WO2024261380A1

WO2024261380A1 - A method for producing propylene

Info

Publication number: WO2024261380A1
Application number: PCT/FI2024/050197
Authority: WO
Inventors: Eveliina MÄKELÄ; Jerick IMBAO; Antti Kurkijärvi; Ferran TORRES MARTÍ; Avelino Corma Canos; Yannick Mathieu
Original assignee: Neste Oyj
Current assignee: Neste Oyj
Priority date: 2023-06-21
Filing date: 2024-04-26
Publication date: 2024-12-26
Anticipated expiration: 2025-12-21
Also published as: FI131329B1; FI20235697A1

Abstract

The present disclosure relates to a method for producing propylene C, in particular to methods comprising catalytic cracking 10 of a feed containing sustainable hydrocarbons A with zeolites having a framework type MTW.

Description

A METHOD FOR PRODUCING PROPYLENE

FIELD

The present disclosure relates to methods for producing propylene, in particular to methods comprising catalytic cracking of a feed containing sustainable hydrocarbons with zeolites having a framework type MTW.

BACKGROUND

Petrochemicals are a growing business area and propylene is the second most important starting product in the petrochemical industry after ethylene. Nearly two thirds of all demand of propylene is used for manufacturing of polypropylene.

There is an increasing demand to use sustainable polyethene and polypropene. However, there is very limited production capacity available for producing the sustainable chemicals that are used in polymerization processes. Renewable ethylene is available for example via carbohydrate fermentation to ethanol followed by catalytic dehydrogenation, but for propylene production routes are more limited. Steam cracking of naphtha cannot meet the demand alone, even if sustainable naphtha was available in large volumes, because steam cracking provides propylene only as a side-product. Fluid catalytic cracking (FCC) is expected to be one of the main sources of propylene. Propylene is also manufactured from propane via dehydrogenation.

WO22096781A1 discloses one-stage catalytic cracking of highly paraffinic feedstocks using relatively low temperature and conventional cracking catalyst such as ZSM-5 to produce propylene and C4 olefin compositions, with an optional recycle of unconverted feedstock.

WO22096782A1 discloses one-stage catalytic cracking, particularly FCC, of highly paraffinic feedstocks using higher temperature and conventional FCC cracking catalyst to produce propylene and gasoline range compositions, with an optional recycle of unconverted feedstock.

WO2021 119610A1 discloses catalytic hydrocracking of a renewable crude product obtained by hydrotreating a renewable feedstock, optionally blended with a fossil crude oil, and distilling the renewable crude or the blend to produce propane and liquid fuels. For conversion of propane to propylene, a further catalytic PDH process was required.

W020091 30392 discloses catalytic cracking of hydrogenated natural fats to C2-C8 hydrocarbons at 250 °C - 450 °C, using a catalyst based on a zeolite and a mesoporous inorganic oxide.

US2014115952 discloses catalytic cracking of partially or fully deoxygenated pyrolysis oil obtained by thermal decomposition of lignin and cellulose, and optional fossil co-feed, aiming at reduced coke formation.

EP2325281 discloses catalytic cracking of partially deoxygenated pyrolysis oil obtained by thermally decomposing lignin and cellulose, containing 10-30 wt.-% oxygen, and a fossil co-feed e.g. VGO or long residue, aiming to avoid excessive formation of coke and dry gas. A zeolite with an amorphous binder could be used as cracking catalyst. The cracking is to be conducted at moderate temperature.

SUMMARY

It is an object of the present disclosure to provide a method for producing propylene, the method comprising the following steps: a) providing a sustainable hydrocarbon feed; b) subjecting a cracking feed comprising the sustainable hydrocarbon feed to catalytic cracking reaction in the presence of a cracking catalyst comprising a zeolite having a framework type MTW and Si/AI molar ratio of 10-200, preferably 20-150, more preferably 30-100, at a temperature in a range from 500 °C to 850 °C, preferably in a range from 520 °C to 820 °C, and a residence time from 0.1 s to 150 s, preferably from 0.2 s to 100 s, more preferably from 0.2 s to 50 s to produce a catalytically cracked stream, and c) separating from the catalytically cracked stream at least a fraction rich in propylene.

The present method provides various benefits as explained hereinafter and/or demonstrated in the experiments.

A number of exemplifying and non-limiting embodiments of the invention are described in accompanied dependent claims. Various exemplifying and non-limiting embodiments of the invention together with additional objects and advantages thereof, will be best understood from the following description of specific exemplifying and non-limiting embodiments when read in connection with the accompanying figure.

The verbs “to comprise” and “to include” are used in this document as open limitations that neither exclude nor require the existence of also un-recited features. The features recited in dependent claims are mutually freely combinable unless otherwise explicitly stated. Furthermore, it is to be understood that the use of “a” or “an”, i.e., a singular form, throughout this document does not exclude a plurality.

BRIEF DESCRIPTION OF DRAWING

Figure 1 shows an exemplary non-limiting schematic overview of production of propylene according to an embodiment of the method of the present disclosure.

Figure 2 shows an exemplary non-limiting schematic overview of processing >C4 hydrocarbons and saturated C1 -C4 hydrocarbons in the method of the present disclosure.

Figure 3 illustrates influence of Si/AI molar ratio of a ZSM-5 catalyst over catalyst activity and propylene selectivity at A: 400°C, B: 500°C, and C: 600°C when cracking hexadecane at total pressure of 1 atm, feed’s partial pressure of 0.33 atm, and varying WHSV. Number in the brackets following the name of the catalyst is the Si/AI ratio of the catalyst.

Figure 4 illustrates influence of cracking temperature over catalyst activity and propylene selectivity for ZSM-5 catalysts having different Si/AI molar ratios of 31 (A and B), 125 (C and D) and 400 (E and F), when cracking hexadecane at total pressure of 1 atm, feed’s initial partial pressure of 0.33 atm, and varying WHSV.

Figure 5 illustrates influence of WHSV over cracking activity in terms of conversion and propylene yield for ZSM-5 (31 ) and ZSM-12 (40) catalysts when cracking n- paraffinic or isomerized feed at 600 °C, atmospheric pressure, and feed’s initial partial pressure of 0.33 atm. Number in the brackets following the name of the catalyst is the Si/AI ratio of the catalyst.

Figure 6 shows comparison of catalytic activity and propylene yield of submicronsized, hierarchical and two standard ZSM-5 catalysts as well as ZSM-12 catalyst when cracking isomerized feed at 600 °C, atmospheric pressure, and feed’s initial partial pressure of 0.33 atm. Number in the brackets following the name of the catalyst is the Si/AI ratio of the catalyst. Also maximum propylene yields are indicated (ZSM-12: 28 wt.-%; other tested catalysts: 21 -22 wt.-%).

Figures 7A and 7B show exemplary non-limiting generalized schematic overviews of catalytically cracking a feed R, followed by recycling a >C5 fraction separated from the catalytically cracked stream to the catalytic cracking in the same cracking reactor (Fig 7A), or followed by re-cracking the >C5 fraction in a separate catalytic cracking reactor (Fig 7B), according to embodiments of the method of the present disclosure.

DESCRIPTION

Figure 1 shows an exemplary process of the present disclosure for producing propylene. In the figure reference numbers and arrows illustrate reactions and streams, respectively.

According to the embodiment shown in the figure, the method comprises the following steps: a) providing a sustainable hydrocarbon feed A, b) subjecting 10 a cracking feed comprising the sustainable hydrocarbon feed to a catalytic cracking reaction to produce a catalytically cracked stream B, and c) separating 20 from the catalytically cracked stream at least a fraction C rich in propylene.

The method yields also one or more propylene depleted fractions D, such as a fraction comprising hydrocarbons having a carbon number of at least C4, preferably at least C5, which are optionally recycled to the cracking feed. Aromatic hydrocarbons present in stream D are preferably separated from the stream prior to the recycling.

The method shown in the figure also includes an optional step of feeding 30 a diluent gas to the cracking feed, which diluent gas may comprise e.g. at least a part of a fraction rich in saturated C1-C4 hydrocarbons optionally separated from the catalytically cracked stream. The key steps of the method according to the present disclosure are discussed below in detail.

As used herein hydrocarbons refer to compounds consisting of carbon and hydrogen, including paraffins, n-paraffins, i-paraffins, monobranched i-paraffins, multibranched i-paraffins, olefins, naphthenes, and aromatics. Oxygenated hydrocarbons refer herein to hydrocarbons comprising covalently bound oxygen.

As used herein paraffins refer to non-cyclic alkanes, i.e. non-cyclic, open chain saturated hydrocarbons that are linear (normal paraffins, n-paraffins) or branched (isoparaffins, i-paraffins). In other words, paraffins refer herein to n-paraffins and/or i-paraffins.

As used herein, cyclic hydrocarbons refer to all hydrocarbons containing cyclic structure(s), including cyclic olefins, naphthenes, and aromatics. Naphthenes refer herein to cycloalkanes i.e. saturated hydrocarbons containing at least one cyclic structure, with or without side chains. As naphthenes are saturated compounds, they are compounds without aromatic ring structure(s) present. Aromatics refer herein to hydrocarbons containing at least one aromatic ring structure, i.e. cyclic structure having delocalized, alternating TT bonds all the way around said cyclic structure.

A fraction “rich in propylene”, means in the context of the present disclosure that the wt.-% amount of the propylene in the fraction, based on the total weight of the fraction, is higher than the wt.-% amount of the propylene in the catalytically cracked stream, based on the total weight of the catalytically cracked stream. Preferably the wt.-% amount of the propylene is higher than the wt.-% amount of any other single compound present in the fraction rich in propylene. In other words, the fraction rich in propylene comprises propylene as the most abundant compound. More preferably the fraction rich in propylene comprises more than 50 wt.-% propylene, based on the total weight of the fraction rich in propylene. Meaning of “rich in” of other fractions is as disclosed above for a fraction rich in propylene mutatis mutandis.

The cracking catalyst may be in ready-to-use state as such, or it may be treated in any customary way to adjust its properties, such as selectivity and/or activity, before or during start-up by subjecting for example to deactivation protocol e.g. steaming, so as to obtain the ready-to-use fresh or ready-to-use regenerated cracking catalyst. As used herein, “cracking catalyst”, and regenerated cracking catalyst, generally refer to the cracking catalyst in its ready-to-use state.

As used herein, “residence time” refers to the time a cracking feed spends in contact with the specified catalyst at the specified temperature.

In the present disclosure, the term “catalytically cracked stream" refers to the effluent from a step of catalytic cracking reaction, more specifically of the catalytic cracking reaction of step b) but excluding the catalyst and possibly formed coke. This applies to all process configurations, whether using fixed or moving catalyst bed(s) such as fluidised catalyst bed. On the other hand, the term “effluent", as in “cracking effluent” or similar expression, is used herein as referring to all materials exiting the reaction step and may hence comprise a catalyst and any formed coke, depending on the process configuration.

Unless otherwise stated, in the context of the present disclosure, for compositions boiling at 36°C or higher at standard atmospheric pressure, contents of n-paraffins, i-paraffins, monobranched i-paraffins, multibranched isoparaffins, naphthenes, and aromatics are expressed as weight % (wt.-%) relative to the degassed weight of the composition in question, or, when so defined, as weight % (wt.-%) relative to the total weight of paraffins, or total weight of i-paraffins of the composition in question. Said contents may be determined by GCxGC-FID/GCxGC-MS method, preferably conducted as follows: GCxGC (2D GC) method was run as generally disclosed in UOP 990-2011 and by Nousiainen M. in the experimental section of his Master's Thesis Comprehensive two-dimensional gas chromatography with mass spectrometric and flame ionization detectors in petroleum chemistry, University of Helsinki, August 2017, with the following modifications. The GCxGC was run in reverse mode, using a semipolar column (Rxi17Sil) first and a non-polar column (Rxi5Sil) thereafter, followed by FID detector, using run parameters: carrier gas helium 31 .7 cm/s (column flow at 40 °C 1 .60 ml/min); split ratio 1 :350; injector 280 °C; Column T program 40 °C (0 min) - 5 °C/min - 250 °C (0 min) - 10 °C/min - 300 °C (5 min), run time 52 min; modulation period 10 s; detector 300 °C with H2 40 ml/min and air 400 ml/min; makeup flow helium 30 ml/min; sampling rate 250 Hz and injection size 0.2 microliters. Individual compounds were identified using GCxGC-MS, with MS-parameters: ion source 230 °C; interface 300 °C; scan range 25 - 500 amu; event time (sec) 0.05; scan speed 20000. Commercial tools (Shimadzu's LabSolutions, Zoex's GC Image) were used for data processing including identification of the detected compounds or hydrocarbon groups, and for determining their mass concentrations by application of response factors relative to n-heptane to the volumes of detected peaks followed by normalization to 100 wt.- %. Olefins were lumped with naphthenes and heteroatomic species with aromatics, unless separately reported. The limit of quantitation for individual compounds of this method is 0.1 wt.-%.

