Process for the manufacture of renewably-sourced n-butanol or 2-ethylhexanol

Abstract

A process for the manufacture of an alcohol selected from n-butanol and 2-ethylhexanol, comprises providing a gasifier feedstock; gasifying the gasifier feedstock by partial oxidation to form a raw product gas stream comprising a plurality of gases comprising methane, hydrogen and carbon monoxide; and recovering syngas from the raw product gas stream. Further, the process comprises obtaining propylene from renewably-sourced ethanol by subjecting renewably-sourced ethanol to dehydration to produce a renewably-sourced ethylene stream; and subjecting the renewably-sourced ethylene stream to an olefin-interconversion. The propylene obtained from renewably-sourced ethanol is mixed with propylene not obtained from renewably-sourced ethanol to form mixed propylene; and the mixed propylene is subjected to a sequence of chemical conversions to obtain n-butanol or 2-ethylhexanol. The process can be easily integrated into an existing production site in which one or more chemicals of interest are manufactured based on a fossil feedstock, allowing for fossil-based propylene to be partially substituted by renewably-sourced propylene.

Description

[0001] Process for the Manufacture of Renewably-Sourced n-Butanol or 2-Ethy lhexanol

[0002] Technical Background

[0003] The present invention relates to a process for the manufacture of an alcohol selected from n-butanol and 2-ethylhexanol. n-Butanol and 2-ethylhexanol are commercially significant organic compounds having use in a wide variety of applications and which are produced in large quantities per year. The most common method for making n-butanol is hydroformylation. Propylene reacts with syngas to produce an aldehyde, which is then hydrogenated to give the alcohol. 2-Ethylhexanol can be produced from an aldol condensation reaction followed by hydrogenation. This method converts butyraldehyde to 2-ethylhexanol.

[0004] Historically, lower olefins such as propylene have been obtained through the catalytic or steam cracking of fossil fuel feedstocks. Applicants have realized that the production of propylene and downstream products of propylene would benefit from the replacement of at least a part of the carbonaceous raw materials of fossil origin by renewable resources, such as carbonaceous matter derived from biomass. Of particular interest is the ethanol feedstock which is produced from renewable resources. Such renewably-sourced ethanol, also referred to as “bioethanol” or “hydrous fuel alcohol” can be prepared in large quantities from organic waste or biomass via fermentation. The different feedstocks for producing ethanol may be sucrose- containing feedstocks, e.g., sugarcane, starchy materials, e.g., corn, starch, wheat, cassava, lignocellulosic biomass, e.g., switchgrass, and / or agricultural waste. The purification or isolation of bioethanol is frequently carried out by complicated, multistage distillation.

[0005] It would be desirable to integrate renewably-sourced ethanol into existing processes designed for the conversion of fossil-derived ethylene or its intermediates. However, some of the impurities may interfere with the downstream processes which use bioethanol as feedstock and which generate chemical products, especially when some of the downstream steps are catalytic conversions.

[0006] If efforts are not made to remove at least some of these impurities, the yield of desired intermediate and final products and efficacy of the overall process may be diminished.

[0007] US 2008 / 0312485 discloses a method for continuously producing propylene by dehydrating ethanol obtained from biomass to obtain ethylene and reacting ethylene with n-butene in a metathesis reaction. The n-butene is made by dimerization of ethylene which is obtained from biomass-derived ethanol.

[0008] WO 2010 / 066830 discloses the transformation of bioethanol to ethylene. The bioethanol is produced by fermentation of carbohydrates or from synthesis gas made by gasification of biomass. The ethylene is subsequently dimerized or oligomerized to, e.g., 1 -butene and / or 1-hexene. The dimeric or oligomeric alphaolefins are transformed into internal olefins that are subsequently subjected to metathesis with ethylene.

[0009] WO 2011 / 085223 discloses an integrated process to prepare renewable hydrocarbons. The process includes dehydrating renewable isobutanol to form a mixture of linear butenes and isobutene and dehydrating renewable ethanol to ethylene. Subsequently the butene mixture and the ethylene are reacted to form one or more renewable C3-C16 olefins.

[0010] Downstream conversion of propylene to alcohols generally comprises a hydroformylation step of propylene in the presence of syngas ("synthesis gas"), i.e., a mixture of hydrogen and carbon monoxide, as well as a hydrogenation step in the presence of hydrogen.

[0011] Applicants have found that the carbon footprint of the production of propylene-derived products from renewably-sourced ethanol may be (further) reduced by providing syngas derived from a gasifier feedstock. In particular, gasifier feedstock may be obtained from materials which cannot be efficiently recycled and would otherwise go to waste.

[0012] In summary, the invention seeks to advise a reaction scheme with a low carbon footprint that provide a renewably-sourced alcohol selected from n-butanol and n-hexanol derived from light olefins, such as ethylene and propylene, which at least partially replaces the light olefins output from a steam cracker These light olefins are used as building blocks for producing a variety of chemicals of interest. It is desirable that the renewably-sourced light olefins can be blended or used interchangeably with a fossil -derived intermediate of the same chemical structure without necessitating adjustments in downstream processes. This includes that the starting olefins of all branches of the value chains, which historically have been served by the steam cracker output, can be supplied at the same time on a renewably-sourced basis. In this way, the greenhouse gases footprint and / or the carbon footprint for the production of a chemical of interest is at least reduced.

[0013] To this effect, the present invention relates to a process for the manufacture of an alcohol selected from n- butanol and 2-ethylhexanol, said process comprising the steps of: a) providing a gasifier feedstock; gasifying the gasifier feedstock by partial oxidation to form a raw product gas stream comprising a plurality of gases comprising methane, hydrogen and carbon monoxide; and recovering syngas from the raw product gas stream; b) obtaining propylene from renewably-sourced ethanol by subjecting renewably-sourced ethanol to dehydration to produce a renewably-sourced ethylene stream; and subjecting the renewably-sourced ethylene stream to an olefin-interconversion, to obtain propylene; the olefin-interconversion comprising (i) and (ii):

[0014] (i) ethylene dimerization to obtain n-butenes;

[0015] (ii) metathesis reaction between n-butenes obtained according to (i) and ethylene to obtain propylene; and c) mixing the propylene obtained from renewably-sourced ethanol with propylene not obtained from renewably-sourced ethanol to form mixed propylene; d) subjecting the mixed propylene to a sequence of chemical conversions to obtain n-butanol or 2-ethylhexanol, the sequence of chemical conversions comprising a) and (3), or a) and y): a) hydroformylation of the propylene in the presence of syngas obtained in a) to obtain n- butyraldehyde; P) hydrogenation of the n-butyraldehyde obtained in a) in the presence of hydrogen to produce n- butanol; y) condensation of the n-butyraldehyde obtained in a) to produce 2-ethyl-3-hydroxyhexanal, and subjecting the 2-ethyl-3-hydroxyhexanal to a hydrogenation reaction in the presence of hydrogen to produce 2-ethylhexanol.

[0016] The key advantage of the process according to the present invention is that it can be easily integrated into an existing production site in which one or more chemicals of interest are manufactured based on a fossil feedstock, in particular naphtha. This means that fossil-based propylene can be partially substituted by renewably-sourced propylene. Likewise, syngas obtained from the gasification of gasifier feedstock is used without the need of adaptions to the existing production. Hereby, one obtains n-butanol or 2-ethylhexanol, the carbon atoms of which are partially based on a renewable-sourced carbon (so-called "green” carbon).

