimensional (3-D) pore structure, such as MFI, wherein the pressure is 2-20 bar for instance 5-10 bar, and the temperature is 150-350°C, for instance 200-300°C, or 250-350°C; or the pressure is 2-30 bar, for instance 2-20 bar or 5-10 bar, and the temperature is 340-400°C, for instance 340-385°C or 360-380°C, and wherein an olefin stream comprising C2-C3 olefins is withdrawn from said olefin stream and used as additional feed stream.

Hence, the pressure may be 2-30 bar and the temperature 150-400°C. The pressure is suitably 2-20 bar, such as 5-10 bar; or 2-30 bar, such as 2-20 bar. The temperature is suitably 150-350°C, such as 200-300°C or 250-350°C; or 340-400°C, for instance 340- 385°C or 360-380°C.

A zeolite with a framework having a 10-ring pore structure means a pore circumference defined by 10 oxygens.

A 3-D pore structure means zeolites containing intersecting pores that are substantially parallel to all three axes of the crystal. The pores preferably extend through the zeolite crystal.

The three letter code, e.g. MFI, for structure types are assigned and maintained by the International Zeolite Association Structure Commission in the Atlas of Zeolite Framework Types, which is at http:// www.iza-structure.org/databases/ or for instance also as defined in “Atlas of Zeolite Framework Types”, by Ch. Baerlocher, L.B. McCusker and D.H. Olson, Sixth Revised Edition 2007.

It would be understood that the term “temperature” means the MTO reaction temperature in an isothermal process, or the inlet temperature to the MTO in an adiabatic process.

The catalyst may be formed by combining the zeolite with a binder, and then forming the catalyst into pellets. The pellets may optionally be treated with a phosphoric reagent to create a zeolite having a phosphorous component between 0.5 and 15 wt % of the treated catalyst. The binder is used to confer hardness and strength on the catalyst. Binders include alumina, aluminum phosphate, silica, silica-alumina, zirconia, titania and combinations of these metal oxides, and other refractory oxides, and clays such as montmorillonite, kaolin, palygorskite, smectite and attapulgite. A preferred binder is an aluminum-based binder, such as alumina, aluminum phosphate, silica-alumina and clays.

It has been found that despite the relatively low temperatures used, i.e. reaction temperatures of 150-480°C, in particular 340-400°C; more particularly 360°C or below, or 350°C or below, as recited above; the catalysts are active in not only suppressing the formation of aromatics, but also in providing a high selectivity for higher olefins as well as full conversion. Particularly in the temperature range 340-400°C a significant increase in higher olefins is observed as well as a sharp decrease in aromatics content, while still fully converting the oxygenates, e.g. methanol.

Hence, the above combination of features enables the production of an olefin stream which is an ideal oligomerization feed for the further conversion to jet fuel, particularly SAF in accordance with ASTM as defined above.

While a suitable oligomerization feed may have some aromatics, for instance 10-20 wt% aromatics, as well as higher and lower olefins, the ideal oligomerization feed is namely substantially free of aromatics and composed of higher olefins, and preferably as little as possible C2- light fraction. The olefin stream may comprise at least 20 wt% C4-C8 olefins, such as above 30 wt% C4-C8 olefins and less than 10 wt% aromatics. The lower the temperature, the higher the content of higher olefins and thereby also the ratio of higher olefins to lower olefins, i.e. the selectivity to higher olefins. Thus, the oligomerization feed complies with the above ASTM requirements stipulating the 50% SAF blending part to be almost aromatic-free, more specifically that the content of aromatics be limited to below 0.5 wt%. The olefin stream can be converted into such jet fuel via oligomerization and hydrogenation in a more efficient overall process due to i.a. less recycling and higher oligomerization yields. In other words, the higher olefins and low selectivity to aromatics simplifies separation steps and increase overall yields of the jet fuel.

By using the moderately high pressure of 2-20 bar, in particular 5-10 bar, it is now possible to further shift the selectivity towards higher olefins. It has namely been found that while higher pressures increase the ratio of higher olefins to lower olefins i.e. higher selectivity to higher olefins, the higher pressures may also decrease the total yield of olefins (i.e. lower conversion of the oxygenate feed to olefins) and also increase the required temperature to achieve full conversion, which in turn creates the risk of less desired cracking reactions taking place. At the pressure range of 5-10 bar, it is now possible to obtain a higher selectivity to higher olefins without requiring increasing the temperatures to high levels for achieving full conversion, thereby also reducing the occurrence of cracking reactions.

At the same time, reducing the temperature to for instance 250°C or 300°C or 350°C, or 360°C, despite this in principle implying a reduction in methanol conversion, in fact significantly assists in the process. Without being bound by any theory, it is believed that the low acid strength of e.g. Ca/Mg-ZSM-5 suppresses the formation of aromatics. However, the low acid strength means that relatively high temperatures are necessary for achieving reasonable methanol conversions and such temperature increase would result in also increasing the rate of olefin cracking, thereby countering the effect of the increased pressure. Accordingly, by the present invention and contrary to the prior art, the pressure is increased, and the temperature lowered, resulting in that it is still possible to maintain substantial methanol conversion whilst at the same time achieving an olefin stream substantially free of aromatics and having a high content of higher olefins.

