DK181970B1

DK181970B1 - Revamping of a gasoline plant into a jet fuel plant

Info

Publication number: DK181970B1
Application number: DKPA202330217A
Authority: DK
Inventors: Dan Palis Sørensen Martin; Neil Burn Jeremy; Thi Minh Nguyen Thoa
Original assignee: Topsoe As
Priority date: 2023-09-19
Filing date: 2023-09-19
Publication date: 2025-04-29
Also published as: DK202330217A1; WO2025061653A1

Abstract

The invention relates to a method of revamping an existing methanol-to-gasoline (MTG) plant into a methanol-to-jet-fuel (MTJ) plant; the method comprises converting the MTG loop section of the MTG plant into a MTH-loop section by at least replacing the catalyst of said MTG reactor (110) with a catalyst active in the conversion of methanol to hydrocarbons boiling in the jet fuel range, such as a catalyst comprising a zeolite in which said zeolite is any of *MRE (ZSM-48), MTT (ZSM-23), TON (ZSM-22), or combinations thereof, thereby said MTG reactor(110) being converted into a first set (R1) of one or more adiabatic fixed bed reaction zones; in which R1 operates as an oxygenateto-olefins reaction zone, such as a methanol-to-olefins (MTO) reaction zone, e.g. a MTO reactor.

Description

DK 181970 B1 1

Revamping of a gasoline plant into a jet fuel plant

FIELD OF THE INVENTION

The present invention relates to the revamping of an existing methanol to gasoline plant into an oxygenate to jet fuel plant, such as a methanol to jet fuel plant.

BACKROUND

To limit the continued increase of global heating and consequently increasing world temperatures, actions must be taken to reduce human consumption of fossil fuels. Cost reductions of renewable energy enable decarbonization of certain specific applications, however aviation is one of the sectors where direct electrification is not feasible. What is needed are methods that efficiently convert carbon neutral feed into distillate boiling range hydrocarbons, here in particular jet fuels.

Conventionally, the conversion of methanol or other oxygenates into distillate boiling range hydrocarbons, such as jet fuel, is conducted in two steps and with no active reg- ulation of the oxygenate slip to control the process. First, the oxygenates are converted into light (short) olefins, typically ethylene, propylene, and butylene, and some other light or intermediate olefins (with carbon number < C9). This first step is conducted at low reactant partial pressures, low residence times (high space velocities), and high temperature, e.g. above 370°C, to facilitate the production of short olefins i.e. olefins with low carbon numbers. In the second step, the stream containing such intermediate olefins is pressurized, before it is fed into a second reactor system operating at an ele- vated pressure and lower temperature to facilitate carbon chain growth. The second re- actor is typically an oligomerization reactor (OLI reactor) where the olefins produced upstream in a methanol to olefins reactor (MTO reactor) are converted into C8-C19 such C8-C17 or C8-C16 hydrocarbons, which correspond to the jet fuel boiling range.

A subsequent hydroprocessing step is typically also conducted. Hence, at least two se- quential steps are required to produce the hydrocarbons in the jet fuel boiling range.

Applicant's co-pending European patent application EP22214261.4 discloses a pro- cess and plant for producing hydrocarbons such as jet fuel.

DK 181970 B1 2

Applicant's WO 2022063992 discloses in the examples oxygenate conversion (metha- nol to olefins, MTO) conducted with ZSM-48 in the temperature range 320-480°C, with the proportion of C8-C17 olefins produced therein being lower than 10 wt%.

Applicant's WO 2022063994 discloses the conversion of an olefin stream to hydrocar- bons boiling in the jet fuel range in a combined oligomerization and hydrogenation step.

In a particular embodiment, a process for oxygenate conversion to said olefin stream using a ZSM-48 zeolite, is disclosed.

Applicant's US 11060036 (corresponding to WO 2019219397) relates to gasoline pro- duction in a fixed bed reactor.

It would be desirable to be able to in a simple manner, with lowest number of modifica- tions, to convert an existing plant for producing gasoline from methanol into a plant for producing jet fuel.

SUMMARY OF THE INVENTION

Accordingly, the present invention is a method of revamping a methanol-to-gasoline (MTG) plant, i.e. an existing MTG plant, into a methanol-to-jet fuel (MTJ) plant (100), the MTG plant comprising: - a methanol to gasoline synthesis loop section (MTG-loop section) arranged to receive a methanol feed stream (101) and provide a raw gasoline stream (111); said MTG loop section comprising: a MTG reactor (110) of one or more adiabatic fixed bed reaction zones comprising a catalyst active in the conversion of methanol to gasoline as a first raw gasoline stream (109); one or more heaters for heating one or more feed streams of the methanol feed stream (101) to an inlet temperature of said MTG reactor (110), one or more conduits for introducing the one or more heated feed streams into inlet of said MTG reactor (110); a product separation unit (112) arranged to receive the first raw gasoline stream (109) and provide: an overhead recycle stream (115, 115”), a pro- cess condensate (113) comprising water, and a raw gasoline stream (111); - a distillation section arranged to receive said raw gasoline stream (111) and provide a stabilized gasoline stream;

DK 181970 B1 3 - optionally, an upgrading section comprising any of a hydroisomerization (HDI) section and/or hydrocracking (HCR) section, arranged to receive said stabilized gasoline stream and provide a gasoline product; the method comprising: - converting said MTG-loop section into a MTH-loop section (100°) by at least replacing the catalyst of said MTG reactor (110) with a catalyst active in the conversion of metha- nol to hydrocarbons boiling in the jet fuel range, the catalyst comprising a zeolite with a framework having a 10-ring pore structure, said 10-ring pore structure being a unidi- mensional (1D) pore structure, such as a catalyst comprising a zeolite in which said ze- olite is any of *MRE (ZSM-48), MTT (ZSM-23), TON (ZSM-22), or combinations thereof, thereby said MTG reactor (110) being converted into a first set (R1) of one or more adiabatic fixed bed reaction zones; in which R1 operates as an oxygenate-to-ole- fins reaction zone, such as a methanol-to-olefins (MTO) reaction zone, e.g. a MTO re- actor; and thereby withdrawing from said separation unit (12, 112) a converted-oxygen- ate product (11) comprising hydrocarbons boiling in the jet fuel boiling range.

It would be understood that instead of methanol (MeOH), the oxygenate dimethyl ether (DME), or a mixture of MeOH and DME, may be provided as the feed in the MTH-loop section.

The revamping of an existing plant for producing gasoline into an improved version of the plant for producing the same product, i.e. gasoline, is not trivial, and much less so the revamping of an existing gasoline plant where the intended product is completely different, more specifically jet fuel.

Surprisingly, it has been found that producing jet fuel occurs at similar conditions as for gasoline synthesis, thus enabling that existing process/plant layout and design condi- tions, including design temperature, design pressure, flow rates, material, for gasoline production can be directly applied to jet fuel production, with a surprisingly high propor- tion of jet fuel hydrocarbons being produced in the revamped MTG reactor (R1), which are withdrawn downstream as said converted-oxygenate product. Thereby, with mini- mal changes, an existing MTG plant is converted i.e. transformed into a MTJ plant, thus significantly also further increasing the commercial viability of the invention. Fig. 1- 2 show how the existing MTG-loop in Fig. 2 is with minor modifications revamped into

DK 181970 B1 4 the MTH-loop 100’ of Fig. 1 and optionally further revamped by installing an OLI-loop 100” section.

For instance, the MTG loop section comprises already a MTG reactor 110 with a fixed bed of a catalyst comprising a 3D zeolite such as ZSM-5 as said catalyst active in the conversion of methanol to gasoline, which is easily replaced by e.g. ZSM-48. Not only are jet fuel hydrocarbons produced, but also in a surprisingly high proportion and fur- ther with a surprisingly high proportion thereof as olefins. The MTG loop section com- prises already a product separation unit 112 as well as recycle compressor 114 for supplying the overheat recycle stream 115, 115’ to the MTG reactor. The MTG loop section comprises also already associated equipment for the cooling of the first raw gasoline stream from the MTG reactor upstream the product separation unit 112, such feed/effluent heat exchangers and air cooler(s), all of which can be directly applied for the MTH-loop section.

The MTJ plant is thus quick to design, install and operate with associated low capital expenses (CAPEX) and low operating expenses (CAPEX) compared to when providing a MTJ plant from scratch, while also being capable of operating with a large production of jet fuel product and thus large plants such as for producing up to 15000 bpd (barrels per day) of jet fuel product; or very large plants producing at least 15000 bpd of jet fuel product, the latter suitably by further incorporating an oligomerization (OLI)-loop sec- tion. For instance, an existing MTG plant of already large capacity, producing 10000- 15000 bpd of gasoline, may be expediently revamped into MTJ plant of similar capacity (1000-15000 bpd jet fuel product), or into a very large MTJ plant producing said at least 150000 bpd jet fuel product, such as 15000-25000 bpd of jet fuel product.

As is well-known in the art, the term "loop section”, herein also simply referred to as "loop” means a process or section comprising a reactor producing a raw product with subsequent separation from the raw product in a product separation unit, from which a recycle stream comprising light hydrocarbons, such as C2-C4 hydrocarbons, is with- drawn and sent via a recycle compressor to the reactor.

Further, for the purposes of the present application:

DK 181970 B1

The term “comprising” includes “comprising only” i.e. “consisting of”.

The term “suitably” is used interchangeably with the term “optional” i.e. an optional em- bodiment.

The term "process/plant” means process and/or plant. 5 The term "MeOH-loop” means methanol synthesis loop” and is a specific embodiment of a methanol synthesis section. The term "OTH-loop” means oxygenate-to-hydrocarbons loop. The term "MTH-loop” means methanol-to-hydrocarbons loop and is a specific em- bodiment of an OTH-loop, thus the oxygenate being e.g. methanol. The term “MTO” means methanol-to-olefins. The term "OLI-loop” means oligomerization loop. The terms "OTH-loop” or “MTH-loop” or "OLI-loop” may be used interchangeably with the terms "OTH-loop section” or “MTH-loop section” or "OLI-loop section”, respectively. The use of the term "section” includes also the provision of process steps therein.

The term "loop” is thus used interchangeably with the term "loop section”, and as already described above, means a process or section comprising a reactor producing a raw prod- uct with subsequent separation of the raw product in a product separation unit into at least a recycle stream comprising light hydrocarbons which is sent via a recycle com- pressor to the reactor.

The term “and/or” means in connection with a given embodiment any of three options.

The term “and/or” may be used interchangeably with the term “at least one of” the three options.

R1 means the OTH-reactor of the OTH-loop, such as the such as the MTH-reactor (R1), e.g. MTO-reactor of the MTH-loop section. R2 means the oligomerization reactor (OLI reactor) of the OLI-loop section. The term “R1” is used when referring to the OTH-loop section or to the MTH-loop section. Again, the MTH-loop section is a specific embodi- ment of the OTH-loop section.

The term MTG means methanol-to-gasoline. Hence, the term MTG-reactor means meth- anol-to-gasoline reactor, and the MTG-loop section (interchangeably, MTG-loop) means methanol-to-gasoline loop section.