[0017] Further benefits occur from the reduction of carbon dioxide emissions. The chemical conversions involved in a reaction route leading to an individual chemical of interest are usually less than 100% selective. The yield losses manifest themselves in the generation of by-products that vary depending on the type of reaction involved. The various species present may undergo a host of side reactions, which generate color forming species, oligomers, and various decomposition products or the like. These are generally removed during work-up, e.g., by distillation, yielding light boiler and / or high boiler fractions in addition to the desired product. The light boiler or high boiler fractions are conventionally used for their calorific value, i.e. combusted as fuel, or exploited as hydrocarbon source, e.g. as steam cracker feed. It should be appreciated that partial replacement of fossil propylene by its renewably-sourced counterpart at the beginning of the processing chain reduces the emission of fossil-based carbon dioxide resulting from the combustion of downstream side-products.

[0018] The expressions "renewable” or "renewably-sourced'' in relation to a chemical compound are used synonymously and mean a chemical compound comprising a quantity of renewable carbon, i.e., having a reduced or no carbon content of fossil origin. Renewable carbon entails all carbon sources that avoid or substitute the use of any additional fossil carbon from the geosphere. Renewable carbon can come from the biosphere, atmosphere or technosphere - but not from the geosphere. Thus, the expression “renewable” or “renewably-sourced” includes, in particular, biomass-derived chemical compounds. It also includes compounds derived from waste such as polymer residues, or from waste streams of chemical production processes.

[0019] Further, the process of the invention envisages producing syngas from gasifier feedstock and allows for obtaining downstream products of propylene having a low product carbon footprint.

[0020] Detailed Description of the Invention

[0021] Gasification

[0022] Reforming processes are often used to make synthesis gas (i.e., a gas mixture having predominant quantities of CO and H2) from natural gas or relatively low-boiling hydrocarbons. The present invention involves gasification-based processes for the conversion of various feedstocks into synthesis gas. In particular, according to step a), the process comprises providing a gasifier feedstock; gasifying the gasifier feedstock by partial oxidation to form a raw product gas stream comprising a plurality of gases comprising methane, hydrogen and carbon monoxide; and recovering syngas from the raw product gas stream.

[0023] Gasification is a process by which either a solid, gaseous or liquid gasifier feedstock is reacted with an oxidant such as air, oxygen, and / or steam. Sufficient energy is provided to produce a raw product gas stream that may be laced with volatile and condensable organic compounds, e.g., tars. The gasifier feedstock is gasified by partial oxidation. Partial oxidation is intended to produce as much CO as possible, with as little CO2 as possible. A fraction of the feedstock may be burned in the process to provide heat and pressure. Carbon dioxide from the combustion may be co-mingled with the product gas.

[0024] Generally, gasification of the gasifier feedstock by partial oxidation forms a raw product gas stream comprising a plurality of gases comprising methane, hydrogen and carbon monoxide. This raw product gas stream is treated to recover syngas. “Synthesis gas” or “syngas” refers to a gaseous mixture that is rich in CO and H2.

[0025] Preferably the gasifier feedstock is at least partially renewably-sourced. Typically at least 10 wt.-%, preferably at least 20 wt.-%, more preferably at least 30 wt.-%, even more preferably at least 40 wt.-%, particularly preferably at least 50 wt.-%, very particularly preferably at least 60 wt.-%, or even at least 70 wt.- %, at least 80 wt.-%, at least 90 wt.-%, or at least 95 wt.-% of the gasifier feedstock, are renewably-sourced, the weight percentages being based on the gasifier feedstock. Any desired weight percentage can be easily adjusted by combining fossil feedstocks such as coal, oil, vacuum residues, or natural gas with a renewably- sourced feedstock. Preferred renewably-sourced feedstocks are specified below.

[0026] In a preferred embodiment, the gasifier feedstock exhibits a heating value in the range of 15,000 to 45,000 J / g, preferably 20,000 to 40,000 J / g, more preferably 22,000 to 39,000 J / g, even more preferably 25,000 to 39,0000 J / g, most preferably 30,000 to 39,000 J / g, the heating value being measured in accordance with DIN 51900.

[0027] Preferably, the sum of the amounts of carbon (C), hydrogen (H), oxygen (O), sulfur (S) and nitrogen (N) in the gasifier feedstock is in the range of from 70 to 100 wt.-%, more preferably of from 80 to 99.9 wt.-%, more preferably of from 90 to 99.5 wt.-%, more preferably of from 95 to 99.5 wt.-%, based on the weight of the gasifier feedstock.

[0028] The gasifier feedstock may comprise a liquid gasifier feedstock and / or a solid gasifier feedstock. The gasifier feedstock may also comprise a gaseous feedstock, such as bio-natural gas

[0029] In a liquid gasifier feedstock, the amount of carbon (C), hydrogen (H), oxygen (O), sulfur (S) and nitrogen (N) is preferably:

[0030] C: 60 to 97 wt.-%, more preferably 70 to 94 wt.-%;

[0031] H: 1 to 15 wt.-%, more preferably 2 to 10 wt.-%;

[0032] O: 0 to 25 wt.-% more preferably 2 to 20 wt.-%;

[0033] S: 0 to 5 wt.-%, more preferably 0.005 to 4 wt.-%; N: 0 to 5 wt.-%, more preferably 0.005 to 4 wt.-%, each based on the weight of the liquid gasifier feedstock.

[0034] The liquid gasifier feedstock may comprise one or more of a pyrolysis oil, which may be derived from end of life tires, solid biomass, municipal solid waste (MSW), and / or industrial waste; and a waste oil, a used oil, a bio-based oil, and / or a bio-based fat

[0035] In a preferred embodiment, the liquid gasifier feedstock comprises one or more of a pyrolysis oil derived from solid biomass, a bio-based oil and a bio-based fat.

[0036] The term “solid biomass" comprises wood, wood pellets, wood chips, straw, lignocellulosic biomass, energy crops, algae, and mixtures thereof. Preferably, solid biomass is selected from wood, wood pellets, wood chips, straw, and lignocellulosic biomass. More preferably, solid biomass is selected from wood, wood pellets, wood chips and straw.

[0037] Preferably, the pyrolysis oil exhibits one or more of the following parameters: a final boiling point in the range of from 190 to 630 °C, as measured in accordance with ASTM D 86; a viscosity in the range of from 1 to 100 mPa s, as measured at 40 °C in accordance with DIN 53019; an ash content in the range of from 30 to 17000 mg / kg, as measured in accordance with ISO 6245.

[0038] In a solid gasifier feedstock, the amount of carbon (C), hydrogen (H), oxygen (O), sulfur (S) and nitrogen (N) preferably:

[0039] C: 40 to 97 wt.-%, more preferably 60 to 97 wt.-%;

[0040] H: 1 to 15 wt.-%, more preferably 1.2 to 10 wt.-%;

[0041] O: 0 to 25 wt.-% more preferably 1 .2 to 20 wt.-%;

[0042] S: 0 to 5 wt.-%, more preferably 0.005 to 4 wt.-%;

[0043] N: 0 to 5 wt.-%, more preferably 0.005 to 4 wt.-%, each based on the weight of the solid gasifier feedstock.

[0044] The solid gasifier feedstock may comprise one or more of solid biomass, municipal solid waste (MSW), refuse-derived fuel (RDF), shredder residues such as car shredder residues (ASR), textiles, plastic waste and packaging waste.