By the invention, an olefin stream comprising C2-C3 olefins is withdrawn from said olefin stream and used (recycled) as additional feed stream, e.g. by combining with the feedstock stream comprising oxygenates. Thereby, the concentration of higher olefins in the olefin stream is further increased while also having full utilization of the less desired lower olefins for conversion into higher olefins. Any undesired cracking of higher olefins in the process is contained by recycling products of such cracking, namely C2- C3 olefins, back to the feed. Furthermore, this recycle further provides a dilution effect on the feedstock stream, since light paraffins may be recycled, including methane, thereby enabling better control of the exothermicity during the conversion to olefins.

Hence, the present invention enables the production of an olefin stream with almost no C2-light fraction, in particular ethylene, and that the C2-light fraction, again particularly ethylene, will be recycled with part of the C3 fraction, particularly C3-olefin, thus said C2-C3 olefins, to the MTO to further reduce the operating temperature therein and with that also reduce the yields of ethylene and increase the lifetime. The feed for the oligomerization step is then virtually free from C2 olefins.

Hence, in an embodiment, the feedstock stream may be combined with a diluent, i. an inert diluent, such as nitrogen or carbon dioxide or a light paraffin such as methane, thereby reducing the exothermicity in the conversion to olefins, which is particularly preferred when the catalyst is arranged as a fixed bed. For instance, where the feedstock stream is methanol, it is diluted with e.g. nitrogen so that the methanol concentration in the feedstock is 2-20 vol.%, preferably 5-10 vol. %.

In an embodiment, said 3-D pore structure is MFI, such as MFI modified with an alkaline earth metal, for instance a Ca/Mg-modified ZSM-5, in particular a Ca-modified ZSM-5.

As used herein, the term “Ca/Mg-modified ZSM-5” means a ZSM-5 modified with Ca and/or Mg.

The catalysts may be prepared by standard methods in the art. For instance, Ca and/or Mg are loaded in a commercially available ZSM-5 zeolite at concentrations of 1- 10 wt.%, such as 2, 4 or 6 wt.%, by ion-exchange e.g. solid-state ion-exchange; or wet impregnation e.g. incipient wetness impregnation or any other suitable impregnation. For instance, impregnation of the final catalyst with binder/matrix, such as in a catalyst that contains up to 30-90 wt% zeolite such as 50-80 wt% zeolite in a matrix/binder comprising an alumina component such as a silica-alumina matrix binder. As an example, the catalyst is 60 wt% zeolite and 40 wt% alumina.

It would be understood that the wt% of zeolite in the binder means the wt% of the zeolite with respect to the catalyst weight, in which the catalyst comprises the zeolite and the binder.

It would also be understood, that for the purposes of the present application, the term “binder” is also referred to as “matrix binder” or “matrix/binder” or “binder/matrix”. In a particular embodiment, the weight hour space velocity (WHSV) is 0.5-12 h’¹, such as 1.5-10, or 4-10, for instance 6, 8, or 10 h’¹. In particular, at the higher values of WHSV, for instance in the range 6-10 h’¹’ despite this normally conveying a reduction of oxygenate conversion, e.g. methanol conversion, it has now also been found that the formation of aromatics may be significantly reduced while not paying a penalty in terms of methanol conversion to particularly higher olefins. Without being bound by any theory, it may appear that two consecutive repetitive cycles take place; a first cycle in which precursor olefin compounds oligomerize, and a second cycle in which higher olefins cyclize and dehydrogenate to form aromatic compounds, which are maintained as active species at the intersection of pores in the zeolite. Due to the high WHSV, i.e. low residence times, the cycles are in a way interrupted already in the first cycle, thus significantly impeding even more the further formation of aromatic compounds. Accordingly, the present invention counterintuitively invites to not only increase the pressure and reduce the temperature, but also optionally to use a high space velocity.

In an embodiment, the feedstock stream comprising oxygenates is derived from one or more oxygenates taken from the group consisting of triglycerides, fatty acids, resin acids, ketones, aldehydes or alcohols or ethers, where said oxygenates originate from one or more of a biological source, a gasification process, a pyrolysis process, Fischer- Tropsch synthesis, or methanol-based synthesis. In a particular embodiment, said one or more oxygenates are hydroprocessed oxygenates. By “hydroprocessed oxygenates” is meant oxygenates such as esters and fatty acids derived from hydroprocessing steps such as hydrotreating and hydrocracking.

