Methanol to Olefins (MTO) process

Abstract

The present application relates to a method for producing an olefin stream, the method comprising passing a feed stream comprising an oxygenate over a catalyst comprising a zeolite having a framework having a 10-ring pore structure, wherein the 10-ring pore structure comprises a one-dimensional (1-D) pore structure, such as a *MRE, at a pressure of 1-25 bar and a temperature of 240-360° C. The olefin stream can be converted into jet fuel, particularly sustainable aviation fuel (SAF), by further oligomerization and hydrogenation.

Description

Technical Field

[0001] The present invention relates to the conversion of a feedstock comprising oxygenates such as methanol and / or dimethyl ether into an olefin stream which is essentially free of aromatics and ethylene (C2=), but has a high content of higher olefins C3=-C8=, in particular (C4=-C8=), and optionally a large amount of isoparaffins. The invention also relates to the subsequent conversion of the olefin stream into hydrocarbons boiling in the jet fuel range, in particular sustainable aviation fuel (SAF), by oligomerization and hydrogenation. Background Art

[0002] Currently, processes that convert oxygenates such as methanol to olefins (MTO) are used to produce ethylene and propylene as the primary olefin products, intended for use as feedstocks in plastics production. When higher hydrocarbons are the desired product, such as in the methanol-to-gasoline (MTG) process, approximately 30% aromatics are typically formed. However, when producing hydrocarbons boiling in the jet fuel range, particularly sustainable aviation fuel (SAF), current requirements prohibit the presence of aromatics in the olefin feed stream.

[0003] The SAF market is expected to grow significantly in the coming decades due to societal concerns about global climate change and the resulting political pressure on the aviation industry. Currently, a small number of biocatalytic and thermocatalytic processes have been approved by ASTM to produce SAF. Therefore, a prerequisite for using any SAF as an aviation turbine fuel is ASTM certification. To date, only a small number of processes for producing SAF or synthetic paraffinic kerosene (SPK) fuel have been approved by ASTM International (ASTM) Method D7566 for blending into jet fuel at levels up to 50%. Therefore, an important general requirement is that the synthetic portion of the SAF (50 vol%) must be essentially free of aromatics, while the final SAF blend can contain up to 26.5 vol% aromatics.

[0004] Until now, the only process expected to be able to produce relevant quantities of SAF is based on Fischer-Tropsch (FT) synthesis from biomass, followed by multi-step and cost-intensive refining of the FT products, with a modest final selectivity towards jet fuel. The present invention uses the well-known methanol-to-olefins (MTO) process as a more attractive route to obtain olefins with higher selectivity. Currently, the proposed process layout for converting methanol to jet fuel is a multi-step process consisting of at least MTO, oligomerization and hydrogenation, which are proven technologies but appear to be inefficient to combine due to high recycle flows and very different process conditions in the individual steps.

[0005] Potential feedstocks for SAF production are generally categorized as (a) oil-based feedstocks, such as vegetable oils, waste oils, algae oils, and pyrolysis oils; (b) solid-based feedstocks, such as lignocellulosic biomass (including wood products, forestry waste, and agricultural waste) and municipal waste (organic fraction); or (c) gas-based feedstocks, such as biogas and synthesis gas (syngas). Synthesis gas, alcohols, sugars, and bio-oils can be further upgraded to jet fuel through a variety of synthetic (fermentation or catalytic processes).

[0006] Currently, there is no viable one-step catalyst / process that can convert a feedstock containing oxygenates such as methanol directly into hydrocarbons boiling in the jet fuel range, i.e., jet fuel, at least not in reasonable yields. To produce jet fuel starting from methanol, a three-step process is typically used, comprising: a) methanol to olefins (MTO), b) oligomerization of olefins, and c) hydrogenation of long-chain olefins. In the late 1970s and early 1980s, conventional methods for converting methanol into gasoline or diesel hydrocarbon products were envisioned. Thus, US 4,021,502, US 4,211,640, US 4,22,7992, US 4,433,185, and US 4,456,779 disclose process layouts based on classical MTO process conditions, i.e., high temperature (e.g., about 500° C.) and moderate pressure (e.g., about 1-3 bar) to achieve efficient conversion of methanol to olefins. However, under these conditions, a large amount of aromatic hydrocarbons (aromatics) is produced, for example 10-30 wt% or 10-35 wt% of aromatics in the olefin stream, which need to be separated, and a relatively large volume of the MTO product effluent must be cooled and treated to separate a C2-light gas stream, which is non-reactive except for ethylene, but its reactivity is very low. The remainder of the olefin stream must be pressurized to the substantially higher pressure in the oligomerization (OLI) reactor.

[0007] Therefore, the so-called Mobil Olefins to Gasoline Distillate (MOGD) process patents (e.g., US 5,177,279) in the early 1990s attempted to address these issues and improve overall operating efficiency through further process integration of MTO and oligomerization, while also reducing investment costs by splitting the methanol stream between the MTO and OLI reactors. This methanol feed split has two advantages: first, it reduces the MTO reactor size for the same overall methanol conversion, and second, only half of the methanol is processed at the high temperature conditions of the MTO reactor, thereby reducing the aromatics and C2 content.

[0008] US 5,177,279 discloses two possible general process layouts: (1) a conventional two-stage process with MTO and OLI / MOGD reactors in series, and (2) a three-stage process including an intermediate "olefin interconversion" (MOI) reactor that converts lighter olefins (C2=-C3=) to heavier olefins (C5=-C9=), thereby increasing the amount of higher (heavier) olefins from 25-35 wt% to 35-70 wt%. The latter process design offers greater flexibility, using only two reactors (MTO+MOI) when gasoline is the desired product, and using all three reactors when distillates are the preferred product.

[0009] US 9,957,449 discloses a process for producing hydrocarbons in the jet fuel range by oligomerization of renewable olefins having 3 to 8 carbons.

[0010] US 8,524,970 discloses a process for producing better quality diesel (ie diesel with a higher cetane number) comprising conversion of oxygenates to olefins, olefin oligomerization and subsequent hydrogenation.

[0011] Yarulina et al., ChemCatChem 8 (2016) 3057–3063, disclose the use of Ca-modified ZSM-5 for methanol-to-olefin conversion with the goal of achieving high propylene selectivity, with the process being carried out at 1 bar and a high temperature of 500° C. The resulting olefin stream showed no aromatics formation and had high selectivity for light olefins (C2=-C3=), but low selectivity for higher olefins (C4=-C8=).

[0012] US 2002 / 0103406 A1 discloses a process for dimerizing or oligomerizing olefin streams using a nickel-based catalyst.

[0013] WO 2011 / 138520 A2, EP 2123736 A1, and WO 2014 / 008337 A1 disclose variations of processes for producing hydrocarbons, including jet fuel, by dehydrating alcohols in the presence of zeolites (particularly ZSM-5) over a wide range of temperatures and pressures to primarily form ethylene. Similarly, US 2018 / 155637 A1 and US 8524970 B2 disclose a process for producing olefin streams from oxygenates over a ZSM-5 catalyst.

[0014] US 2020 / 0290940 and US 2020 / 290940 A1 disclose a process for isomerizing olefins by using ZSM-5, ZSM-23, ZSM-35, ZSM-11, ZSM-12, ZSM-48, ZSM-57, and mixtures or combinations thereof, wherein the microporous crystalline aluminosilicate has a SiO2 / Al2O3 molar ratio of less than or equal to about 100.

[0015] US 7,482,300 discloses a composition comprising ZSM-48 crystals having a silicon:aluminum molar ratio of 110 or less, or at least 70, the composition being free of non-ZSM-48 seed crystals and free of ZSM-50. The composition is used for catalytic dewaxing.

[0016] US 2019 / 0176136 by the applicant discloses the use of ZSM-23 zeolite as a catalyst for the conversion of methanol to olefins in a process step carried out at atmospheric pressure (about 1 bar) and 400° C., thereby producing a hydrocarbon stream with less than 5 wt% aromatics. The catalyst life is increased by providing the catalyst with a specific size in the direction of the channel system.

[0017] US 2017 / 0121237 A1 discloses a process for converting a feedstock containing oxygenates into gasoline and distillates, wherein the methanol conversion catalyst is selected from a variety of zeolites, including ZSM-48, and the process is carried out at a pressure of 15 to 90 psig and a temperature above 450°C.

[0018] US Pat. No. 4,476,338 discloses a process for converting methanol and / or dimethyl ether to olefins containing a significant amount of light olefins at moderate temperatures and atmospheric pressure, comprising contacting the feed with a crystalline zeolite catalyst designated as ZSM-48. The reference (Examples 1-2, Table 2) teaches the use of ZSM-48 having a silicon-to-aluminum ratio (SAR) greater than 110, more specifically 113 or 180, with methanol being converted over the zeolite catalyst at atmospheric pressure and a moderate temperature of 370°C. Aromatic yields are substantial, ranging from 10 to 12 wt%. Summary of the Invention

[0019] As used herein, "MTO" (methanol to olefins) refers to the conversion of oxygenates, such as methanol, into olefins.

[0020] As used herein, "OLI" means oligomeric.