In the context of this disclosure, CX+ hydrocarbons, paraffins, or similar, refer to hydrocarbons, paraffins, or similar, respectively, having a carbon number of at least X, where X is any feasible integer. It is understood that every compound falling within the definition is not necessarily present.

The sustainable hydrocarbon feed

The sustainable hydrocarbon feed used in the present method has typically a T5 temperature (5 vol-% recovered, EN ISO 3405-2019) at least 180°C, preferably at least 190°C, more preferably at least 200°C, and T95 (95 vol-% recovered, EN ISO 3405-2019) at most 500°C, preferably at most 450°C, more preferably at most 430°C. Preferred hydrocarbon feeds have both T5 and T95 temperatures, or even initial and final boiling points, within 180 °C - 500°C, preferably within 190 °C - 450°C, more preferably within 200 °C - 430°C (EN ISO 3405-2019). Preferably the hydrocarbon feeds have a difference between the T95 and T5 temperatures at most 300°C, preferably at most 200°C, more preferably at most 150°C, even more preferably at most 100°C or at most 80°C. Hydrocarbon feeds having relatively uniform composition in terms of boiling points are preferred as they may allow easier optimisation of the process conditions for the majority of the feed molecules, and in that way enhance formation of desired products and suppress formation of less desired products or side products. A well-defined boiling range is advantageous also for controlled and essentially complete evaporation of the feed ensuring that the feed is in gas phase in the catalytic cracking step, providing future benefits as discussed in the following.

In embodiments, the sustainable hydrocarbon feed comprises, based on the total weight of the sustainable hydrocarbon feed, at least 50 wt.-%, preferably at least 60 wt.-%, more preferably at least 70 wt.-% paraffins, preferably having a carbon number of at least C12. In preferred embodiments, the sustainable hydrocarbon feed comprises, based on the total weight of the sustainable hydrocarbon feed, at least 40 wt.-%, preferably at least 50 wt.-%, more preferably at least 60 wt.-% paraffins having a carbon number of at least C14, and/or at least 25 wt.-%, preferably at least 35 wt.-%, more preferably at least 45 wt.-% paraffins having a carbon number of at least C16. Highly paraffinic sustainable hydrocarbon feeds are preferred as they help to control or reduce coke formation in the catalytic cracking step, and paraffins, particularly long paraffins having a carbon number of at least C12 or more, tend to crack more easily compared to e.g. cyclic hydrocarbons or to shorter paraffins, making these feeds desired for the present catalytic cracking process.

In embodiments, the sustainable hydrocarbon feed comprises, based on the total weight of the sustainable hydrocarbon feed, at least 90 wt.-%, preferably at least 95 wt.-%, more preferably at least 98 wt.-% C10-C40 hydrocarbons. In preferred embodiments, the sustainable hydrocarbon feed comprises, based on the total weight of the sustainable hydrocarbon feed, at least 80 wt.-%, preferably at least 85 wt.-%, more preferably at least 90 wt.-% C12-C30 hydrocarbons, and/or at least 75 wt.-%, preferably at least 80 wt.-%, more preferably at least 85 wt.-% C12-C25 hydrocarbons. Sustainable hydrocarbon feeds having relatively uniform composition are preferred as they allow optimising the process conditions for majority of the feed molecules, and in that way may enhance formation of desired products and suppress formation of undesired products.

E.g. in WO2022096781 , where highly paraffinic feeds were cracked with ZSM-5 catalyst at relatively low temperature, at 400 °C, it was observed that the higher the isoparaffin content in the feed, the higher the propylene and C4-olefin yields, and the lower the C5-C9 hydrocarbon formation. It was also observed that the higher the multibranched isoparaffin content in the feed, the higher the methane and hydrogen formation. For all the tested feeds the conversion was modest, around 30%, and the ratio of produced propylene to ethylene was high, over 15. In WO2022096782, where highly paraffinic feeds were cracked with a conventional FCC catalyst at relatively high temperature, at 650 °C, it was observed that the higher the isoparaffin content in the feed, the lower the total C4 formation and the higher the yield of standard gasoline range hydrocarbons. It was also observed that the higher the multibranched isoparaffin content in the feed, the higher the methane formation, while hydrogen formation remained about the same. For all the tested feeds the ratio of produced propylene to ethylene remained low, around 3. The conversion of the tested feeds varied significantly, being about 40% for the feed having highest degree of isomerisation, and about 80% for the highly n-paraffinic feed.

Contrary to the previous findings, the present inventors found that in the method of the present disclosure both a mainly n-paraffinic hydrocarbon feed and a highly isoparaffinic hydrocarbon feed provide very similar product slates, high propylene and C4 mono-olefin yields, and low formation of methane, hydrogen, aromatics, and coke, while the conversion levels are high, at least about 80%. It is highly beneficial, and not common, that similar product slate is obtainable with different feeds, providing flexibility to the process e.g. depending on feedstock availability and/or catalyst life cycle, with minimal or no changes required to the processing of the catalytically cracked stream.

Hence, in certain preferred embodiments, the sustainable hydrocarbon feed comprises, based on the total weight of the sustainable hydrocarbon feed, at least 50 wt.-%, preferably at least 60 wt.-%, more preferably at least 70 wt.-% paraffins, preferably having a carbon number of at least C12, of which paraffins 1 - 99 wt.-%, preferably 3 - 98 wt.-%, more preferably 5 - 95 wt.-% are isoparaffins.

The benefits of the present process are obtainable even with feeds having significant content of bulky molecules, such as isoparaffins, or even multibranched isoparaffins. Hence, in embodiments, the sustainable hydrocarbon feed comprises, based on the total weight of the sustainable hydrocarbon feed, at least 20 wt.-%, preferably at least 40 wt.-%, more preferably at least 60 wt.-% or even at least 80 wt.-% isoparaffins; and/or at least 10 wt.-%, preferably at least 20 wt.-%, more preferably at least 30 wt.-% or even at least 40 wt.-% multibranched isoparaffins.

However, highly isoparaffinic hydrocarbon feeds may provide slightly higher aromatics and/or coke formation compared to mainly n-paraffinic feeds. Hence, in certain preferred embodiments, the sustainable hydrocarbon feed comprises, based on the total weight of the sustainable hydrocarbon feed, at most 20 wt.-%, preferably at most 15 wt.-%, more preferably at most 10 wt.-% isoparaffins. Aromatics do not form or form only very little light olefins. Accordingly, they may accumulate in an optional recycle loop, unless specifically removed from the recycled stream, and tend to increase coke formation. Naphthenes, on the other hand, tend to decrease formation of light olefins, compared to paraffins, and naphthenes comprising a C6 ring are precursors for aromatics. Hence, in certain preferred embodiments, the sustainable hydrocarbon feed comprises, based on the total weight of the sustainable hydrocarbon feed, at most 15 wt.-%, preferably at most 10 wt.-%, more preferably at most 5 wt.-%, even more preferably at most 1 wt.-% aromatics, and/or at most 30 wt.-%, preferably at most 25 wt.-%, more preferably at most 10 wt.-%, even more preferably at most 5 wt.-%, further more preferably at most 3 wt.-% naphthenes.

Low oxygen content is desired for the hydrocarbon feed so as to avoid presence of oxygen-containing hydrocarbon species in the catalytically cracked stream, as oxygen-containing cracking products may be difficult to remove from the product fraction(s) containing the desired cracking product(s). Hence, in certain preferred embodiments, the sustainable hydrocarbon feed comprises, based on the total weight of the sustainable hydrocarbon feed, at most 3 wt.-%, preferably at most 2 wt.-%, more preferably at most 1 wt.-%, even more preferably at most 0.5 wt.-% oxygenated hydrocarbons, expressed as elemental oxygen.

Sustainable hydrocarbon feeds suitable for the present method may be prepared from renewable and/or circular feedstocks or hydrotreatment feeds.

Chemically, the renewable or fossil origin of any organic compound, including hydrocarbons, can be determined by suitable method for analysing the content of carbon from renewable sources e.g. DIN 51637 (2014), ASTM D6866 (2020), or EN 16640 (2017). Said methods are based on the fact, that carbon atoms of renewable or biological origin comprise a higher number of unstable radiocarbon (¹⁴C) atoms compared to carbon atoms of fossil origin. Therefore, it is possible to distinguish between carbon compounds derived from renewable or biological sources and carbon compounds derived from non-renewable or fossil sources by analysing the ratio of ¹²C and ¹⁴C isotopes. Thus, a particular ratio of said isotopes can be used as a “tag” to identify a renewable carbon compound and differentiate it from non- renewable carbon compounds. The isotope ratio does not change in the course of chemical reactions. Therefore, the isotope ratio can be used for identifying renewable carbon compounds and distinguishing them from non-renewable, fossil carbon compounds in feedstocks, hydrotreatment feeds, co-feeds, products or compositions, or various blends thereof. Numerically, the biogenic carbon content can be expressed as the amount of biogenic carbon in the material as a weight percent of the total carbon (TC) in the material (in accordance with ASTM D6866 (2020) or EN 16640 (2017)). In the present context, the term renewable preferably refers to a material having a biogenic carbon content of more than 50 wt.-%, especially more than 60 wt.-% or more than 70 wt.-%, preferably more than 80 wt.- %, more preferably more than 90 wt.-% or more than 95 wt.-%, even more preferably about 100 wt.-%, based on the total weight of carbon in the material (EN 16640 (2017)).

As used herein, the term circular in connection with materials such as feedstocks, hydrotreatment feeds, products or compositions refers to materials based on reused and/or recycled carbon from any source available.

As used herein, the term fossil refers to materials such as co-feeds, products or compositions that are obtainable, derivable, or originating from naturally occurring non-renewable compositions, such as crude oil, petroleum oil/gas, shale oil/gas, natural gas, or coal deposits, and the like, and combinations thereof, including any hydrocarbon-rich deposits that can be utilized from ground/underground sources.

Renewable, circular, and fossil materials, feedstocks, hydrotreatment feeds, cofeeds, products, or compositions are considered differing from one another based on their origin and impact on environmental issues. Therefore, they may be treated differently under legislation and regulatory framework. Typically, renewable, circular, and fossil materials etc. are differentiated based on their origin and information thereof provided by the producer.

The renewable and/or circular feedstock may be converted to the sustainable hydrocarbon feed using methods known in the art. The suitable methods are dependent on the nature of the feedstock.

In preferred embodiments, the sustainable hydrocarbon feed is provided by subjecting a renewable and/or circular feedstock, in the context of hydrotreatment also referred to as a renewable and/or circular hydrotreatment feed, to at least one or more hydrotreatment(s). By hydrotreatment, sometimes also referred to as hydroprocessing, is meant herein a catalytic process of treating organic material by means of molecular hydrogen. The hydrotreatment reactions may include removal of oxygen from organic oxygenates as water i.e. hydrodeoxygenation (HDO), sulphur from organic sulphur compounds as dihydrogen sulphide (H2S), i.e. hydrodesulphurisation, (HDS), nitrogen from organic nitrogen compounds as ammonia (NH3), i.e. hydrodenitrogenation (HDN), halogens, for example chlorine from organic chloride compounds as hydrochloric acid (HCI), i.e. hydrodechlorination (HDCI), and/or metals by hydrodemetallization; hydrogenation of olefinic bonds to saturated bonds and/or of aromatics to naphthenes; and/or hydrocracking, including ring-opening. Depending on the composition of the renewable and/or circular hydrotreatment feed, different reactions may occur and/or prevail in the hydrotreatment. Hydrotreatment is a preferred conversion method for providing the sustainable hydrocarbon feed because it is capable of converting hydrotreatment feeds of varying compositions to highly paraffinic materials, and at the same time reduce content of heteroatoms, metals, olefins, aromatics and/or other impurities in the hydrotreatment feed.

Hence, in preferred embodiments, providing the sustainable hydrocarbon feed of step a) of the present method comprises a1 ) providing a renewable and/or circular hydrotreatment feed, a2) subjecting the hydrotreatment feed, optionally mixed with a hydrocarbon liquid diluent, to hydrotreatment reaction(s) to produce a hydrotreated stream, a3) subjecting the hydrotreated stream to a gas-liquid separation to produce at least a gaseous stream and a hydrotreated liquid stream, a4) optionally subjecting at least a part of the hydrotreated liquid stream to a hydroisomerization step to produce a hydroisomerized stream; a5) separating the sustainable hydrocarbon feed from the hydrotreated liquid stream, from the hydroisomerized stream and/or from a combination thereof, by dividing, stabilizing and/or fractionating.