[0045] In a preferred embodiment, the solid gasifier feedstock comprises solid biomass.

[0046] Optionally, the gasifier feedstock is pre-treated before being subjected to gasification in a gasifier. A suitable pre-treatment method or combination of pre-treatment methods in a pre-treatment unit should provide a sufficiently homogeneous carbon-based feedstock to the gasification reaction and likewise enable the continuous production of syngas by gasification of a feedstock. A pre-treatment method or a combination of more than one pre-treatment methods in a pre-treatment unit preferably results in a homogenization of the physical and / or chemical properties of the gasifier feedstock and / or the requirement(s) for a specific type of gasifier.

[0047] The pre-treatment method for the gasifier feedstock is preferably selected from the group comprising drying, comminution, classification, sorting, agglomeration, thermochemical methods, and biological methods.

[0048] Suitable gasifiers comprise counter-current fixed bed reactors, co-current-fixed bed reactors, bubbling fluidized bed reactors, circulation fluidized bed reactors, and downdraft or updraft entrained flow reactors. The selection of size and reactor type depends on several parameters, including the composition of the gasifier feedstock, demand of products, moisture content and availability of the gasifier feedstock. Preferably, the gasifier is an “oxygen blown” gasifier, i.e., oxygen is preferably used as the oxidant.

[0049] The gasification reaction in a gasifier is typically carried out at a temperature of greater than 400 °C, such as greater than 700 °C, in the presence of a sub-stoichiometric amount of an oxidant such as oxygen, air, steam, supercritical water, CO2, or a mixture of the aforementioned. Preferably, the gasification is carried out at a temperature in the range of greater than 700 to 1500 °C, more preferably, 850 to 1400 °C, even more preferably 1100 to 1500 °C.

[0050] Typically, the gasification is conducted at an absolute pressure of greater than 1 bar. Preferably, it is conducted at an absolute pressure in the range from 2 to 80 bar, more preferably 2 to 50 bar.

[0051] In one embodiment, the gasifier feedstock comprises a liquid gasifier feedstock, and gasifying the liquid gasifier feedstock comprises subjecting the liquid gasifier feedstock to partial oxidation in an entrained flow reactor to obtain the raw product gas stream. The gasifier feedstock may further comprise a gaseous feedstock, such as bio-natural gas.

[0052] In this embodiment, the entrained flow reactor is preferably operated at a temperature above 400 °C, such as 1000 to 2000 °C, more preferably 1250 to 1500 °C, and at a pressure of 1 bar(abs) or more, such as 5 to 200 bar(abs), preferably 10 to 100 bar(abs), more preferably 11 to 50 bar(abs). In particular, gasifying the organic matter feedstock comprises introducing the liquid gasifier feedstock, oxygen, and optionally steam, into the entrained flow reactor being operated at a temperature above 400 °C and at a pressure of 1 bar(abs) or more, and bringing the liquid gasifier feedstock into contact with oxygen, and optionally the steam, in said reactor The atomized liquid gasifier feedstock is gasified with oxygen in co-current flow.

[0053] In one embodiment, the gasifier feedstock comprises a solid gasifier feedstock, and gasifying the solid gasifier feedstock comprises subjecting the solid gasifier feedstock to partial oxidation in a fluidized bed reactor so as to obtain an intermediate gasification product, and directing the intermediate gasification product to gasification in an entrained flow reactor to obtain the raw product gas stream. The gasifier feedstock may further comprise a gaseous feedstock, such as bio-natural gas.

[0054] In this embodiment, the fluidized bed reactor is preferably operated at a temperature in the range of from 350 to 1000 °C, such as 600 to 1000 °C, and at a pressure in the range of from 1 to 200 bar(abs), such as 1 to 10 bar(abs), preferably 1.5 to 8 bar(abs), more preferably 2 to 5 bar(abs);and the entrained flow reactor is operated at a temperature above 400 °C and at a pressure of 1 bar(abs) or more, preferably at a temperature in the range of from 1000 to 1700 °C, such as 1100 to 1450 °C, and at a pressure in the range of from 1 to 200 bar(abs), such as 1 to 50 bar(abs), preferably 3 to 45 bar(abs), more preferably 8 to 40 bar(abs).

[0055] Moreover in this embodiment, the process may additionally comprise directing a liquid gasifier feedstock into the entrained flow reactor.

[0056] In another embodiment, gasifying the gasifier feedstock comprises partial oxidation in a plasma reactor. In particular, the gasifier feedstock may be contacted in a plasma reactor with a stream comprising one or more of O2, CO2 and steam, and subjected to partial oxidation in the presence of a plasma. The partial oxidation is preferably conducted at a temperature in the range from 1000 to 2000 °C, such as 1100 to 1450 °C, and a pressure of at least 1 bar(abs), such as 1.0 to 6 bar(abs), preferably 1.5 to 6 bar(abs), more preferably 1 .75 to 3 bar(abs).

[0057] Preferably, the plasma is obtained by a process comprising generating plasma at at least one plasma torch comprised in the plasma reactor by applying an electric voltage in the range of from 0.5 to 70 MW, more preferably in the range of from 5 to 70 MW, more preferably in the range of from 20 to 60 MW.

[0058] Plasma processes can be performed as disclosed in Ullmann’s Encyclopedia of Industrial Chemistry, Waste, 7. Thermal treatment, Ch. 3.2.6, Ed. 2021 Wiley-VCH Verlag GmbH & Co. KGaA.

[0059] Oxygen is the most common oxidant used for gasification because of its easy availability and low cost The hydrogen to carbon monoxide molar ratio ("the molar ratio of H2 to CO") depends on the composition of the gasifier feedstock and the amount of steam used in the gasification. The molar ratio of H2 to CO as required for the hydroformylation can for instance be adjusted by choosing an appropriate amount of steam in the gasification.

[0060] The conversion of a gasifier feedstock in the gasifier produces a raw product gas stream which consists primarily of H2, CO, CO2, methane, other hydrocarbons, and impurities. Said syngas has a dedicated molar ratio of H2 to CO when leaving the gasifier which ranges from about 0.1 : 1 to about 3 : 1 and depends on the type of solid and / or liquid feedstocks used, the oxidant and other reaction conditions applied such as temperature and / or residence time of the reactants in the gasifier.

[0061] When steam acts as oxidant, the syngas has a higher molar ratio of H2 to CO than when air is used as oxidant. For example, a typical molar ratio of “air to combined feedstock” ranges from 0.3 to less than 1.

[0062] Another possibility to adjust the molar ratio of H2 to CO is to separate CO from the syngas, as further discussed below

[0063] Purification of the Raw Product Gas Stream

[0064] Typical impurities in the raw product gas stream obtained from the gasification reaction comprise chlorides, sulfur-containing organic compounds such as sulfur dioxide, trace heavy metals (e.g., as respective salts) and particulate residues. Various chemical and / or physical methods for removal of such impurities from said raw syngas such as filtration, scrubbing, hydrotreatment and ab- / adsorption are known and can be chosen and adapted according to the type and respective concentration of the impurities in said raw syngas and the tolerance to such impurities in the successive process steps.

[0065] The gasification reaction usually results in further reaction products such as solid and / or highly viscous carbonaceous residues (e.g., char and / or tar) which can be further treated in separate steps not relevant for the systems and methods according to the present invention.