[0021] As used herein, "Hydro" refers to hydrogenation.

[0022] As used herein, "Hydro / OLI" refers to a single combined step comprising hydrogenation and oligomerization.

[0023] As used herein, "MTJ" refers to methanol to jet fuel and is interchangeable with the term "integrated process" or "integrated process and apparatus," which refers to a process / apparatus that combines MTO, OLI, and Hydro to convert a feedstock containing oxygenates, such as methanol, into jet fuel.

[0024] As used herein, the terms "jet fuel" and "hydrocarbons boiling in the jet fuel range" are used interchangeably to mean a mixture of C8-C16 hydrocarbons boiling in the range of about 130-300°C at atmospheric pressure.

[0025] As used herein, "SAF" refers to sustainable aviation fuel or aviation turbine fuel that complies with ASTM D7566 and ASTM D4054.

[0026] As used herein, the terms "methanol" and "dimethyl ether" are used interchangeably with the terms MeOH and DME, respectively. "MeOH / DME" refers to MeOH and / or DME.

[0027] As used herein, "olefin stream" refers to an olefin-rich hydrocarbon stream comprising higher and lower olefins, and optionally aromatics, paraffins, isoparaffins, and cycloparaffins, wherein the combined content of higher and lower olefins is at least 25 wt%, e.g., 30 wt% or 50 wt%.

[0028] As used herein, the term "higher olefins" refers to olefins having three (3) or more carbons (C3+ olefins), particularly C3-C8 olefins (C3=-C8=), including olefins having four (4) or more carbons (C4+ olefins), particularly C4-C8 olefins.

[0029] As used herein, the term "lower olefin" refers to an olefin having two carbons, ie, ethylene (C2-olefin or synonymously C2= or ethylene).

[0030] As used herein, the term "high content of higher olefins" means that the weight ratio of higher olefins to lower olefins in the olefin stream is higher than 1, suitably higher than 10, such as 20-90, for example 70-80.

[0031] As used herein, the term "selectivity to higher olefins" refers to the weight ratio of higher olefins to lower olefins, i.e., the weight ratio of higher olefins to ethylene. "High selectivity to higher olefins" or "higher selectivity to higher olefins" means that the weight ratio of higher olefins to ethylene is greater than 10.

[0032] As used herein, the term "C2-lights" refers to C2= and C1-2 hydrocarbons.

[0033] As used herein, the term "lower hydrocarbons" refers to C1-2 (eg, methane, ethane), optionally also C2 =. The term may also be used interchangeably with the term "light paraffins."

[0034] As used herein, the term "substantially free of ethylene" or "ethylene-free" means 1 wt% or less.

[0035] As used herein, the terms "essentially free of aromatics," "substantially free of aromatics," "aromatic-free," or "low aromatics" refer to less than 5 wt%, such as 1 wt% or even less than 1 wt%. Aromatics include benzene (B), toluene (T), xylene (X), and ethylbenzene.

[0036] As used herein, the term "partial conversion of oxygenates" or "partially converting oxygenates" refers to an oxygenate conversion rate of 20-80%, such as 40-80% or 50-70%.

[0037] As used herein, the term "full conversion of the oxygenates" or "fully converting the oxygenates" means that the conversion of the oxygenates is greater than 80%, such as 90% or 100%.

[0038] As used herein, the term "substantial methanol conversion" is used interchangeably with the term "complete conversion of the oxygenate," wherein the oxygenate is methanol.

[0039] As used herein, the terms "catalyst comprising a zeolite" and "zeolite catalyst" are used interchangeably.

[0040] As used herein, the term "silicon to aluminum ratio (SAR)" refers to the molar ratio of SiO2 to Al2O3.

[0041] As used herein, the term "substantial amounts of paraffins" refers to 5 to 20 wt%, such as 10 to 15 wt%, of the olefin stream.

[0042] The object of the present invention is to provide a process for the conversion of oxygenates such as methanol to olefins (MTO) which is capable of more efficiently producing an olefin stream as a feedstock for oligomerization, in particular an olefin stream (olefin feed for oligomerization) which is low in aromatics, high in higher olefins, essentially free of ethylene and optionally has a substantial amount of isoparaffins.

[0043] It is another object of the present invention to provide such an olefin feed for oligomerization while still maintaining complete conversion of the oxygenates.

[0044] It is another object of the present invention to be able to achieve any of the above objects while also being able to increase the life of the catalyst while also performing the conversion of oxygenates to olefins.

[0045] The present invention addresses these and other objects.

[0046] Thus, in a first aspect, the present invention is a method for producing an olefin stream, the method comprising: passing a feed stream comprising an oxygenate through a catalyst active in the conversion of oxygenates at a pressure of 1-25 bar, e.g., 2-25 bar or 1-15 bar, and a temperature of 240-360°C, e.g., 260-360°C or 300-360°C, e.g., 300-340°C, wherein the catalyst comprises a zeolite having a framework having a 10-ring pore structure, wherein the 10-ring pore structure comprises a one-dimensional (1-D) pore structure; and wherein the 1-D pore structure is any one of *MRE (ZSM-48), MTT (ZSM-23), TON (ZSM-22), or a combination thereof.

[0047] A zeolite having a framework with a 10-ring pore structure means that the pore perimeter is defined by 10 oxygens.

[0048] A 1-D pore structure refers to a zeolite containing non-intersecting pores that are substantially parallel to one crystal axis. The pores preferably extend through the zeolite crystal.

[0049] The three letter codes for structure types, e.g. *MRE, are assigned and maintained by the International Zeolite Association Structure Committee in the Atlas of Zeolite Framework Types, located at http: / / www.iza-structure.org / databases / or as defined, for example, in "Atlas of Zeolite Framework Types", Ch. Baerlocher, L.B. McCusker and D.H. Olson, 2007, 6th revised edition.

[0050] It will be appreciated that the term "ZSM-48" may be used interchangeably with the term "EU-2."

[0051] It will be understood that the term "temperature" refers to the MTO reaction temperature in an isothermal process, or the inlet temperature of the MTO in an adiabatic process.

[0052] In one embodiment, the catalyst includes a binder. The catalyst is suitably formed by combining the zeolite with the binder and then forming the catalyst into particles. The particles can optionally be treated with a phosphoric acid reagent to produce a treated catalyst comprising a zeolite having a phosphorus component of 0.5 wt% to 15 wt%; that is, the phosphorus in the catalyst is 0.5 wt% to 15 wt%. Phosphorus provides stability to the catalyst. The binder is used to give the catalyst hardness and strength. The binder includes alumina, aluminum phosphate, silica, silica-alumina, zirconia, titania and combinations of these metal oxides, as well as other refractory oxides, and clays such as montmorillonite, kaolin, palygorskite, montmorillonite and attapulgite. Preferred binders are aluminum-based binders, such as alumina, aluminum phosphate, silica-alumina and clay.

[0053] In one embodiment, the catalyst comprises up to 30-90 wt% of the zeolite and a binder, suitably 50-80 wt%, the binder suitably comprising an alumina component, such as silica-alumina. For example, the method comprises mixing the catalyst with the binder (e.g., impregnating the catalyst) such that the catalyst comprises up to said 50-80 wt% of the zeolite and the binder, suitably comprising an alumina component, such as silica-alumina, thereby forming a silica-alumina binder. As an example, the catalyst is 60 wt% zeolite and 40 wt% alumina.

[0054] It should be understood that the wt% of zeolite in the binder refers to the wt% of zeolite relative to the weight of the catalyst, wherein the catalyst includes the zeolite and the binder.

[0055] It should also be understood that for the purposes of this application, the term "adhesive" may also be referred to as "matrix adhesive" or "matrix / adhesive" or "adhesive / matrix."

[0056] Binders impart hardness and strength to the catalyst. However, it has now been discovered that the use of binders in catalysts containing an alumina component, when operated at temperatures above 360°C, also has the undesirable effect of perpetuating the cracking of MeOH / DME in the MTO reaction, thereby producing methane as an undesirable byproduct. This requires, for example, the need to dispose of the methane by combustion or ignition, which increases the carbon footprint of the process and equipment. Furthermore, the yield of the desired olefin stream (olefin product stream), comprising, for example, higher olefins, is reduced because some of the feed is converted to the undesirable methane byproduct. It will be appreciated that in the MTO reaction, methanol can be initially converted to DME by methanol dehydration.

[0057] Thus, it is now surprising that at the low MTO temperatures of the present invention (360°C or less), the catalyst binder does not promote MeOH / DME cracking and thus does not produce undesirable methane, thereby enabling an increase in the yield of the desired olefin stream, as described above. Without being bound by any theory, it is believed that by the present invention, DME is rapidly reacted to form olefins before being cracked to methane.

[0058] Thus, the present invention provides the benefits associated with having a binder, including better stability of the catalyst, while eliminating its disadvantage of promoting the cracking of MeOH / DME to undesirable methane.

[0059] In one embodiment, the zeolite has a silicon to aluminum ratio (SAR) of up to 240. In a particular embodiment, the zeolite has a SAR of up to 110, for example up to 100. In another particular embodiment, the zeolite has a SAR of greater than 10, for example 15 or 20, or 30, 40, 50, 60, 70, 80, 90, 100.