Exemplary renewable feedstocks or hydrotreatment feeds include vegetable oils, animal fats, microbial oils, and other fatty materials comprising fatty acids, mono-, di- or triglycerides, resin acids, and/or alkyl esters of fatty acids and/or resin acids; and pyrolysis oils from lignocellulosics, lignin or similar. Exemplary circular feedstocks or hydrotreatment feeds include non-catalytically such as thermally (including hydrothermally and by pyrolysis) liquefied organic waste and residues and catalytically (including thermo-catalytically) liquefied organic waste and residues, wherein the waste and residues may comprise waste plastics, end of life tires, municipal solid waste, and other similar materials; and gas-to-liquid (GTL) hydrocarbons obtained e.g. by Fischer-Tropsch conversion of syngas obtained by gasification of waste and residues or by capturing carbon dioxide e.g. from emissions from hydrocarbon production, steel production, coal or natural gas power plants and generating H2 e.g. electrolytically from water.

Preferably the renewable and/or circular feedstock or hydrotreatment feed comprises at least one or more of vegetable oils, animal fats, microbial oils, and/or thermally or catalytically liquefied organic waste and residues. These materials are readily available in quantities and qualities desired for hydrotreatment, various established pre-treatment techniques exist for purifying these materials, and by hydrotreatment the optionally purified materials may be converted to sustainable hydrocarbon feeds having at least one or more of the desired characteristics specified in the present disclosure.

If the renewable and/or circular hydrotreatment feed includes amounts or species of impurities that are not tolerated or preferred in the hydrotreatment, their content in the hydrotreatment feed may be reduced to acceptable limits using methods known in the art. Exemplary pre-treatment methods suitable for the present disclosure comprise treating with mineral acids, degumming, treating with hydrogen, heat treating, deodorizing, washing with water, treating with base, demetallation, distillation, removal of solids, bleaching, and any combinations thereof.

The renewable and/or circular hydrotreatment feed subjected to hydrotreatment reaction(s) may further comprise a hydrocarbon liquid diluent. This may be attained by mixing the renewable and/or circular hydrotreatment feed with a hydrocarbon liquid diluent, e.g. separately or by co-feeding the hydrocarbon liquid diluent to the hydrotreatment. This may be beneficial for example for controlling exotherm, i.e. heat released by hydrotreatment reaction(s) and/or for increasing solubility of hydrogen in the hydrotreatment reaction mixture. By suitably selecting the feed to diluent ratio, hydrotreatment reactions may be controlled more efficiently, pressures may be lowered still without reducing the amount of hydrogen in solution, and/or hydrotreatment catalyst deactivation may be reduced. Exemplary hydrocarbon liquid diluents include product recycle such as a portion of the hydrotreated liquid stream, the hydroisomerized stream or a combination thereof, and/or various fossil crude oil streams such as fossil crude oil distillate(s) and/or fossil crude oil (hydro)crackate(s).

Typically, hydrotreated streams obtained by subjecting the renewable and/or circular hydrotreatment feed to hydrotreatment reaction(s) comprise mainly hydrocarbons that are non-gaseous at NTP, providing upon gas-liquid separation high yields of hydrotreated liquid stream, wherefrom the sustainable hydrocarbon feed may be separated. While reference is made to a hydrotreated liquid stream (or hydrocarbon liquid diluent), it is to be understood that some or all of the molecules in the hydrotreated liquid stream (or product recycle) may actually be solid at NTP.

The hydrotreatment reaction(s) are typically carried out at conditions comprising at least one or more of a temperature in the range from 120 °C to 500 °C, a pressure in the range from 10 bar to 200 bar, a WHSV in the range from 0.1 h’¹ to 10 h’¹, a H2 flow of from 50 to 2000 N-L H2/L feed, and/or a hydrotreatment catalyst, preferably a sulphided hydrotreatment catalyst, comprising at least one or more metals from Group VIII of the Periodic Table and/or from Group VIB of the Periodic Table, preferably at least one or more of Ni, Mo, W, and/or Co, even more preferably at least one or more of Ni and/or Co and Mo and/or W, such as NiMo, C0M0, NiCoMo, NiW, and/or NiMoW, preferably on a support.

The above mentioned hydrotreatment catalysts are efficient, readily available, and tolerate typical impurities of fatty feedstocks well. If using a catalyst having hydrodewaxing properties, such as a catalyst containing NiW, as the hydrotreatment catalyst or as a co-catalyst, sustainable hydrocarbon feeds with somewhat elevated isoparaffins content may be attained.

According to a particular embodiment, the hydrotreatment reaction(s), especially involving effective HDO of a renewable hydrocarbon feed comprising fatty materials, are carried out at conditions comprising temperature in the range from 200 °C to 500 °C, pressure in the range from 10 bar to 200 bar, a WHSV in the range from 0.1 h^-1 to 10 h^-1, H2 flow of from 50 to 2000 N-L H2/L feed, and a sulphided hydrotreatment catalyst. The sulphided state of the sulphided hydrotreatment catalyst may be maintained during the hydrotreatment step for example by adding a sulphur compound to the hydrotreatment feed and/or to the hydrogen stream or by using a hydrotreatment feed or a co-feed comprising sulphur compound(s). Sulphur may be deliberately added within a range from 50 w-ppm (ppm by weight) to 20 000 w-ppm, preferably within a range from 100 w-ppm to 1000 w-ppm, based on the weight of the hydrotreatment feed, when using hydrotreatment catalysts requiring a sulphided form for operation.

Hydrotreatment conditions effective for HDO may reduce the oxygen content to less than 1 wt.-%, such as to less than 0.5 wt.-% or to less than 0.2 wt.-%, based on the total weight of the hydrotreated liquid stream or the sustainable hydrocarbon feed separated therefrom.

In certain other embodiments, the hydrotreatment reaction(s), especially involving hydrotreatment of a circular hydrocarbon feed comprising thermally liquefied waste plastics and/or end of life tires, are carried out at conditions specified in Fl 130219B.

In the gas-liquid separation the hydrotreated stream is separated to produce at least a gaseous stream and a hydrotreated liquid stream. The gas-liquid separation may be conducted in a conventional manner, for example as an integral step within the respective hydrotreatment reactor, separately or as part of a fractionation system. Typically, the gas-liquid separation is conducted at a temperature within a range from 0 °C to 500 °C, such as from 15°C to 300°C, or from 15 °C to 150 °C, preferably from 15 °C to 65 °C, such as from 20 °C to 60 °C, and preferably at the same pressure as that of the hydrotreatment reactor. Typically, the pressure during the gas-liquid separation(s) may be within a range from 0.1 MPa to 20 MPa, preferably from 1 MPa to 10 MPa, or from 3 MPa to 7 MPa.

Exemplary compounds retained in the gaseous stream in the gas-liquid separation may include at least one or more of residual hydrogen, carbon monoxide, carbon dioxide, water, hydrogen disulphide, ammonia, and/or light hydrocarbons. The gaseous stream from the gas-liquid separation may be subjected to conventional treatments, depending on the composition of the gaseous stream, such as to sweetening, recovery of hydrogen, and/or recovery of light hydrocarbons such as C1 -C3 hydrocarbons. At least part of the optionally recovered hydrogen may be recycled to the hydrotreatment step, and at least part of the optionally recovered light hydrocarbons may be recycled to the cracking feed, serving as a diluent gas. At least part of the optionally recovered light hydrocarbons, particularly ethane and/or propane, may be subjected to catalytic dehydrogenation to produce further ethylene and/or propylene.

Optional hydroisomerisation step

In certain embodiments, at least part of the hydrotreated liquid stream and/or a (nonhydrotreated) gas to liquid (GTL) hydrocarbon stream is subjected to a hydroisomerization step to produce a hydroisomerized stream. The optional hydroisomerisation step can be conducted in a conventional hydroisomerisation unit, such as those depicted in W02007068795A1 , WO2016062868A1 or EP2155838B1. The hydroisomerisation is conducted in the presence of added hydrogen.

In embodiments, the optional hydroisomerization step is conducted at a temperature in the range from 200 °C to 500 °C, preferably from 250 °C to 450 °C; a pressure in the range from 1 to 10 MPa, preferably from 2 to 8 MPa; a WHSV in the range from 0.1 h^-1 to 10 h^-1, preferably 0.2 h^-1 to 8 h’¹, and a H2 flow of from 10 to 2000 N-L H2/L feed, preferably from 50 to 1000 N-L H2/L feed, in the presence of an hydroisomerisation catalyst comprising at least one or more Group VIII metal, preferably Pd, Pt and/or Ni, and at least one or more acidic porous material selected from zeolites and/or zeolite-type materials, and optionally at least one or more of alumina, silica, amorphous silica-alumina, titanium, alumina, titania, and/or zirconia.

The hydroisomerisation step converts at least a certain amount of n-paraffins in the hydrotreated liquid stream and/or in the (non-hydrotreated) GTL hydrocarbon stream to i-paraffins. Depending on the isomerization degree, that may be controlled by adjusting severity of the hydroisomerization, more of the n-paraffins can be converted to i-paraffins, and mono-branched i-paraffins to multibranched i-paraffins, such as di-branched and/or tri-branched i-paraffins, even i-paraffins comprising more than three branches. Also some cracking reactions may occur during the hydroisomerisation. The severity of the hydroisomerisation may be increased e.g. by at least one or more of: decreasing WHSV, increasing temperature, and/or increasing pressure. When using fresh hydroisomerisation catalyst, high severity hydroisomerisation conditions may be reached at lower temperature and/or pressure, and/or using higher WHSV, than towards the end of the hydroisomerisation catalyst lifetime.

There may be further steps included either combined with the hydroisomerization step, or thereafter. These may comprise e.g. gas-liquid separation, hydropolishing, dearomatizing, just to name a few. Typically, such additional process steps aim at better control of desired properties of the hydroisomerized stream.

In an embodiment including a hydroisomerization step, the sustainable hydrocarbon feed may be separated from the hydroisomerized stream by dividing, stabilizing and/or fractionating.

Catalytic cracking

Catalysts with higher hydrogen transfer index (HTI, expressed in the present disclosure as a weight ratio of unsaturated C4: s to saturated C4: s) are thought to generate fewer secondary reactions, hence preserving a greater quantity of C5+ olefins, and controlling formation of aromatics. Generally it has been thought that zeolites having smaller pore size provide steric constraints, that are beneficial for suppressing formation of bulky bimolecular reaction intermediates involved in hydrogen transfer reactions. E.g. ZSM-5, which is one of the most widely used zeolites in cracking catalysts, has three-dimensional channel system of 10-ring channels.

Surprisingly, the present inventors found that by utilising a cracking catalyst comprising a zeolite of a low to medium Si/AI molar ratio and a framework type MTW, it is possible to convert feeds rich in n-paraffins or even bulky iso-paraffins to propylene with excellent conversion levels, and with less secondary reactions occurring compared to zeolites conventionally used for catalytic cracking. This is unexpected as zeolites having a framework type MTW have a one-dimensional channel system, and larger channels e.g. compared to ZSM-5.

Hence, according to the present method, the catalytic cracking reaction is performed in the presence of a cracking catalyst comprising a zeolite having a framework type MTW and Si/AI molar ratio of 10-200, preferably 20-150, more preferably 30-100. The Si/AI molar ratio of the zeolite having a framework type MTW may be selected or adjusted within the specified ranges to ensure sufficient resistance to deactivation during the cracking and/or good cracking activity. Adjusting the Si/AI molar ratio may be conducted with any known technique, including pre-modification and/or postmodification techniques. For example, the zeolite having a framework type MTW may be subjected to dealumination and/or desilication treatment to further adjust the properties of the zeolite.

Typically zeolites having a framework type MTW are characterised by a onedimensional system of channels parallel to [010] direction with somewhat elliptical 12-ring apertures. The crystallographic free diameters of the channels may be e.g. 5.6 x 6.0 A. Preferably the zeolite having a framework type MTW, and the specified Si/AI molar ratio is selected from a group consisting of ZSM-12, Nll-13, TPZ-12, Theta-3, and combinations thereof. Particularly advantageously the zeolite is ZSM- 12.

In embodiments, the zeolite having a framework type MTW may comprise crystals having a submicron crystal size. Such catalysts may be beneficial as typically they have higher specific surface area (SSA).

In preferred embodiments, the zeolite having a framework type MTW, and a Si/AI molar ratio as specified above has one or more of the following properties: an aluminium content from 0.1 wt.-% to 1.5 wt.-%, preferably from 0.6 wt.-% to 1.3 wt.-%, determined by inductively coupled plasma - optical emission spectroscopy (ICP-OES); a BET surface area from 200 m²/g to 400 m²/g, preferably from 250 m²/g to 380 m²/g, determined by nitrogen physisorption.