DK 181970 B1 6

The term “MTG plant” means methanol-to-gasoline plant. It would be understood that the term “MTG-loop section” is a specific section of a MTG-plant. The term “MTJ plant” means methanol-to-jet fuel plant.

The term “oxygenate compounds” is used interchangeably with the term "oxygenates”.

For instance, an oxygenate compound, i.e. an oxygenate, is methanol (MeOH). For in- stance, an oxygenate compound, i.e. an oxygenate, is dimethyl ether (DME).

The term “present invention” or simply “invention” are used interchangeably with the term “present application” or simply “application”.

The term “reaction zone” means a physically delimited space where a catalytic reaction takes place and thus comprising a catalyst. For instance, an adiabatic fixed bed, or a reactor comprising an adiabatic fixed bed.

The term “distillate boiling range hydrocarbons” or “distillate boiling range hydrocarbon product” means, for the purposes of the present application, C5-C30 hydrocarbons and comprises hydrocarbons boiling in the naphtha boiling range, hydrocarbons boiling in the jet fuel boiling range, hydrocarbons boiling in the diesel boiling range; optionally, a heavy hydrocarbon fraction i.e. maritime fuel.

The term “hydrocarbons boiling in the gasoline boiling range” may be used interchange- ably with the term “gasoline” and means C5-C12 hydrocarbons boiling in the range 30- 210°C.

The term “hydrocarbons boiling in the naphtha boiling range may be used interchange- ably with the term “naphtha” and means C5-C9 hydrocarbons boiling in the range 30- 160°C, such C5-C8 hydrocarbons, e.g. C5-C8 olefins. The term “naphtha” is sometimes used interchangeably with the term “naphtha stream”.

The term “hydrocarbons boiling in the diesel boiling range” may be used interchangeably with the term “diesel” and means C8-C25 hydrocarbons boiling in the range 120-360°C, for instance 160-360°C.

The term “hydrocarbons boiling in the jet fuel range” may be used interchangeably with the term “jet fuel hydrocarbons” or “jet fuel range hydrocarbons”, or respectively, “jet fuel” or “jet fuel range”. The term means C8-C19 hydrocarbons, such as C8-C17 or C8-C16

DK 181970 B1 7 hydrocarbons, boiling in the range 130-300°C. The term “C8-C19” hydrocarbons is also used interchangeably with the term "C8+ hydrocarbons”. The term "C13+ hydrocarbons” means C13-C19 hydrocarbons, such as C13-17 hydrocarbons. For instance, the jet fuel is sustainable aviation fuel (SAF) in compliance with ASTM D7566 and ASTM D4054.

For instance, the jet fuel is in compliance with ASTM D7566.

The term boiling in a given range, shall be understood as a hydrocarbon mixture of which at least 80 wt% boils in the stated range.

The term “constant” is used interchangeably with the term “stable” and means within 10% of a given flow level, such as within 10% of a given flow rate.

Unless otherwise stated, when percentages are provided for a given stream it is meant vol.%. It would be understood that vol.% are normally used for gas streams, while wt% are normally used for liquid streams.

The term “at least a portion” of a stream means a portion of the stream or the entire stream.

The article “a” or “an” in connection with an item, such as a unit or process step, means “one or more”. For instance, a fractionation unit in said separation section means one or more fractionating units, such as one or more distillation units.

The term “conduit” means a process line such as a pipe carrying a given process stream.

Oher definitions are provided in connection with one or more of above or below embod- iments.

In an embodiment, the method of revamping further comprises: in said MTH-loop section (100): - converting in the one or more adiabatic fixed bed reaction zones of R1 the one or more heated feed streams to a raw converted-oxygenate product (9); suitably, said raw converted-oxygenate product (9) comprising C5+ hydrocarbons of which at least 20 wt%, or at least 30 wt% are said hydrocarbons boiling in the jet fuel boiling range; or at least 20 wt%, or at least 30 wt% of said converted-oxygenate product (9) are said hydrocarbons boiling in the jet fuel range;

DK 181970 B1 8 - withdrawing from the one or more adiabatic fixed bed reaction zones the raw con- verted-oxygenate product (9); - separating in said separation unit (12) the raw converted-oxygenate product (9) into: an overhead recycle stream (15, 15’) as overhead recycle stream of the MTH-loop sec- tion (100°); a process condensate (13) comprising water; and said converted oxygenate product (11); - recycling said overhead recycle stream (15, 15°) to said first set (R1) of one or more adiabatic fixed bed reaction zones, for instance by admixing with said feed stream (1) of one or more oxygenate compounds.

It would be understood that the recycling of the overhead recycle stream is provided by a recycle compressor, i.e. a dedicated recycle compressor for the MTH-loop section, and this recycle compressor is the existing recycle compressor of the MTG plant. This overhead recycle stream is also referred to as “overhead recycle stream of the MTH- loop section” or simply “recycle stream of the MTH-loop”. The recycle compressor is a highly costly, not only in terms of capital expenses (CAPEX) but also in terms of oper- ating expenses (OPEX), so not needing to replace this equipment is highly advanta- geous.

This overhead recycle stream suitably comprises light hydrocarbons, such as C4- hy- drocarbons, thus providing dilution of the feed stream to R1 for control of the exother- micity therein. Suitably, the C4- hydrocarbons, i.e. the C4-hydrocarbon fraction, is rich in C1-C4 paraffins and/or olefins.

The oxygenate conversion in R1 is operated in a dynamic manner that maintains both oxygenate conversion and yields of hydrocarbons boiling in the jet fuel range stable over time, irrespectively of catalyst deactivation. By converting an oxygenate such as methanol over the conversion catalyst having a 10-ring 1D pore structure zeolite and adjusting the reactor inlet temperature to maintain a certain oxygenate (e.g. methanol) slip at outlet R1, there is a constant, i.e. stable, and high yield of hydrocarbons boiling in the jet fuel boiling range. Astonishingly, at least 30 wt% of the total C5+ hydrocar- bons produced in the raw converted-oxygenate product, or at least 30 wt% such as at least 40 wt% of the raw converted-oxygenate product, are already hydrocarbons boiling in the jet fuel boiling range, i.e. C8+ hydrocarbons, such as C8-C17 hydrocarbons. For

DK 181970 B1 9 instance, at least 50 wt% of the total C5+ hydrocarbons produced are already C8-C17 hydrocarbons i.e. a C8-C17 hydrocarbon fraction. For instance, at least 45 wt% of the raw converted-oxygenate product are already C8-C17 hydrocarbons i.e. a C8-C17 hy- drocarbon fraction. Furthermore, a high proportion of at least the C8+ hydrocarbons, is also already provided as olefins.

The skilled person would understand that this is a highly significant and completely un- expected deviation from producing hydrocarbons in the gasoline boiling range, the lat- ter meaning C5-C12 hydrocarbons. More specifically, in an embodiment, the C8+ yield over R1 is already about 40 wt%, i.e. the raw converted-oxygenate product exiting R1 comprises already about 40 wt% C8+ hydrocarbons (jet fuel), or higher, such as 45 wt%. This yield is further increased by providing the OLI-loop section as described far- ther below.

In connection thereto, it would also be appreciated by the skilled person that the prod- uct pool for gasoline is very different from that of jet fuel, as is shown in appended Fig. 8 which depicts typical carbon numbers for gasoline, jet fuel and diesel.

In an embodiment, the MTG plant further comprises in the MTG-loop section: - means for continuously conducting said step of withdrawing from the one or more adi- abatic fixed bed reaction zones of MTG reactor (110) the first raw gasoline stream (109); - means, such as an analysis apparatus and/or controller, for determining at outlet of the one or more adiabatic fixed bed reaction zones of the MTG reactor (110) an amount of one or more unconverted oxygenate compounds in the withdrawn first raw gasoline stream (109); - means, such as regulating valve, for continuously adjusting the inlet temperature of the one or more feed streams to maintain a constant amount of the one or more uncon- verted-oxygenate compounds as determined at said outlet; the method further comprising: - continuously conducting said step of withdrawing from the one or more adiabatic fixed bed reaction zones of R1 the raw converted-oxygenate product (9); further, determining at outlet of the one or more adiabatic fixed bed reaction zones an amount of one or more unconverted oxygenate compounds in the withdrawn raw converted-oxygenate

DK 181970 B1 10 product (9); and continuously adjusting the inlet temperature of the one or more feed streams to maintain a constant amount of the one or more unconverted-oxygenate compounds as determined at said outlet of between 10 and 4000 ppmv.

The revamping enables therefore the use of these existing means in the MTG-loop section of the MTG-plant for a completely different purpose, namely the production of jet fuel hydrocarbons.

It has also been found that when low temperatures, e.g. below 375°C, such as below 350°C, in the fixed bed of the reactor (R1) of the MTH-loop are maintained, deactivation of catalysts comprising a zeolite having a 1D 10 ring pore structure, such as ZSM-48, may increase due an increasing impact of coking: at lower temperatures the oxygenate conversion rate is lower, which means that the impact of coking is accentuated and thus more visible on the measured oxygenate conversion, whereby the oxygenate conversion reaction will extinguish quicker. Thus, the apparent catalyst stability decreases when the temperature decreases, thereby leading to an instable process for producing hydrocar- bons. On the other hand, operating at higher temperature also conveys catalyst deacti- vation due to a higher degree of dealumination of the catalyst via the so-called steaming, i.e. exposure of the catalyst to vapor water thereby dealuminating the catalyst. The cat- alyst comprises a zeolite having a given silica-to-alumina ratio (SAR) and is desirable to minimize dealumination of the catalyst for prolonging its longevity, i.e. lifetime. Coke (pre- graphitic and graphitic type) deposited in the catalyst at higher reaction temperatures also conveys catalyst deactivation and is therefore removed by periodical burning off, thereby regenerating the catalyst. This normally requires i.a. a high energy input to pro- vide high oxidation temperatures and/or high oxygen environment, thereby generating

CO» emissions. The deposited carbon, due to the higher reaction temperatures during oxygenate conversion, has a lower oxidation reactivity, thus requiring higher tempera- tures for its burning during regeneration of the catalyst.

Hence, while reducing the operating temperature during oxygenate conversion in R1 may at first glance appear to overcome the challenge of catalyst dealumination, coking of the catalyst at the low temperatures still occurs, thereby leading to an instable process.

The provision of said strategy in the process according to this embodiment , see e.g. also appended Fig. 3-4, by which the process also comprises said continuous

DK 181970 B1 11 adjustment of the inlet temperature to R1 to maintain a constant amount of e.g. metha- nol at said outlet of between 10 and 4000 ppmv, enables a process and plant that pro- duces distillate boiling range products such as jet fuel in a stable manner despite cata- lyst coking, while at the same time reducing deactivation of the oxygenate conversion catalyst due to dealumination, thus also increasing catalyst longevity. Furthermore, there is a lower energy input during catalyst regeneration of R1. In connection thereto, there is a reduction in carbon intensity (Cl), i.e. there is a lower carbon footprint by re- ducing CO; emissions during catalyst regeneration by burning off coke deposited in the catalyst., via reduction of the consumption of hydrocarbon fuel utilized during the re- generation, such as when utilizing a fired heater and feeding natural gas as fuel for the burning.