[0060] In a typical MTO, which typically operates at temperatures above 360°C and uses ZSM-5 zeolite in the catalyst, the catalyst life is proportional to the SAR, so the higher the SAR value, the longer the catalyst life, which is desirable for MTO operation. However, high SAR values, such as 120 or higher, can also lead to reduced catalytic activity. Therefore, higher catalyst activity is found at lower SAR values, which can also reduce catalyst life. Through the present invention, it has been found that catalyst life is inversely proportional to the SAR, so as mentioned above, lower SAR values (e.g., 110 or lower) lead to longer catalyst life without sacrificing catalytic activity.

[0061] In one embodiment, pressure is 2-25 bar, for example 2,5,10 or 12 or 17 or 20 or 22 bar.Have been found that, although higher pressure (higher than 25 bar) has increased the ratio of higher olefins and lower olefins, i.e. to the higher selectivity of higher olefins, higher pressure also may reduce the total yield of olefins, and namely oxygenate is fed to the conversion efficiency of olefins lower, and has also improved and realized the required temperature of complete conversion, and this may produce the risk of the cracking reaction that does not too expect again.Within above-mentioned pressure range, for example 2-10 bar, 5-10 bar, 1-15 bar or 20-25 bar, can obtain the higher selectivity to higher olefins now, and do not need temperature to be raised to high level to realize complete conversion, thereby also reduced the generation of cracking reaction.In addition, the formation of ethene is suppressed, and the formation of aromatics is also suppressed.Therefore, olefin stream comprises higher olefins C3-C8 and isoparaffins.Especially, carrying out this method under the pressure higher than atmospheric pressure has the effect of making a certain amount of diluent as " heat tank (heat sink) " of exothermic reaction. See also below for specific embodiments, for example, for a feed stream combined with a diluent, wherein the feed stream is, for example, methanol, and is diluted to a methanol concentration in the feed of 2-20 vol%, preferably 5-10 vol%.

[0062] Therefore, in one embodiment, the feed stream is mixed with a diluent (i.e., an inert diluent such as nitrogen or carbon dioxide or a light paraffin such as methane) to reduce the exothermicity during the conversion to olefins. This is particularly preferred when the catalyst is arranged as a fixed bed. For example, when the feed stream is methanol, it is diluted with, for example, nitrogen so that the methanol concentration in the feed is 2-20 vol%, preferably 5-10 vol%.

[0063] It should be understood that the lower the methanol concentration in the feed, the higher the pressure required to maintain high methanol conversion because the methanol partial pressure (P MeOH ) is the actual relevant parameter to be tracked during the process operation. For example, when the methanol concentration in the feed is 10 vol%, and the MTO is at a P of 0.3 or 0.5 bar, MeOH This corresponds to a pressure of 3 or 5 bar in the MTO, respectively, when operated at P = 0. MeOH This corresponds to a pressure of 6 or 10 bar in the MTO when operating at low pressure.

[0064] Pressures at the higher end of this range, such as 15 bar, 20 bar, or 25 bar, can better match the pressure of the subsequent oligomerization or Hydro / OLI, thereby significantly saving compression energy, as will also be explained further below. Thus, the present invention also provides a process in which the pressure of the MTO can now be closely matched to the pressure of the subsequent oligomerization or Hydro / OLI, while still maintaining high conversion and olefin flow (olefin product flow), which is ideal for subsequent oligomerization and / or Hydro / OLI.

[0065] It has now also been found that at low temperatures in the MTO carried out according to the invention, the partial pressure (P MeOH ) does not play a decisive role in the selectivity to aromatics. This is contrary to the common understanding that methanol partial pressure plays a decisive role in the selectivity to aromatics and paraffins. This means that compared to operating the MTO at temperatures above 360 °C, where P MeOH It has a great influence on the production of aromatics. MeOH Will produce significantly higher aromatic content; by the present invention, wherein at 360 ℃ or lower temperature MTO, for example 320 ℃, change P MeOH did not show any significant effect, as the aromatic content was kept below 2 wt% or below 1 wt%, and regardless of P MeOH Anyway, all at similar values.

[0066] For example, when running MTO at 400°C, MeOH When the P is 0.3 and 0.5, the aromatic content is about 2 wt% and 6 wt%, respectively. MeOH When the P is 0.3 and 0.5, the aromatic content is about 1 wt% and 2 wt%, respectively. MeOH When the values are 0.3 and 0.5, the MeOH The aromatic content is about 1 wt%.

[0067] Therefore, P MeOH The higher the temperature, the higher the aromatic content. According to the present invention, by lowering the temperature, the P MeOH , thereby increasing the pressure (total pressure) without producing more aromatics.

[0068] Therefore, by the present invention, at lower MTO temperatures, for example at 350°C or lower, such as 340°C or 320°C or 300°C, the aromatic content relative to P MeOHThe higher independence of the process also enables operation at higher pressures, such as 15, 20, or 25 bar. As mentioned above, these pressures are not only better matched to downstream operations (such as oligomerization or Hydro / OLI), but they are also closer to the pressures used in upstream processes, especially methanol synthesis, which operates at high pressures, typically around 50-100 bar. This results in higher energy savings in terms of lower compression energy, as well as a reduction in equipment size.

[0069] Despite the relatively low temperatures employed, i.e., reaction temperatures of 240-360°C, e.g., 260-360°C or 260-340°C, the catalyst is active not only in suppressing the formation of aromatics, but is also active in providing high selectivity to higher olefins C3=-C8=, no ethylene formation, significant isoparaffin formation, complete conversion, while also exhibiting extended catalyst life.

[0070] For example, it has been found that when used at the low temperatures described, ZSM-48 converts methanol into an olefin stream that is ideal for further oligomerization into jet fuel, particularly SAF according to ASTM as defined above. The olefin product is essentially free of ethylene and aromatics (which are considered undesirable), while the yield of C3-C8 olefins is 70-80%, combined with 10-15% isoparaffins, a fact that makes this product an ideal feed for further oligomerization into SAF.

[0071] Compared to the prior art according to US Pat. No. 4,476,338, in which MTO is carried out on ZSM-48 having a SAR of 113 or 180 and a temperature of 370° C. (Examples 1 and 2 therein), in the present invention, when MTO is carried out at a temperature of 360° C. or lower using, for example, ZSM-8 having a lower SAR, the yield of total olefins (e.g., C2-C8 olefins) is higher; the yield of ethylene is lower, for example, the ethylene content is now below 1 wt %; the yield of aromatics is lower, for example, the content of aromatic compounds is now below 1 wt %; and optionally, the yield of isoparaffins is higher, for example, now 10-15 wt %. In addition, the catalyst life is increased, as described below.

[0072] Operating the conversion of oxygenates using, for example, ZSM-48 with a SAR as high as 110 in combination with lower temperatures (300-360°C) results in at least three highly beneficial effects:

[0073] a) In either case, the selectivity to ethylene and aromatics drops to below 1 wt%.

[0074] b) A large amount of isoparaffins can be formed that can be used in the process. Isoparaffins, as well as C3-C8 olefins, can also be oligomerized so that isoparaffins can be formed as a desired by-product. The isoparaffins can optionally be separated for alkylation to increase octane and then blended into the gasoline pool, or simply used as part of the olefin stream for downstream oligomerization.

[0075] c) Due to the lower operating temperature, the overall catalyst life (number of cycles) is increased due to the lower dealumination rate (affected by the combination of high temperature and water vapor generated during the reaction). Furthermore, the catalyst life during each cycle (i.e., cycle time) is significantly increased. Without being bound by theory, this may be due to the lower selectivity for aromatics caused by fewer hydrogen transfer reactions. High catalyst life, both in terms of overall life (number of cycles) and cycle time, is crucial for achieving practical commercial use.

[0076] For the purposes of this application, the terms "catalyst longevity" and "catalyst lifetime" are used interchangeably.

[0077] Furthermore, if a binder is included in the catalyst, which is relevant for commercial applications, the present invention will also enable higher yields of desired products (e.g., C3-C8 olefins) since no or limited cracking of MeOH / DME to methane occurs.

[0078] Thus, the features of the present invention work synergistically to produce a superior process that is commercially applicable for converting oxygenates to olefins for subsequent downstream steps, such as oligomerization.

[0079] While suitable oligomerization feeds typically may have some aromatics, e.g., 10-20 wt% aromatics, as well as higher olefins and ethylene, ideal oligomerization feeds are essentially free of aromatics and comprised of higher olefins, preferably with as little C2-light ends as possible, and more specifically, free of ethylene. The olefin stream may comprise at least 70 wt% C3-C8 olefins, e.g., 75-85 wt% C3-C8 olefins, e.g., 80 wt% C2-C8 olefins, 1 wt% or less aromatics, 1 wt% or less ethylene, and 10-15 wt% isoparaffins. Lower temperatures increase the higher olefin content, and thus the higher olefin to ethylene ratio, i.e., the higher the selectivity for higher olefins. Furthermore, lower temperatures decrease the ethylene content, and thus the olefin stream is essentially free of ethylene, while the isoparaffin content increases. Thus, the oligomerization feed meets the aforementioned ASTM requirements, which stipulate that the 50% SAF blending portion is virtually free of aromatics, and more specifically, the aromatic content is limited to less than 0.5 wt%. Olefin streams can be converted to such jet fuels through oligomerization and hydrogenation in a more efficient overall process due to, for example, less recycle and higher oligomerization yields. In other words, the higher olefins and low selectivity to aromatics and ethylene simplify the separation steps and increase the overall yield of jet fuel.