Typically, the cracking catalyst further comprises at least one or more supports. In this context the expression support comprises any binders and/or fillers. The support may comprise amorphous silica, alumina, amorphous silica-alumina, zirconia, a clay material, such as kaolin or bentonite, that may be further thermally or chemically treated. Depending on the type of reactor used in the catalytic cracking, the support needs to meet different specifications related to mechanical strength and formability. For example, a catalyst with a silica support deactivates slowly which has an advantage when the cracking reaction is performed in the fixed or moving bed reactor. Dopants such as zinc or gallium may also be included, but preferably the cracking catalyst does not contain dopants. In some embodiments the cracking catalyst may also comprise minor amounts of other components, such as conventional catalyst additives, and/or other zeolites than those having a framework type MTW.

The catalytic cracking step of the present disclosure is performed at a temperature in a range from 500 °C to 850 °C, preferably in a range from 520 °C to 820 °C, and a residence time typically from 0.1 s to 150 s, preferably from 0.2 s to 100 s such as from 0.2 s to 75 s, more preferably from 0.2 s to 50 s or even from 0.2 s to 35 s, to produce a catalytically cracked stream. Generally, the shorter the residence time the less there is time for the desired cracking products to react further. The catalytic cracking process conditions comprise typically a pressure from range from 50 kPa to 150 kPa (absolute), preferably form 80 kPa to 120 kPa (absolute), more preferably at atmospheric pressure. The present process is preferably conducted at a pressure not significantly higher than ambient pressure so as to decrease risk of the formed light olefins, such as the desired propylene, to react further. Furthermore, typically when the reaction is performed above atmospheric pressure the yields of the light olefins are decreased compared to when the reaction is performed at atmospheric pressure.

Since the present process is for producing light olefins, particularly propylene, the catalytic cracking reactions are performed without added hydrogen. However, a minor amount of hydrogen may result from the reactions taking place during cracking. Additionally, minor amount of hydrogen may be carried over as part of the sustainable hydrocarbon feed and/or the catalytically cracked stream if recycling a part thereof to the cracking feed.

In preferred embodiments, the cracking feed comprises, based on the total weight of the cracking feed, at least 40 wt.-%, preferably at least 50 wt.-%, more preferably at least 60 wt.-%, even more preferably at least 90 wt.-% of the sustainable hydrocarbon feed, or may even consist essentially thereof.

In preferred embodiments, the cracking feed comprises, based on the total weight of the cracking feed, from 0.5 wt.-% to 50 wt.-%, preferably from 0.5 wt.-% to 30 wt.- %, more preferably from 1 wt.-% to 20 wt.-% a diluent gas. In this context by the diluent gas it is meant steam or compounds that are gaseous at NTP. Typically the diluent gas is non-reactive or essentially non-reactive under the conditions that prevail in the catalytic cracking step. Preferably the diluent gas comprises at least one or more of steam, methane, ethane, propane, nitrogen, carbon monoxide, and/or carbon dioxide. By incorporating a diluent gas to the cracking feed, the partial pressure of the hydrocarbons in the catalytic cracking feed may be decreased, thus favouring formation of light olefins.

In embodiments, the cracking feed may additionally comprise at least one or more recycled stream, preferably comprising hydrocarbons having a carbon number of at least C4 or at least C5, separated from the catalytically cracked stream, optionally followed by subjecting to a selective hydrogenation of diolefins to mono-olefins and/or to aromatics removal by extractive distillation or solvent extraction. Selective hydrogenation of diolefins is disclosed e.g. in WO 2022/144490.

In certain preferred embodiments, the sum amount of the sustainable hydrocarbon feed, and a diluent gas and/or a recycled stream in the cracking feed is at least 60 wt.-%, preferably at least 70 wt.-%, more preferably at least 80 wt.-%, even more preferably at least 90 wt.-% or 100 wt-%, based on the total weight of the cracking feed.

In embodiments, the cracking feed may also comprise at least one or more fossil crude oil distillate and/or fossil crude oil (hydro)crackate as a co-feed, preferably of similar boiling range as the sustainable hydrocarbon feed.

In preferred embodiments, the present method further comprises vaporizing the cracking feed before subjecting to the catalytic cracking reaction. In this way it is possible to ensure that the cracking feed is in gas phase when subjected to the catalytic cracking reaction, allowing higher feed rates with less pressure drop, less limitations to the mass transfer and better mixing characteristics in general, compared to processes where the feed is not fully or at all in gas phase.

In embodiments, in step b) of the present method the cracking catalyst is arranged at least in one or more catalyst bed(s), in one or more reactor(s), preferably at least in one or more fluidised catalyst bed reactor(s) and/or fixed catalyst bed reactor(s). Process configurations using fixed catalyst bed(s) are more simple and less expensive, while process configurations using moving catalyst, such as fluidised, ebullated, or slurry catalysts, are more flexible regarding coke-formation and catalyst regeneration.

The catalyst regeneration may be conducted in any conventional manner, e.g. continuously or batch-wise. Typically, in a fixed bed reactor, the cracking catalyst is not regenerated during the catalytic cracking, but in separate regeneration cycles during which catalytic cracking cannot be performed in the fixed bed reactor being regenerated. When using fixed catalyst beds, in order to reduce or avoid down-time, a swing reactor concept may be utilised wherein one reactor is removed from service for regeneration and a freshly regenerated reactor is simultaneously returned to service. When using moving solid catalyst reactors, it is possible to regenerate the cracking catalyst during the catalytic cracking, e.g. in a separate coke burning reactor. Since the operating temperature of the present process is close to typical catalyst regeneration temperatures, the regenerated catalyst may not require significant cooling after the regeneration, before re-introducing to the catalytic cracking reactor. This may provide shorter catalyst regeneration cycle time.

Separation

The catalytically cracked stream may be separated into different fractions using known separation methods. For example, the catalytically cracked stream may be removed from the used reactor via an overhead line, cooled and sent to fractionation, such as a fractionator tower, for recovering of the various cracking products. The separation of the fractions and recovery of the various cracking products may be conducted in several steps.

The separation and/or recovery may comprise for example at least one or more of dividing, stabilising, fractionating, distilling, evaporating, flash-separating, membrane separating, extracting, using extractive-distillation, using chromatography, using molecular sieve adsorbents, using thermal diffusion, complex forming, crystallising, preferably at least dividing, stabilizing and/or fractionating. The separation and/or recovery may comprise multiple unit operations in parallel and/or succession. Usable processing units for separation of the fractions and/or recovery of the cracking products may include de-methanizers, deethanizers, de-propanizers, de-butanizers, propane-propylene splitters, and/or ethane-ethylene splitters, just to name a few. According to the method at least a fraction rich in propylene is separated from the catalytically cracked stream. One of the benefits of the present method is a very high propylene content in the C3 fraction of the catalytically cracked stream, herein also referred to as C3 olefin icity. Hence, a separated C3 fraction may have a propylene content of a chemical grade propylene already as such, and it requires less efforts, less expensive equipment, and less energy to be refined to polymer grade purity compared to fractions having lower propylene content in the C3 fraction.

According to an exemplary embodiment at least one or more of the following fractions are also separated: a fraction rich in ethylene, a fraction rich in one or more of C4 mono olefins, a fraction of hydrocarbons having a carbon number of at least C4, preferably at least C5, a fraction rich in saturated C1 -C4 hydrocarbons, and/or a fraction rich in aromatics.

The present method provides a high ethylene content in the C2 fraction of the catalytically cracked stream, herein also referred to as C2 olefinicity. Hence, a separated C2 fraction may be usable even as such, and it requires less efforts, less expensive equipment, and less energy to be refined to chemical grade or even polymer grade purity compared to fractions having lower ethylene content in the C2 fraction.

The present method may further comprise separating from the catalytically cracked stream a fraction rich in one or more of C4 mono olefin(s). The present method provides relatively high yields of C4 mono olefins, that are useful e.g. as (co)monomers as such or as further derivatised, in the production of butyl rubber by polymerisation, iso-octane by alkylation, diisobutylene by dimerization, and methyl and ethyl tert-butyl ethers by reacting with methanol or ethanol, respectively.

Overall yield of propylene can be increased by cracking at least part the fraction of hydrocarbons having a carbon number of at least C4, preferably at least C5 and/or at least part of the fraction rich in saturated C1 -C4, preferably saturated C2-C4 hydrocarbons. In certain preferred embodiments, shown in figure 2, the method includes recycling at least a part of the fraction of hydrocarbons having a carbon number of at least C4, preferably at least C5, marked as > C4 in the figure, to the cracking feed for recracking 10, optionally after subjecting to a selective hydrogenation of diolefins to mono olefins and/or to aromatics removal by extractive distillation or solvent extraction, and separating 20 propylene from the cracked stream.

Recycling at least a part of the fraction of hydrocarbons having a carbon number of at least C4 to the same reactor may lead to accumulation of C5+ isoparaffins, that are desired for fuels and solvents. Typically the fraction of hydrocarbons having a carbon number of at least C4, preferably at least C5, contains unconverted or uncracked components of the cracking feed, but also C4+ olefins or C5+ olefins, which is particularly beneficial when recycling: one C4+ or C5+ olefin molecule forms 2 lighter olefins when cracked, while one C4+ or C5+ paraffin molecule forms 1 lighter olefin and 1 lighter saturated hydrocarbon. C4+ or C5+ olefins are also easier to crack than the paraffins having the same carbon number. Value of this fraction may be enhanced by subjecting to a selective hydrotreatment, such as selective hydrogenation of diolefins to mono-olefins and/or to aromatics removal, for example by extractive distillation or solvent extraction, so as to reduce content of components increasing risk of coking and/or side product formation when recycled, and to avoid accumulation of such components in a recycle loop.

In certain further preferred embodiments also shown in figure 2 the method includes subjecting 40 at least part of the fraction of hydrocarbons having a carbon number of at least C4, preferably at least C5, optionally subjected to a selective hydrogenation of diolefins to mono olefins and/or to aromatics removal by extractive distillation or solvent extraction, and the further cracking reaction is a catalytic cracking reaction in the presence of a cracking catalyst of the method of the present invention, to produce a further cracked stream, and separating 50 from the further cracked stream at least a fraction rich in propylene.

In certain further preferred embodiments the fraction rich in saturated C1 -C4, preferably saturated C2-C4 hydrocarbons is subjected to a further cracking reaction 60 to produce a further cracked stream and separating 70 from the further cracked stream at least a fraction rich in propylene. For cracking a fraction rich in saturated C1 -C4, preferably saturated C2-C4 hydrocarbons the cracking is thermal cracking, preferably steam cracking.

In certain embodiments, the method comprises separating a fraction rich in saturated C1 -C4 hydrocarbons, and recycling at least a part thereof to the cracking feed. The fraction rich in saturated C1 -C4 hydrocarbons may act as a diluent gas, decreasing partial pressure of the hydrocarbons in the catalytic cracking feed, thus promoting formation of light olefins, including propylene. This embodiment is shown in figure 2 as dotted lines.

Regarding the cracking products typically present in the catalytically cracked stream in varying amounts, the C2-C4 paraffins, may also be used e.g. as a component in a catalytic dehydrogenation feed, to further produce light olefins. Methane, ethane, propane and/or C4-paraffins may also be used as a component in a steamreforming feed to produce H2, and/or in a fuel gas composition.

EXPERIMENTAL

Sustainable hydrocarbon feeds

Three different sustainable hydrocarbon feeds were prepared. R1 and R2 are renewable hydrocarbon feeds prepared by catalytic hydrodeoxygenation of conventionally purified glyceridic feed of animal fat/vegetable oil origin mixed with a hydrotreated liquid stream (product recycle) as diluent, followed by gas-liquid separation. Hydrocarbon feed R2 was additionally hydroisomerized followed by gasliquid separation and stabilisation. C1 was a circular hydrocarbon feed prepared by catalytic hydrotreatment of conventionally purified liquefied waste plastic (obtained by thermal degradation/pyrolysis), followed by gas-liquid separation and fractionation. The characteristics of the thus obtained sustainable hydrocarbon feeds R1 , R2 and C1 were then analysed and are reported in Table 1 .

Table 1. Characteristics of the hydrotreated sustainable hydrocarbon feeds. Chemical composition was determined by GCxGC-FID/GCxGC-MS.