[0066] Bulk particulate impurities can be removed from the raw syngas by a cyclone and / or filters, fine particles, and chlorides by wet scrubbing, trace heavy metals, catalytic hydrolysis for converting sulfur-containing organic compounds to H2S and acid gas removal for extracting sulfur-containing gases such as H2S. Bulky and fine particles in the syngas may also be removed with a quench in a soot water washing unit.

[0067] In one embodiment, recovering syngas from the raw product gas stream comprises subjecting the raw product gas stream to a washing step to remove particulate solids and to a drying step to remove water.

[0068] In one embodiment, recovering syngas from the raw product gas stream further comprises sweetening the raw product gas stream (removing acid gases such as carbon dioxide and hydrogen sulfide) to form a sweetened product gas stream.

[0069] In one embodiment, the raw product gas stream is subjected to a first purification stage, whereby a purified product gas stream is obtained, the purified product gas stream comprising CO, H2 and CH4 and being depleted in CO2, H2O and solid particles, wherein the first purification stage comprises: a') subjecting the raw product gas stream obtained to a washing step in a washing unit, more preferably the washing unit being a column having a spray nozzle or an atomizer located at the top of the column, obtaining a washed product gas stream depleted in particulate solid compared to the raw product gas stream and comprising CO, H2, CO2, H2O, CH4, and optionally H2S; a") subjecting the washed product gas stream obtained according to a') to a drying step in a drying unit, obtaining a dried product gas stream depleted in H2O compared to the raw product gas stream and the washed product gas stream; and a'") optionally, subjecting the washed product gas stream obtained according to a") to an acid gas removal step in a CO2 / H2S adsorption unit, more preferably the CO2 / H2S adsorption unit is an absorption column, obtaining a sweetened product gas stream.

[0070] Preferably, a') comprises contacting the raw product gas stream with water in the washing unit. For example, the washing step can be performed as disclosed in WO 2023 / 161302A1. Preferably, the washed product gas stream has a temperature in the range of from 80 to 100°C.

[0071] Preferably, the drying unit used in a") is a water separation unit. Preferably, a") comprises cooling the washed product gas stream obtained according to a'), obtaining the dried product gas stream, a gaseous stream, separated from water condensate. For example, the drying step can be performed as disclosed in WO 2023 / 161302 A1. Preferably, the acid gas removal step according to a'") is amine scrubbing. Preferably, o'") comprises using an amine, such as monoethanolamine (MEA). Diethanolamine (DEA), methyl-diethanolamine (MDEA) or diglycolamine (DGA), in the CO2 / H2S adsorption unit, more preferably the absorption column. For example, the acid gas removal step can be performed as disclosed in WO 2023 / 161302 A1.

[0072] Preferably, o'") comprises subjecting the dried product gas stream obtained according to a") to an acid gas removal step in a CO2 / H2S adsorption unit, more preferably the CO2 / H2S adsorption unit is an absorption column, obtaining the sweetened product gas stream and a CO2-comprising stream.

[0073] Preferably, the CO2-comprising stream is recycled in the process of the present invention. Still more preferably, the CO2-comprising stream is recycled in p") as described below as at least a portion of the source of CO2.

[0074] As discussed above, the molar ratio of H2 to CO in the obtained syngas varies. In one embodiment, the molar ratio of H2 to CO may be adjusted by separating CO from the syngas. CO can be separated from the syngas in a syngas separation unit which is downstream of and fluidly connected to a syngas producing unit. CO can be separated from syngas by cryogenic separation methods, commonly referred to as a "cold box” which makes use of the different boiling points of CO and H2. H2 can be separated using Fh-selective membranes thorough which H2 permeates and is thereby separated from a syngas stream. It is also possible to fully separate CO and H2 via cryogenic separation. The resulting CO and H2 can be used to create a syngas having the desired ratio.

[0075] In one embodiment, recovering syngas from the raw product gas stream further comprises adjusting the hydrogen to carbon monoxide molar ratio in the raw product gas stream or the sweetened gas stream, respectively, to obtain an adjusted gas stream comprising CO and H2 at a molar ratio different from the molar ratio in the raw product gas stream or the sweetened gas stream, respectively; and, optionally, separating the adjusted gas stream to provide a CO-comprising gas stream, a F -comprising gas stream and a CHi-comprising gas stream.

[0076] More specifically, the molar ratio of H2 to CO in the purified product gas stream, i.e., in the washed product gas stream or the sweetened product gas stream, may be adjusted, obtaining a modified product gas stream comprising CO and H2and having a H2 to CO molar ratio differing from the H2 to CO molar ratio of the purified product gas stream.

[0077] Adjusting the hydrogen to carbon monoxide molar ratio may involve shift reactions such as a water gas shift reaction according to

[0078] CO + H2O CO2+ H2or a reverse water gas shift reaction according to

[0079] H2+ C02CO + H20. In one embodiment, the H2 to CO molar ratio in the purified product gas stream is adjusted by: p') passing and contacting water with the purified product gas stream obtained in the first purification stage into a reaction unit RU(1) and subjecting to a water gas shift reaction in RU(1), obtaining a CO-depleted product gas stream depleted in CO compared to the purified product gas stream and comprising CO, H2, CH4 and CO2; or p") passing and contacting CO2, optionally the CO2-comprising stream from o'"), with the purified product gas stream into a reaction unit RU(2) for a reverse water gas shift reaction, obtaining a CO-enriched product gas stream enriched in CO compared to the purified product gas stream and comprising CO, H2 and CH4; or

[0080] P'") adding H2 to the purified product gas stream, obtaining a Fh-enriched product gas stream enriched in H2 compared to the purified product gas stream and comprising CO, H2 and CH4.

[0081] The water gas shift reaction is preferably performed according to known processes in the art, such as for example those defined in Wei-Hsin Chen, et al., "Water gas shift reaction for hydrogen production and carbon dioxide capture”, Applied energy 258 (2020) 114078, https: / / doi.Org / 10.1016 / j.apenergy.2019.114078. The reverse water gas shift reaction is preferably performed according to known processes in the art such as for example those defined in E. Rezaei, S. Dzuryk "Techno-economic comparison of reverse water gas shift reaction to steam and dry methane reforming reactions for syngas production", Chemical Engineering Research and Design, Vol 144 (2019), S. 354-369, EP2175986, CN103183346 and US8946308.

[0082] Preferably, when p') is performed, the process further comprises passing CO-depleted product gas stream through an acid gas removal unit, obtaining a sweetened CO-depleted product gas stream depleted in CO2 compared to the CO-depleted product gas stream and comprising CO, H2 and CH4.

[0083] Preferably, in p'") at least a portion of the added H2 is renewably sourced H2.

[0084] In one embodiment, the process additionally comprises recovering hydrogen from the raw product gas stream or the syngas obtained therefrom, and directing the hydrogen at least partially to p) and / or y).

[0085] In one embodiment, the purified product gas stream or the modified product gas stream, i.e., the CO-depleted product gas stream, the CO-enriched product gas stream, the Ftenriched product gas stream or the sweetened CO-depleted product gas stream, are subjected to a second purification stage.

[0086] In particular, the purified product gas stream or the modified product gas stream are subjected to a second purification stage, comprising subjecting the purified product gas stream or the modified product gas stream to cryogenic separation, e.g., in a cold box, obtaining a CO-comprising gas stream, a hb-comprising gas stream and a methane-comprising gas stream.