[0080] By using the medium high pressure of 2-25 bar, for example 2,10,15 or 20 bar, can further selectivity be transferred to higher olefins now.That is to say, have found that, although higher pressure has increased the ratio of higher olefins and lower olefins, promptly to the higher selectivity of higher olefins, higher pressure also may reduce the total yield of olefins (being that the conversion efficiency of oxygenate feed to olefins is lower) and also improved and realized fully converted required temperature, this produces the risk of the cracking reaction that does not too expect again.Under the pressure range of 1-25 bar, can obtain the higher selectivity to higher olefins now, and do not need to be increased in temperature to for example higher than 360 ℃ to realize fully converted, thereby also reduced the generation of cracking reaction.By carrying out this method under the pressure higher than 1 bar (for example, 2 bar or higher), can also provide a certain amount of diluent to be used for " hot groove " of exothermic reaction, as further explained above.The present invention has realized the high degree of flexibility that pressure is selected, for example at the higher end of pressure range, for example 20 or 25 bar, as further explained above.

[0081] At the same time, the temperature is lowered to, for example, 300°C or 320°C or lower, which, although in principle means a reduction in methanol conversion, in practice still maintains oxygenate conversion close to 100%, while keeping the aromatics content below 1 wt% and the ethylene content below 1 wt%.

[0082] The catalyst may be prepared by standard methods in the art, such as ZSM-48 as disclosed in US 4,476,338. Suitably, the final catalyst is mixed with a binder / matrix, for example in a catalyst containing up to 50-80 wt% zeolite in a matrix / binder comprising an alumina component (e.g. a silica-alumina matrix binder).

[0083] In one embodiment, the process further comprises recycling a portion of the olefin-containing stream to the feed stream, said portion of the olefin-containing stream being suitably a stream comprising C2-C3 olefins or a C3 olefin stream (propylene stream) withdrawn from said olefin stream.

[0084] In other words, the process also includes recycling a portion of the olefin stream to the feed stream and using it as an additional feed stream, i.e., as a co-feed, which may include recycling light paraffins (including methane) acting as diluents and reducing the adiabatic temperature increase, for example, by combining with the feed stream containing oxygenates. Thus, the operating temperature in the oxygenate conversion can be further reduced, thereby improving selectivity and catalyst life.

[0085] The concentration of higher olefins in the olefin stream is further increased, while less desirable lower olefins are also fully utilized for conversion to higher olefins. By recycling these cracked products (i.e., C2-C3 olefins) back into the feed, the cracking of any undesirable higher olefins in the process can be suppressed. In addition, the recycle also suitably contains light paraffins and optionally isoparaffins, which further provide a dilution effect on the feed stream, thereby enabling better control of the exothermicity during the conversion to olefins.

[0086] The co-feed stream (ie, recycle stream) is 1 to 20 times, such as 2 to 10 times, the volume of the feed stream (eg, methanol feed stream).

[0087] In particular embodiments, the recycle stream contains 0.5-10% or 1-10% mol propylene and the concentration of methanol in the feed is 10 vol%.

[0088] In particular, it has been found that:

[0089] - Co-feeding lower (light) olefins, such as C3 olefins, can initiate MTO at lower temperatures. Therefore, it is possible to further reduce the temperature, for example to 350°C, 340°C, 320°C or 300°C, or even lower, to 280°C, 260°C or even 240°C under adiabatic conditions. This has the benefit of further reducing hydrogen transfer, thereby reducing the production of paraffins, isomers and aromatics, while providing greater freedom in terms of higher olefin chain lengths, i.e., higher olefins, and total (or methanol partial) pressure; for example, the pressure can be increased, which may be advantageous, as mentioned above.

[0090] - Furthermore, it was surprisingly found that when a co-feed is provided, the catalyst lifetime is also significantly increased compared to when no co-feed is provided, ie by using pure methanol feed.

[0091] It will be understood that the term "catalyst life" or "catalyst life period" includes not only the total life of the catalyst (number of cycles), but also the life during each cycle, i.e., the cycle time. The term "cycle time," also known as "cycle length," is the length of time a catalyst exhibits adequate catalytic activity, which is typically measured in hours on stream (HOS). From a process perspective, this is very beneficial because the recycle stream of C2-C3 olefins will also make it easier to control the exothermicity of the MTO.

[0092] In one embodiment, the weight hourly space velocity (WHSV) is 0.1-3h -1 , for example 1-2h -1 The higher the WHSV value, the lower the methanol conversion rate.

[0093] In one embodiment, the feedstream comprising oxygenates is derived from one or more oxygenates selected from triglycerides, fatty acids, resin acids, ketones, aldehydes, or alcohols or ethers, wherein the oxygenates are derived from one or more of a biosource, a gasification process, a pyrolysis process, a Fischer-Tropsch synthesis, or a methanol-based synthesis. In a particular embodiment, the one or more oxygenates are hydrotreated oxygenates. "Hydrotreated oxygenates" refers to oxygenates, such as esters and fatty acids, derived from a hydrotreating step such as hydrotreating and hydrocracking.

[0094] In one embodiment, the oxygenate is selected from methanol (MeOH), dimethyl ether (DME), or a combination thereof. These are particularly advantageous oxygenate feedstocks because they are widely commercially available. DME is more reactive than methanol, and therefore the MTO step can be operated at a lower temperature, thereby increasing the selectivity for higher olefins. In addition, the conversion of DME releases only half the amount of water (steam) compared to methanol, thereby reducing the (irreversible) deactivation rate caused by steam dealumination of the zeolite catalyst.

[0095] Suitably, water is removed from the olefin stream produced in the MTO as its presence may be undesirable when subjected to downstream oligomerisation.

[0096] In one embodiment, methanol is produced from synthesis gas (ie, methanol syngas) produced using electricity from renewable energy sources (eg, wind or solar), such as eMethanol. TM. Thus, in one embodiment, syngas is suitably prepared by combining air separation, autothermal reforming or partial oxidation and water electrolysis, as disclosed in applicant's WO 2019 / 020513 A1, or from syngas produced by electrothermal reforming, as disclosed, for example, in applicant's WO 2019 / 228797. Thus, an even more sustainable method for producing jet fuel, in particular SAF, is achieved. Methanol can be produced from many primary resources, including biomass and waste, and where wind and solar power generation are low cost, e-methanol TM The production of PET enables even more sustainable front-end solutions.

[0097] In a particular embodiment, the process of the present invention further comprises, prior to passing the oxygenate-containing feedstock stream over a catalyst active in oxygenate conversion (where the oxygenate-containing feedstock stream is a methanol stream, i.e., a methanol feed stream):

[0098] The methanol feed stream is produced by methanol synthesis from methanol synthesis gas, wherein the methanol synthesis gas is produced by steam reforming of a hydrocarbon feed such as natural gas, and / or at least in part by electrolysis of water and / or steam.

[0099] Thus, in another specific embodiment, the methanol feed stream is produced from methanol synthesis gas, which is produced by using water electrolysis in an alkaline or PEM electrolysis unit or using steam in a solid oxide electrolysis cell (SOEC) unit, thereby producing a hydrogen stream, and using a CO2-rich stream in the SOEC unit to produce a stream comprising carbon monoxide and carbon dioxide, and then combining the hydrogen stream and the stream comprising carbon monoxide and carbon dioxide to produce the methanol synthesis gas, as disclosed, for example, in the applicant's co-pending European patent application No. 20216617.9. The methanol synthesis gas is then converted into a methanol feed stream by a methanol synthesis reactor, as is well known in the art.

[0100] Methanol synthesis gas, as is well known in the art, is a mixture comprising primarily hydrogen and carbon monoxide, which mixture is suitable for methanol synthesis, i.e., a methanol synthesis gas having a module M = (H2-CO2) / (CO+CO2). Methanol synthesis gas for methanol synthesis is usually described in terms of the module M because when M = 2, the synthesis gas is in equilibrium for the methanol reaction.

[0101] Thus, an alternative highly sustainable front-end solution for generating a methanol feed stream (ie methanol synthesis gas) is provided, whereby electrolysis alone is utilized to generate the methanol synthesis gas and thereby methanol.

[0102] Thus, it should be understood that the term "process," as used herein, may also include the prior (front-end) production of the methanol feed stream, as described above.

[0103] In one embodiment, the catalyst is contacted with a stream of hydrogen. Hydrogen can increase methanol conversion by at least slightly reducing the rate of catalyst deactivation, thereby increasing catalyst life. However, when performing this method, hydrogen is not added because this would carry the risk of hydrogenating some olefins, thereby reducing olefin yield.