The revamped MTG reactor, thus R1, is operated on purpose with no complete conver- sion of oxygenates, such as with incomplete conversion of methanol as the oxygenate.

Instead, the oxygenate conversion is kept at a level corresponding to the unconverted oxygenate at the outlet, herein also referred to as “oxygenate slip” or for more simplicity "methanol slip”, of 10-4000 ppmv, thereby capturing the optima for C5+ hydrocarbons yield and yield of the desired hydrocarbons boiling in the jet fuel range, such as C8-17 hydrocarbons. Yet again, it has been highly surprising, that this high amount of hydro- carbons already in the desired jet fuel boiling range could be produced under these conditions, and also completely unexpected, that a reactor operation method similar to that of applicant's US 11060036 (WO 2019219397) utilized for a completely different purpose, namely for the conversion of oxygenates to C5+ hydrocarbons boiling in the gasoline boiling range under the presence of a 3D pore structure zeolite (ZSM-5, typi- cal for gasoline production), also provides optimal conditions for production of distillate boiling range hydrocarbons, in particular the desired hydrocarbons boiling in the jet fuel boiling range, e.g. C8-17 hydrocarbons. This takes place directly from an oxygenate such as methanol, in a fixed bed reactor carrying out the oxygenate conversion, without a subsequent oligomerization step, the latter typically being required for increasing the number of carbons of the hydrocarbon product resulting from the oxygenate conversion and thus providing the hydrocarbons in the jet fuel boiling range.

DK 181970 B1 12

The lower inlet temperature required for using this method which is similar to the tem- peratures required for gasoline production using the method ensure that expensive re- builds of the heating elements are not necessary. This has at least two advantages: it significantly reduces the cost of the revamp; it allows for a more cost-effective return to gasoline production if this is desirable at a later date.

Providing a raw converted-oxygenate product comprising C5+ hydrocarbons of which already at least 30 wt%, such as at least 40 wt%, are said hydrocarbons boiling in the jet fuel boiling range; or where the C8+ yield over R1 is at least 30 wt%, such as al- ready about 40 wt% of the raw converted-oxygenate product, enables that the load to the OLI-reactor R2, see farther below, for producing hydrocarbons in the jet fuel range is significantly reduced e.g. throughput of said second reactor set (R2).

The operating conditions of e.g. R1 are dynamically adjusted, either manually or from automated feed-back control, suitably the latter, to maintain the process at the optimal point of operation by counteracting the continuous effect of catalyst deactivation caused by coking and dealumination, and maximizing the yield of distillate boiling range hydrocarbons, in particular hydrocarbons in the jet fuel boiling range. This dy- namic adjustment of R1 upstream the separation unit of the MTH-loop section, e.g. high-pressure separator, and upstream a separation section for i.a. producing interme- diate jet fuel product and intermediate naphtha product, results in stable operation of the corresponding units, as the different flows therethrough are kept at a steady level, i.e. constant. Conventionally, to obtain constant product for fast deactivating processes such as catalytic deactivation, fluid bed or moving bed operation is needed. A fluid bed, for instance, is much more complicated to operate and to scale up. In contrast, a fixed bed is easier to operate, more reliable and significantly simpler to scale up, thus also more directly applicable to industrially relevant conditions and not least commercially viable.

Inventors across the world have been looking for ways to decrease the cost and main- tain the necessary stability in jet fuel production for years without success. By using the oxygenate slip method of the present invention, it is possible to lessen the effects of the zeolite deactivation in the first reactor R1. That is to say that between regenerations of the zeolite, the same amount of all the product streams (C8+, C5-C7 and C4-) with the

DK 181970 B1 13 same composition will be produced. This stability of R1 is a huge advantage for the fur- ther processing of the product, such as in downstream fractionation units. Other appli- cants have used complex technologies such as fluid bed to get a similar effect, how- ever, again, this is more complicated and makes the product more expensive.

Further, as mentioned before, by the invention the energy input required for regenera- tion of the conversion catalyst is significantly reduced. Also completely unexpected, the properties of the coke generated in the conversion catalyst according to the present in- vention are different than e.g. applicant's US 11060036 (WO 2019219397). In particu- lar, the oxidation reactivity and thus burning rate of coke generated in the conversion catalyst according to the present invention is higher than e.g. applicants US 11060036 (WO 2019219397), thus enabling to start and finish at lower temperatures the regener- ation of the catalyst by burning off the coke. Furthermore, not only is the content of ole- fins in the jet fuel range hydrocarbon fraction, such as the C8-C17, higher than in US 11060036 (WO 2019219397), but also the content of olefins in the C5-C7 hydrocarbon fraction. The latter is thus made readily available, already in R1, for oligomerization into jet fuel range hydrocarbons.

Suitably, said C5+ hydrocarbons are at least 75 wt%, such as at least 80 wt%, or at least 90 wt%, such as 80-90 wt%, of the raw converted-oxygenate product (9). Hence, as for instance shown in appended Fig. 5, there is a high yield of C5+ hydrocarbons.

Suitably, at least 40 wt%, or at least 45 wt%, or at least 50 wt% of said C5+ hydrocar- bons in the raw converted-oxygenate product (9) are said hydrocarbons boiling in the jet fuel boiling range; suitably C8-C17 hydrocarbons. For instance, 40, 50, 60, 70, 80 or 90 wt%; such as 40-60 wt%, or 45-55 wt of said C5+ hydrocarbons are said hydrocar- bons boiling in the jet fuel boiling range. Suitably, said hydrocarbons boiling in the jet fuel boiling range are C8-C19 hydrocarbons, such as C8-C17 or C8-C16 hydrocarbons.

Hence, there is a high yield of C8+ hydrocarbons, e.g. C8-C17, in the C5+ hydrocarbon fraction.

Suitably, as recited above, at least 35 wt%, such as at least 40 wt%, or at least 45 wt%, or at least 50 wt% of the raw converted-oxygenate products, are said hydrocarbons boiling in the jet fuel boiling range (C8+ hydrocarbons).

DK 181970 B1 14

The rate of raw converted-oxygenate product being produced is not only stable, but also shows already a product distribution comprising hydrocarbons boiling in the jet fuel boiling range, such as C8-C17 or C8-C16 hydrocarbons, for instance as olefins. Hence, there is a high selectivity towards the desired product: the hydrocarbons boiling in the jet fuel range, as for instance shown in appended Fig. 5, or alternatively by the C8+ yield in the raw converted-oxygenate product, thus at exit of R1, being already about 40 wt% or higher.

Again, conventionally, such hydrocarbons in the jet fuel boiling range are produced first after conducting a subsequent oligomerization of e.g. C4-C8 olefins produced during a prior oxygenate conversion. For instance, as described earlier, applicant's WO 2022063992 discloses in the examples oxygenate conversion (methanol to olefins,

MTO) conducted with ZSM-48 in the temperature range 320-480°C, with the proportion of C8-C17 olefins produced therein being lower than 10 wt%. See also comparative ex- ample 2.

The yields represented by the above wt% (weight percentages) are as measured from

MeOH on water-free basis.

Suitably, at least 40 wt%, such as at least 50 wt%, or at least 60 wt%, of said hydrocar- bons boiling in the jet fuel range, i.e. of the C8+ hydrocarbons (C8+ product fraction), suitably of said C8-C17 hydrocarbons, are olefins. For instance, the olefin content in

C8+ is 45, 50, 55, 60, 65, 70, 75, 80 wt%. Hence, as for instance shown in appended

Fig. 9, not only there is a high yield of C8+ hydrocarbons, e.g. C8-C17 hydrocarbons, but also there is a high concentration of olefins. It is advantageous yet counterintuitive that there is a significant concentration of olefins in the C8+ hydrocarbons. These ole- fins, being already in the jet fuel range, require later hydrogenation into their saturated form. Yet, the higher the content of olefins the better since it allows further upgrading, i.e., further yield improvements. The present invention thus enables dynamically coun- teracting the effect of coking by temperature adjustments, allowing operation at the op- timum operation point that maximize the C8+ fraction, as well as its olefinic fraction.

DK 181970 B1 15

Suitably, at least 4 wt%, such as at least 5 wt% or at least 6 wt% of said hydrocarbons boiling in the jet fuel range, i.e. of the C8+ hydrocarbons (C8+ product fraction), suita- bly of said C8-C17 hydrocarbons, are C13+ hydrocarbons. Hence, as for instance shown in appended Fig. 10, there is a shift towards the heavier C13+ hydrocarbons of the jet fuel range hydrocarbons, thus away from e.g. gasoline. The desired heavier compounds for jet fuel such as C13+ are produced in a surprisingly high proportion. As shown in connection with comparative example 3 (and associated Fig. 10), the concen- tration of C13+ in C8+ product fraction, thus in the heavy end of jet fuel as illustrated in the typical product distributions of Fig. 8, is not only clearly visible but also highly signif- icant compared to the prior art where values near zero wt% were measured. For in- stance, by the invention, the content of C13+ in said hydrocarbons boiling in the jet fuel range is at least 6, 7, 8, 9 or 10 wt%.

Suitably, said hydrocarbons boiling in the jet fuel range are C8-C17 hydrocarbons; in the raw converted-oxygenate product (9), said C5+ hydrocarbons, i.e. the C5+ hydro- carbon fraction, comprises C5-C7 hydrocarbons; said C8-C17 hydrocarbons and said

C5-C7 hydrocarbons in said C5+ hydrocarbons add up to 100 wt%.

It would be understood that in a particular embodiment, “C5+ hydrocarbons” means a hydrocarbon fraction comprising only C8-C17 hydrocarbons and C5-C7 hydrocarbons.

Suitably, at least 40 wt%, such as at least 50 wt%, or at least 60 wt%, of said C5-C7 hydrocarbons, are olefins. Hence, of the C5-C7 hydrocarbon fraction, which are not useful directly as jet fuel, in an embodiment, at least 40 wt% are olefins. These olefins are suitably later oligomerized, as explained farther above.

Suitably, the one or more adiabatic fixed bed reaction zones R1 are provided as sepa- rate conversion reactors R1, i.e. each with its own pressure shell. Hence, a reaction zone is a reactor; an adiabatic fixed bed reaction zone is an adiabatic fixed bed conver- sion reactor. For instance, one or more feed streams of one or more oxygenate com- pounds are introduced into the inlet of one or more adiabatic fixed bed conversion re- actors (R1), i.e. one or more downstream adiabatic fixed bed conversion reactors (R1).

For instance, one or more adiabatic fixed bed conversion reactors is provided as a first set (R1) of one or more adiabatic fixed bed conversion reactors operating in parallel.

DK 181970 B1 16

The parallel arrangement and operation enable periodic regeneration of the catalyst by burning off the coke deposited therein.