[0087] The cryogenic separation can be performed by method known in the art, such as disclosed in Ullmann's Encyclopedia of Industrial Chemistry, Carbon Monoxide, Chapter 4.3.2, p.685-686. Renewably-Sourced Ethanol

[0088] Bioethanol is a preferred form of renewably-sourced ethanol, although the scope of the invention is not limited to the use of bioethanol.

[0089] In the present invention, bioethanol refers to the ethanol obtained from a biomass feedstock, such as plant or non-crop feedstock containing a carbon source that is convertible to ethanol, for example by microbial metabolism. Typical carbon source examples are starch, sugars like pentoses or hexoses, such as glucose, fructose, sucrose, xylose, arabinose, or degradation products of plants, hydrolysis products of cellulose or juice of sugar canes, beet and the like containing large amounts of the above components.

[0090] Biomass feedstock can originate from several sources. The term “biomass” includes wood, wood pellets, wood chips, straw, lignocellulosic biomass, energy crops, algae, biobased-oils, biobased-fats, and mixtures thereof. Bioethanol production may be based on food crop feedstocks such as corn and sugar cane, sugarcane bagasse, cassava (first generation biofeedstock)

[0091] Another source of biomass feedstock is lignocellulosic materials from agricultural crops (second-generation biofeedstock). Potential feedstocks include agricultural residue by-products such as rice, straw (such as wheat, oat and barley straw), rice husk, and corn stover. Biomass feedstock may also be waste material from the forest products industry (wood waste) and saw dust or produced on purpose as an ethanol crop. Switchgrass and napier grass may be used as on-purpose crops for conversion to ethanol.

[0092] The first-generation bioethanol is produced in four basic steps:

[0093] (1) Enzymatic saccharification or hydrolysis of starch into sugars

[0094] (2) Microbial fermentation of sugars

[0095] (3) Purification by distillation to give hydrous ethanol

[0096] (4) Dehydration (water removal) to produce anhydrous ethanol

[0097] Second-generation feedstocks are considered as renewable and sustainable carbon source. Pretreatment of this feedstock is an essential prerequisite before it is subjected to enzymatic hydrolysis, fermentation, distillation, and dehydration. Pretreatment involves milling and exposure to acid and heat to reduce the size of the plant fibers and hydrolyze a portion of the material to yield fermentable sugars. Saccharification utilizes enzymes to hydrolyze another portion to sugar. Finally, fermentation by bioengineered microorganisms converts the various sugars (pentoses and hexoses) to ethanol. The production of bioethanol is well-known and carried out on an industrial large scale.

[0098] Renewably-sourced ethanol can also be obtained from carbon-containing waste materials like waste products from the chemical industry, garbage and sewage sludge. The production of ethanol from waste materials can be done by gasification to syngas and catalytic conversion thereof the ethanol, see for example Recent Advances in Thermo-Chemical Conversion of Biomass, 2015, Pages 213-250, https: / / doi.org / 10.1016 / B978-0-444-63289-0.00008-9, and Nat Commun 11, 827 (2020), https: / / doi.Org / 10.1038 / S41467-020-14672-8. Dehydration of Renewably-Sourced Ethanol

[0099] The process involves the dehydration of renewably-sourced ethanol. The production of ethylene by catalytic dehydration of ethanol is a well-known process. The reaction is commonly carried out at 300 to 400 °C and moderate pressure in the presence of a catalyst. Catalytic effects are reviewed in Ind & Eng Chem Research, 52, 28, 9505-9514 (2013), Materials 6, 101-115 (2013) and ACS Omega, 2, 4287-4296 (2017). Examples for catalysts are activated alumina or silica, phosphoric acid impregnated on coke, heteropoly acids (HPA salts), silica-alumina, molecular sieves such as zeoliths of the ZSM-5 type or SAPO-11 type, other zeolites or modified zeolites of various molecular structures with zeoliths and HPA salts being preferred.

[0100] Ethanol dehydration is, for example described in WO 2009 / 098268, WO 2010 / 066830, WO 2009 / 070858 and the prior art discussed therein, WO 2011 / 085223 and the prior art discussed therein, US 4,234,752, US 4,396,789, US 4,529,827 and WO 2004 / 078336.

[0101] The ethanol dehydration reaction is in general carried out in the vapor phase in contact with a heterogeneous catalyst bed using either fixed bed or fluidized bed reactors. For fixed bed reactors, the operation can be either isothermal (with external heating system) or adiabatic (in the presence of a heat carrying fluid). The feedstock is vaporized and heated to the desired reaction temperature; the temperature drops as the reaction proceeds in the reactor. Multiple reactor beds are usually used in series to maintain the temperature drop in each bed to a manageable range. The cooled effluent from each bed is further heated to bring it to the desired inlet temperature of the subsequent beds. Moreover, a portion of the water is recirculated along with fresh and unreacted ethanol. The presence of water helps in moderating the temperature decrease in each bed.

[0102] Prior to dehydration, the renewably-sourced ethanol feedstock may be sent to a pretreatment section to remove mineral contaminants, which would otherwise be detrimental to the downstream catalytic reaction. The pretreatment may involve contacting the renewably-sourced ethanol feedstock with cation and / or anion exchange resins. After a certain period of operation, the resins may be regenerated by passing a regenerant solution through the resin bed(s) to restore their ion exchange capacity. Two sets of beds are preferably operated in parallel to maintain continuous operation. One set of resin beds is suitably regenerated while the other set is being used for pretreatment.

[0103] In the isothermal design, the catalyst is placed inside the tubes of multitubular fixed-bed reactors which arranged vertically and surrounded by a shell (tube and shell design). A heat transfer medium, such as molten salts or oil, is circulated inside the shell to provide the required heat. Baffles may be provided on the shell side to facilitate heat transfer. The cooled heating medium is heated externally and is recirculated. The temperature drop on the process side can be reduced as compared to the adiabatic reactor. A better control on the temperature results in increased selectivity for the ethylene formation and reduction in the amount of undesirable by-products. The temperature is maintained at approximately constant levels within the range of 300° to 350°C. Ethanol conversion is between 98 and 99%, and the selectivity to ethylene is between 94 and 97 mol%. Because of the rate of coke deposition, the catalyst must be regenerated frequently. Depending on the type of catalyst used, the cycle life is between 3 weeks and 4 months, followed by regeneration, for example for 3 days. In the adiabatic design, the endothermic heat of reaction is supplied by a preheated inert diluent such as steam. Three fixed-bed reactors may typically be used, with intermediate furnaces to reheat the ethanol / steam mixed feed stream to each reactor. Feeding steam with ethanol results in less coke formation, longer catalyst activity, and higher yields.

[0104] A further process is a fluidized-bed process. The fluidized-bed system offers excellent temperature control in the reactor, thereby minimizing by-product formation. The heat distribution rate of the fluidized bed operation approaches isothermal conditions. The endothermic heat of reaction is supplied by the hot recycled silica-alumina catalyst returning from the catalyst regenerator. Thus, external heating of the reactor is not necessary.