[0104] In one embodiment, the catalyst is arranged as a fixed bed.

[0105] In one embodiment, the method comprises: using a first reactor group comprising a single reactor or a plurality of reactors (preferably arranged in parallel with each other) for partially or completely converting the oxygenate. Thus, large feedstocks comprising one or more oxygenates can be processed simultaneously.

[0106] It is to be understood that the term "with respect to each other" refers to a parallel arrangement between the reactors of a reactor group, for example between the reactors of the first reactor group.

[0107] In a particular embodiment, the process further comprises the use of a second reactor group comprising a single reactor or a plurality of reactors (preferably arranged in parallel with each other) to further convert the oxygenates and a phase separation stage located between the first reactor group and the second reactor group to form an olefin stream.

[0108] As used herein, the term "using a first reactor train" means passing the feedstock comprising oxygenates through the first reactor train. As used herein, the term "using a second reactor train" means passing the feedstock or a portion thereof through the second reactor train and through a separation stage after partial or complete conversion of the oxygenates.

[0109] Thus, large feedstocks containing one or more oxygenates can be processed simultaneously and lower temperatures can be used in both reactor trains, which increases the lifetime conversion capacity of the catalyst and also improves the selectivity to higher olefins due to less cracking.

[0110] In one embodiment, the entire feed stream is passed through the first reactor train, ie, there is no substantial splitting of the feed stream.

[0111] As used herein, the term "entire feedstock" refers to at least 90 wt% of the feedstock.

[0112] In another specific embodiment, the method comprises:

[0113] - passing a feed stream comprising oxygenates through a first reactor group under conditions that partially convert the oxygenates (e.g., 40-80%, such as 60-70% conversion), thereby forming a crude olefin stream comprising unconverted oxygenates and C2-C8 olefins, for example, the crude olefin stream may comprise water, methanol and C2-C8 olefins;

[0114] - passing the crude olefin stream through said phase separation stage to produce:

[0115] a first olefin stream rich in lower olefins;

[0116] a separated oxygenate stream comprising unconverted oxygenates, for example, the separated oxygenate stream may comprise water and methanol;

[0117] a second olefin stream rich in higher olefins;

[0118] - combining the first olefin stream with a separated oxygenate stream comprising unconverted oxygenates, thereby forming a combined stream comprising lower olefins and unconverted oxygenates;

[0119] - passing the resulting combined stream comprising lower olefins and unconverted oxygenates through a second reactor train, e.g., into the first reactor of the second reactor train, under conditions such that the unconverted oxygenates and lower olefins are fully converted (e.g., 85%, 90%, 95% or more) to a third olefin stream rich in higher olefins;

[0120] - combining the second olefin stream (which may be considered as a by-pass stream of the second reactor train) with the third olefin stream, thereby forming said olefin stream which is preferably rich in higher olefins and substantially free of aromatics.

[0121] For example, the olefin stream, ie, the olefin product stream, contains less than 1 wt% aromatics. The olefin stream also suitably contains no ethylene, eg, less than 1 wt%, while having a significant amount of isoparaffins, eg, 10-15 wt%.

[0122] Therefore, the operational flexibility of increase is achieved, particularly when processing large-scale feed streams, and it is not necessary, for example, before entering the first reactor to convert oxygenates, feed streams are separated and transferred to a separate olefin interconversion reactor, as disclosed in, for example, US 5,177,279. In addition, the temperature in all reactor groups can be even further reduced, for example, down to 250-350 ℃, but still achieve the complete conversion of oxygenates, for example, conversion rate is up to 100%. In addition, further reducing temperature also can increase catalyst life. In addition, in processing various raw materials comprising oxygenates (comprising the fatty acid in renewable feed) or oxygenates derived from one or more bio-sources and optionally in two different reactor groups, processing different types of catalysts, there is the flexibility of increase.

[0123] In another specific embodiment, the first reactor group and the second reactor group use a catalyst having a one-dimensional (1-D) pore structure, such as a *MRE, eg, ZSM-48.

[0124] It is also possible to recycle C2-C3 olefins as co-feed. For example, in one embodiment, the method also includes recycling a portion of an olefin stream (that is, an olefin product stream from the second reactor group) to the combined stream comprising lower olefins and unconverted oxygenates, and feeding it to the second reactor group, the olefin stream portion preferably being an olefin stream comprising C2-C3 olefins, more preferably a C3 olefin stream taken out from the olefin stream. The same related benefits related to recycling C2-C3 olefins are also obtained.

[0125] In another specific embodiment, the first reactor group consists of 2-4 reactors (eg, 3 reactors), and the second reactor group consists of 1-3 reactors (eg, 2 reactors).

[0126] In the first and second reactor groups, the reactors are preferably arranged in parallel with each other. In large MTO plants processing large feed streams, several reactors (e.g. five (5) reactors) are usually operated in parallel. By means of the present invention, it is possible to replace five reactors in parallel by, for example, a first reactor group consisting of three reactors and a second reactor group consisting of two reactors. Thus, full conversion can be achieved by operating the first three reactors, for example, at a conversion rate of only 70%, and then further converting the unconverted oxygenates (e.g. methanol) together with the C2-C3 olefins to 100% in two reactors arranged in series with the first three. Likewise, the temperature of all five reactors is reduced, but full conversion is achieved. Flexibility is also improved by taking one reactor out of operation for regeneration.

[0127] It will be appreciated that the first reactor group and the second reactor group are arranged in series.

[0128] In another specific embodiment, the reactors in the first and second reactor trains are operated at 1-25 bar (eg, 1-15 bar or 2-25 bar) and 240-360°C (eg, 300-360°C, such as 320°C or 340°C).

[0129] In another specific embodiment, the weight hourly space velocity (WHSV) is 0.1-3h -1 , for example 1-2h -1 , such as 6, 8 or 10 hours -1In yet another particular embodiment, the weight hourly space velocity (WHSV) in the first reactor train is higher than in the second reactor train. For example, in the first reactor train intended for partial conversion of the oxygenate feedstock, the WHSV is suitably 3 h -1 or 6 hours -1 , while in the second reactor group where complete conversion is intended, the WHSV is suitably 2h -1 .

[0130] In one embodiment, the method further comprises:

[0131] - Separation of an isoparaffin stream from an olefin stream.

[0132] In one embodiment, the method further comprises:

[0133] - passing at least a portion of the olefin stream through an oligomerization step over an oligomerization catalyst, and optionally followed by a separation step, such as after separation of the isoparaffin stream, thereby generating an oligomerization stream.

[0134] Isoparaffins, as well as C3-C8 olefins, can also be oligomerized. Thus, rather than producing unwanted aromatics as a by-product, the present invention now provides isoparaffins as a desired product, which can optionally be separated for use as an alkylation feed, thereby increasing the octane rating of the gasoline optionally produced in the process. Providing an isoparaffin stream separation step also increases flexibility in selecting the zeolite structure used in the oligomerization step.

[0135] In one embodiment, preferably after withdrawing said olefin stream comprising C2-C3 olefins therefrom, the entire olefin stream is passed through the oligomerization step. As used herein, the term "entire olefin stream" refers to at least 90 wt% of the stream.

[0136] In one embodiment, the olefin stream, for example the entire olefin stream after separation of the isoparaffin stream, is fed directly to the oligomerization step, i.e., the olefin stream is in direct fluid communication with the oligomerization step or the combined oligomerization and hydrogenation step, as further described below. Thus, there is no fractionation of the olefin stream prior to feeding it to the oligomerization step, thereby further simplifying the process and apparatus.

[0137] The oligomerization step is preferably carried out by conventional methods, including the use of an oligomerization catalyst, such as solid phosphoric acid ("SPA"), an ion exchange resin or a zeolite catalyst, such as conventional *MRE, BEA, FAU, MTT, TON, MFI and MTW catalysts, and a pressure of 30-100 bar, such as 50-100 bar, and a temperature of 100-350° C. The products from the oligomerization reaction can then be separated in a separation step (e.g., distillation) to extract a lighter hydrocarbon stream (e.g., naphtha, which contains C5-C7 hydrocarbons) and an oligomerized stream (which contains C8+ hydrocarbons).

[0138] In one embodiment, the method further comprises:

[0139] At least a portion of the oligomerization stream is passed to a hydrogenation step over a hydrogenation catalyst and optionally followed by a separation step, thereby producing a hydrocarbon stream comprising hydrocarbons boiling in the jet fuel range.

[0140] The hydrogenation step is preferably carried out by conventional methods, including the use of a hydrotreating or hydrogenation catalyst, such as a catalyst comprising one or more metals, such as Pd, Rh, Ru, Pt, Ir, Re, Co, Mo, Ni, W, or a combination thereof, in the presence of hydrogen, at a pressure of 60-70 bar and a temperature of 50-350° C. The C8+ hydrocarbons in the oligomerized stream are thereby saturated to form the corresponding paraffins. These can then be separated in a separation step (e.g., a distillation step), thereby removing any hydrocarbons boiling in the diesel range from hydrocarbons boiling in the jet fuel range (i.e., jet fuel).