Suitably, said constant amount of the one or more unconverted-oxygenate compounds is between 1000 and 3000 ppmv, such as between 1500 and 2500 ppmv.

Within these ranges of the unconverted oxygenate compounds, such as a methanol slip at outlet (exit or effluent) R1, the highest yield of the desired jet fuel range hydro- carbons e.g. C8-C17 in C5+ hydrocarbons, are often observed. For instance, a metha- nol slip between 1500 and 2500 ppmv provides a normalized yield of above 98%, as shown in appended Fig. 3. Thus, the maximum jet fuel yield in C5+ hydrocarbons and the maximum C5+ yield can be found located at a conversion level corresponding to a methanol slip between 1500 to 2500 ppmv.

The term ”normalized yield” means that the values are normalized by the highest yield value obtained.

The yield is provided as weight percentage (wt%) of a given hydrocarbon fraction, and is measured on water-free basis, thus as defined earlier, measured from oxygenate, e.g. MeOH on water-free basis.

Suitably, the process further comprises maintaining a constant level of conversion of the one or more oxygenate compounds of between 93 and 99.9%.

Suitably also, there is a constant recycle to oxygenate ratio, such as recycle-to-oxygen- ate ratio of 5-15 w/w, for instance a recycle-to-methanol ratio of 5-15 w/w. It would be understood that this recycle is said overhead recycle of the MTH-loop. The process of the invention operates at a broader range of oxygenate conversion while producing a high proportion of hydrocarbons boiling in the jet fuel boiling range, compared to appli- cants US 11060036 (WO 2019219397) where the corresponding oxygenate conver- sion range is 95-99.9%, apart from the latter being directed to gasoline production, which again is completely different to producing jet fuel. Higher flexibility in the process of the present invention is thereby achieved. The yields go down at lower oxygenate conversion because of loss of oxygenate, e.g. methanol, which goes unconverted through R1, thus a higher methanol (MeOH) slip. The unconverted methanol may be

DK 181970 B1 17 reclaimed from the process condensate (water rich stream) withdrawn downstream, and routed back, e.g. via a methanol feed tank arranged upstream R1.

Suitably, the composition of the raw converted-oxygenate product comprising C5+ hy- drocarbons, i.e. said converted-oxygenate product, is: paraffins (P): 4-11 wt%, iso-par- affins (1): 5-30 wt%, olefins (O): 40-75 wt%, naphthenes (N): 6-15 wt%, aromatics (A): 4-20 wt%, in which the sum of P+I+O+N+A (PIONA) is 100 wt%.

As is well-known in the art, to determine n-paraffins, iso-paraffins, olefins, naphthenes and aromatics (i.e. PIONA composition) in a hydrocarbon mix containing hydrocarbons having boiling points in between -42 and 350°C, analytical techniques based on gas chromatography are available, e.g. in accordance with ASTM D8071 - VUV-PIONA; or

GCxGC-FID.

It would be understood that the art of chemistry and here in particular the art of produc- ing hydrocarbons over zeolite-based catalysts is highly unpredictable, as the product distribution and yields is determined from a large reaction network involving hundreds of different chemical reactions, and where the choice of zeolite catalyst (pore network, pore size, acidity, etc.) and the reactor operating conditions influence the conversion and selectivity. The above PIONA distribution is, astonishingly, significantly different from the PIONA obtained when using a similar strategy to maximize yield of hydrocar- bons boiling in the gasoline boiling range according to applicant's US 11060036 (WO 2019219397). Suitably also, the hydrocarbons boiling in the jet fuel boiling range are

C8-C17 hydrocarbons. For instance, the content of olefins according to this citation is much lower than in the present invention, as for instance shown in appended Fig. 9.

In an embodiment, the method further comprises: - installing an oligomerization loop section (OLI-loop section, 100”) by at least sepa- rately providing: an OLI reactor (R2) comprising an oligomerization catalyst and ar- ranged to receive an olefinic feed stream (17°) and provide a raw olefin product (19); further a separation unit (16) arranged to receive the raw olefin product (19) from said

OL! reactor (R2) and provide: an overhead recycle stream (5, 5) to said OLI reactor, a process condensate stream (21) comprising water, and an olefin product (23);

DK 181970 B1 18 - installing a separation section (20) arranged to receive the converted-oxygenate prod- uct and/or the olefin product (11, 23, 25) and provide: an intermediate jet fuel product (27) comprising said hydrocarbons boiling in the jet fuel boiling range; an intermediate naphtha product (17, 17”) comprising hydrocarbons boiling in the naphtha boiling range; optionally, an intermediate diesel product (29) comprising hydrocarbons boiling in the diesel boiling range; - installing a conduit for recycling at least a portion (17°) of said intermediate naphtha product (17, 17”) to said OLI-loop section (1007), as said olefinic feed stream (17°).

It has been found that the yield of the desired hydrocarbons already produced in the

OTH-section, such as said MTH-loop section, in particular the C8+ yield, is further in- creased by the provision of the OLI-loop section being fed with a naphtha intermediate stream separated downstream. Traditionally, an MTH, in particular a methanol-to-ole- fins (MTO) is conducted at conditions for producing short chained olefins, which in a sequential step are oligomerized to heavier hydrocarbons within the jet fuel range. The present invention produces a significant proportion of hydrocarbons already in the jet fuel range in the OTH-loop section in said converted-oxygenate product, while the sep- arate OLI-loop section instead of directly being supplied with said converted-oxygenate product from the OTH-loop section, is advantageously supplied with said intermediate naphtha stream separated downstream. Furthermore, C5-C7 hydrocarbons in the con- verted-oxygenate product, suitably as olefins, are also recovered as part of the inter- mediate naphtha product being recycled to the OLI reactor (R2).

Further, it has been found that the thus revamped plant of the invention (MTJ plant), with the separate OLI-loop section and recycle of naphtha thereto, in particular the re- cycle of the major portion of intermediate naphtha product, to the OLI-loop, is at least highly advantageous for large capacity plants in which the desired production of hydro- carbons boiling in the jet fuel range is at least 15000 bpd, as equipment sizing may be- come limiting, for instance for the recycle compressor of the MTH-loop. The required total recycle is distributed in the two loops in any rate, e.g. by providing two recycle compressors rather than a single one of large size, as the size of the recycle compres- sor of the MTH-loop may become a limiting factor. This is highly advantageous for the revamping a traditional or existing methanol-to-gasoline plants into a plant for produc- tion of jet range hydrocarbons (e.g. suitable for use as SAF) with naphtha and

DK 181970 B1 19 optionally also LPG (liquified petroleum gas) as valuable side products, again by utiliz- ing a significant part of already existing equipment of the methanol to gasoline plant, hence with minimal change or modification thereof.

As recited farther above, the term “hydrocarbons boiling in the naphtha boiling range may be used interchangeably with the term “naphtha” and means C5-C9 hydrocarbons boiling in the range 30-160°C, such C5-C8 hydrocarbons, e.g. C5-C8 olefins. The term “naphtha” is sometimes used interchangeably with the term “naphtha stream”. Hence, by the invention, naphtha withdrawn from the downstream separation section as said intermediate naphtha. This intermediate naphtha comprises olefins, such as C5-C8 ole- fins, which are advantageously recycled to the OLI-loop section as the olefinic feed stream thereto. The intermediate naphtha stream is advantageously also not being hy- drogenated, thereby maintaining the olefins therein unsaturated. This enables a high yield of desired olefins then being produced in the OLI reactor of the OLI-loop, such as

C8-C19 olefins e.g. C8-C17 or C8-C16 olefins, thus in the jet fuel range.

The OLI-loop section enables also relieving the MTH-loop section from performing all the work of producing hydrocarbons in the jet fuel range. The MTH-loop is highly sur- prisingly capable for producing a high proportion of hydrocarbons already in the jet fuel range, thus also requiring a minor OLI-loop for “topping-up” the amount of jet fuel finally produced by oligomerization of the intermediate naphtha.

In any of the MTH-loop and OLI-loop sections, the separating step of, respectively, raw converted-oxygenate product or raw olefin product, is conducted in a separation unit, suitably a 3-phase separator or a high-pressure separator, the latter may also be re- ferred to as high-pressure product separator.

For instance, in the MTH-loop, the converted-oxygenate product withdrawn from the separation unit may be regarded as a crude distillate boiling range hydrocarbon prod- uct. The overhead recycle stream is suitably a C4-hydrocarbon fraction, for instance rich in C1-C4 paraffins and/or olefins. For instance, the first C4- hydrocarbon fraction is any of: methane, ethane, ethene (ethylene), propane, propene (propylene), butane, bu- tene (butylene), and combinations thereof, optionally also comprising a minor portion of hydrogen, such as 1-10%. The overhead recycle stream is released at the top of e.g. a

DK 181970 B1 20 high-pressure separator, compressed by a recycle compressor and sent back to the in- let of R1. There are at least two benefits associated with this: first, the first recycle gas stream dilutes the reactor feed(s) and thereby controls the adiabatic temperature in- crease therein, which is suitably maintained in the range 30-100°C; second, the first re- cycle gas stream contains unconverted light olefins such as ethene, and by recycling these components the yield of hydrocarbons boiling in the jet fuel boiling range is fur- ther increased. The process condensate may be separated into water and unconverted oxygenates which may be recycled to inlet of R1, optionally also to inlet of R2.

Suitably, the MTH-loop section (100°) and the OLI-loop section (100”) operate at the same pressure, such as 15-30 barg, e.g. 20-25 barg.

Suitably, the MTH-loop section (100) and the OLI-loop section (100) operate at differ- ent pressure; optionally, wherein the pressure of the OLI-loop is higher than the pres- sure of the MTH-loop, such as 15-30 barg in the MTH-loop, e.g. 20-25 barg in the

MTH-loop, and 25-70 barg in the OLI-loop, e.g. 30-50 barg.

Hence, as an example, both loops may be operated at 20 barg, or one loop at 20 barg and the other at 70 barg, or both at 50 barg.

By having two separate loops it is possible to operate them at different pressures, thus significantly increasing flexibility during operation.

It would be understood that the pressure of the MTH-loop corresponds also to the pres- sure of R1. It would be understood that the pressure of the OLI-loop corresponds also to the pressure of R2.

For instance, in another embodiment, the pressure of R1 and/or R2 is at least 10 barg, such as 10-120 barg, e.g. 15-120 barg, 20-100 barg, or 20-80 barg. For instance, the pressure of R1 and/or R2 is any of: 10, 15, 20, 25, 30, 35, 40, 45, 50, 55, 60, 65, 70, 75 barg.

In particular for R1, while changing, e.g. increasing, pressure may have an effect on downstream separation and thereby the components being recycled, also highly

DK 181970 B1 21 surprisingly, when operating the MTH-loop in particular R1 therein according to the pre- sent invention (see also Fig. 3-4), changing the pressure turns out not having any sig- nificant impact on the oxygenate conversion as well as selectivity and composition of the hydrocarbons boiling in the jet fuel range, for instance as measured by its PIONA composition. The process is thus highly pressure insensitive which is highly advanta- geous for the operation of the process/plant, since the flexibility is significantly in- creased: the selected pressure may be accommodated to at least closely match the operating pressure either upstream oxygenate production, i.e. upstream R1, such as the pressure associated with an upstream methanol synthesis which is typically con- ducted at high pressures e.g. 80 barg or higher. The prior art teaches to keep the pres- sure low in the reactors, e.g. R1, presumably because there is a desire to keep the aro- matic slip low, whereas in the present invention the aromatic content turns out to be largely independent of pressure.