In an embodiment, the oxygenates are selected from methanol (MeOH), dimethyl ether (DME), or combinations thereof. These are particularly advantageous oxygenate feedstocks, as these are widely commercially available. DME is more reactive than methanol and thus enables running the MTO step at lower temperatures, thereby increasing the selectivity for higher olefins. Furthermore, conversion of DME, releases only half the amount of water (steam) compared to methanol, thereby reducing the rate of (irreversible) deactivation due to steam-dealumination of the zeolite catalyst.

Suitably, water is removed from the olefin stream produced in the MTO, since its presence may be undesirable when conducting the downstream oligomerization. In an embodiment, the methanol is made from synthesis gas prepared by using electricity from renewable sources such as wind or solar energy, e.g. eMethanol™. Hence, the synthesis gas is prepared by combining air separation, autothermal reforming or partial oxidation, and electrolysis of water, as disclosed in Applicant’s WO 2019/020513 A1, or from a synthesis gas produced via electrically heated reforming as for instance disclosed in Applicant’s WO 2019/228797. Thereby, an even more sustainable approach for the production of jet fuel, in particular SAF, is achieved. While methanol can be produced from many primary resources (including biomass and waste), in times of low wind and solar electricity costs, the production of e-methanol™ enables a sustainable front-end solution.

In a particular embodiment, the process of the invention further comprises, prior to passing the feedstock stream comprising oxygenates over a catalyst active in the conversion of oxygenates, in which the feedstock comprising oxygenates is a methanol stream i.e. methanol feed stream: producing said methanol feed stream by methanol synthesis of a methanol synthesis gas, wherein the methanol synthesis gas is generated by: steam reforming of a hydrocarbon feed such as natural gas, and/or at least partly by electrolysis of water and/or steam.

Hence, in another particular embodiment, the methanol feed stream is produced from methanol synthesis gas which is generated by combining the use of water electrolysis in an alkaline or PEM electrolysis unit, or steam in a solid oxide electrolysis cell (SOEC) unit, thereby generating a hydrogen stream, together with the use of a CO2- rich stream in a SOEC unit for generating a stream comprising carbon monoxide and carbon dioxide, then combining the hydrogen stream and the stream comprising carbon monoxide and carbon dioxide for generating said methanol synthesis gas, as e.g. disclosed in Applicant’s co-pending European patent application No. 20216617.9. The methanol synthesis gas is then converted into the methanol feed stream via a methanol synthesis reactor, as is well-known in the art.

The methanol synthesis gas, as is also well-known in the art, is a mixture comprising mainly hydrogen and carbon monoxide tailored for methanol synthesis i.e. by the methanol synthesis gas having a module M=(H2-CO2)/(CO+CC>2). The methanol synthesis gas used for the methanol synthesis is normally described in terms of said module M, since the synthesis gas is in balance for the methanol reaction when M=2.

Thereby, an alternative highly sustainable front-end solution for generating the methanol feed stream, i.e. methanol synthesis gas, is provided, whereby only electrolysis is utilized for generating the methanol synthesis gas and thereby the methanol.

It would thus be understood, that as used herein, the term “process” may also encompass the prior (front-end) production of the methanol feed stream, as recited above.

In an embodiment, the process is conducted under the presence of hydrogen. The hydrogen improves the methanol conversion by at least slightly decreasing the rate of deactivation of the catalyst, thereby increasing catalyst lifetime. Yet, when conducting the process, there is no addition of hydrogen, since this conveys a risk of hydrogenating some olefins and thereby decrease the olefin yield.

In an embodiment, the catalyst is arranged as a fixed bed.

In an embodiment, the process comprises: using a first reactor set including a single reactor or several reactors, preferably mutually arranged in parallel, for the partial or full conversion of the oxygenates. Thereby, large feedstocks comprising one or more oxygenates can be handled simultaneously.

It would be understood, that the term “mutually” means in between the reactors of a reactor set, e.g. arranged in parallel in between the reactors of the first reactor set.

In a particular embodiment, the process further comprises using a second reactor set including a single reactor or several reactors, preferably mutually arranged in parallel, for the further conversion of the oxygenates, and a phase separation stage in between the first reactor set and the second reactor set for thereby forming the olefin stream.

As used herein, the term “using a first reactor set” means passing the feedstock comprising oxygenates through the first reactor set. As used herein, the term “using a second reactor set” means passing the feedstock or a portion thereof through the second reactor set after the partial or full conversion of the oxygenates and passage through the separation stage.

Thereby, large feedstocks comprising one or more oxygenates can be handled simultaneously and lower temperatures may be used in both reactor sets, which improves the lifetime conversion capacity of the catalyst and also improves the selectivity to higher olefins due to less cracking.

In an embodiment, the entire feedstock stream passes through the first reactor set, i.e. there is no substantial splitting of the feedstock stream.

As used herein, the term “entire feedstock” means at least 90 wt% of the feedstock.