As can be seen from Table 1 , all the hydrocarbon feeds R1 , R2 and C1 were highly paraffinic, had >C16 paraffins as the major constituent and very low aromatics content, and were essentially free from oxygen and olefins. All three feeds had desired characteristics for use in the present method. The renewable hydrocarbon feeds R1 and R2 even had very low total content of cyclic hydrocarbons (summed amount of aromatics and naphthenes). Hydrocarbon feeds R1 and R2 were used in the catalytic cracking experiments as described in the following. Catalysts

MFI type zeolite, such as ZSM-5, a highly siliceous aluminosilicate microporous zeolite with an intersecting and three-dimensional channel system is a shape selective catalyst characterized by a 10-membered ring (10-MR) pore structure having straight channels with pore openings of 5.2 x 5.7 A connected by sinusoidal channels of 5.3 x 5.6 A and intersection cavities of around 9 A. Three different ZSM- 5 catalysts having the following textural properties were used in the experiments:

ZSM-5 (31 ) having Si/AI molar ratio 31 measured through ICP-OES analysis; BET area 389 m²/g; micropore area 343 m²/g; external area 46 m²/g; micropore volume 0.17 cm³/g; mesopore volume 0.07 cm³/g determined through nitrogen physisorption using BET and t-plot method; crystal size 0.3-0.7 microns estimated through SEM analysis.

ZSM-5 (125) having Si/AI molar ratio 125; BET area 368 m²/g; micropore area 325 m²/g; external area 42 m²/g; micropore volume 0.16 cm³/g; mesopore volume 0.11 cm³/g; crystal size 0.5-1 .0 microns.

ZSM-5 (400) having Si/AI molar ratio 400; BET area 417 m²/g; micropore area 378 m²/g; external area 39 m²/g; micropore volume 0.18 cm³/g; mesopore volume 0.05 cm³/g; crystal size 0.2-0.3 microns.

MTW type zeolite such as ZSM-12, a silica rich microporous zeolite with unidimensional 12-membered ring channel system and pore openings of 5.7 x 6.1 A, has slightly larger pore diameter than a usual cracking catalyst ZSM-5. The inhouse synthetized ZSM-12 used in the experiments had the following textural properties: Si/AI molar ratio 40; BET area 313 m²/g; micropore area 241 m²/g; external area 71 m²/g; micropore volume 0.12 cm³/g; mesopore volume 0.14 cm³/g; crystal size 0.1 microns.

Catalytic cracking

Experiments were carried out using a continuous plug flow reactor system including a 19 mm diameter quartz tube where one or various catalytic beds can be stacked and heated by a 3 independent heating zones furnace controlled by internal thermocouples. In a typical experiment, previously calcined commercial or in-house synthetized catalysts were pelletized to obtain a desired particle size (i.e. 50-100 pm) to avoid any diffusion limitations and were diluted with silicon carbide to reach a constant volume of 5 cm³ for each one of the tested catalytic beds. At the top of the catalytic bed, an additional volume of 2 cm³ of pure silicon carbide was added, helping for the good vaporization of the injected feedstock before to reach the catalytic zone. Before each reaction, the system was preheated at the reaction temperature and was purged with 50 mL/min of nitrogen flow during at least 30 min. During the reaction step, feedstock was injected at a controlled flow to adjust the Weight Hourly Space Velocity (WHSV) of the reaction considering the amount of catalyst previously loaded. At the same time, a controlled nitrogen flow is introduced in the reactor to adjust the initial hydrocarbon partial pressure. Typically, nitrogen flow was adjusted to obtain a hydrocarbon partial pressure of 0.33 atm. Samplings at different Time-On-Stream (TOS) to determine the evolution of the activity and selectivity of the tested catalyst(s) were carried out. Liquid products were collected at the exit of the reactor in glass receivers kept at 289 K by means of a cold bath, while gaseous products were collected in a gas bag. Nitrogen was used as internal standard for the quantification of the gaseous fraction. Gaseous products were analysed using a Shimadzu GC- 2014 chromatograph equipped with three detectors: two thermal conductivity detectors (TCD) for analysis of H2 and N2, after separation in a 6 ft MS 5A and 2.5 ft MS 13X molecular sieve respectively, and a flame ionization detector (FID) for C1-C6 hydrocarbons which were separated in a 164 ft plot/A^Os column. Liquid products were weighted and analysed through comprehensive two-dimensional gas chromatography (GCxGC). An Agilent 7890A GC equipped with a capillary flow technology GCxGC flow modulator (Agilent Technologies), a split injector at 250°C (250/1 ) (1.0 pL) and a flame ionization detection (FID) system (300°C, 200 Hz). The column sets consisted of a combination of a non-polar column (HP-5 30 m x 0.25 mm x 0.5 pm, Agilent technologies) and a polar column (HP-INNOWax 5 m x 0.25 mm x 0.15 pm, Agilent technologies), placed in different ovens. The system was operated at constant flow of 0.5 ml/min for the first column and 30 ml/min for the second one, and at programmed temperatures, going from 50 °C (1 min) to 260 °C (79 min) at 3 °C/min. High purity hydrogen was used as a carrier gas at constant flow rate. Modulation period was set to 4.5 s and data were processed using the GC image™ v2.1 software developed by GC Image™, LLC. At the end of the reaction, stripping was carried out during 15 minutes with a high nitrogen flow to remove all the hydrocarbons of the system before to carry out the regeneration step where 100 mL/min of air was injected at a temperature of 550°C. Coke was estimated by measuring the CO2 produced during the catalyst regeneration step, by means of an infrared analyser. This procedure was used in the following Examples. 1 - Influence of Si/AI molar ratio and

nd liqht olefins

alkanes, as illustrated with

ZSM-5

different Si/AI molar ratios

Experiments were conducted at 400°C, 500°C and 600°C, atmospheric pressure, and constant initial partial pressure of hydrocarbons of 0.33 atm through nitrogen dilution, space velocity was varied in order to obtain different conversion levels and to compare product selectivity with special emphasize in propylene. Influence of WHSV conditions over the cracking activity in terms of conversion and propylene yield for ZSM-5 catalysts having different Si/AI molar ratio and similar other textural properties for the processing of hexadecane feed at different temperature is shown in Figures 3 and 4. Figure 3 illustrates influence of the Si/AI molar ratio of the ZSM- 5 catalyst over catalyst activity at 400°C (A), 500°C (B) and 600°C (C), when cracking pure hexadecane at total pressure of 1 atm, feed’s partial pressure of 0.33 atm, and varying WHSV. Figure 4 illustrates influence of the cracking temperature over catalyst activity and propylene selectivity for the ZSM-5 catalysts having different Si/AI molar ratio of 31 (A and B), 125 (C and D) and 400 (E and F) when cracking pure hexadecane at total pressure of 1 atm, feed’s partial pressure of 0.33 atm, and varying WHSV.

From Figure 3 it can be seen that the lower the Si/AI molar ratio of the used ZSM-5 catalyst, the more active the catalysts were for the cracking of pure hexadecane.

On the other hand, cracking temperature has also an impact. When cracking hexadecane in the range of 400-600°C, it can be seen from Figure 4; A, C and E that the higher the temperature the higher the conversion at a given space velocity, independently of the used ZSM-5 catalyst. Propylene selectivity is also positively affected when the cracking temperature is increased (Fig. 4; B, D and F).

Based on the results of these tests it can be concluded that an elevated cracking temperature, such as 600°C, a solid acid catalyst having limited Si/AI molar ratio, and an elevated space velocity (limited residence time), appear to be beneficial contributors for reaching high conversion levels and for maximizing the yields of light olefins and propylene. Example 2 - Cracking product yields obtained by once-throuqh single cracking of n- paraffinic feed R1 and isomerized feed R2 using ZSM-12 and comparison to ZSM- 5

Catalytic cracking tests with highly n-paraffinic feed R1 and highly isoparaffinic feed R2 were conducted using ZSM-12 or ZSM-5(125) at 600°C and at different space velocities. The catalytic cracking reaction conditions and products of each experiment are shown in Tables 2 and 3. The WHSV was varied to obtain different conversion levels. This was done by adjusting the feed rate and/or the catalyst weight. Table 2. R1 , n-paraffinic feed (mass-balance normalized)

Table 3. R2, isomerized feed (mass-balance normalized)

Tables 2 and 3 show product yields for a once-through single cracking reactor process at varying WHSV. It can be seen that ZSM-12 has high activity, providing conversion of n-paraffinic feed of about 90% or more, and of isomerized feed of about 70% or more, typically about 80% or more. The conversions were almost 20 wt. percentage points higher with n-paraffinic feed and over 30 wt. percentage points higher with isomerized feed, compared to the ZSM-5 at comparable WHSV. Additionally, the conversion provided by ZSM-12 was of similar magnitude for both feeds, while ZSM-5’s conversion of n-paraffinic feed was about 1.5 times the conversion of isoparaffinic feed.

It was also surprising that the yields of propylene, C4 mono-olefins and ethylene provided by ZSM-12 were of similar level for both feeds, and propylene and C4 mono-olefins yields were much higher than those provided by ZSM-5 at comparable WHSV. C3 olefinicity and C2 olefinicity were very high for both catalysts, at least about 90 wt.-% and 75 wt.-%, respectively. While both zeolites generated more propylene than ethylene, the propylene to ethylene weight-ratios were clearly higher with ZSM-12 compared to ZSM-5. Regarding C4 mono olefins, an interesting benefit of the present process were elevated isobutylene yields, that were of similar level for both feeds. Isobutylene is a valuable chemical, usable e.g. in the production of butyl-rubber by polymerisation, iso-octane by alkylation with butane, diisobutylene by dimerization, and methyl and ethyl tert-butyl ethers by reacting with methanol or ethanol, respectively.

Yields of less desired products like methane, a strong greenhouse gas, and 1 ,3- butadiene, a coke-precursor, were minimal by ZSM-12 for both feeds. Also the coke yields obtained by ZSM-12 were low.

Example 3 - ZSM-12 vs ZSM-5: Maximum propylene yields in a once-throuqh single catalytic cracking reactor

Experiments were conducted using ZSM-5 (31 ) or ZSM-12 (40) zeolites, at 600 °C, atmospheric pressure and 0.33 atm of initial partial pressure of hydrocarbons (diluted in N2), space velocity was varied in order to obtain different conversion levels. Influence of WHSV conditions over cracking activity in terms of conversion and propylene yield for both catalysts using n-paraffinic and isomerized feeds is shown in Figure 5.

After checking that, in the case of the n-paraffinic feed R1 , the catalytic activity is not controlled by diffusion limitations, it can be clearly observed in (A) of Figure 5 that ZSM-5 catalyst is a more active material than ZSM-12 to crack high molecular weight linear paraffins as higher conversion levels can be reached under the same severity conditions. Despite that both materials have similar Si/AI molar ratio, and comparable acidic properties, the higher cracking activity of ZSM-5 catalyst is very likely due to a more constrained environment enhancing the confinement effect and consequently the cracking rate of hydrocarbons. On the other hand, it is also worth to observe that, in the case of linear paraffins, higher propylene yields and light olefins in general may be obtained using ZSM-5 as seen in (B) of Figure 5. This is believed to be due to the shape selectivity of the ZSM-5 zeolite which limit the occurrence of bimolecular reactions and, in particular, secondary hydrogen transfer reactions limiting the production of light olefins, as naphthenes formed by dimerization-cyclization of olefins donate hydrogen to other olefins and become aromatics while the olefins are hydrogenated into paraffins. Paraffins are more stable than olefins, partially explaining the lower rate of re-cracking of the ZSM-12 catalyst, even if part of the lower cracking activity of the ZSM-12 catalyst could be also explained by diffusion effects. In addition, re-cracking of saturated hydrocarbons compared to olefins is less selective to light olefins explaining also partially the slightly lower selectivity to propylene of ZSM-12 zeolite compared to ZSM-5.

However, in the case of the iso-paraffinic feedstock R2, ZSM-12 having onedimensional channel system of 12-ring channels (5.7 x 6.1 A) is more active compared to ZSM-5 zeolite, which has three-dimensional channel system of 10-ring channels (5.2 x 5.7 and 5.3 x 5.6 A channels), highlighting the importance of mass transfer issues and accessibility when branched hydrocarbons are processed ((A) in Figure 5). Although the maximum propylene yield provided by ZSM-5 was slightly higher with the n-paraffinic feed in a once-through set-up compared to ZSM-12, in turn, for isomerized feed ZSM-12 provides significantly higher maximum propylene yield than ZSM-5 ((B) in Figure 5). The low catalytic activity of ZSM-5 when branched hydrocarbons are processed implies to apply much lower WHSV or higher residence times to reach high conversion levels and propylene yields, leading to an important increase of the kinetic rate of secondary mechanisms and in particular of hydrogen transfer reactions explaining, in a large extent, the much lower light olefins and propylene selectivity obtained when ZSM-5 is used to crack the isomerized feed. On the contrary, in the case of the ZSM-12, as lower mass transfer issues occur, very similar WHSV conditions may be applied to obtain high conversion levels and propylene yields, independently of the processed feedstock, explaining why ZSM-12 is a more flexible catalyst in terms of feedstock processability to efficiently and selectively produce light olefins and propylene from linear and branched hydrocarbons cracking in a fixed-bed reactor.