It would be understood that is generally advantageous to conduct oligomerization at high pressures, as higher yields are obtained. The present invention, by virtue of the provision of separate loops, each with its dedicated recycle compressor, enables to in- dependently conduct the oligomerization of R2 at a higher pressure, optionally also at higher temperature than in R1.

In an embodiment, the method of revamping further comprises: - installing a conduit for recycling a portion (17°) of said intermediate naphtha product (17, 17°) to said MTH-loop section (100°); and/or a conduit for supplying a portion (7, 7”) of said feed stream (1) to said OLI-loop section (100).

For instance, the portion (17””) of said intermediate naphtha product (17, 17”) is com- bined with said feed stream (1) of one or more oxygenate compounds, and/or com- bined with overhead recycle stream (15, 15’) of the MTH-loop section (100).

For instance, the portion (7, 7”) of said feed stream (1) is combined with said olefinic feed stream (17°), and/or combined with overhead recycle stream (5, 5) of the OLI- loop section (1007).

The combination e.g. mixing of streams is provided in a mixing point, such as mixing unit or juncture.

DK 181970 B1 22

Accordingly, the MTJ plant is suitably operated in parallel with respect to at least one of: the intermediate naphtha product, or the feed stream of one or more oxygenate compounds to said OLI-loop, for instance by admixing with said olefinic feed stream.

The intermediate naphtha stream comprises hydrocarbons with a significant number of carbons, thereby providing a high heat capacity and thereby also being highly advanta- geous for reducing the exothermicity of the OTH-reactor (R1) of the OTH-loop section, such as the MTH-reactor (R1) of the MTH-loop section.

A portion of distillates from the separation section is withdrawn as the intermediate naphtha product, which may for instance comprise C5-C8 hydrocarbons, suitably as

C5-C8 olefins, as already described. It would be desirable to convert this naphtha stream into valuable products. The present invention provides further synergistic inte- gration, since it has also been found that at least a portion, i.e. at least a fraction, of the intermediate naphtha product can also be recycled to inlet of R1 suitably operating ac- cording to the operating strategy discussed herein below (and illustrated in Figure 3-4), to increase the overall conversion of naphtha without any significant loss of catalyst stability in R1. It has also been found that at least a portion of the intermediate naphtha product, such as C5-C8 hydrocarbons therein, is converted into distillate boiling range hydrocarbons when fed into the oligomerization reactor (OLI reactor, R2) of the OLI- loop section. For instance, by operating R2 at > 20 barg, a weight hourly space velocity between 0.5 to 2 kg feed/kg cat./h, and a reaction temperature between 175-400°C, such as 180-350°C. R1 works as primary reactor with jet fuel yield of e.g. at least 30 wt%, such as at least 40 wt% of feed on hydrocarbon basis (exclusive water); thus, for instance, at least 40 wt% in the raw converted-oxygenate product exiting R1 are hydro- carbons boiling in the jet fuel range (C8+ hydrocarbons), such as C8-C17 hydrocar- bons.

Accordingly, to further increase the yield, the intermediate naphtha product, comprising for instance said C5-C8 olefins, may further react by recycling it either back to R1 as a methanol conversion reactor, or to R2 as a dedicated olefin conversion reactor located in a separate loop. R2 may thus for instance work as a “clean-up reactor’ converting the reactive species of e.g. the C5-C8 hydrocarbons in the intermediate naphtha prod- uct being recycled. R1 is suitably co-fed with naphtha to increase overall conversion,

DK 181970 B1 23 and R2 is suitably co-fed with methanol to improve catalyst stability and overall conver- sion, as illustrated in appended Fig. 1.

Yet again, a portion of the intermediate naphtha product is recycled to R2 and con- verted under reaction conditions tuned to favour oligomerization and facilitate produc- tion of long chained components, for instance by, in an embodiment, providing lower temperatures and lower weight hourly space velocities in R2 compared to R1, or an oli- gomerization catalyst which is different from the catalyst of R1. Limited amounts of oxy- genates such as methanol may be introduced together with the naphtha feed in R2, since the formation of water inhibits coking and thereby provide even longer cycle lengths before regeneration is necessary. For instance, 10% or less, such as 1-10% e.g. 3, 5, 7%, of the one or more feed streams of one or more oxygenate compounds, is introduced to R2. A fraction of the naphtha may not be converted because of accu- mulation of low reactivity species such as paraffins outside the distillate boiling range.

A fraction of the naphtha in R2 may be converted to larger hydrocarbons such as C17+ hydrocarbons suitably for e.g. diesel, and the amount will depend on the operation con- ditions and catalyst deactivation. To counteract the effect of catalyst deactivation of R1, the inlet temperature is adjusted/increased over time, resulting in a more stable product distribution over a catalyst cycle, as explained farther below.

The naphtha is recycled to the OLI-loop and the reactive species in the naphtha stream are converted in R2 thus further increasing the yield of the distillate boiling range frac- tion. More specifically, the intermediate naphtha product, which again, for instance comprises C5-C8 hydrocarbons, suitably C5-C8 olefins, is recycled to R1 and/or R2.

To ensure full conversion of the reactive part of the naphtha, the naphtha stream being recycled may be divided between R1 and R2 suitably operating at different conditions.

A portion of the naphtha is for instance co-fed with the feed of oxygenates, e.g. a meth- anol feed, to increase the yield of long chained molecules in R1. A portion, suitably a major portion, such as at least 50, 60, 70, 80, 90 wt%, of the naphtha being recycled i.e. the portion of said intermediate naphtha product, is recycled to R2 and converted under reaction conditions tuned to favour oligomerization of the e.g. C5-C8 hydrocar- bons, and facilitate production of long chained components, for instance by providing a lower temperature and lower weight hourly space velocity in R2. R1 is suitably a meth- anol to olefins (MTO) reactor and R2 is an oligomerization (OLI) reactor. R1 and R2 are

DK 181970 B1 24 thus advantageously arranged in parallel and thereby operate in parallel. Limited amounts of methanol feed may be introduced together with the naphtha being recycled into R2.

In an embodiment, the method further comprises: - installing a mixing point, such as a mixing unit or juncture, arranged to combine the converted-oxygenate product (11) and the olefin product (23), said mixing point being arranged upstream said separation section (20).

A significant portion of the hydrocarbons from the OTH-loop, such as MTH-loop, are al- ready in the jet fuel range, as so is the olefin product from the OLI-loop, so these streams are advantageously combined prior to being supplied to the separation sec- tion, thereby simplifying the process. The converted-oxygenate product 11 from the

MTH-loop section is thus suitably combined with olefin product 23 from the OLI-loop section into combined stream 25 prior to supplying to the separation section 20 com- prising one or more fractionation units, such as one or more distillation columns. This enables a simpler construction as a single conduit transporting the combined stream 25 is thereby provided.

In an embodiment, the method further comprises: - installing a hydroprocessing section comprising any of a hydrogenation reactor or a hydrocracking reactor, arranged to receive any of said intermediate jet fuel product (27), intermediate naphtha product (17, 17”), and intermediate diesel product (29), and provide, respectively, a jet fuel product, a naphtha product, and a diesel product.

The term “intermediate”, such as "intermediate naphtha product” or "intermediate jet fuel product” means that these streams may be further upgraded into a final product, for instance by a hydroprocessing step, such as a hydrogenation step and/or a hy- drocracking step, thereby producing a final product, here respectively, a naphtha prod- uct or jet fuel product. Thereby, for instance by hydrogenation and/or hydrocracking, optionally hydroisomerization, of hydrocarbons boiling in the jet fuel range, such as C8-

C17 hydrocarbons, of the intermediate jet fuel product, which may be predominantly present as olefins, are converted to the corresponding saturated compounds as the jet

DK 181970 B1 25 fuel product, suitably for use as sustainable aviation fuel (SAF) in compliance with

ASTM D7566 and ASTM 1655.

The term “hydroprocessing” means any of hydrotreating, hydrocracking, hydrogenation (herein also referred to as hydrogenating), hydroisomerization, or combinations thereof.

The hydroprocessing step is conducted in a hydroprocessing reactor i.e. a hydropro- cessing unit, comprising a catalyst under the presence of hydrogen. For instance, hy- drotreating is conducted over a hydrotreating catalyst for the removal of sulfur, oxygen, nitrogen, and metals from the hydrocarbons; hydrocracking is conducted over a hy- drocracking catalyst for the cracking of hydrocarbons; hydrogenation is conducted over a hydrogenation catalyst to hydrogenate hydrocarbons. The hydrogenation conditions are well-known in the art, for instance as described in applicant's co-pending patent ap- plication EP 22152691.6. Hydrotreating and hydrocracking are also well-known in the art. For instance, applicant's WO 2021180805 discloses associated catalysts and oper- ating conditions.

In an embodiment, the method further comprises: - installing a conduit for recycling at least a portion (3', 3”) of an off-gas (3) withdrawn form separation section (20) to any of said overhead recycle stream (15, 15°) of the

MTH-loop section (100), said overhead recycle stream (5, 5’) of the OLI-loop section (1007), said feed stream (1) of one or more oxygenate compounds of the MTH-loop section (100°), said olefinic feed stream (17°) of the OLI-loop section (100), or combi- nations thereof

In the separation section (20) an off-gas stream (3), suitably comprising C4- hydrocar- bons, is withdrawn. The off-gas stream (3) may be removed at the top of the fractiona- tion unit(s) of the separation section downstream R1, R2, e.g. in one or more distillation columns, and accordingly represents a light gas. The benefits of recycling such light gasses (C4- hydrocarbon fraction) have been discussed earlier. A portion 3” of the off- gas 3 may be withdrawn as a waste gas for other uses.

For the purposes of the present application, an off-gas stream may comprise a by-prod- uct stream rich in paraffins and/or olefins, for instance a C3 and/or C4 fraction, which is known as liquified petroleum gas, LPG. For the purposes of the present application, the

DK 181970 B1 26 off-gas may comprise streams comprising CO», Hz, CHa, higher hydrocarbons which are typically produced and withdrawn as waste gas stream(s). Accordingly, by the present invention an off-gas stream may be regarded as a by-product stream in accordance with the former, and as a waste gas stream in accordance with the later; or a combination thereof.

BRIEF DESCRIPTION OF THE FIGURES

Fig. 1 shows a simplified block type diagram of the MTJ plant after revamp, comprising a MTH-loop section and OLI-loop section according to an embodiment of the invention.

Fig. 2 shows (stippled-line section) the MTG-loop section of an existing MTG plant which is revamped into a MTH-loop section of a MTJ plant.