In another particular embodiment, the process comprises:

- passing the feedstock stream comprising oxygenates through the first reactor set under conditions for partly converting, e.g. 40-80% such as 60-70% conversion, the oxygenates, thereby forming a raw olefin stream comprising unconverted oxygenates and C2-C8 olefins, e.g. the raw olefin stream may comprise water, methanol and C2-C8 olefins;

- passing the raw olefin stream through said separation stage, for producing: a first olefin stream, which is rich in lower olefins, particularly C2-C3 olefins; a separated oxygenate stream comprising the unconverted oxygenates, e.g. the separated oxygenate stream may comprise water and methanol; a second olefin stream, which is rich in higher olefins, particularly C4-C8 olefins;

- combining the first olefin stream with the separated oxygenate stream comprising the unconverted oxygenates, thereby forming a combined stream comprising lower olefins, particularly C2-C3 olefins, and the unconverted oxygenates;

- passing the resulting combined stream comprising lower olefins and unconverted oxygenates through the second reactor set, e.g. to the first reactor of the second reactor set, under conditions for fully converting, e.g. 85%, 90%, 95% or higher, the unconverted oxygenates and the lower olefins, into a third olefin stream which is rich in higher olefins, particularly C4-C8 olefins; - combining the second olefin stream (which may be regarded as a by-pass stream of the second reactor set) with the third olefin stream, thereby forming said olefin stream, which preferably is rich in higher olefins and substantially free of aromatics.

Thereby increased flexibility in operation is achieved, particularly when handling large feedstock streams, without needing to e.g. divide the feedstock stream prior to entering a first reactor for conversion of oxygenates and pass it to a separate olefin interconversion reactor, as for instance disclosed in US 5,177,279. Furthermore, the temperatures in all reactor sets can be lowered even further, for instance down to 250-350°C or 340- 400°C, yet still achieving full conversion, e.g. up to 100% conversion of the oxygenates. In addition, further lowering the temperature increases also catalyst lifetime.

Moreover, there is increased flexibility in the handling of a variety of feedstocks comprising oxygenates, including fatty acids in renewable feeds, or oxygenates originating from one or more of a biological source, as well as the handling of, optionally, different types in catalysts in the two different sets of reactors.

In another particular embodiment, the first reactor and second reactor set use a catalyst having a three-dimensional (3-D) pore structure, such as MFI.

The recycle of C2-C3 olefins as a co-feed may also be conducted. For instance, in an embodiment, the process further comprises recycling a portion of the olefin stream, i.e. the olefin product stream from the second reactor set, to said combined stream comprising lower olefins and the unconverted oxygenates and which is fed to the second reactor set, said portion of the olefin stream preferably being an olefin stream comprising C2- C3 olefins, more preferably a C3-olefin stream, which is withdrawn from said olefin stream. The same associated benefits recited above in connection with the recycling C2-C3 olefins are also obtained.

In another particular embodiment, the first reactor set consists of 2-4 reactors, such as 3 reactors, and the second reactor set consists of 1-3 reactors, such as 2 reactors. In the first and second reactor set, the reactors are preferably mutually arranged in parallel. In large MTO plants handling large feedstock streams, normally several reactors are run in parallel, e.g. five (5) reactors. By the present invention it is possible to replace the 5 reactors in parallel by for instance the first reactor set consisting of three reactors, and the second reactor set consisting of two reactors. Thereby it is possible to run at full conversion by operating the first three reactors at e.g. only 70% conversion, and then further convert the unconverted oxygenates, e.g. methanol, together the C2- C3 olefins, to 100% in two reactors arranged in series to the first three. Again, the temperature in all five reactors is lowered, yet full conversion is achieved. Flexibility is also improved, by enabling that one reactor may be taken out of service for regeneration.

It would be understood that the first reactor set and second reactor set are arranged in series.

In another particular embodiment, a reactor in the first reactor set and second reactor set operates at 2-30 bar, such as 5-15 bar, and at 150-480°C such as 150-350°C or 200-300°C.

In another particular embodiment, the weight hour space velocity (WHSV) is 0.5-12 h-

1 , such as 1.5-10 h-1, or 4-10 h-1 , for instance 6, 8, or 10 h-1. In yet another particular embodiment, the weight hour space velocity (WHSV) in the first reactor set is higher than in the second reactor set. For instance, in the first reactor set where partial conversion of the oxygenate feedstock is intended, the WHSV is suitably 3 IT¹ or 6h'¹ while in the second reactor set where full conversion is intended the WHSV is suitably 2 h’¹.

In an embodiment, the process further comprises: passing at least a portion of the olefin stream trough an oligomerization step over an oligomerization catalyst, and optionally subsequently conducting a separation step, for thereby producing an oligomerized stream.

In an embodiment, the entire olefin stream passes through the oligomerization step, preferably after said olefin stream comprising C2-C3 olefins is withdrawn from the olefin stream. As used herein, the term “entire olefin stream” means at least 90 wt% of the stream.