One possibility to tackle with the observed limitations of the microporous ZSM-5 as cracking catalyst of branched hydrocarbons would be to improve the accessibility and molecular traffic of this latter. Different well-known strategies can be applied including the incorporation of mesoporosity and the use of small crystal size. The efficiency of such strategies has been evaluated preparing and testing hierarchical and submicron-sized ZSM-5 catalysts. The prepared submicron-sized catalyst has the following properties: Si/AI molar ratio 76; BET area 387 m²/g; micropore area 330 m²/g; external area 57 m²/g; micropore volume 0.17 cm³/g; mesopore volume 0.09 cm³/g; crystal size <0.05 microns. The prepared ZSM-5 hierarchical catalyst has the following properties: Si/AI molar ratio 86; BET area 365 m²/g; micropore area 302 m²/g; external area 63 m²/g; micropore volume 0.16 cm³/g; mesopore volume 0.21 cm³/g; crystal size 0.4-0.8 microns. Experiments were conducted at 600 °C, atmospheric pressure and 0.33 atm of initial partial pressure of hydrocarbons, space velocity was varied to obtain different conversion levels. Influence of WHSV conditions over cracking activity and selectivity in terms of conversion and propylene yield processing the isomerized feeds is shown in Figure 6. According to results, no significant boost of conversion using submicron-sized or hierarchical ZSM-5 catalysts instead of standard microporous ZSM-5 catalysts has been highlighted ((A) in Figure 6), suggesting that the enhancement of acid sites accessibility and molecular traffic bring a limited advantage in terms of cracking activity, since the reactant shape selectivity does not change. The controlling factors appears still to be the diffusion through micropores. Similar finding regarding propylene yield and selectivity may be done ((B) in Figure 6), as it is necessary to apply quite long residence time to reach high conversion levels when working with both, submicron-sized and hierarchical ZSM-5 catalysts. In such conditions, the preponderant importance of the reactions occurring at the external surface mainly control the extent of hydrogen transfer mechanisms leading to limited light olefins maximum yields and explaining why the use of submicron-sized or hierarchical zeolites does not seem to significantly boost the light olefins selectivity of ZSM-5.

Altogether, it can be concluded that ZSM-12 zeolite is thus, far beyond, the most active and selective catalyst to produce propylene and light olefins from the cracking of high molecular weight branched hydrocarbons in a once-through catalytic cracking. As it can be observed in Table 4, ZSM-12 catalyst brings the advantage to reach high conversion levels and maximum propylene yield independently of the processed feedstock, unlike ZSM-5. ZSM-5’s conversion of n-paraffinic feed is about 1 .5-2 times the conversion of isoparaffinic feed at the same WHSV as already observed in (A) of Figure 5. Even if both zeolites generated more propylene than ethylene, the propylene to ethylene weight-ratios are clearly higher with ZSM-12 compared to ZSM-5. ZSM-5 zeolite promotes the olefinicity of C2 fraction, i.e. ratio of ethylene to total C2, for both feedstocks at maximum propylene yield, around 0.9 versus 0.8 for ZSM-5 and ZSM-12 catalyst, respectively as seen in Table 4, while ZSM-12 catalyst promotes the olefinicity of the C3 fraction, i.e. ratio of propylene to total C3, around 0.9 versus 0.8 for ZSM-12 and ZSM-5 catalyst, respectively. Furthermore, the yields of propylene, C4 mono-olefins and ethylene provided by ZSM-12 were of similar level for both feeds, and propylene and C4 mono-olefins yields were much higher than those provided by ZSM-5 at maximum propylene yield. Similar low coke yields at the maximum propylene yield are obtained for both catalysts and feedstocks.

Table 4. Product yields at maximum propylene yield when cracking n-paraffinic or isoparaffinic feedstock in a fixed-bed reactor at Trx = 600 °C, P_tot = 1 atm, P°Feed = 0.33 atm and using ZSM-5 (31 ) or ZSM-12 (40) zeolites. _

ZSM-5 (31 ) ZSM-12 (41 )

Processed feedstock n-paraffinic iso-paraffinic n-paraffinic iso-paraffinic

Cracking temperature (°C) 600 600 600 600

Conversion (wt.%) 97.8 72.0 93.4 84.4

C1 -C4 paraffins (wt.%) 13.5 13.8 9.1 8.8

| Methane (wt.%) 0.9 1.1 2.1 1.5

Ethane (wt.%) 1.4 1.2 2.5 1.2

Propane (wt.%) 6.3 6.2 2.1 2.7

| Butanes (wt.%) 4.9 5.3 2.4 3.4

C2-C4 olefins (wt.%) 65.9 40.7 63.4 54.9

| Ethylene (wt.%) 12.1 10.9 8.1 5.9

Propylene (wt.%) 34.0 20.2 28.3 28.0

Butenes (wt.%) 19.8 9.6 27.0 21.0

C2=/C3= 0.36 0.54 0.29 0.21

C2 olefinicity 0.90 0.91 0.76 0.83

C3 olefinicity 0.84 0.77 0.93 0.91

HTI (C4s=/C4s) 4.0 1.8 12.2 6.4

C5+ (wt.%) 20.2 45.1 27.0 35.8

| n-Paraffins (wt.%) 10.1 19.6 11.6 6.0

| iso-Paraffins (wt.%) 1.2 15.2 0.9 18.8

Olefins (wt.%) 5.0 3.1 12.8 6.9

| Cycloalkanes (wt.%) 1.0 1.0 0.6 1.1

| Cyclic olefins (wt.%) 0.8 0.6 0.3 0.5

| Aromatics-BTX (wt%) 1.8-1.4 5.4-4.5 0.6-0.2 2.3-1.3

| Unknowns (wt.%) 0.3 0.2 0.2 0.2

C5+ Olefinicity 0.27 0.08 0.49 0.21

Coke (wt.%) 0.4 0.4 0.5 0.5 Typically, the relative activity of cracking catalysts for generating secondary reactions is estimated using a hydrogen transfer index (HTI). Catalysts with higher HTI (expressed in the present disclosure as a weight ratio of unsaturated C4 to saturated C4) are thought to generate fewer secondary reactions, preserving a greater quantity of C5+ olefins, that can be subsequently recycled and cracked to lighter olefins. In order to maximise propylene, it is important to suppress hydrogen transfer by maximizing the availability of olefin precursors. Hydrogen transfer reactions may involve the formation of bulky bimolecular reaction intermediates, and are believed to be controlled by steric constraints, due to the space available inside the micropores of the zeolites. They can also occur on the outer surface of the zeolite particles. Generally, it has been thought that the smaller the pore size of the zeolite, the greater the extent of the suppression of the hydrogen transfer reactions of the alkenes, meaning that the HTI should increase with decreased pore size of the zeolite. However, in the present invention it has been found that ZSM-12 is able to convert both the n-paraffinic and the highly iso-paraffinic feeds to propylene with excellent conversion levels, and with less secondary reactions than ZSM-5 as shown clearly in Table 4 where ZSM-12 provides much higher HTI values than ZSM- 5 at maximum propylene yield. Without wishing to be bound to a theory it is believed that even though the smaller channel size of ZSM-5 may help to suppress side reactions, the crossing points in its 3D interconnecting channel system provide more space for side reactions, compared to ZSM-12 1 D channel system. It is also interesting to notice that the yield of C5+ olefins provided by ZSM-12 is high. These unexpected benefits were achieved by utilising a zeolite of a low to medium Si/AI ratio and a framework type MTW having one-dimensional channel system and slightly larger channels than in conventionally used zeolites, at the specified operating conditions.

Optionally it is possible to subject at least one or more fractions separated from the catalytically cracked stream, such as a fraction rich in saturated C2-C4 hydrocarbons, to a thermal cracking reaction, preferably to steam cracking, in a separate thermal cracking reactor to further increase the yield of light olefins, transforming partially C2-C4 paraffins into C2-C4 olefins. Yields of thermal cracking are highly dependent of the processed feedstock but typically e.g. steam cracking provides 80 wt.-% of ethylene, 2 wt.-% of propylene and 3 wt.-% of butenes from ethane, 45 wt.-% of ethylene, 15 wt.-% of propylene and 3 wt.-% of butenes for propane and 37 wt.-% of ethylene, 18 wt.-% of propylene and 8 wt.-% of butenes for butanes. Considering such steam cracking yields, it may be estimated that it is possible to obtain an additional yield of light olefins of around 5 wt.-% including 1 wt.-% of propylene for the proposed process using ZSM-12 zeolite, independently of the processed feed.

Additionally, the composition of the recycle stream has significant impact on its usability. E.g. an elevated C5+ olefins to paraffins ratio is desired for a recycle stream, due to the benefits discussed in the foregoing. Also, a low total content of cyclics, i.e. aromatics and naphthenes, particularly a low total content of aromatics, is desired for a recycle stream as naphthenes convert relatively easily to aromatics, and aromatics may increase coking and are relatively inert and hence not likely to convert to the desired products. Compared to ZSM-5, ZSM-12 allows to limit the production of C1-C4 paraffins and aromatics which are difficult to recycle to further increase the selectivity to the desired light olefins and can act as coke precursors. It is worth to mention that, already obtaining around 55-60 wt.-% of light olefins, about one-third of the feed can be still recycled as C5+ hydrocarbons when ZSM- 12 is used as catalyst in a once-through cracking reactor. On the other hand, obtaining only a slightly higher light olefins yield (i.e. 65 wt.-%), very few C5+ hydrocarbons can be recycled when ZSM-5 is used as catalyst to crack the n- paraffinic feed (i.e. less than 20 wt.-%). On the contrary, a huge fraction of C5+ hydrocarbons would need to be recycled when ZSM-5 is used as catalyst to crack the isoparaffinic feed (i.e. about 40 wt.-%). Moreover, compared to ZSM-5, ZSM-12 provided with both feeds higher yield of C5+ olefins, and significantly higher C5+ olefin to C5+ paraffins weight-ratio, which is highly beneficial when recycling C5+ fraction, as C5+ olefins are easier to crack compared to paraffins having same carbon number, and upon cracking each C5+ olefin molecule produces 2 molecules of lighter olefins, while a paraffin molecule produces only 1 molecule of light olefin and light paraffin. Altogether, owing the high flexibility and product slate obtained when using ZSM-12 instead of ZSM-5, as shown in the foregoing Examples, it seems thus interesting to maximize the yield of light olefins using ZSM-12 zeolite in a first catalytic cracking and subsequently re-cracking or recycling at least part of a fraction of C5+ hydrocarbons, as studied in the following Examples.

Example 4 - ZSM-12 and ZSM-5 zeolites to maximise propylene yield in a once- through catalytic cracking process involving separating and recycling a C5+ stream to the same reactor

One possibility for the intensification of the propylene production is the recycling of at least part of the catalytically cracked stream which has not been cracked enough during the first pass (i.e. C5+ fraction) to give additional C2-C4 products and especially light olefins. It is worth mentioning that in this case also, such intensification may not be done in a single reactor increasing the residence time as it is believed that the increase of the residence time during the catalytic cracking of the feed would promote the occurrence of secondary reactions such as hydrogen transfer reactions leading to decreased yields of light olefins at the benefit of light paraffins and aromatic compounds. It is thus of interest to separate the light olefins stream from the C5+ fraction at the exit of a first cracking reactor working under operative conditions which maximize the production of propylene and light olefins before recycling the C5+ fraction in the same reactor to re-crack the C5+ fraction into more desired olefinic products. Figure 7A depicts a generalized schematic diagram of an embodiment of a catalytic conversion system including recycling of C5+ fraction separated 200 to the catalytic cracking 100.