[0105] After dehydration, the reaction mixture is subjected to a separation step. The general separation scheme consists of quickly cooling the reaction gas, for example in a water quench tower, which separates most of the by-product water and the unreacted ethanol from ethylene and other light components which, for example exit from the top of the quench tower. In one type of separation scheme, the water-washed ethylene stream is immediately caustic-washed, for example in a column, to remove traces of CO2. The gaseous stream may enter a compressor directly or pass to a surge gas holder first and then to a gas compressor. After compression, the gas is cooled with refrigeration and then passed through an adsorber with, for example activated carbon, to remove traces of heavy components, (e.g., C4s), if they are present. The adsorber is followed by a desiccant drying and dust filtering step before the ethylene product leaves the plant. This separation scheme produces 99%+ purity ethylene. If desired, the ethylene is further purified by caustic washing and desiccant-drying, and fractionated in a low-temperature column to obtain the final product.

[0106] Several commercial processes are currently in operation, developed by Braskem, Chematur, British Petroleum (BP), and Axens together with Total and IFPEN. The processes differ, e.g., in their process conditions, catalysts and adopted heat integration scheme. The process by BP (now Technip) is called Hummingbird. In this process, a heteropoly acid is used as catalyst, and the reactor operates at 160 to 270 °C and 1 to 45 bar The unreacted ethanol in recirculated to the reactor. The process developed by Axens is called Atol. Two fixed bed adiabatic reactors, operating at 400 to 500 °C, are used. Chematur’s process operates with four adiabatic tubular reactors. Syndol catalysts, with the main components of AI2O3- MgO / SIC>2, are employed in this process that was developed by American Halcon Scientific Design, Inc. in the 1980s. In the Braskem process, the adiabatic reactor feed is diluted with steam to a large extent. In such a process, the reactor operates at 180 to 600 °C, preferably 300 to 500 °C, and at 1.9 to 19.6 bar. An alumina or silica-alumina catalyst is used. The Braskem process is described in more detail in US 4,232,179. A process control in accordance with the Braskem process is particularly preferred

[0107] Dimerization of Ethylene

[0108] The process of the invention involves an ethylene-dimerization to obtain n-butenes in accordance with step b)-(i). Any known method can be used for ethylene dimerization to produce n-butenes. A review on dimerization and oligomerization chemistry and technology is given in Catalysis Today, vol. 14(no. 1), April 10, 1992. Expediently, step b)-(i) comprises: contacting the renewably-sourced ethylene stream with a dimerization catalyst in a dimerization zone; operating said dimerization zone at conditions effective to produce an effluent consisting essentially of n-butenes, a stream consisting essentially of heavier olefins, and optionally an unconverted ethylene stream; and fractionating the effluent to recover a stream consisting essentially of n-butenes, a stream consisting essentially of heavier olefins, and an optional ethylene stream.

[0109] The dimerization catalyst may be homogeneous or heterogeneous. Typical dimerization catalysts are titanium or nickel compounds activated with alkyl aluminum compounds. In general, the Ti(IV) valency is stabilized by selecting the appropriate ligands, alkyl aluminum compound, the solvent polarity and the Al / Ti ratio. Nickel compounds that can catalyze the selective production of butenes are typically based on cationic nickel salts stabilized with phosphine and activated with alkyl aluminum compounds.

[0110] In one embodiment, the oligomerization of ethylene is implemented in the presence of a catalytic system in the liquid phase comprising a nickel compound and an aluminum compound. Such catalytic systems are described in the documents FR 2 443 877 and FR 2794 038. The Dimersol E™ process is based on this technology and leads to the industrial production of olefins.

[0111] Thus, in one embodiment, the oligomerization of ethylene is implemented in the presence of a catalytic system comprising: i) at least one bivalent nickel compound, ii) at least one hydrocarbyl aluminum dihalide of formula AIRX2, in which R is a hydrocarbyl radical comprising 1 to 12 carbon atoms, such as alkyl, aryl, aralkyl, alkaryl or cycloalkyl, X is a chlorine or bromine atom, and iii) optionally a Bronsted organic acid.

[0112] As the bivalent nickel compound, nickel carboxylates of general formula (R1COO)2Ni are preferably used, where R1is an optionally substituted hydrocarbyl radical, for example alkyl, cycloalkyl, alkenyl, aryl, aralkyl, or alkaryl, containing up to 20 carbon atoms, preferably a hydrocarbyl radical of 5 to 20 carbon atoms, preferably 6 to 18 carbon atoms. Suitable bivalent nickel compounds include: chloride, bromide, carboxylates such as octoate, 2-ethylhexanoate, decanoate, oleate, salicylate, hydroxydecanoate, stearate, phenates, naphthenates, and acetyl acetonates. Nickel 2-ethylhexanoate is preferably used.

[0113] The hydrocarbyl aluminum dihalide compound corresponds to the formula AIRX2, in which R is a hydrocarbyl radical comprising 1 to 12 carbon atoms, such as alkyl, aryl, aralkyl, alkaryl or cycloalkyl, and X is a chlorine or bromine atom. As examples of such compounds, it is possible to mention ethylaluminum sesquichloride, dichloroethyl aluminum, dichloroisobutyl aluminum, chlorodiethyl aluminum or mixtures thereof

[0114] According to a preferred method, a Bronsted organic acid is used. The Bronsted acid compound corresponds to the formula HY, where Y is an organic anion, for example carboxylic, sulfonic or phenolic. Halocarboxylic acids of formula R2COOH in which R2is a halogenated alkyl radical are preferred, in particular those that contain at least one alpha-halogen atom of the group — COOH with 2 to 10 carbon atoms in all. Preferably, a haloacetic acid of formula CXPH3-P— COOH is used, in which X is fluorine, chlorine, bromine or iodine, with p being an integer from 1 to 3. By way of example, it is possible to cite the trifluoroacetic, difluoroacetic, fluoroacetic, trichloroacetic, dichloroacetic, and chloroacetic acids. It is also possible to use arylsulfonic, alkylsulfonic, and fluoroalkylsulfonic acids, and picric acid and nitroacetic acid. Trifluoroacetic acid is preferably used.

[0115] The three components of the catalytic formula can be mixed in any order. However, it is preferable first to mix the nickel compound with the Brpnsted organic acid, and then next to introduce the aluminum compound. The molar ratio of the hydrocarbyl aluminum dihalide to the nickel compound, expressed by the Al / Ni ratio, is 2 / 1 to 50 / 1, and preferably 2 / 1 to 20 / 1. The molar ratio of the Brpnsted acid to the nickel compound is 0.25 / 1 to 10 / 1 , and preferably 0.25 / 1 to 5 / 1.

[0116] According to a preferred method, the hydrocarbyl aluminum dihalide can be enriched with an aluminum trihalide, the mixture of the two compounds then corresponding to the formula AIRnX3-n, in which R is a hydrocarbyl radical comprising 1 to 12 carbon atoms, such as alkyl, aryl, aralkyl, alkaryl or cycloalkyl, X is a chlorine or bromine atom, and n is a number between 0 and 1. Suitable mixtures include: dichloroethyl aluminum enriched with aluminum chloride, the mixture having a formula AIEto.gCh.i; dichloroisobutyl aluminum enriched with aluminum chloride, the mixture having a formula AliBuo.gClzi; and dibromoethyl aluminum enriched with aluminum bromide, the mixture having a formula AIEto.9Br2.1-

[0117] The reaction for oligomerization of ethylene can be implemented at a temperature of -20 to 80 °C, preferably 40 to 60 °C, under pressure conditions such that the reagents are kept at least for the most part in the liquid phase or in the condensed phase. The pressure is generally between 0.5 and 5 MPa, preferably between 0.5 MPa and 3.5 MPa. The time of contact is generally between 0.5 and 20 hours, preferably between 1 and 15 hours.