In an embodiment, the olefin stream, e.g. the entire olefin stream, is passed directly to the oligomerization step, i.e. the olefin stream is in direct fluid communication with the oligomerization step, or combined oligomerization and hydrogen step, as explained farther below. Thereby, there is no fractionation of the olefin stream prior to entering the oligomerization step, thus further simplifying the process and plant.

The oligomerization step is preferably conducted by conventional methods including the use of an oligomerization catalyst such as solid phosphoric acid (“SPA”), ion-ex- change resins or a zeolite catalyst, for instance a conventional *MRE, BEA, FAU, MTT, TON, MFI and MTW catalyst, at a pressure of 30-100 bar, such as 50-100 bar, and a temperature of 100-350°C. The products from the oligomerization reaction may be subsequently separated in the separation step, such as distillation, thereby withdrawing a lighter hydrocarbon stream such as naphtha, which comprises C5-C7 hydrocarbons, and the oligomerized stream, which comprises C8+ hydrocarbons.

In an embodiment, the process further comprises: passing at least a portion of the oligomerized stream through a hydrogenation step over a hydrogenation catalyst, and optionally subsequently conducting a separation step, for thereby producing a hydrocarbon stream comprising hydrocarbons boiling in the jet fuel range.

The hydrogenation step is preferably conducted by conventional methods, including under the presence of hydrogen the use of a hydrotreating or hydrogenation catalyst, for instance a catalyst comprising one or more metals, e.g. Pd, Rh, Ru, Pt, Ir, Re, Co, Mo, Ni, W or combinations thereof, at a pressure of 60-70 bar and a temperature of 50- 350°C. The C8+ hydrocarbons of the oligomerized stream are thereby saturated to form the corresponding paraffins. These may be subsequently separated in a separation step, for instance a distillation step, whereby any hydrocarbons boiling in the diesel range are withdrawn and thereby separated from the hydrocarbons boiling in the jet fuel range i.e. jet fuel.

In an embodiment, the entire oligomerized stream passes through the hydrogenation step. As used herein, the term “entire oligomerized stream” means at least 90 wt% of the stream. In a particular embodiment, the hydrocarbon stream comprising hydrocarbons boiling in the jet fuel range is SAF, i.e. a sustainable aviation fuel in compliance with ASTM D7566 and ASTM D4054.

In an embodiment of the invention, the oligomerization step and hydrogenation step are combined in a single hydro-oligomerization step (Hydro-OLI), e.g. by combining the steps in a single reactor. In other words, by passing at least a portion of the olefin stream trough an oligomerization step and hydrogenation step which are combined in a single hydro-oligomerization step, and optionally subsequently conducting a separation step, for thereby producing a hydrocarbon stream comprising said hydrocarbons boiling in the jet fuel range. This results in a much simpler process/plant layout.

As used herein, the term “single hydro-oligomerization step” or more generally “single step” or “single stage” means a section of the process in which no stream is withdrawn. Typically, a single stage does not include equipment such as compressors, by which the pressure is increased.

In an embodiment of the invention, the oligomerization step is dimerization, optionally also trimerization, i.e. by conducting the oligomerization at conditions suitable for dimerization and/or trimerization. Thereby the single reactor is preferably operated at a relatively low pressure, such as 15-60 bar, for instance 20-40 bar. The oligomerization reaction is very exothermic per oligomerization step and much less heat is produced, ,- since there is only dimerization, optionally also trimerization -, instead of higher oligomerization. The lower heat produced favors approaching equilibrium, i.e. higher conversion of olefins.

Normally, the oligomerization step converts the olefins to a mixture of mainly dimers, trimers and tetramers; for instance, a C6-olefin will result in a mixture comprising C12, C18, C24 products and probably also higher hydrocarbons. By conducting the oligomerization step at conditions suitable for dimerization, a more selective and direct conversion of the higher olefins (C4-C8 olefins) to the jet fuel relevant hydrocarbons, namely C8-C16, is obtained. The dimerization and optional trimerization step comprises the use of lower pressures than in conventional oligomerization processes, thereby also reducing compression requirements which translates into higher energy efficiency as well as reduced costs, e.g. reduced costs of the oligomerization reactor and attendant equipment, as well as reduced operating costs due to less need of separating C16+ olefins otherwise formed in conventional OLI reactors. Accordingly, the pressure of the Hydro/OLI can be adapted to better match the pressure of the previous oxygenate conversion step.

Moreover, instead of using a dedicated separation such as distillation in the OLI step for separating naphtha and another dedicated separation in the hydrogenation step for separating diesel from the jet fuel, only one subsequent separation stage, if any, will be needed. Thereby a simpler process for oligomerization and hydrogenation is obtained and consequently also a simpler overall process and plant.

The hydrogenation or ^-addition is conducted in the same reactor, for instance by adjusting the activity of the hydrogenation component e.g. nickel. In an embodiment, the single hydro-oligomerization step is conducted in a single reactor having a stacked reactor bed where a first bed comprises an oligomerization catalyst, e.g. zeolite catalyst, and a subsequent bed comprises a hydrogenation catalyst.