Table 8. Product yields at maximum propylene yield when cracking n-paraffinic feedstock in a fixed-bed reactor using ZSM-5 (31 ) or ZSM-12 (40) and recycling C5+ fraction in the same fixed bed cracking reactor. _

1^st pass cracking Cracking of C5+ Recycling _ fraction _ equilibrium

Catalyst ZSM-5 ZSM-12 ZSM-5 ZSM-12 ZSM-5 ZSM-12

Conversion (wt.-%) 97.8¹ 93.4¹ 63.6² 61.0² 100.0¹ 100.0¹ C1 -C4 paraffins (wt.%) 13.6 8.9 7.8 8.6 13.2 14.2 | Methane (wt.-%) 0.1 2.0 1.4 0.7 1.7 1.1 | Ethane (wt.-%) 1.5 2.5 I .7 0.9 2.5 1.2 | Propane (wt.-%) 6.3 2.1 2.6 3.3 4.9 5.1 | Butanes (wt.-%) 4.9 2.3 2.1 3.7 4.1 6.8

C2-C4 olefins (wt.%) 65.8 63.4 55.3 51.8 86.1 85.1

| Ethylene (wt.-%) 12.1 8.1 I I .5 8.2 14.4 8.7 | Propylene (wt.-%) 34.0 28.3 28.6 26.0 45.0 43.8 | Butenes (wt.-%) 19.7 27.0 15.2 17.6 26.7 32.6 C2=/C3= 0.36 0.29 0.37 0.31 0.32 0.20 C2 olefinicity 0.90 0.76 0.87 0.90 0.85 0.88 C3 olefinicity 0.84 0.93 0.92 0.88 0.90 0.89 HTI (C4s=/C4s) 4.0 11.7 8.4 4.7 6.5 4.8 C5+ (wt.%) 20.2 27.2 36.4 38.9

| n-Paraffins (wt.-%) 10.2 11.6 18.7 15.3 | iso-Paraffins (wt.-%) 1.2 0.9 4.1 6.8 | Olefins (wt.-%) 5.1 12.8 5.8 8.8 | Cycloalkanes (wt.-%) 1.0 0.6 1.3 1.2 | Cyclic olefins (wt.-%) 0.7 0.3 0.9 0.8 | Aromatics-BTX (wt.-%) 1.8-1.4 0.9-0.2 5.5-3.9 5.6-3.9 | Unknowns (wt.-%) 0.2 0.1 0.1 0.4 C5+ Olefinicity 0.25 0.47 0.16 0.23 Coke (wt.%) _ 0.4 0.5 0.5 0.7

¹ Defined as C1-C12 + coke yields. ² Defined as C1-C4 + coke yields Table 9. Product yields at maximum propylene yield when cracking iso-paraffinic feedstock in a fixed-bed reactor using ZSM-5 (31 ) or ZSM-12 (40) and recycling C5+ fraction in the same fixed bed cracking reactor.

1^st pass Cracking of Recycling cracking C5+ fraction equilibrium

Catalyst ZSM-12 (40) ZSM-12 (40) ZSM-12 (40)

WHSV (h’¹) 40 40 40

Cracking temperature (°C) 600 600 600

Recycling (wt.%)

85.3

Conversion (wt.%) 84.4¹ 47.2²

100.0¹

C1-C4 paraffins (wt.%) 8.9 5.0 11.7

| Methane (wt.-%) 1.5 1.1 2.2

| Ethane (wt.-%) 1.3 0.9 1.7

| Propane (wt.-%) 2.7 1.3 3.5

| Butanes (wt-%-) 3.4 1.7 4.3

C2-C4 olefins (wt.%) 54.9 41.7 87.2

| Ethylene (wt.-%) 5.9 6.3 10.6

| Propylene (wt.-%) 28.0 21.3 46.8

| Butenes (wt.-%) 21.0 14.1 29.8

C2=/C3= 0.21 0.29 0.23

C2 olefinicity 0.83 0.88 0.86

C3 olefinicity 0.91 0.94 0.93

HTI (C4s=/C4s) 6.4 8.2 6.9

C5+ (wt.%) 35.7 52.7

| n-Paraffins (wt.-%) 6.0 13.4

| iso-Paraffins (wt.-%) 18.8 23.7

| Olefins (wt.-%) 6.9 8.5

| Cycloalkanes (wt.-%) 1.1 1.3

| Cyclic olefins (wt.-%) 0.5 0.7

| Aromatics-BTX (wt.-%) 2.3-1.3 4.9-3.2

| Unknowns (wt.-%) 0.1 0.2

C5+ Olefinicity 0.19 0.16

Coke (wt.%) _ 0.5 0.6

¹ Defined as C1-C12 + coke yields. ² Defined as C1-C4 + coke yields. To estimate the interest of the recycling strategy, in the case of the n-paraffinic feed, a once-through run was conducted with ZSM-5 (31 ) or ZSM-12 (40) under operative conditions maximizing the production of light olefins and propylene, i.e. 600 °C, atmospheric pressure, 0.33 bar of hydrocarbon initial partial pressure and WHSV of 250 and 25 h^-1 for respectively the ZSM5 and ZSM-12 zeolites. In both cases, C5+ fraction was selectively recovered and accumulated and a representative recycling stream including fresh feedstock and recycled C5+ fraction was prepared and recracked under the same operating conditions. The previously described procedure was repeated the necessary number of times to reach the equilibrium state where the C5+ fraction recycling rate stays constant, and no significant changes of products slate can be appreciated. Results can be seen in Table 8. As it can be clearly observed, it is possible to reach an equilibrium state including very similar recycling rate of 50-55 wt.-% for both catalysts indicating clearly that it is possible to further transform C5+ fraction compounds into more desired C1 -C4 compounds without significant accumulation of stable products such as short paraffins (mainly C5) and mono-aromatics even if potential accumulation of undesired cracking products in the recycling loop could be handled by removing a part of the recycled stream continuously or occasionally.

From the results of this test run, it can be seen that the yield of C2-C4 olefins, especially propylene, can be increased significantly by arranging ZSM-12 catalyst in a single reactor, separating C4+ fraction, preferably C5+ fraction as in this example, from the catalytically cracked stream of the reactor as a recycled stream, and re-cracking the recycled stream in the same reactor mixed with the fresh feed. At the equilibrium and for both zeolites, the proposed process scheme is very selective to light olefins representing an estimated yield of around 85 wt.-% including almost 45 wt.-% of propylene without producing an extensive quantity of light paraffins (i.e. less than 15 wt.-%) and very low amounts of coke (i.e. less than 1 wt.%). In all cases, the C2=/C3= ratio (i.e. 0.2-0.3) is far better than the one typically obtained in a steam cracking process when naphtha range feedstocks are processed (i.e. 2-3) allowing to produce more selectively propylene than ethylene and adjusting the production to the market demand. To boost even more the production of light olefins, optionally it would be possible to couple the proposed catalytic reactor with a small steam cracking unit in order to partially transform C2- C4 paraffins into C2-C4 olefins. Steam cracking yields is highly dependent of the processed feedstock but typically provides 80 wt.-% of ethylene, 2 wt.-% of propylene and 3 wt.-% of butenes from ethane, 45 wt.-% of ethylene, 15 wt.-% of propylene and 3 wt.-% of butenes for propane and 37 wt.-% of ethylene, 18 wt.-% of propylene and 8 wt.-% of butenes for butanes. Considering such steam cracking yields, it can be estimated that it would be possible to obtain an additional yield of light olefins of around 8 wt.-% including 2 wt.-% of propylene for both zeolites. It would be thus possible to reach light olefins yields as high as 90-95 wt.-% processing the n-paraffinic feedstock derived from hydrogenated vegetal and animal fats. By using this kind of direct recycling strategy it is possible to increase the light olefins yield without the necessity to use a separate additional reactor thereby reducing investment and operating costs. Concerning the iso-paraffinic feed, as mentioned in the foregoing, ZSM-5 catalyst is not the best suitable catalyst to maximize conversion and light olefins yield due to important diffusion limitations. In this case, it makes much more sense to use exclusively ZSM-12 catalyst. Thus, in the same way that before mentioned, the recycling strategy has been experimentally tested and a once-through run was conducted with ZSM-12 (40) under operative conditions maximizing the production of light olefins and propylene, i.e. 600 °C, atmospheric pressure, 0.33 atm of hydrocarbon initial partial pressure and WHSV of 40 h’¹. C5+ fraction was selectively recovered and accumulated and a representative recycling stream including fresh feedstock and recycled C5+ fraction was prepared and re-cracked under the same operating conditions. Such procedure was repeated the necessary number of times to reach the equilibrium state where the C5+ fraction recycling rate stays constant, and no significant changes of products slate can be appreciated. Results can be seen in Table 9. According to results, due to the more recalcitrant behaviour of the iso-paraffinic feedstock, it is necessary to apply a higher recycling rate to reach the equilibrium than in the case of n-paraffinic feedstocks. Nevertheless, it is of extremely high interest to observe that it is possible to reach very similar and extremely high light olefins and propylene yields of respectively 87 wt.-% and 47 wt.-% when processing the iso-paraffinic feedstock and applying the recycling strategy. In this case also the proposed conversion scheme is very selective to light olefins producing few light paraffins and coke (i.e. 12 wt.-% and 1 wt.-% respectively). As explained in the foregoing, the coupling of the present process with a small thermal cracking unit, such as steam cracking unit, when processing the isoparaffinic could boost the light olefins yield of around 6 wt% including 1 additional point of propylene.

Such results clearly indicate the huge interest to use ZSM-12 zeolite instead of ZSM- 5 zeolite as it is possible to obtain very similar performances in the case of the n- paraffinic feedstocks and much better results in the case of the iso-paraffinic feedstock, bringing a clear advantage in terms of feedstock flexibility to the proposed processing scheme.

Yet another option to intensify the yield of light olefins involves running the process as disclosed above, but re-cracking 300 the separated C5+ fraction in a separate catalytic cracking reactor as depicted in Figure 7B, instead of recycling it to the same reactor as depicted in Figure 7A.

Compared to the direct recycling strategy (Fig 7A) as disclosed above, re-cracking in a separate catalytic cracking reactor (Fig 7B) may allow reaching even higher increase in the light olefins yield and/or reduction of side reactions. Such enhancement may not be achieved in a single reactor merely by increasing the residence time, as this is foreseen to promote the occurrence of secondary reactions such as hydrogen transfer reactions leading to decreased yields of light olefins at the benefit of light paraffins and aromatic compounds. Compared to direct recycling (Fig 7A), re-cracking in separate catalytic cracking reactor (Fig 7B) brings more flexibility in terms of operating conditions including the possibility to tune the cracking temperature and space velocity both in the first catalytic cracking stage and in the re-cracking stage of the C5+ fraction, allowing to deeper maximize the production of light olefins in each reactor.

By using ZSM-12 as the cracking catalyst both in the first reactor 100 and in the second reactor 200, brings the advantage to produce less light paraffins and aromatics and to be more selective to light olefins, keeping at the same time a higher yield of recyclable C5+ fractions, which could be recycled to the first or second reactor to further increase the light olefin yields.

Similar considerations apply both the n-paraffinic and the iso-paraffinic feed. Conversion of the isomerised feed into C1 -C4 hydrocarbons is more refractory than the n-paraffinic feed, leading to higher yield of remaining C5+ fraction even after the secondary cracking. Recycling the remaining C5+ fraction to the first 100 or to the second reactor 300 is thus an interesting option when processing the iso-paraffinic feedstock to further increase the light olefin yields.

As supported by the results, the benefits of the present method may be summarized as follows.

• The conversion-normalised total yield of C2-C4 olefins is high, typically at least 50 wt.-%, often at least 55 wt.-%, even at least 60 wt.-%. The high conversion levels may be achieved even in once-through process i.e. without recycling.

• Conversion-normalised total yield of C1 -C4 paraffins remain controlled. In the experiments the yield was about 9 wt.-%. C1-C4 paraffins are beneficial in an optional recycle, acting as a diluent gas decreasing partial pressure of the hydrocarbons in the cracking feed, thereby favouring formation of light olefins. The formed C2-C4 paraffins are also usable in steam cracker feeds and in feeds for catalytic dehydrogenation for producing light olefins. Additionally or alternatively, the formed C1 -C4 paraffins could be used in steam-reforming to produce H2, e.g. for use in a hydrotreatment for providing a sustainable hydrocarbon feed for the present method. Sustainable methane is valuable also in various conventional uses of methane, such as in methane-to-methanol conversion.

• The present process using ZSM-12 provides high C3=/C2= weight ratio, significantly higher than obtainable e.g. by steam cracking (typically far below 1 ) which is the current industry standard for producing fossil propylene, even if obtained only as a by-product to fossil ethylene. High C3=/C2= weight ratio is especially beneficial as using currently available technologies it is more difficult to produce renewable propylene than renewable ethylene.

• The present process using ZSM-12 facilitates easy recovery of a sustainable propylene composition having high propylene content and very low propane content. A C3 fraction separated from the catalytically cracked stream of the present process has excellent C3=/total C3 ratio (total C3 being the sum of propane and propene), typically having propylene content of at least 90 wt.-%, even 95 wt.-%. Moreover, it was surprisingly found out that the benefits may be achieved both with highly n-paraffinic feed and with highly isoparaffinic feed, making the present process flexible regarding the feed. It is highly beneficial, and not common, that similar product slate is obtainable with different feeds. High C3=/total C3 share is beneficial since propane and propylene have similar molecular size and physical properties, which makes their separation challenging. This separation is mostly carried out in distillation columns generally having more than 150 theoretical plates, and operating with very high reflux ratios, often 10-20, and at a high pressure, typically of about 16-26 atm. The separation process requires high capital cost and very high energy consumption. The propylene purity will affect the grade and value of the propylene product: for refinery grade 50-70 % purity may suffice, while for chemical grade 90-95 % purity is typically required, and for polymer grade even 99.5 % purity or higher. A typical setting for obtaining the polymer grade purity involves using a distillation column, also called a “splitter”, constituted by 152 theoretical plates and operating initially with a reflux of 24.1 in order to separate propane/propylene mixture until the purity of 99.5% is obtained. Logically, the higher the propylene/total C3 share in the beginning of the propane/propylene separation, the lower the energy needed for reaching the polymer grade purity. Also less complex/expensive equipment may suffice.