[0141] In one embodiment, the entire oligomeric stream is passed to the hydrogenation step.As used herein, the term "entire oligomeric stream" refers to at least 90 wt% of the stream.

[0142] In certain embodiments, the hydrocarbon stream comprising hydrocarbons boiling in the jet fuel range is SAF, ie, sustainable aviation fuel conforming to ASTM D7566 and ASTM D4054.

[0143] In one embodiment of the present invention, the oligomerization step and the hydrogenation step are combined in a single hydro-oligomerization step (Hydro-OLI), for example by combining these steps in a single reactor. In other words, a hydrocarbon stream comprising hydrocarbons boiling in the jet fuel range is produced by passing at least a portion of the olefin stream through an oligomerization step and a hydrogenation step combined in a single hydro-oligomerization step, optionally followed by a separation step. This allows for a simpler process / apparatus layout.

[0144] As used herein, the term "single hydrooligomerization step" or more generally "single step" or "single stage" refers to a section of a process from which no stream is withdrawn. Typically, a single stage does not include equipment such as a compressor by which pressure is increased.

[0145] The oligomerization step is a dimerization and, optionally, a trimerization, i.e., the oligomerization is carried out under conditions suitable for dimerization and / or trimerization. Therefore, the individual reactors are preferably operated at relatively low pressures, such as 15-60 bar, for example, 20-40 bar. The oligomerization reaction is significantly exothermic in each oligomerization step, and much less heat is generated because only dimerization and, optionally, trimerization occur, rather than higher-order oligomerizations such as tetramerization or even pentamerization. The lower heat generated facilitates a near-equilibrium state, i.e., a higher olefin conversion rate.

[0146] Typically, the oligomerization step converts olefins into a mixture of primarily dimers, trimers, and tetramers or even pentamers; for example, C6 olefins will produce a mixture comprising C12, C18, C24 products and possibly also higher hydrocarbons. By carrying out the oligomerization step under conditions suitable for dimerization, and optionally trimerization, a more selective and direct conversion of higher olefins (C3-C8 olefins, including C4-C8 olefins) to jet fuel-related hydrocarbons (i.e., C8-C16) is achieved. The dimerization and optional trimerization steps involve the use of lower pressures than conventional oligomerization processes, thereby also reducing compression requirements, which translates into higher energy efficiency (due to lower compression energy) and reduced costs, such as reduced costs for the oligomerization reactor and ancillary equipment, and reduced operating costs due to the reduced need to separate the C16+ olefins otherwise formed in conventional OLI reactors. Therefore, the pressure of the Hydro / OLI can be adjusted to better match the pressure of the previous oxygenate conversion step.

[0147] Furthermore, only one subsequent separation stage, if any, is required, compared to using a dedicated separation such as distillation in the OLI step to separate the naphtha (which can be upgraded to a gasoline product) and another dedicated separation in the hydrogenation step to separate the diesel from the jet fuel. This results in a simpler oligomerization and hydrogenation process and, therefore, a simpler overall process and apparatus.

[0148] The hydrogenation or H addition is carried out in the same reactor, for example by adjusting the activity of the hydrogenation component (e.g., nickel). In one embodiment, the single hydro-oligomerization step is carried out in a single reactor having stacked reactor beds, wherein the first bed comprises an oligomerization catalyst, such as a zeolite catalyst, and subsequent beds comprise a hydrogenation catalyst.

[0149] The hydrooligomerization step is carried out by reacting the olefin stream over a catalyst comprising a zeolite and a hydrogenation metal (e.g., a hydrogenation metal selected from the group consisting of Pd, Rh, Ru, Pt, Ir, Re, Co, Cu, Mo, Ni, W, and combinations thereof) in the presence of hydrogen, e.g., after separation of the isoparaffin stream, preferably at a pressure of 15-60 bar (e.g., 20-40 bar) and a temperature of 50-350° C. (e.g., 100-250° C.). In a particular embodiment, the catalyst comprises a zeolite having a structure selected from the group consisting of MFI, MEL, SZR, SVR, ITH, IMF, TUN, FER, EUO, MSE, *MRE, MWW, TON, MTT, FAU, AFO, AEL, and combinations thereof, preferably a zeolite having a framework with a 10-ring pore structure (i.e., a pore perimeter defined by 10 oxygens), e.g., a zeolite having a structure selected from the group consisting of TON, MTT, MFI, *MRE, MEL, AFO, AEL, EUO, FER, and combinations thereof. These zeolites are particularly suitable due to the confined space of the zeolite pores, which causes dimerization to favor larger molecules. Optionally, the weight hourly space velocity (WHSV) is 0.5-6h -1 , for example 0.5-4h -1 .

[0150] The lower pressure corresponding to operation under dimerization (optionally also trimerization) conditions is in particular 15-50 bar, for example 20-40 bar. This is also significantly lower than the pressures usually used for oligomerization, which are usually in the range of 50-100 bar.

[0151] The present invention purposefully uses conditions that result in gentle hydrogenation. Particularly suitable catalysts are catalysts comprising NiW (e.g., sulfide NiW (NiWS)) or Ni (e.g., Ni supported on a zeolite with a FAU or MTT structure (e.g., Y zeolite or ZSM-23)). Catalysts active for oligomerization and hydrogenation can, for example, contain up to 50-80 wt% zeolite in a matrix / binder comprising an alumina component. A hydrogenation metal can then be incorporated by impregnation on the catalyst. The hydrogenation metal is selected to provide moderate activity, thereby better controlling the exothermicity of the oligomerization step by primarily hydrogenating the dimer being formed when oligomerization occurs, thereby interrupting the formation of higher oligomers.

[0152] Thus, rather than having separate reactors and accompanying separation units for oligomerization and subsequent hydrogenation (each with its own catalyst), the present invention enables the use of less equipment in a single hydro-oligomerization step, such as a single reactor, one type of catalyst, and optionally a single separation stage downstream to obtain jet fuel. This results in a more efficient and simpler overall process and apparatus for converting oxygenates (e.g., methanol) into jet fuel, particularly SAF.

[0153] In one embodiment, a stream comprising C8 hydrocarbons resulting from cracking C9-C16 hydrocarbons is withdrawn from the hydrocarbon stream comprising hydrocarbons boiling in the jet fuel range and fed to another process. For example, the process according to the invention can be integrated with a refinery (or process), particularly a biorefinery, and the stream comprising C8 hydrocarbons is added to the gasoline pool in a separate process for producing gasoline at the refinery. Optionally, a stream comprising C8 hydrocarbons resulting from cracking C9-C16 hydrocarbons is withdrawn from the hydrocarbon stream comprising hydrocarbons boiling in the jet fuel range and fed to (recycled to) the oligomerization step or a single hydrooligomerization step as an additional feed stream.

[0154] In a second aspect of the present invention, there is also provided an MTO method for producing an olefin stream, the method comprising passing a feed stream comprising oxygenates through a catalyst active in the conversion of oxygenates at a pressure of 1-25 bar (e.g., 2-25 bar), wherein the catalyst comprises a zeolite having a skeleton having a 10-ring pore structure, wherein the 10-ring pore structure is a one-dimensional (1-D) pore structure, wherein the feed stream is combined with a diluent, the feed stream is methanol, and is diluted to a methanol concentration in the feed of 2-20 vol.%, preferably 5-10 vol.%, and wherein the 1-D pore structure is any one of *MRE (ZSM-48), MTT (ZSM-23), TON (ZSM-22) or a combination thereof.

[0155] It should be understood that the methanol concentration range and pressure range correspond to the following methanol partial pressure (P MeOH ) operation: 0.02-5 bar, preferably 0.05-2.5 bar. For example, at a pressure of 1 bar (total pressure) and a methanol concentration of 2 vol%, P MeOH is 0.02 bar, while at a pressure of 25 bar (total pressure) and a methanol concentration of 2 vol%, P MeOH = 0.5 bar (= 25 * 0.02). At a pressure of 1 bar (total pressure) and a methanol concentration of 20 vol%, P MeOH is 0.2 bar, while at a pressure of 25 bar (total pressure) and a methanol concentration of 20 vol%, P MeOH It is 5 bar (=25*0.2).

[0156] By combining the feed with a diluent (e.g. nitrogen or carbon dioxide) or a recycle stream of C2-C3 olefins or C3 olefins (including, for example, light paraffins, including methane used as a diluent), the exothermicity of the conversion to olefins is reduced by using the light paraffins as a heat sink; this is particularly preferred when the catalyst in the MTO is arranged as a fixed bed.

[0157] Any embodiments and associated benefits of the first aspect of the invention may be applied to the second aspect of the invention, and vice versa.

[0158] BRIEF DESCRIPTION OF THE DRAWINGS

[0159] Figure 1 is a simplified diagram showing the conversion of oxygenates to olefins and optional further conversion to jet fuel according to one embodiment of the present invention.

[0160] Figure 2 is a simplified diagram showing one specific embodiment of the present invention for converting a feedstock comprising oxygenates into olefins and optionally further into jet fuel.

[0161] Figure 3 A graph showing methanol conversion and C3-C8 selectivity as a function of temperature according to Example 1 is shown.