[0118] The oligomerization stage can be implemented in a reactor with one or more reaction stages in a series, with the ethylene feedstock and / or the catalytic composition that is preferably pre-conditioned in advance being introduced continuously, either in the first stage, or in the first stage and any other one of the stages. At the outlet of the reactor, the catalyst can be deactivated, for example by injection of ammonia and / or an aqueous solution of soda and / or an aqueous solution of sulfuric acid. The unconverted olefins and alkanes that are optionally present in the feedstock are then separated from the oligomers by a separation stage, for example by distillation or washing cycles by means of caustic soda and / or water.

[0119] The conversion per pass is generally 85 to 98%. The selectivity of n-butenes that are formed is generally between 50 and 80%. The n-butenes consist of butene-2 (cis- and trans-) and butene-1.

[0120] The effluent generally contains less than 0.2% by weight of isobutene, or even less than 0.1 % by weight of isobutene. Separation of a Stream Rich in n-Butenes

[0121] The effluent that is obtained by dimerization of ethylene is suitably subjected to a separation stage in such a way as to obtain an n-butene-enriched fraction.

[0122] The separation can be carried out by evaporation, distillation, extractive distillation, extraction by solvent or else by a combination of these techniques. These processes are known by one skilled in the art. Preferably, a separation of the effluent that is obtained by oligomerization of ethylene is carried out by distillation.

[0123] Preferably, the effluent of the oligomerization is sent into a distillation column system comprising one or more columns that makes it possible to separate, on the one hand, n-butenes from ethylene, which can be returned to the oligomerization reactor, and heavier olefins with 5 carbon atoms and more.

[0124] Such heavier olefins can be used as a gasifier feedstock according to step a) of the process according to the present invention. The heavier olefins can constitute the entire gasifier feedstock or can form a part thereof. Preferably, the heavier olefins form part of the gasifier feedstock, meaning they are mixed with suitable other materials to form the gasifier feedstock.

[0125] This constitutes a synergistic advantage of the combination of the gasification as per step a) and the dimerization according to step b). Step b) inevitably results in the formation of the heavier olefins stream as specified above. This stream is typically considered a waste stream and subjected to incineration, resulting in a loss of renewable carbon atoms for synthesis. By the combination of steps a) and b), the renewable carbon can be recycled to the synthesis of n-butanol and 2-ethy I hexanol.

[0126] Alternatively, the heavier olefins may be subjected to hydrogenation so as to obtain renewably-sourced naphtha. "Renewably-sourced naphtha" shall mean naphtha produced from renewable sources. It is a hydrocarbon composition, consisting of mainly paraffins. The molecular weight of this renewably-sourced naphtha may range from hydrocarbons having 5 to 8 carbon atoms. Renewably-sourced naphtha can be used as a feedstock in steam-cracking to produce renewably-sourced light olefins, dienes and aromatics.

[0127] Hence, in an embodiment, step b)-(i) comprises:

[0128] - contacting the renewably-sourced ethylene stream with a dimerization catalyst in a dimerization zone;

[0129] - operating said dimerization zone at conditions effective to produce an effluent consisting essentially of n-butenes, a stream consisting essentially of heavier olefins, and optionally an unconverted ethylene stream;

[0130] - fractionating the effluent to recover a stream consisting essentially of n-butenes, a stream consisting essentially of heavier olefins, and an optional ethylene stream; and

[0131] - optionally directing the stream consisting essentially of heavier olefins to step a) as a gasifier feedstock or subjecting the stream consisting essentially of heavier olefins to hydrogenation so as to obtain renewably-sourced naphtha. Metathesis of Ethylene with n-Butenes

[0132] Ethylene is able to undergo metathesis with n-butenes to produce propylene. Step b)-(ii) comprises a metathesis reaction between n-butenes obtained according to step (i) and ethylene to obtain propylene. The n-butenes obtained according during ethylene dimerization (i) are a mixed stream including 1 -butene and 2-butenes. Essentially only the 2-butenes react in a metathesis reaction, while 1 -butene is essentially inert.

[0133] In one embodiment, 1-butene is removed from the mixed stream of 1-butene and 2-butenes and directed to a use elsewhere in the plant. Thus, in one embodiment, step b)-(ii) comprises removal of 1-butene from the mixed stream to obtain a stream rich in 2-butenes, and subjecting the stream rich in 2-butenes to the metathesis reaction. A stream rich in 2-butenes may comprise at least 90 wt.-% of 2-butenes, based on the total amount of n-butenes.

[0134] Alternatively, 1-butene may be converted to 2-butene by double bond isomerization. Double bond isomerization is an equilibrium-limited reaction It is thus advantageous to subject the mixed stream of n-butenes to metathesis so as to react 2-butene with ethylene prior to double bond isomerization of 1-butene. Hence, in one embodiment the n-butenes are a mixed stream including 1 -butene and 2-butenes, and b)-(ii) comprises b)-(iia) subjecting the mixed stream to the metathesis reaction to obtain propylene and unreacted 1-butene; b)-(iib) subjecting the unreacted 1 -butene to double bond isomerization to obtain 2-butenes; and b)-(iic) recycling the 2-butenes obtained in step b)-(iib) to step b)-(iia).

[0135] In another embodiment, it is possible to convert 1-butene to 2-butene simultaneously with the metathesis reaction. For this purpose, a metathesis catalyst and an isomerization catalyst may be physically mixed or provided as distinct layers to allow both reactions to proceed simultaneously. Thus, in one embodiment, step b)-(ii) is carried out by passing the mixed stream through a metathesis / isomerization zone comprising both a metathesis catalyst and an isomerization catalyst. As 2-butene is consumed due to the metathesis reaction over the metathesis catalyst, it is thus replenished by isomerization of 1 -butene to 2-butene over the isomerization catalyst.

[0136] The reaction is carried out in the presence of a metathesis catalyst on the basis of a metal which is selected from tungsten, molybdenum, rhenium, niobium, tantalum, vanadium, ruthenium, rhodium, iridium, osmium and nickel and the like. Tungsten, molybdenum and rhenium are preferred and tungsten is particularly preferred. Typically, tungsten catalysts are supported on silica, molybdenum and rhenium are supported on alumina based carriers. Especially preferred metathesis catalysts are WOs-based catalysts, for example silica-supported WOg in the form of granules.

[0137] Suitable isomerization catalysts include magnesium-based catalysts such as MgO-based catalysts, for example tableted MgO.

[0138] Metathesis is carried out under conditions effective to produce an effluent comprising propylene, unconverted ethylene, and optionally 1-butene Unconverted ethylene and / or unconverted n-butenes may be recycled and combined with fresh ethylene and n-butenes to provided the metathesis feedstock.

[0139] The reaction may be conducted at 340 - 375°C, 25-40 bar, a weight hourly space velocity (WHSV) of 7.5-30 hr1, and an ethylene to 2-butene molar ratio of 3:1 to 10:1.

[0140] The reactor effluent may be sent to a deethenizer to remove C2 and lighter material. The bottoms from the deethenizer are sent to the depropenizer. High-purity, polymer-grade propylene (> 99.9% molar purity) is recovered from the depropenizer overhead. The lighter material from the deethenizer and heavier C4+ material from the depropenizer are partly recycled to the reactors. Purge streams are provided for the lighter and heavier material to prevent buildup of inerts.