• Also the ethylene content in a C2 fraction separated from the catalytically cracked stream of the present process is high, typically at least 70%, even close to 90%, providing similar benefits as discussed above for propylene.

• The C4 mono olefins, when separated from the catalytically cracked stream of the present process, may find use as (co)monomers, as such or after derivatisation. In the experiments it was observed that the conversion normalised yields of iso-butene were good, around 10 wt.-%, regardless of the process conditions and the choice of feed. Isobutenes are desired and usable e.g. in etherification with methanol and/or ethanol to produce methyl t-butyl ether (MTBE) and/or ethyl t-butyl ether (ETBE), that are valuable high-octane gasoline components. In the experiments it was also observed that regardless of the process conditions and the choice of feed, butadiene was formed only little, which is beneficial, as it may cause fouling and coking.

• Low aromatics formation in the present process is beneficial particularly if recycling part of the catalytically cracked stream, as aromatics may increase coking, and they are relatively inert and hence not likely to be converted to the desired products even upon recycling. With higher feed rates the aromatics formation was observed to decrease further.

• Compared to typical cracking processes, in the present process the formation of coke is low, which is beneficial as then only minimal amounts of the valuable sustainable carbon is lost to the coke. Low coke formation at the specified operating temperatures is surprising, and contrary to persistent belief in the field. For example in WO22096781A1 it was warned that at reaction temperatures above 450 °C formation of coke and aromatics may start to increase, so that an undesired high level of aromatics in the catalytically cracked stream is reached. • The present process converts both n-paraffinic and isomerised feeds well, which provides flexibility to the process e.g. depending on feedstock availability and/or catalyst life cycle. Isoparaffins in the feed may lead to a bit higher formation of aromatics, which may lead to more coke on the catalyst. Hence it may be better to feed less or non-isomerised feed in the beginning and higher isomerised feed towards the end of the catalyst life cycle. Generally it is highly beneficial that the product slate remains similar even when the feed composition is varied, requiring minimal or no changes to the processing of the catalytically cracked stream, including product recovery, separation, and/or purification.

• The catalytically cracked stream obtained by the present process has very high share of C5+ olefins in the C5+ stream, which is beneficial when recycling: one C5+ olefin molecule forms 2 lighter olefins when cracked, while one C5+ paraffin molecule forms 1 lighter olefin and 1 lighter paraffin. C5+ olefins are also easier to crack than the paraffins having same carbon number.

• In some further experiments recycling part of the catalytically cracked stream, accumulation of C5+ isoparaffins was observed. If desired, these could be recovered e.g. for use in solvents or as gasoline components.

• Zeolites of MTW framework type are available with the specified Si/AI molar ratio, and desired porosity and acidity characteristics.

• Desired highly paraffinic sustainable hydrocarbon feeds with limited contents of olefinic and cyclic hydrocarbons, especially aromatics, and with very low or no oxygen content are readily available e.g. by HVO technology (hydrotreatment of vegetable oils and/or animal fats).

• The present process may utilise conventional reactor systems, without a need for tailor-made equipment.

• Due to the relatively high operating temperature of the present process, both the cracking feed and the products are in gas phase allowing higher feed rates with less pressure drop, less limitations to the mass transfer and better mixing characteristics in general, compared to processes where the feed is not fully or at all in gas phase.

• Additionally the relatively high operating temperature of the present process is close to the typical catalyst regeneration temperatures, so that after the regeneration, the regenerated cracking catalyst does not require significant cooling before re-introducing to the catalytic cracking reactor. This may provide shorter catalyst regeneration cycle time.

The specific examples provided in the description given above should not be construed as limiting the scope and/or the applicability of the appended claims. Lists and groups of examples provided in the description given above are not exhaustive unless otherwise explicitly stated.

Claims

1 . A method for producing propylene, the process comprising the following steps: a) providing a sustainable hydrocarbon feed; b) subjecting a cracking feed comprising the sustainable hydrocarbon feed to catalytic cracking reaction in the presence of a cracking catalyst comprising a zeolite having a framework type MTW and Si/AI molar ratio of 10-200, preferably 20-150, more preferably 30-100, at a temperature in a range from 500 °C to 850 °C, preferably in a range from 520 °C to 820 °C, and a residence time from 0.1 s to 150 s, preferably from 0.2 s to 100 s, more preferably from 0.2 s to 50 s, to produce a catalytically cracked stream, and c) separating from the catalytically cracked stream at least a fraction rich in propylene.

2. The method according to claim 1 , wherein the sustainable hydrocarbon feed comprises, based on the total weight of the sustainable hydrocarbon feed, at least 50 wt.-%, preferably at least 60 wt.-%, more preferably at least 70 wt.-% paraffins, preferably having a carbon number of at least C12; wherein 1 - 99 wt- %, preferably 3 - 98 wt.-%, more preferably 5 - 95 wt.-% of the paraffins are isoparaffins.

3. The method according to claim 1 or 2, wherein the sustainable hydrocarbon feed has a T5 temperature (EN ISO 3405-2019) at least 180°C, preferably at least 190°C, more preferably at least 200°C, and T95 temperature (EN ISO 3405- 2019) at most 500°C, preferably at most 450°C, more preferably at most 430°C; and/or T5 and T95 temperatures (EN ISO 3405-2019), preferably initial boiling point and final boiling point, within 180°C - 500°C, preferably within 190°C - 450°C, more preferably within 200°C - 430°C.

4. The method according to any one of claims 1 to 3 wherein the sustainable hydrocarbon feed has a difference between the T95 and T5 temperatures (EN ISO 3405-2019) at most 300°C, preferably at most 200°C, more preferably at most 150°C, even more preferably at most 100°C, still more preferably at most 80°C.

5. The method according to any one of claims 1 to 4 wherein the zeolite is selected from a group consisting of ZSM-12, Nll-13, TPZ-12, Theta-3, and combinations thereof, preferably ZSM-12.

6. The method according to any one of claims 1 to 5 wherein the zeolite is ZSM- 12, and the zeolite has one or more of the following properties an aluminium content from 0.1 wt.-% to 1 .5 wt.-%, preferably from 0.6 wt.- % to 1 .3 w.t-%, determined by ICP-OES; a BET surface area from 200 m²/g to 400 m²/g, preferably from 250 m²/g to 380 m²/g, determined by nitrogen physisorption.

7. The method according to any one of claims 1 to 6 wherein the catalytic cracking reaction is conducted at a pressure in a range from 50 kPa to 150 kPa (absolute), preferably form 80 kPa to 120 kPa (absolute), more preferably at atmospheric pressure.

8. The method according to any one of claims 1 to 7, wherein the cracking feed comprises, based on the total weight of the cracking feed, from 0.5 wt.-% to 50 wt.-%, preferably from 0.5 wt.-% to 30 wt.%, more preferably from 1 wt.-% to 20 wt.-% a diluent gas.

9. The method according to claim 8 wherein the diluent gas comprises at least one or more of steam, methane, ethane, propane, nitrogen, carbon monoxide, and/or carbon dioxide.

10. The according to any one of claims 1 to 9 wherein step c) comprises separating from the catalytically cracked stream at least one or more further fractions selected from a fraction rich in ethylene, a fraction rich in one or more of C4 mono olefins, a fraction of hydrocarbons having a carbon number of at least C4, preferably at least C5, a fraction rich in saturated C1 -C4 hydrocarbons, and/or a fraction rich in aromatics.

11. The method according to claim 10 comprising recycling at least a part of the fraction of hydrocarbons having a carbon number of at least C4, preferably at least C5, to the cracking feed optionally after subjecting to a selective hydrogenation of diolefins to mono-olefins and/or to aromatics removal by extractive distillation or solvent extraction.

12. The method according to claim 10 or 11 comprising i. subjecting at least part of at least one or more further fraction(s) to a further cracking reaction in a separate cracking reactor to obtain further cracked stream(s) wherein the further fraction is the fraction of hydrocarbons having a carbon number of at least C4, preferably at least C5, optionally subjected to a selective hydrogenation of diolefins to mono-olefins and/or to aromatics removal by extractive distillation or solvent extraction, and the further cracking reaction is a catalytic cracking reaction in the presence of a cracking catalyst according to any one of the preceding claims; and/or wherein the further fraction is a fraction rich in saturated C2-C4 hydrocarbons, and the further cracking reaction is a thermal cracking reaction, preferably steam cracking; and ii. separating from the further cracked stream(s) at least a fraction rich in propylene.

13. The method according to any one of claims 1 to 12 comprising vaporizing the cracking feed of claim 1 before subjecting to the catalytic cracking reaction.

14. The method according to any one of claims 1 to 13 wherein for providing the sustainable hydrocarbon feed of step a) of claim 1 the method comprises a1 ) providing a renewable and/or circular hydrotreatment feed, preferably comprising at least one or more of vegetable oils, animal fats, microbial oils, and/or thermally or catalytically liquefied organic waste and residues; a2) subjecting the hydrotreatment feed, optionally mixed with a hydrocarbon liquid diluent, to hydrotreatment reaction(s) to produce a hydrotreated stream; a3) subjecting the hydrotreated stream to a gas-liquid separation to produce at least a gaseous stream and a hydrotreated liquid stream; a4) optionally subjecting at least a part of the hydrotreated liquid stream to a hydroisomerization step to produce a hydroisomerized stream; and a5) separating the sustainable hydrocarbon feed from the hydrotreated liquid stream, from the hydroisomerized stream or from a combination thereof, by dividing, stabilizing and/or fractionating.

15. The method according to claim 14 wherein the hydrotreatment reaction(s) are carried out at conditions comprising at least one or more of: a temperature in the range from 120 °C to 500 °C, a pressure in the range from 10 bar to 200 bar, a WHSV in the range from 0.1 h^-1 to 10 h’¹, a H2 flow of from 50 to 2000 N-L H2/L feed, and/or a hydrotreatment catalyst, preferably a sulphided hydrotreatment catalyst, comprising at least one or more metals from Group VIII of the Periodic Table and/or from Group VIB of the Periodic Table, preferably at least one or more of Ni, Mo, W, and/or Co, even more preferably at least one or more of Ni and/or Co and Mo and/or W, such as NiMo, C0M0, NiCoMo, NiW, and/or NiMoW, preferably on a support.

16. The method according to claim 14 or 15 wherein the hydroisomerization step is conducted at a temperature in the range from 200 °C to 500 °C, preferably from 250 °C to 450 °C; a pressure in the range from 1 to 10 MPa, preferably from 2 to 8 MPa; a WHSV in the range from 0.1 h^-1 to 10 h’¹, preferably 0.2 h^-1 to 8 h’¹, and a H2 flow of from 10 to 2000 N-L H2/L feed, preferably from 50 to 1000 N-L H2/L feed, in presence of a hydroisomerisation catalyst comprising at least one or more Group VIII metal, preferably Pd, Pt and/or Ni, and at least one or more acidic porous material selected from zeolites and/or zeolite-type materials, and optionally at least one or more of alumina, silica, amorphous silica-alumina, titanium, alumina, titania, and/or zirconia.

17. The method according to any one of claims 1 to 16 wherein the sustainable hydrocarbon feed comprises, based on the total weight of the sustainable hydrocarbon feed, at most 15 wt.-%, preferably at most 10 wt.-%, more preferably at most 5 wt.-%, even more preferably at most 1 wt.-% aromatics, and/or at most 30 wt.-%, preferably at most 25 wt.-%, more preferably at most 10 wt.-%, even more preferably at most 5 wt.-%, further more preferably at most 3 wt.-% naphthenes.

18. The method according to any one of claims 1 to 17, wherein the sustainable hydrocarbon feed comprises, based on the total weight of the sustainable hydrocarbon feed, at most 3 wt.-%, preferably at most 2 wt.-%, more preferably at most 1 wt.-%, even more preferably at most 0.5 wt.-% oxygenated hydrocarbons, expressed as elemental oxygen.

19. The method according to any one of claims 1 to 18, wherein the cracking feed comprises, based on the total weight of the cracking feed, at least 40 wt.-%, preferably at least 50 wt.-%, more preferably at least 60 wt.-%, even more preferably at least 90 wt.-% the sustainable hydrocarbon feed.

20. The method according to any one of claims 1 to 19, wherein in step b) the cracking catalyst is arranged at least in one or more catalyst bed(s), in one or more reactor(s), preferably at least in one or more fluidised catalyst bed reactor(s) and/or fixed catalyst bed reactor(s).