In an embodiment of the invention, the hydro-oligomerization step is conducted by reacting, under the presence of hydrogen, the olefin stream over a catalyst comprising a zeolite and a hydrogenation metal, such as a hydrogenation metal selected from Pd, Rh, Ru, Pt, Ir, Re, Co, Cu, Mo, Ni, W and combinations thereof, and preferably at a pressure of 15-60 bar such as 20-40 bar, and a temperature of 50-350°C, such as 100- 250°C. In a particular embodiment, the catalyst comprises a zeolite having a structure selected from MFI, MEL, SZR, SVR, ITH, IMF, TUN, FER, EUO, MSE, *MRE, MWW, TON, MTT, FAU, AFO, AEL, and combinations thereof, preferably a zeolite with a framework having a 10-ring pore structure i.e. pore circumference defined by 10 oxygens, such as zeolites having a structure selected from TON, MTT, MFI, *MRE, MEL, AFO, AEL, EUO, FER, and combinations thereof. These zeolites are particularly suitable due to the restricted space of the zeolite pores, thereby enabling that the dimerization is favored over larger molecules. Optionally, the weight hour space velocity (WHSV) is 0.5-6 h’¹, such as 0.5-4 h’¹. Lower pressures corresponding to the operating at conditions for dimerization are in particular 15-50 bar, such as 20-40 bar. This, again, is significantly lower than the pressures normally used in oligomerization, which typically are in the range 50-100 bar.

The present invention purposefully uses conditions that result in a mild hydrogenation. Particularly suitable catalysts are catalysts comprising NiW, for instance sulfide NiW (NiWS), or Ni such as Ni supported on a zeolite having a FAU or MTT structure, for instance a Y-zeolite, or ZSM-23. The catalyst which is active for oligomerization and hydrogenation may for instance contain up to 50-80 wt% zeolite in a matrix/binder comprising an alumina component. The hydrogenation metal may then be incorporated by impregnation on the catalyst. The hydrogenation metals are selected so as to provide a moderate activity and thereby better control of the exothermicity of the oligomerization step by mainly hydrogenating the dimers being formed as the oligomerization takes place, thereby interrupting the formation of higher oligomers.

Hence, rather than having separate reactors and attendant separation units for conducting oligomerization and subsequent hydrogenation, each with its own catalyst, the present invention enables in a single hydro-oligomerization step the use of less equipment e.g. one single reactor, one type of catalyst, optionally a single separation stage downstream for obtaining the jet fuel. A more efficient and simpler overall process and plant for the conversion of oxygenates such as methanol to jet fuels, particularly SAF, is thereby achieved.

In an embodiment, a stream comprising C8-hydrocarbons resulting from cracked COCI 6 hydrocarbons, is withdrawn from said hydrocarbon stream comprising hydrocarbons boiling in the jet fuel range and added to other processes. For instance, the process according to the first aspect of the invention cooperates with a refinery plant (or process), in particular a bio-refinery, and the stream comprising C8-hydrocarbons is added to the gasoline pool in a separate process for producing gasoline of said refinery. Optionally, a stream comprising C8- hydrocarbons resulting from cracked C9-C16 hydrocarbons, is withdrawn from said hydrocarbon stream comprising hydrocarbons boiling in the jet fuel range and used (recycled) as additional feed stream to the oligomerization step or the single hydro-oligomerization step. BRIEF DESCRIPTION OF THE DRAWINGS

Fig. 1 is a simplified figure showing the conversion of oxygenates to olefins and optional further conversion to jet fuel in accordance with an embodiment of the invention.

Fig. 2 is a simplified figure showing a particular embodiment of the invention for the conversion a feedstock comprising oxygenates to olefins and optional further conversion to jet fuel.

Fig. 3 shows plots of methanol conversion (upper chart) and aromatics content (lower chart) in the olefin stream as a function of the temperature in degrees Celsius.

Fig. 4 shows plots of weight ratio of higher olefins to lower olefins (upper chart) and the weight content of higher olefins in the olefin stream as a function of the temperature in degrees Celsius.

DETAILED DESCRIPTION

With reference to Fig. 1 , a feedstock comprising oxygenates 100, such as methanol and/or DME, is directed together with an optional hydrogen stream 102 and an olefin stream 104 comprising C2-C3 olefins which is withdrawn from the olefin stream 106 formed in oxygenate conversion section 200. The oxygenate conversion section 200, for instance a MTO section, converts the oxygenates over a zeolite catalyst such as Ca-modified-ZSM-5 at e.g. 5-15 bar and 150-350°C or 340-400°C. The resulting olefin stream 106 at these conditions is rich in higher olefins (C4=-C8=) and low in aromatics and is optionally further converted (as shown by stippled lines) to a hydrocarbon stream 112 comprising hydrocarbons boiling in the jet fuel range (C8-C16).