Fig. 3 shows the normalized yields (%) of C5+ hydrocarbons and jet fuel hydrocarbons (C5-C17) of R1 at any given methanol slip of R1. Solid curve: trend line for C5+ hydro- carbons. Stippled line: trend line for C8-C17 hydrocarbons.

Fig. 4 shows an automated inlet temperature strategy to maintain methanol slip constant of R1. R1 is an adiabatic fixed bed reactor, as illustrated in Fig. 1.

Fig. 5 shows the R1 yields of liquid product (C5+) and distillate boiling range hydrocarbon product as fraction of the liquid product (C8-17 in C5+). R1 is an adiabatic fixed bed reactor operated in once-through mode in a pilot scale plant.

Fig. 6 shows the loss of catalyst activity due to dealumination for R1 loaded with ZSM- 48. R1 is an adiabatic fixed bed reactor. The activity decay has been calculated from a kinetic model taking the effect of temperature into account.

Fig. 7 show parity plot of deactivation model used in Fig 6, and actual deactivation data from loss of activity of the zeolite during pilot plant operation, where the reactor oper- ates under commercially relevant conditions.

Fig. 8 shows a typical carbon number distribution of gasoline, jet fuel and diesel.

Fig. 9 shows a comparative example of olefin content in the C8+ product fraction (jet fuel hydrocarbons) from the oxygenate conversion reactor.

DK 181970 B1 27

Fig. 10 shows a comparative example of the concentration of the heavier molecules

C13+ in the C8+ product fraction (jet fuel hydrocarbons) from the oxygenate conversion reactor.

DETAILED DESCRIPTION

Fig. 1 shows a schematic layout of an embodiment of the process/plant 100 for produc- ing hydrocarbons in the jet fuel boiling range (MTJ plant 100), after revamping an exist- ing plant for producing gasoline (MTG plant, see also MTG-loop section of Fig. 2). In

Fig. 1, the process/plant 100 comprises a MTH-loop section 100’, optionally an OLI- loop section 100”. A feed source of oxygenate(s) such as methanol 1 is supplied to the

MTH-loop section 100'. The MTH-loop section comprises a MTH reactor, e.g. MTO re- actor (R1); a separation unit 12 and overhead recycle stream 15, 15' which is recycled to R1 via dedicated recycle compressor 14. The OLI-loop section 100” comprises an

OLI-reactor (R2), a separation unit 16, and overhead recycle stream 5,5' which is recy- cled to R2 via dedicated recycle compressor 18.

In the MTH-loop section, the methanol stream 1 is combined, e.g. admixed, with over- head recycle stream 15' directed by the recycle compressor 14 and olefinic feed stream 17” into feed stream 7. The olefinic feed stream 17”” is a portion of intermedi- ate naphtha stream 17, 17” from downstream separation section 20. The methanol stream 1 is optionally combined with a portion 3”” of an off-gas 3 comprising light hy- drocarbons (e.g. C4- hydrocarbons, such as LPG) which is withdrawn from the down- stream separation section 20. A minor portion 7” of the combined feed stream 7 is sup- plied to the OLI-section, more specifically to R2. A major portion 7” of the combined feed stream 7 is then introduced to first set (R1) of one or more adiabatic fixed bed re- action zones, here illustrated as one adiabatic fixed bed reactor comprising one fixed catalyst bed, and which is tuned to facilitate methanol conversion and formation of dis- tillate boiling range hydrocarbons. The adiabatic fixed bed reactor R1 is operated such that the reaction temperature is maintained as low as possible for as long time as pos- sible. This is achieved by combining frequent measurements of methanol in effluent 9 at R1-outlet (e.g. by GC) and a programmable feed-back action adjusting the inlet tem- perature to the reactor. In that way, the effect of catalyst coking is continuously

DK 181970 B1 28 counteracted by automated temperature adjustments. Since the catalyst of the reactor set R1 is regenerated regularly, for instance once a month, several reactors operating in parallel are necessary during continuous production, whereby the reactor shifts be- tween the reaction mode and regeneration mode. A raw converted-oxygenate product, already comprising a high proportion of hydrocarbons boiling in the jet fuel range, is withdrawn as reactor effluent 9, cooled down in an air cooler and heat exchangers (not shown) and introduced to separation unit 12 of the MTH-loop section, such as a high- pressure separator. Water and unconverted methanol are withdrawn from the separa- tion unit 12 in process condensate stream 13, while overhead recycle stream 15 e.g. a

C4-hydrocarbon fraction, is released at the top, compressed by recycle compressor 14 and sent back as recycle gas 15’ to the inlet of R1 as described above. From the sepa- ration unit 12 a converted-oxygenate product stream 11 comprising said hydrocarbons boiling in the jet fuel range, is withdrawn.

The converted-oxygenate product 11 from the MTH-loop section is combined with ole- fin product 23 from the OLI-loop section into stream 25 prior to supplying to the separa- tion section 20 comprising one or more fractionation units, such as one or more distilla- tion columns. From the separation section 20 an intermediate naphtha product stream 17 is withdrawn and at least a portion 17’ thereof is provided as the olefinic feed stream to the OLI-loop section 100”, more specifically to R2. A portion 17°” of the intermediate naphtha product stream 17 is suitably also combined with the feed to R1 in the MTH- loop section 100' as described above. A portion 17” of the intermediate naphtha prod- uct is also withdrawn.

In the OLI-loop section, the olefinic feed stream 17’ is combined, e.g. admixed, with overhead recycle stream 5’ of the OLI-loop section directed by the recycle compressor 18, along with the portion 7” of feed stream to R1, optionally also with a portion 3’ of the off-gas stream 3 comprising C4- hydrocarbons withdrawn from the downstream separation section 20, and then introduced to the OLI-reactor (R2) which comprises one or more adiabatic fixed bed reaction zones, here illustrated as one adiabatic fixed bed reactor comprising one fixed catalyst bed. From R2 a raw olefin product 19 is with- drawn, cooled down in an air cooler and heat exchangers (not shown) and supplied to separation unit 16 of the OLI-loop section 100”. A process condensate stream 21 is withdrawn therefrom as so is overhead recycle stream 5 and an olefin product stream

DK 181970 B1 29 23. The overhead recycle stream is 5 is compressed by recycle compressor 18 and sent back to the inlet of R2 as the recycle gas 5'.

As inferred from above, R1 and R2 are suitably arranged in parallel and operate in par- allel. R1 and R2 are arranged and operated in parallel, for instance also by virtue of: re- cycling a portion 17°” of said intermediate naphtha product 17, 17’ to said MTH-loop section 100’, and/or by supplying a portion 7” of said feed stream 7 of one or more oxy- genate compounds also to said OLI-loop section 100”. In order to convert the reactive species in the naphtha range fraction for increasing the yield of the distillate boiling range fraction, in particular the jet fuel fraction, the naphtha fraction is recycled to the

OLI-loop section 100”. Thus, more specifically, the portion 17’ of the intermediate naphtha product 17, which for instance comprises C5-C8 hydrocarbons, suitably C5-C8 olefins, is recycled to R1 and/or R2. To ensure full conversion of the reactive part of the naphtha fraction, the naphtha being recycled may be divided between R1 and R2 suita- bly operating at different conditions. Hence, a portion 17°” of the naphtha 17, 17’ can be recycled and co-fed together with the methanol feed 1 to increase the yield of long chained molecules in R1. At least a portion 17° of the naphtha 17 is recycled to R2 and converted under reaction conditions tuned to favour oligomerization of the e.g. C5-C8 hydrocarbons, and facilitate production of long chained components, for instance by providing a lower temperature and lower weight hourly space velocity in R2; or due to being in a separate loop also at a pressure which may be different from that of R1 in the MTH-loop section 100’.

The separation section 20 comprising one or more fractionation units, such as one or more distillation columns, and provides an intermediate jet fuel product 27, the interme- diate naphtha product 17, 17” optionally an intermediate diesel product 29, optionally also the off-gas stream 3 comprising for instance LPG. The intermediate products, such as the intermediate jet fuel product 27 can be e.g. hydrogenated directly in a downstream hydrogenation section (not shown), thereby providing a final jet fuel prod- uct.

With reference to Fig. 2, conduits or process streams 103”” (off-gas from downstream separation section 20) and 107” (oxygenate feed to R2), although shown in this figure, are not envisaged as part of the existing MTG-loop. By the invention it has surprisingly

DK 181970 B1 30 been found that operation with a different catalyst, yet at conditions similar as for gaso- line synthesis produces a high proportion of hydrocarbons boiling in the jet fuel range already in R1. Hence, an existing methanol-to-gasoline (MTG) plant, in particular the

MTG-loop section thereof, is advantageously revamped into a MTH-loop 100’ and fur- ther to a methanol-to-jet fuel (MTJ) plant 100. The stippled-line section of Fig. 2 serves to illustrate a conventional MTG-loop section (MTG-loop) in which the reactor 110 therein is a MTG-reactor. The existing MTG-loop section is arranged to receive metha- nol feed stream 101 and provide a raw gasoline stream 111. The MTG reactor 110 comprises a catalyst active in the conversion of methanol to gasoline as a first raw gas- oline stream 109, such as a catalyst comprising a zeolite with a framework having a 10- ring pore structure, said 10-ring pore structure being a three-dimensional (3D) pore structure, typically ZSM-5; a product separation unit 112 of the MTG-loop section ar- ranged to receive the first raw gasoline stream 109 and provide an overhead recycle stream 115, 115’, a process condensate 113 containing water, and the raw gasoline stream 111. Recycle compressor 114 of the MTG-loop section supplies overhead recy- cle stream 115’ to the MTG reactor 110 by combining with the methanol stream 101, thereby providing feed stream 107. The feed stream 107, 107’ is introduced to the

MTG-reactor 110. The existing MTG-plant may further comprise (not shown): a distilla- tion section arranged to receive said raw gasoline stream and provide a stabilized gas- oline stream; optionally, an upgrading section comprising any of a hydroisomerization (HDI) section and/or hydrocracking (HCR) section, arranged to receive said stabilized gasoline stream and provide a gasoline product.

The method of revamping comprises: converting said MTG loop section (Fig. 2) into a

MTH-loop section 100’ (Fig. 1) by at least replacing the catalyst of said MTG reactor 110 with a catalyst active in the conversion of methanol to hydrocarbons boiling in the jet fuel range, such as catalyst comprising a zeolite in which said zeolite is any of *MRE (ZSM-48), MTT (ZSM-23), TON (ZSM-22). Thereby MTG reactor 110 of Fig. 2 is re- vamped into R1 of Fig. 1. The method of revamping may also comprise applying means for continuously adjusting the inlet temperature of R1 to maintain a constant amount of at outlet of R1 of oxygenate, e.g. methanol between 10 and 4000 ppmv, as shown in appended Fig. 3-4. A high proportion of jet fuel hydrocarbons is thereby ob- tained, as shown in Fig. 5, 10, and further a high proportion being produced as olefins, as shown in Fig. 9. The method of revamping may also comprise installing an OLI-loop

DK 181970 B1 31 section as shown in Fig. 1-2 and which corresponds to process streams and associ- ated units already shown in Fig. 1 including installing a conduit 17’ for recycling at least a portion of said intermediate naphtha product 17 to said OLI-loop section 100”, as the olefinic feed stream.