[0162] Figure 4 A graph showing methanol conversion and yield as a function of catalyst cycle time (hours on stream, HOS) at a specific temperature of 360° C. according to Example 1 is shown.

[0163] Figure 5 The results show that the aromatic content in the olefin stream varies with temperature and methanol partial pressure (P MeOH ) changes in the graph.

[0164] Figure 6 Shown is a graph of methanol conversion as a function of temperature for a feed comprising propylene as a co-feed and a pure methanol feed according to Example 4.

[0165] Detailed description

[0166] refer to Figure 1 A feedstock comprising oxygenates 100 (e.g., methanol and / or DME) is directed together with an optional hydrogen stream 102 and an olefin stream 104 comprising C2-C3 olefins, which is withdrawn from an olefin stream 106 formed in an oxygenate conversion section 200. The oxygenate conversion section 200 (e.g., an MTO section) converts the oxygenates over a zeolite catalyst (e.g., ZSM-48 having a SAR of, for example, up to 110) at, for example, 1-15 bar and 260-360° C. (e.g., 300-360° C. or 300-350° C.). The olefin stream 106 obtained under these conditions is rich in higher olefins (C3=-C8=) and small amounts of aromatics, and is optionally further converted (as indicated by the dashed line) into a hydrocarbon stream 112 comprising hydrocarbons boiling in the jet fuel range (C8-C16).

[0167] This further conversion is carried out in the downstream oligomerization and hydrogenation section 300, which is preferably combined into a single hydrogenation-oligomerization step, for example in a single reactor. The olefin stream 106, after appropriate removal of its water content, is mixed with an optional oligomerized olefin stream 110, which comprises C8-hydrocarbons and is obtained from cracked C9-C16 hydrocarbons, which are taken from a hydrocarbon stream 112 comprising hydrocarbons boiling in the jet fuel range. The resulting mixed stream is then directed to the section 300 and converted into a hydrocarbon stream 112 comprising hydrocarbons boiling in the jet fuel range in the presence of hydrogen fed as stream 108, over a catalyst (e.g. Ni supported on a zeolite having a FAU or MTT structure, such as Y zeolite or ZSM-23), at, for example, 20-40 bar and 50-350° C. Under these conditions, particularly at lower pressures, the single reactor in section 300 operates so that the oligomerization is dimerization and, optionally, trimerization, with hydrogenation activity also present. Due to the higher olefins, isoparaffins, low aromatics (e.g., less than 1 wt%), and low ethylene (e.g., less than 1 wt%) of olefin stream 106, the hydrocarbons boiling in the jet fuel range (i.e., jet fuel) in stream 112 can be used as SAF.

[0168] refer to Figure 2A feed stream 100 comprising oxygenates such as methanol and / or DME passes through a first reactor group 200', for example three reactors arranged in parallel, to achieve a methanol conversion rate of 50-70% and produce a crude olefin stream 105 comprising water, methanol and olefins (e.g., C2-C8 olefins). The crude olefin stream 105 is separated in a three-phase separator 200" to produce a first olefin stream 105a rich in lower olefins, in particular C2-C3 olefins or mainly C2 olefins (ethylene), a separated oxygenate stream 105b containing unconverted oxygenates (e.g., unconverted methanol) and a second olefin stream 105c rich in higher olefins, in particular C3-C8 olefins, including C4-C8 olefins. The first olefin stream 105a is combined with the separated oxygenate stream 105b containing unconverted oxygenates to form a combined stream 105d comprising lower olefins, in particular C2-C3 olefins or mainly ethylene and unconverted oxygenates. The combined stream is pressurized and fed to a second reactor group 200'" arranged downstream, which may, for example, include two reactors arranged in parallel to achieve a complete conversion rate of, for example, 85% or 90% or more. Thus, the first reactor group 200' and the second reactor group 200"' are arranged in series. A third olefin stream 105e rich in higher olefins, particularly C3-C8 olefins, is produced. Finally, the second olefin stream 105c (bypass stream) is combined with the third olefin stream 105e to form the olefin stream 106 which may have been pressurized. By the above arrangement of the MTO section 200, the reactors of the first reactor group and the second reactor group can be operated at low temperatures, for example, 250-350°C or 260-360°C, appropriately at a lower temperature than using Figure 1 The embodiment of the present invention is operated at a lower temperature, which helps to improve the lifetime conversion capacity of the catalyst used and improve the selectivity to higher olefins due to less cracking. The resulting olefin stream 106, after appropriate removal of its water content, is optionally further converted in a downstream oligomerization and hydrogenation section 300 (as shown by the dashed line), which is combined into a single hydrogenation-oligomerization step, for example in a single reactor, thereby producing a hydrocarbon stream 112 containing hydrocarbons boiling in the jet fuel range (C8-C16), in particular SAF, as described with respect to Figure 1 Explanatory. Example

[0169] Example 1 Product Selectivity in MTO

[0170] The MTO test was run in a fixed catalyst bed (fixed bed) reactor with a zeolite catalyst ZSM-48 (EU-2) having a one-dimensional pore structure and a silicon to aluminum ratio (SAR) of 110, and was run under the following conditions: zeolite catalyst loading: 250 mg cat / 750 mg SiC, pressure = 1 barg (2 bar), space velocity (WHSV) = 2 h -1 , total flow rate = 3.5NL / h (59mL / min); methanol concentration in the feed (C MeOH ) = 10% (based on volume), with nitrogen as the diluent. Therefore, P MeOH The pressure is 0.2 bar and the operating temperature range is 320-360°C.

[0171] Figure 3 The change of methanol conversion with temperature is shown. It is observed that at 320°C, the conversion is almost 100%, and at 360°C, the conversion is 100%. Aromatics are formed, but remain at a low level, i.e., below 1wt%, more specifically, at 360°C, it is about 0.5wt%, at 320°C, it is close to 1wt%. Therefore, the selectivity for aromatics formation is low. Simultaneously, in the temperature window of 320-360°C, the content of isoparaffins in the olefin stream increases by 10-15wt% along with the reduction of temperature, while the content of C2 olefins (ethylene, represented as O2 in the figures) decreases with temperature and becomes 1wt% or lower in the same temperature window, thereby providing an olefin stream that does not contain ethylene. For example, at 320°C, the ethylene content is as low as 0.2wt%. The content of higher olefins (C3-C8, represented as O3-O8 in the figures) remains at a high level, i.e., within the 70-80wt% range of the olefin stream. Olefins having more than nine (9) carbons (denoted as O9+ in the figure) are also kept at low levels.

[0172] The table below shows the olefin distribution (in wt%) in the olefin stream.

[0173]

[0174] *Undetectable

[0175] Figure 4 Methanol conversion and yield (Y-axis, mass %) are shown as a function of operating time (X-axis, hours) at a specific temperature of 360° C. It is observed that the catalyst cycle time remains relatively long, thus making it suitable for commercial, i.e., industrial, applications.

[0176] Compared to the prior art according to US Pat. No. 4,476,338, in which MTO is carried out on ZSM-48 with a SAR of 113 or 180 and at 370° C. (Examples 1 and 2 therein), the olefin stream according to the present invention achieves a higher total olefin yield (e.g., C2-C8 olefins); a lower ethylene yield, e.g., the ethylene content is now below 1 wt. %; a lower aromatics yield, e.g., the content of aromatic compounds is now below 1 wt. %; and a higher isoparaffin yield, e.g., now 10-15 wt. %. In addition, the catalyst life is increased.

[0177] Example 2 Effect of Binder on Catalyst in MTO

[0178] The same test using ZSM-48 (EU-2) was carried out under the following conditions: WSHV = 2h -1 , methanol concentration in the feed (C MeOH ) = 8% (based on volume), nitrogen as diluent, pressure of 5 bar, using a catalyst with a binder: 60 wt% zeolite and 40 wt% alumina. The table below shows the effect of adding a binder.

[0179] As shown in the table below, at high temperatures, i.e., above 360°C (e.g., 400, 440, or 480°C), a significant amount of paraffins is formed, almost entirely in the form of methane, which is likely the result of MeOH / DME cracking (i.e., cracking of MeOH and / or DME). At lower temperatures (360°C) corresponding to the present invention, cracking is negligible, resulting in efficient conversion of DME / MeOH to olefins.

[0180]

[0181] Example 3 Effect of methanol partial pressure in MTO

[0182] At WSHV = 2h -1 and methanol concentration in the feed (C MeOH )=10% (based on volume), the same test was carried out using ZSM-48(EU-2). Figure 5 The results show the aromatic content (total aromatics, wt%) measured as benzene (B), toluene (T), xylene (X) and ethylbenzene (total aromatics) as a function of temperature and methanol partial pressure P at different temperatures. MeOH At a given temperature, for example 320°C, each column is assigned a P MeOH :The column on the left is P MeOH = 0.2 bar, the middle column P MeOH = 0.3 bar, and the right column P MeOH= 0.5 bar. When the methanol concentration in the feed is 10%, the pressures (total pressure) at each given partial pressure are: 2 bar, 3 bar and 5 bar respectively.