This further conversion is conducted in downstream oligomerization and hydrogenation section 300, which preferably is combined as a single hydro-oligomerization step, for instance in a single reactor. The olefin stream 106, suitably after removing its water content, is mixed with optional oligomerization olefin stream 110 comprising C8- hydrocarbons and resulting from cracked C9-C16 hydrocarbons withdrawn from said hydrocarbon stream 112 comprising hydrocarbons boiling in the jet fuel range. The resulting mixed stream is then directed to section 300 and converted, under the presence of hydrogen being fed as stream 108, over a catalyst such as Ni supported on a zeolite having a FAU or MTT structure, for instance Y-zeolite, or ZSM-23, at e.g. 20-40 bar and 50-350°C, to the hydrocarbon stream 112 comprising hydrocarbons boiling in the jet fuel range. At these conditions, particularly the lower pressures, the single reactor in section 300 operates such that the oligomerization is dimerization and at the same time there is hydrogenation activity. Due to the higher olefins and low aromatics of the olefin stream 106, the hydrocarbons in stream 112 boiling in the jet fuel range i.e. jet fuel, can be used as SAF.

With reference to Fig. 2, a feedstock stream 100 comprising oxygenates such as methanol and/or DME passes through a first reactor set 200’, for instance three reactors arranged in parallel, for thereby achieving 50-70% conversion of the methanol and producing a raw olefin stream 105 comprising water, methanol and C2-C8 olefins. The raw olefin stream 105 is subjected to separation in 3-phase separator 200” thereby producing a first olefin stream 105a, which is rich in lower olefins, particularly C2-C3 olefins, a separated oxygenate stream 105b comprising the unconverted oxygenates, e.g. unconverted methanol, and a second olefin stream 105c which is rich in higher olefins, particularly C4-C8 olefins, e.g. having about 60-70% C4-C8 olefins. The first olefin stream 105a is combined with the separated oxygenate stream 105b comprising the unconverted oxygenates, thereby forming a combined stream 105d comprising lower olefins, particularly C2-C3 olefins, and the unconverted oxygenates. This combined stream is pressurized and fed to a second reactor set 200”’ arranged downstream, and which may for instance include two reactors arranged in parallel, for thereby achieving full conversion e.g. 85% or 90% or higher. The first reactor set 200’ and second reactor set 200’” are thereby arranged in series. A third olefin stream 105e is produced which is rich in higher olefins, particularly C4-C8 olefins. Finally, the second olefin stream 105c (bypass stream) is combined with the third olefin stream 105e, thereby forming said olefin stream 106 which may have been pressurized. By the above arrangement of the MTO section 200, the rectors of the first and second set can be operated at low temperature, e.g. 250-350°C or 340-400°C, preferably at a lower temperature than when using the embodiment of Fig. 1 , which improves the life-time conversion capacity of the catalysts used and improve the selectivity to higher olefins due to less cracking. The resulting olefin stream 106, suitably after removing its water content, is optionally further converted (as shown by the stippled lines) to the hydrocarbon stream 112 comprising hydrocarbons boiling in the jet fuel range (C8-C16), particularly SAF, as explained in connection with Fig. 1.

EXAMPLE

Tests were run in a fixed catalyst bed (fixed bed) reactor with the zeolite catalyst Ca- ZSM-5, having a 3-D pore structure, and at the following operating conditions: zeolite catalyst load: 250 mg cat/250 mg SiC, pressure = 10 barg, space velocity (WHSV)= 4 h’¹, total flow = 7.0 NL/h (117 mL/min); methanol concentration in the feed (CM₆OH) = 10% (volume basis) with nitrogen as the diluent. The temperature used is in the range 320-480°C, with tests running in the order 480-440-400-360-320°C, and subsequently in reverse order in order to evaluate the effect of any catalyst deactivation. Note: pressures are in barg, i.e. absolute pressure minus atmospheric pressure.

Fig. 3 shows in the upper chart the methanol conversion as a function of the temperature. It is observed, that already at 340°C, there is 90% or more conversion, and at 360°C, there is 100% conversion. Aromatics are formed, as shown in the lower chart of Fig. 3, the level of which increases with temperature, yet there is low selectivity towards formation of aromatics, which are maintained at a low level of about 10 wt% and below 10 wt% throughout or even well below 5 wt% at about 350°C or lower temperatures. In the upper chart, the curve marked by squares (CONV*k) shows the results as the temperature is changed from 480°C to 320°C and the curve marked with traingles (CONV^) when the temperature subsequently is changed from 320 to 480°C. In the lower chart, the curves are almost indistinguishable.