EXAMPLES

EXAMPLE 1

The conversion from oxygenates into distillate boiling range hydrocarbons in R1 is highly surprisingly achieved in a single oxygenate conversion step utilizing reaction synergies discovered from experimental tests in a 0.6 bpd (barrels per day) pilot plant.

The findings were made during pilot tests where the zeolite catalyst ZSM-48 was tested under industrially realistic reactor operating conditions. The reactor diameter-to-catalyst particle ratio in the experimental setup was about 40, which is more than sufficient to avoid undesired wall effects. In the pilot, the feed flow rate to the reactor was set to match the industrial relevant mass fluxes, thereby ensuring the correct mass and en- ergy transfer between the bulk fluid phase and the surface of the catalyst particles. Ap- plication of commercial catalyst particle dimensions combined with industrial Reynolds numbers thus provide an overall transport limitation, from the bulk of the fluid to the ac- tive catalytic sites, like what is obtained for the commercial reactors. The pilot simu- lated the entire oxygenate conversion loop including recycle of light gasses to the inlet of the reactor. Various pressure levels, space velocities, recycle-to-make up gas ratios, and temperature levels were tested. The adiabatic catalyst bed temperature was meas- ured along the entire bed length with a 4 mm diameter multipoint Rosemount K-type thermoelement (273.15 K up to 1373.15 K) in direct contact with the catalyst bed mate- rial. Inlet, exit, and recycle stream rates are measured by Coriolis mass flow meters.

These together with composition analysis, form full mass balance closure by error smoothing of the entire oxygenate conversion loop.

From the experimental pilot results, it was i.a. found, that when operating on methanol feed at high pressure (e.g. at least 20 barg) and low reaction temperature such as out- let temperature below 370°C, a large fraction, i.e. about 50 wt% or more, of the C5+ hydrocarbons being produced, was astonishingly already in the jet fuel boiling range.

This high yield of larger hydrocarbon species already in product range was completely

DK 181970 B1 32 unexpected with this catalyst from numerous laboratory scale experiments as well as from the open literature. The total liquid yield (C5+) at these conditions was about 85 wt% on hydrocarbon basis excluding water, i.e. on water-free basis, see Fig. 5. The gas phase product (C4-) was mainly butane and propane and some light olefins. If the temperature was increased the result was an observable sharp decrease in both liquid yield and the fraction of the distillate boiling range hydrocarbons which include said hy- drocarbons boiling in the jet fuel range. Measured yields at both temperatures (low and high temperature) are provided in below Table 1. On the other hand, if the temperature remained at the low level, such as 290°C, the continuous coking of the catalyst would extinguish the chemical reactions thereby leading to an unstable process. Conse- quently, to keep stable reactor operation, the inlet temperature to the catalyst bed was adjusted/increased over time to counteract the effects of that mechanism of catalyst deactivation. Coking is counteracted by continuous and appropriate increase of inlet temperature, yet if the temperature becomes too high, catalyst deactivation by other mechanism, namely dealumination of the zeolite, takes place at an accelerated rate.

The reactor inlet temperature needed to balance the methanol slip exiting the reactor at the same constant value resulted in a fairly constant liquid product yield (C5+ hydrocar- bons) as well the yield of the distillate boiling range hydrocarbons, as also shown in

Fig. 5.

The normalized yields plotted vs the methanol slip are shown in Fig. 3. The term ”nor- malized yield” means that the values are normalized by the highest yield value ob- tained (highest yield is for instance here about 55 wt%). The yield of C5+ hydrocarbons of which for instance about 45 wt% are jet fuel hydrocarbons (in Fig. 3 represented as

C8-C17) — see also Fig 5 and below Table 1 - is a function of the methanol conversion and selectivity to said C5+ hydrocarbons. The temperature increase and corresponding methanol slip are shown in Figure 4 and the obtained product yields are provided in

Figure 5. Operating with a constant methanol slip of between 10 and 4000 ppmv, as shown in Fig. 3, thus using the temperature strategy shown in Fig. 4, allows a stable high production of final range distillate product already in the methanol conversion re- actor R1. Fig. 3 shows, along with Fig. 4 and 5, the advantage in terms of yields of e.g.

C8-C17 hydrocarbons, by not operating with complete conversion of methanol (com- plete conversion means methanol slip = 0 ppmv). By maintaining a given methanol slip in the range 10-4000 pppmyv, the optimum yield for C8-C17 is captured.

DK 181970 B1 33

Table 1: Measured yields from methanol on hydrocarbon basis exclusive water (i.e. water-free basis) from pilot plant operation. Low temperature operation in the methanol conversion reactor favours production of distillate boiling range hydrocarbons. These yields are valid for “once- though” operation, i.e. no recycle of heavier hydrocarbon species. oo Low T operation High T operation

Hydrocarbon fraction Yields, [wt%]

Olefin content in C8-C17 hydrocarbon fraction, Low T operation: 62 wt%.

In connection with Table 1, in the pilot experiments during the “Low T operation”, the inlet temperature was increased as per invention, and at a point in time where the inlet temperature was about 310°C, a step change was made in the temperature up to 350°C (“High T operation”). A decrease in yield was observed, as shown in the table.

It is observed from the Low T operation column (invention) that, for instance, the C5+ hydrocarbons, i.e. the C5+ hydrocarbon fraction, is 85 wt% of said raw converted-oxy- genate product, while the remaining 15 wt% is C4- hydrocarbons, i.e. the C4- hydrocar- bon fraction. It is also observed that the C5+ hydrocarbons, for instance here 85 wt% of the raw converted-oxygenate product, contains 45 wt% hydrocarbons in the jet fuel range, here specifically 45 wt% C8-C17 hydrocarbons. This represents more than 50 wt% of the C5+ hydrocarbons. (45/85 x 100 = 53 wt%). The remaining C5-C7 hydrocar- bons, i.e. the C5-C7 hydrocarbon fraction, of the C5+ hydrocarbons, represents thereby less than 50 wt%, more specifically 47 wt%. In contrast, when allowing the operation to complete conversion of methanol, the Hight T operation column not only shows lower yield of C5+ hydrocarbons, but also the proportion of C5-C7 in the C5+ hydrocarbons is higher than the proportion of the more desirable C8-C17 hydrocarbons.

DK 181970 B1 34

Alternatively, it is also observed from Table 1, that the jet fuel range from R1 is e.g. 45 wt% of the raw converted-oxygenate product; here specifically 45 wt% C8-C17 hydro- carbons (Low T operation column of Table 1).

In additional pilot tests, the produced liquid olefin rich product was after separation pumped back to the inlet reactor. Two tests were made. In one test the olefin rich prod- uct was introduced into the synthesis loop together with methanol feed, and in another test the olefin rich product was introduced alone. From the experiments it was found that a fraction corresponding to at least all the hydrocarbons in the naphtha range can be recycled to inlet of the methanol conversion reactor operating according to the strat- egy discussed and illustrated in Fig. 3, 4 to increase the overall conversion of naphtha without any significant loss of catalyst stability. It was further found, that at least a frac- tion of the naphtha range hydrocarbons can be converted into distillate boiling range hydrocarbons when fed into an adiabatic reactor (corresponding to R2) operating at high pressure, e.g. at least 20 barg, a weight hourly space velocity between 0.5 to 2 kg feed/kg cat./h, and a reaction temperature between 180-350°C. It would be understood that the reaction temperature lies within the range of the inlet and outlet temperature of the adiabatic fixed bed reactor, for which the adiabatic temperature rise is e.g. 30- 100°C.

Fig. 6 shows a plot of irreversible deactivation of the zeolite catalyst via dealumination in R1. The plot shows relative activity of the catalyst as a function of the normalized time. The upper solid line shows the relative activity (loss of activity) according to the new method, i.e. present invention, thus with active regulation of methanol slip as per

Fig. 3, 4, whereas the lower stippled line shows the loss of activity according to the standard method, i.e. whereby the temperature is fixed, for instance temperature at in- let of 380°C, temperature at outlet 450°C, without active regulation of methanol slip.

The activity decay shown in Fig. 6 has been calculated from a kinetic model taking the effect of temperature into account. Fig. 7 (X-axis: measured activity; Y-axis: calculated activity) shows that the kinetic model - kinetic evaluation of the loss of Brønsted acidity (catalyst activity) -, in the figure denoted as “model” is reliable.

DK 181970 B1 35

Fig. 8 shows a typical carbon number distribution of gasoline (left hand curve in the fig- ure), jet fuel (center curve in the figure) and diesel (right hand curve in the figure), ac- cording to the prior art, as retrieved from: https:/www.researchgate.net/figure/Carbon-number-distribution-of-petroleum- fuels fig1 267420915

This figure serves to show that the product pool for gasoline is very different from that of jet fuel. The production of jet fuel relates to a completely different field: different catalyst, different chemistry and kinetics, as well as a different product.

EXAMPLES — COMPARATIVE

COMPARATIVE EXAMPLE 1:

A comparative example was conducted to show the highly unexpected results of the present invention with respect to the prior art. Hence, Fig. 9 shows the olefin content in the C8+ product fraction (more specifically the C8-C17 hydrocarbon fraction) from the oxygenate conversion reactor according to the present invention with respect to the afore mentioned prior art (D1: applicant's WO 2019219397; D2: applicant's WO 2022063994 with oxygenate conversion using the same ZSM-48 zeolite).

Fig. 9 shows that there is a surprising cumulative effect in the present invention that is not inferred from D1 and D2, alone or in combination. By the present invention, not only it is now possible to obtain a high C8+ yield, e.g. C8-17 yield, with high olefin content (at least 30 wt% of said C5+ hydrocarbons, such as at least 50 wt% of said C5+ hydro- carbons, as measured from MeOH on water-free basis), but also with a high concentra- tion of olefins: at least 40 wt% olefins in said C8+ product fraction, such as above 60 wt% or above 62 wt% olefins as shown in the right-hand column of the figure. This C8+ olefin content (wt% olefins in C8+) is next to nothing in D1, namely < 5 wt%). D2 does not use temperature adjustments to control the methanol conversion as in the present invention, and the wt% content of olefins in C8+ is significantly lower.

DK 181970 B1 36

COMPARATIVE EXAMPLE 2:

As recited earlier, the product distribution of the present invention is also completely different to applicant's WO 2022063992. This citation discloses in the examples oxy- genate conversion (methanol to olefins, MTO) conducted with the same ZSM-48 in the temperature range 320-480°C, with the proportion of C8-C17 olefins produced therein being lower than 10 wt%. A subsequent oligomerization of e.g. C4-C8 olefins produced from the oxygenate conversion is thus required to increase the yield of product into the desired olefins in the jet fuel range.