[0183] It was observed that at low temperatures of 360° C. and below, the MeOH partial pressure does not appear to have any significant effect on the aromatics content in the resulting olefin stream. Thus, the (negative) effect of high MeOH partial pressure leading to higher aromatics yields appears to be significantly reduced at 360° C. and more or less disappears at temperatures below about 350° C. This enables an increase in the pressure used to carry out MTO, which provides benefits, for example, in terms of higher throughput in MTO and reduced equipment size in equipment for producing SAF.

[0184] like Figure 5 As shown, at lower temperatures, regardless of P MeOH The aromatic content is kept below 2 wt% and at similar values. For example, when running MTO at 400 °C, the aromatic content is kept below 2 wt% and at similar values. MeOH When the P is 0.3 and 0.5, the aromatic content is about 2 wt% and 6 wt%, respectively. MeOH When P is 0.3 and 0.5, the aromatic content is about 1 wt% and 2 wt%, respectively. MeOH At 0.3 and 0.5, in both P MeOH The aromatic content is about 1 wt%.

[0185] When a co-feed of light olefins such as C3 olefins (propylene) is provided, see Example 4, this allows for further reductions in temperature, which in turn leads to further reductions in hydrogen transfer (e.g., reduced formation of aromatics), higher olefin chain lengths, and, in addition, greater freedom with respect to total pressure (or methanol partial pressure), extending the ultimate catalyst life.

[0186] Example 4 Effect of lower olefin co-feed in MTO

[0187] The effect of adding propylene (propene) to methanol feed is as follows: Figure 6 shown.

[0188] In MTO, under the same reaction conditions, the same zeolite ZSM-48 zeolite (EU-2) was used to compare the case with 1 mol% propylene and without (i.e. pure methanol feed). Operating conditions: zeolite catalyst loading: 250 mg cat / 750 mg SiC, pressure = 2 barg (3 bar), space velocity (WHSV) = 2h -1 , total flow rate = 3.5NL / h E (59mL / min); methanol concentration in the feed (C MeOH )=10% (based on volume), nitrogen as diluent.

[0189] Figure 6 It is shown that adding propylene as a co-feed (upper line in the figure) significantly promotes the start or startup of oxygenate (methanol) conversion. In the operation of MTO, there will be a large amount of light olefins, i.e. C2-C3 olefins, in the recycle stream, which are suitable as a part of the olefin stream, and due to its exothermicity, it can be used for temperature control in MTO in any case. Adding lower olefins to the methanol feed, for example as a recycle stream, makes it possible to significantly reduce the inlet temperature of MTO, thereby reducing the content of aromatics and paraffins (as used herein, also including methane), while increasing the average olefin chain length, thereby increasing the content of higher olefins. In addition, the co-feed with lower olefins significantly increases catalyst life, such as measured by catalyst cycle time. In addition, hydrogen transfer reactions are minimized, not to mention that the lower temperature of MTO (e.g., 320°C) enables operation under a higher pressure range, which may also be advantageous, as described in conjunction with Example 3.

Claims

1. A method for producing an olefin stream, wherein the olefin stream comprises at least 70 wt% of C3-C8 higher olefins and less than 5 wt% of aromatics, the method comprising passing a feed stream comprising oxygenates through a catalyst active in the conversion of oxygenates at a pressure of 1-25 bar and a temperature of 240-360°C, wherein the catalyst comprises a zeolite having a framework having a 10-ring pore structure, wherein the 10-ring pore structure is a one-dimensional pore structure; and wherein the one-dimensional pore structure is at least one selected from *MRE-type ZSM-48, MTT-type ZSM-23 and TON-type ZSM-22; and wherein the method further comprises recycling a portion of the stream comprising C2-C3 olefins to the feed stream.

2. The method according to claim 1, wherein the catalyst comprises a binder, and the binder is at least one selected from the group consisting of alumina, aluminum phosphate, silica, silica-alumina, zirconia, titania, and clay.

3. The method according to claim 2, wherein the catalyst contains 30-90 wt% of zeolite and a binder comprising an alumina component.

4. The process according to any one of claims 1 to 3, wherein the zeolite has a silicon to aluminum ratio of up to 240.

5. The process of any one of claims 1 to 3, wherein the feed stream is combined with a diluent, the feed stream is methanol, and the feed stream is diluted to a methanol concentration in the feed of 2 to 20 vol%.

6. The process according to any one of claims 1 to 3, wherein the portion of the stream comprising C2-C3 olefins is an olefin stream comprising propylene.

7. A method according to any one of claims 1 to 3, wherein the feedstream comprising oxygenates is derived from one or more oxygenates selected from triglycerides, fatty acids, resin acids, ketones, aldehydes, alcohols or ethers, wherein the oxygenates are derived from one or more of a biological source, a gasification process, a pyrolysis process, a Fischer-Tropsch synthesis or a methanol-based synthesis.

8. The method according to any one of claims 1 to 3, wherein the oxygen-containing compound is selected from one or more of methanol (MeOH) and dimethyl ether (DME).

9. The method according to any one of claims 1 to 3, comprising: A first reactor group comprising a single reactor or a plurality of reactors arranged in parallel with one another is used for partial or complete conversion of the oxygenates.

10. The process of claim 9 further comprising using a second reactor group comprising a single reactor or a plurality of reactors to further convert the oxygenate, wherein the plurality of reactors are arranged in series with each other and the phase separation stage is located between the first reactor group and the second reactor group, thereby forming the olefin stream.

11. The method according to claim 10, comprising: - passing a feed stream comprising oxygenates through the first reactor group under conditions such that the oxygenates are partially converted, thereby forming a crude olefin stream comprising unconverted oxygenates and C2-C8 olefins; - passing the crude olefin stream through the phase separation stage to produce: a first olefin stream rich in lower olefins; a separated oxygenate stream comprising said unconverted oxygenates; a second olefin stream rich in higher olefins; - combining the first olefin stream with the separated oxygenate stream comprising unconverted oxygenates, thereby forming a combined stream comprising lower olefins and said unconverted oxygenates; - passing the resulting combined stream comprising lower olefins and the unconverted oxygenates through the second reactor group under conditions such that the unconverted oxygenates and the lower olefins are completely converted to form a third olefin stream rich in higher olefins; - combining the second olefin stream with the third olefin stream, thereby forming the olefin stream.

12. The method according to claim 1, further comprising: - Separating an isoparaffin stream from said olefin stream.

13. The method according to claim 12, further comprising: - passing at least a portion of said olefin stream through an oligomerization step over an oligomerization catalyst and subsequently to a separation step, thereby producing an oligomerization stream.

14. The process according to claim 13, wherein the entire olefin stream is fed directly to the oligomerization step.

15. The method according to claim 13, further comprising: At least a portion of the oligomerization stream is passed to a hydrogenation step over a hydrogenation catalyst and subsequently subjected to a separation step, thereby producing a hydrocarbon stream comprising hydrocarbons boiling in the jet fuel range.

16. The process according to claim 15, wherein the oligomerization step and the hydrogenation step are combined in a single hydrogenation-oligomerization step, wherein the oligomerization step is a dimerization and / or trimerization performed by the hydrogenation-oligomerization step, wherein the hydrogenation-oligomerization step is performed by reacting the olefin stream, after separation of the isoparaffin stream, over a catalyst comprising a hydrogenation metal incorporated in a zeolite, the hydrogenation metal being at least one selected from the group consisting of Pd, Rh, Ru, Pt, Ir, Re, Cu, Co, Mo, Ni and W in the presence of hydrogen and at a pressure of 15 to 60 bar and a temperature of 50 to 350°C.

17. The method according to claim 1, further comprising: - passing at least a portion of said olefin stream through an oligomerization step over an oligomerization catalyst and subsequently to a separation step, thereby producing an oligomerization stream.

18. The method according to claim 1, wherein the method is carried out at a pressure of 2-25 bar and a temperature of 300-360°C.

19. The method according to claim 2, wherein the clay is at least one selected from the group consisting of montmorillonite, kaolin, palygorskite, montmorillonite, and attapulgite.

20. The method of claim 2, wherein the adhesive is an aluminum-based adhesive.

21. The method according to claim 2, wherein the binder is at least one selected from the group consisting of alumina, aluminum phosphate, silica-alumina, and clay.

22. The process of claim 3, wherein the catalyst contains 50-80 wt% zeolite.

23. The method of claim 3, wherein the binder comprises silica-alumina.

24. The method of claim 4, wherein the zeolite is ZSM-48 having a silicon to aluminum ratio of up to 110.

25. The method of claim 4, wherein the zeolite is ZSM-48 having a silicon to aluminum ratio of up to 100.

26. The process of claim 5, wherein the feed stream is diluted to a methanol concentration of 5-10 vol% in the feed.

27. The method of claim 8, wherein the methanol is produced from synthesis gas.

28. The method of claim 27, wherein the synthesis gas is produced by using electricity from a renewable energy source.

29. The method of claim 28, wherein the renewable energy is wind energy or solar energy.

30. The process of claim 16, wherein the hydrogenation-oligomerization step is carried out at a pressure of 20-40 bar and a temperature of 100-250°C.

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