Fig. 4 shows in the upper chart the selectivity (OH/OL) towards higher olefins as a function of the temperature. The curves, as the temperature is changed from 480°C to 320°C (squares, OH/OL'k) and from 320 to 480°C (triangles, OH/OL^) are indistinguishable in the upper chart. It is observed, that there is a significant increase in the ratio of higher to lower olefins as the temperature is decreased and particularly in the range 340-400°C. While at 450 or 500°C the ratio is about 1 or below, at 400°C the ratio is already above 2 and at about 350°C close to 3. In the lower chart, the concentration of higher olefins C4+ (04+) is observed to increase as the temperature is lowered. At 400°C the content of higher olefins is significantly higher than the level achieved at 450°C or 500°C, and at about 360°C or 350°C a maximum is obtained, with a higher olefins C4+ (04+) content close to 25 wt%.

Claims

25 CLAIMS

1. A process for producing an olefin stream, said process comprising passing a feedstock stream comprising oxygenates over a catalyst active in the conversion of oxygenates, in which the catalyst comprises a zeolite with a framework having a 10-ring pore structure, in which said 10-ring pore structure comprises a three-dimensional (3-D) pore structure, such as MFI, wherein the pressure is 2-20 bar for instance 5-10 bar, and the temperature is 150-350°C, for instance 200-300°C, or 250-350°C; or the pressure is 2-30 bar, for instance 2-20 bar or 5-10 bar, and the temperature is 340-400°C, for instance 340-385°C or 360-380°C; and wherein an olefin stream comprising C2-C3 olefins is withdrawn from said olefin stream and used as additional feed stream.

2. Process according to claim 1 , wherein said 3-D pore structure is MFI, such as MFI modified with an alkaline earth metal, for instance a Ca/Mg-modified ZSM-5, in particular a Ca-modified ZSM-5.

3. Process according to any of claims 1-2, wherein the weight hour space velocity (WHSV) is 0.5-12 h’¹, such as 1.5-10, or 4-10, for instance 6, 8, or 10 h’¹.

4. Process according to any of claims 1-3, wherein the feedstock stream comprising oxygenates is derived from one or more oxygenates taken from the group consisting of triglycerides, fatty acids, resin acids, ketones, aldehydes or alcohols or ethers, where said oxygenates originate from one or more of a biological source, a gasification process, a pyrolysis process, Fischer-Tropsch synthesis, or methanol-based synthesis.

5. Process according to any of claims 1-4, wherein the oxygenates are selected from methanol (MeOH), preferably methanol made from synthesis gas prepared by using electricity from renewable sources such as wind or solar energy; dimethyl ether (DME); or combinations thereof.

6. Process according to any of claims 1-5, wherein the process is conducted under the presence of hydrogen.

7. Process according to any of claims 1-6, comprising: using a first reactor set including a single reactor or several reactors, preferably mutually arranged in parallel, for the partial or full conversion of the oxygenates.

8. Process according to any of claims 1-8, further comprising using a second reactor set including a single reactor or several reactors, preferably mutually arranged in series, for the further conversion of the oxygenates, and a phase separation stage in between the first reactor set and the second reactor set, for thereby forming the olefin stream.

9. Process according to any of claims 7-8, comprising:

- passing the feedstock stream comprising oxygenates through the first reactor set under conditions for partly converting the oxygenates, thereby forming a raw olefin stream comprising unconverted oxygenates and C2-C8 olefins;

- passing the raw olefin stream through said separation stage, for producing: a first olefin stream, which is rich in lower olefins, particularly C2-C3 olefins; a separated oxygenate stream comprising the unconverted oxygenates; a second olefin stream, which is rich in higher olefins, particularly C4-C8 olefins;

- passing the resulting combined stream comprising lower olefins and the unconverted oxygenates through the second reactor set under conditions for fully converting the unconverted oxygenates and the lower olefins, into a third olefin stream which is rich in higher olefins, particularly C4-C8 olefins;

- combining the second olefin stream with the third olefin stream, thereby forming said olefin stream.

10. Process according to any of claims 1-9, further comprising: passing at least a portion of the olefin stream trough an oligomerization step over an oligomerization catalyst, and optionally subsequently conducting a separation step, for thereby producing an oligomerized stream.

11. Process according to claim 10, wherein the olefin stream is passed directly to the oligomerization step.

12. Process according to any of claims 10-11 , further comprising: passing at least a portion of the oligomerized stream through a hydrogenation step over a hydrogenation catalyst, and optionally subsequently conducting a separation step, for thereby producing a hydrocarbon stream comprising hydrocarbons boiling in the jet fuel range.

13. Process according to any of claims 10-12, wherein the oligomerization step and hydrogenation step are combined in a single hydro-oligomerization step (Hydro-OLI), wherein the oligomerization step is dimerization, optionally also trimerization, by the the hydro-oligomerization step being conducted by reacting, under the presence of hydrogen, the olefin stream over a catalyst comprising a hydrogenation metal, such as a hydrogenation metal selected from Pd, Rh, Ru, Pt, Ir, Re, Cu, Co, Mo, Ni, W and combinations thereof, preferably incorporated in a zeolite, and preferably at a pressure of 15-60 bar such as 20-40 bar, and a temperature of 50-350°C, such as 100-250°C.