In contrast thereto, the rate of converted-oxygenate product being produced by the pre- sent invention is not only stable, but also shows already a product distribution compris- ing a higher proportion of hydrocarbons boiling in the jet fuel boiling range. There is a high selectivity towards these, again as shown in appended Fig. 4. Further, as shown in above Table 1 — Low T operation, the corresponding content of C8-C17 olefins of the present invention is 28 wt% (45x62%=28 wt% - see also Fig. 9). Hence, almost a factor of 3 higher than according to WO 2022063992.

COMPARATIVE EXAMPLE 3:

Fig. 10 shows a comparative example of the concentration of the desired heavier mole- cules C13+, hence in the heavy end of jet fuel, according to the present invention and the prior art (D1: applicant's WO 2019219397; D2: applicant's WO 2022063994). This figure clearly shows that the concentration of heavier molecules which are in the heavy end of jet fuel — see for instance Fig. 8 - is next to nothing in D1, where a content of about 0.02 wt% C13+ in the C8+ product fraction is measured, while in D2 a content of 0 wt% was measured (below detection limit) even for C10+. In contrast thereto, the pre- sent invention shows here at least 6 wt% C13+ in the C8+ product fraction. The C13+ are hydrocarbons boiling in the jet fuel range, i.e. C13-C19, such as C13-C17.

The invention provides at least the following benefits: - Revamping of an existing methanol to gasoline (MTG) plant to a plant for producing a completely different product from gasoline, namely jet fuel, with minor number of modi- fications. For instance, the lower inlet temperature to R1 required for using this method

DK 181970 B1 37 which is similar to the temperatures of the MTG reactor required for gasoline produc- tion using the method, ensure that expensive rebuilds of the heating elements are not necessary. This has at least two advantages: it significantly reduces the cost of the re- vamp; it allows for a more cost-effective return to gasoline production if this is desirable at a later date. - Revamping into a MJT plant capable of operating with a large production of jet fuel hydrocarbons of up to 15000 bpd or a very large production of at least 15000 bpd. - Revamping into a MTJ plant with high integration of process streams produced therein, such as recycle of naphtha intermediate product to OLI-loop to increase yield of jet fuel, and utilization of hydrogen-rich stream withdrawn from the methanol synthe- sis section, suitably a MeOH loop, for at least providing hydrogen for hydroprocessing downstream, e.g. hydrogenation of the intermediate jet fuel product into a final jet fuel product. - Constant, i.e. stable, product composition even though fixed bed reactors (R1) in the

MTH-loop are applied. Conventionally, to obtain constant product for fast deactivating processes fluid bed operation is needed. - The invention enables dynamically counteracting the effect of coking by temperature adjustments, allowing operation at the optimum operation point that maximize the C8+ product fraction, e.g. C8-C7, and its olefinic fraction. The higher the content of olefins the better since it allows further upgrading, i.e., further yield improvements. - Significant prolonged catalyst lifetime. The mild temperature operation in R1 reduces significantly the catalyst deactivation rate due to coking and dealumination.

Further to the latter, the mild or low temperature operation results also in deposition of coke with higher oxidation reactivity compared to coke deposited at higher reaction temperature. This implies, that regeneration of the coke deposited on the catalyst when operating according to the invention is more easily burned off at lower regeneration temperatures, resulting in less degree of dealumination of the zeolite during regenera- tion, and again prolonging the catalyst lifetime. A kinetic evaluation of the loss of

Bronsted acidity (catalyst activity) for operating ZSM-48 according to this new process compared to the conventional MTO operation at constant and high temperatures is pro- vided in Fig. 6-7. The loss of activity of the catalyst is significantly lower when operating according to the new method of the invention, as shown in Fig. 3 and 4, compared to a conventional (standard) method in which a fixed temperature is applied.

DK 181970 B1 38 - Another benefit of the present invention is the increased commercial viability of the re- vamped plant as the infrastructure of already existing MTG plants is utilized.

Much quicker installation of plants for jet fuel production, crucial for the green energy transi- tion, instead of providing a MTJ plant from scratch is thereby possible.

- There is also increased commercial viability by the flexibility in terms of loop pres- sures.

The prior art teaches operation of fixed bed reactors at different pressures re- quiring more expenditure, or at lower pressures giving reduced effectivity particularly in the oligomerization (OLI reactor, R2). By providing a fixed bed reactor - based pro- cess/plant that is pressure independent in the methanol conversion (MTO reactor, R1),

a much superior solution than process/plants according to the prior art is achieved.

Claims

DK 181970 B1 39 PATENT CLAIM

1. A method for upgrading a methanol-to-gasoline (MTG) plant to a methanol-to-jet fuel (MTJ) plant (100), the MTG plant comprising: - a methanol-to-gasoline synthesis loop section (MTG loop section) adapted to receive a methanol feed stream (101) and provide a crude gasoline stream (111), the MTG loop section comprising: an MTG reactor (110) having one or more adiabatic fixed bed reaction zones comprising a catalyst active in the conversion of methanol to gasoline as a first crude gasoline stream (109), one or more heaters for heating one or more feed streams of the methanol feed stream (101) to an inlet temperature of said MTG reactor (110), one or more channels for feeding the one or more heated feed streams into said MTG reactor (110) inlet, a product separation unit (112) adapted to receive the first crude gasoline stream (109) and provide: an overhead recycle stream (115, 115"), a process condensate (113) comprising water, and a crude gasoline stream (111), - a distillation section adapted to receive said crude gasoline stream (111) and provide a stabilized gasoline stream, - optionally an upgrading section comprising a section of any one of a hydroisomerization (HDI) section and/or hydrocracking (HCR) section and adapted to receive said stabilized gasoline stream and provide a gasoline product, the method comprising: - converting said MTG loop section to a MTH loop section (100') by at least replacing the catalyst of said MTG reactor (110) with a catalyst active in the conversion of methanol to hydrocarbons having a boiling point in the jet fuel range, said catalyst comprising a zeolite having a framework having a 10-ring pore structure, said 10-ring pore structure being a one-dimensional (1D) pore structure, such as a catalyst comprising a zeolite, said zeolite being any of *MRE (ZSM-48), MTT (ZSM-23), TON (ZSM-22) or combinations thereof, whereby said MTG reactor (110) is converted to a first set (R1) of one or more adiabatic fixed bed reaction zones, wherein R1 functions as an oxygenate-to-olefins reaction zone, such as a methanol-to-olefins (MTO) reaction zone, e.g. an MTO reactor, and whereby a converted oxygenate product (11) comprising hydrocarbons boiling in the jet fuel boiling point range.

DK 181970 B1 40

2. The method of claim 1, further comprising: in said MTH loop section (100): - converting, in said one or more adiabatic reaction zones of R1, said one or more heated feed streams to a crude converted oxygenate product (9), suitably said crude converted oxygenate product (9) comprising C5+ hydrocarbons, of which at least 20% by weight or at least 30% by weight are said hydrocarbons boiling in the jet fuel boiling range, or at least 20% by weight or at least 30% by weight of said crude converted oxygenate product (9) are said hydrocarbons boiling in the jet fuel boiling range, - withdrawing from said one or more adiabatic fixed bed reaction zones said crude converted oxygenate product (9), - separating in said separation unit (12) the crude converted oxygenate product (9) to: an overhead recycle stream (15, 15') as an overhead recycle stream of the MTH loop section (100'), a process condensate (13) comprising water, and said converted oxygenate product (11), - recycling said overhead recycle stream (15, 15") to said first set (R1) of one or more downstream adiabatic fixed bed reaction zones, for example by mixing with said feed stream (1) of one or more oxygenate compounds.

3. A method according to any one of claims 1-2, wherein the MTG plant further comprises, in the MTG loop section: - means for continuously performing said step of withdrawing, from the one or more adiabatic fixed bed reaction zones of the MTG reactor (110), the first crude gasoline stream (109), - means, such as an analyzer and/or a controller, for determining, at the outlet of the one or more adiabatic fixed bed reaction zones of the MTG reactor (110), an amount of one or more unconverted oxygenate compounds in the withdrawn first crude gasoline stream (109), - means, such as a control valve, for continuously adjusting the inlet temperature of the one or more feed streams in order to maintain a constant amount of the one or more unconverted oxygenate compounds as determined at said outlet,

DK 181970 B1 41 the method further comprising: - continuously carrying out said step of extracting, from the one or more adiabatic fixed bed reaction zones of R1, the crude converted oxygenate product (9), further determining, at the outlet of the one or more adiabatic fixed bed reaction zones, an amount of one or more unconverted oxygenate compounds in the extracted crude converted oxygenate product (9), and continuously adjusting the inlet temperature of the one or more feed streams in order to maintain a constant amount of the one or more unconverted oxygenate compounds as determined at said outlet of between 10 and 4000 ppmv.

4. A method according to any one of claims 1-3, further comprising: - installing an oligomerization loop section (OLI loop section, 100") by at least separately providing: an OLI reactor (R2) comprising an oligomerization catalyst and arranged to receive an olefin feed stream (17') and provide a crude olefin product (19), further a separation unit (16) arranged to receive the crude olefin product (19) from said OLI reactor (R2) and provide: an overhead recycle stream (5, 5') to said OLI reactor, a process condensate stream (21) comprising water and an olefin product (23), - installing a separation section (20) arranged to receive the converted oxygenate product and/or the olefin product (11, 23, 25) and provide: a jet fuel a substance intermediate (27) comprising said hydrocarbons boiling in the jet fuel boiling range, a naphtha intermediate (17, 17") comprising hydrocarbons boiling in the naphtha boiling range, optionally a diesel oil intermediate (29) comprising hydrocarbons boiling in the diesel boiling range, - installing a channel for recycling at least a portion (17') of said naphtha intermediate (17, 17") to said OLI loop section (100") as said olefin feed stream (17).

5. A method according to claim 4, further comprising: - installing a channel for recycling a portion (17'"") of said naphtha intermediate (17, 17") to said MTH loop section (100'), and/or a channel for feeding a portion (7, 7") of said feed stream (1) to said OLI loop section (100").

DK 181970 B1 42

6. A method according to any one of claims 4-5, further comprising: - installing a mixing point, such as a mixing unit or a junction, arranged to combine the converted oxygenate product (11) and the olefin product (23), said mixing point being located upstream of said separation section (20).

7. A method according to any one of claims 4-6, further comprising: - installing a hydrogenation section comprising any one of a hydrogenation reactor or a hydrocracking reactor or a hydroisomerization reactor adapted to receive any one of said jet fuel intermediate (27), naphtha intermediate (17, 17") and diesel oil intermediate (29), and providing a jet fuel product, a naphtha product and a diesel product, respectively.

8. A method according to any one of claims 4-7, further comprising: - installing a channel for recycling at least a portion (3', 3"") of a waste gas (3) extracted from the separation section (20) to any of said overhead recycle stream (15, 15") of the MTH loop section (100'), said overhead recycle stream (5, 5) of the OLI loop section (100"), said feed stream (1) of one or more oxygenate compounds of the MTH loop section (100'), said olefin feed stream (17) of the OLI loop section (100"), or combinations thereof.