[0141] It should be noted that propane is not produced during the metathesis reaction. Consequently, polymer- grade propylene can be produced from the process, without the need for an expensive propylene-propane superfractionator.

[0142] Commercial processes for producing polymer-grade propylene by metathesis from ethylene and butenes feedstock are available from CB&I / Lummus (tradnemame OCT™) and from LyondellBasell.

[0143] Moreover, in an embodiment, the process comprises: blending the renewably-sourced ethylene with complementary ethylene prior to step b), the complementary ethylene not being obtained from renewably-sourced ethanol in accordance with step a); and / or blending the renewably-sourced n-butenes with complementary n-butenes prior to step b)-(ii), the complementary n-butenes not being obtained from renewably-sourced ethanol in accordance with steps a) and b)-(i).

[0144] Examples for complementary ethylenes are ethylenes obtained by steam cracking of fossil based feeds, like naphtha, natural gas or crude oil. Examples for complementary propylenes are propylenes obtained by steam cracking of fossil based feeds, like naphtha, natural gas or crude oil. Examples for complementary n-butenes are n-butenes obtained by steam cracking of fossil based feeds, like naphtha, natural gas or crude oil.

[0145] Mixing Propylene Obtained from Renewably-Sourced Ethanol with Further Propylene

[0146] According to step c), the propylene obtained from renewably-sourced ethanol is mixed with propylene not obtained from renewably-sourced ethanol to form mixed propylene.

[0147] Preferably, the mixed propylene comprises 5 to 95 wt.-% of propylene obtained from renewably-sourced ethanol, more preferably 10 to 90 wt.-% of propylene obtained from renewably-sourced ethanol, based on the total weight of the mixed propylene.

[0148] The process comprises mixing the propylene obtained from renewably-sourced ethanol with propylene not obtained from renewably-sourced ethanol to form mixed propylene as step c). This can ensure the efficient utilization of downstream processes, e.g. , for transitional periods when supply of propylene obtained from renewably-sourced ethanol is limited. The complementary may be fossil-based, partially renewably-sourced or renewably-sourced by another production route.

[0149] Hydroformylation of Propylene

[0150] Hydroformylation of the mixed propylene produces n-butyraldehyde, isobutyraldehyde or a mixture thereof. The produced aldehydes can be separated by fractionation.

[0151] Hydroformylation or the oxo process is an important large-scale industrial process for preparing aldehydes from olefins, carbon monoxide and hydrogen. These aldehydes can be hydrogenated with hydrogen in the same operation or subsequently in a separate hydrogenation step, to produce the corresponding alcohols. In general, hydroformylation is carried out in the presence of catalysts which are homogeneously dissolved in the reaction medium. Catalysts used are generally the carbonyl complexes of metals of transition group VIII, in particular Co, Rh, Ir, Pd, Pt or Ru, which may be unmodified or modified with, for example, amine- containing or phosphine-containing ligands. A summarizing account of the processes practiced on a large scale in industry is found in J. Falbe, “New Syntheses with Carbon Monoxide”, Springer Verlag 1980, p. 162 ff, US 3,527,809; 3,917,661; 4,148,830; 4,742,178, 4,769,984; 4,885,401; 6,049,011.

[0152] Propylene is preferably hydroformylated using ligand-modified rhodium carbonyls as the catalyst. Hydroformylation of propylene can be carried out at temperatures in the range of 50 °C to 200 °C, preferably 60 °C to 150 °C, and more preferably 70 °C to 120 °C.

[0153] In one embodiment, the hydroformylation reaction is conducted at a low pressure, e.g., a pressure in the range of 0.05 to 50 MPa (absolute), and preferably in the range of about 0.1 MPa to 30 MPa, most preferably at a pressure below 5 MPa. Desirably, the partial pressure of carbon monoxide is not greater than 50% of the total pressure.

[0154] The proportions of carbon monoxide, hydrogen, and propylene in the hydroformylation reaction medium can be selected within a wide range. In some embodiments, based on the total amount of CO, hydrogen, and propylene, CO is from about 1 to 50 mol-%, preferably about 1 to 35 mol-%; H2 is from about 1 to 98 mol-%, preferably about 10 to 90 mol-%; and propylene is from about 0.1 to 35 mol-%, preferably about 1 to 35 mol-%.

[0155] The hydroformylation reaction preferably takes place in the presence of both liquid and gas phases. The reactants generally are in the gas phase. The catalyst typically is in the liquid phase. Because the reactants are gaseous compounds, a high contact surface area between the gas and liquid phases is desirable to enhance good mass transfer. A high contact surface area between the catalyst solution and the gas phase may be provided in any suitable manner. In a batch process, the batch contents are thoroughly mixed during the course of the reaction. In a continuous operation the reactor feed gas can be contacted with the catalyst solution in, for example, a continuous-flow stirred autoclave where the gas is introduced and dispersed at the bottom of the vessel, preferably through a perforated inlet (e.g., a sparger). High contact between the catalyst and the gas feed may also be provided by dispersing the solution of the Rh catalyst on a high surface area support, a technique well known in the art as supported liquid phase catalysis, or providing the Rh as part of a permeable gel.

[0156] The reaction may be conducted either in a batch mode or, preferably, on a continuous basis. One or more reactors may be used in continuous modes to carry out the reaction in one or more stages.

[0157] The ratio of H2to CO in the syngas used for hydroformylation is desirably in the range from 1.1 : 1 to 1.01: 1 , preferably 1.06:1 to 1.02: 1. Often, syngas may be made or otherwise initially provided in a manner such that the ratio of hydrogen to CO is much higher than this The excess hydrogen can be separated and used in other reaction stages as desired. For example, the excess hydrogen may be used to reduce n-butyraldehyde to n-butanol. In some modes of practice, syngas in the practice of the present invention is anhydrous.

[0158] Hydrogenation of n-Butyraldehyde

[0159] In one embodiment, the obtained n-butyraldehyde is hydrogenated in the presence of hydrogen. The hydrogenation of n-butyraldehyde to n-butanol is a well-known reaction and can be conducted by any suitable known process.

[0160] In one embodiment, the hydrogenation is carried out with hydrogen in the liquid or gas phase in the presence of a hydrogenation catalyst. Homogeneous or heterogeneous catalysts can be used. Copper catalysts have proved to be the most suitable. Typically, the reaction is carried out in the liquid phase on fixed-bed catalysts at 20 to 200 °C and pressures of up to 30 MPa. Hydrogenation in the gas phase is preferably carried out continuously. Further details can be taken from Ullmann’s Encyclopedia of Industrial Chemistry, 5thedition, vol. A1 , 1984.

[0161] Condensation of n-Butyraldehyde to Produce 2-Ethyl-3-Hydroxyhexanal

[0162] In one embodiment, the obtained n-butyraldehyde is condensed to produce 2-ethyl-3-hydroxyhexanal, and the 2-ethyl-3-hydroxyhexanal is subjected to a hydrogenation reaction produces 2-ethylhexanol.

[0163] An aldol condensation is a well-known condensation reaction in which an enol or an enolate ion reacts with a carbonyl compound to form a |3-hydroxyaldehyde or p-hydroxyketone (an aldol reaction) in the presence of an acid or base catalyst, followed by dehydration to give a conjugated enone and hydrogenation to the corresponding alcohol. In the present case, n-butyraldehyde is reacted in a self-aldol condensation to obtain 2-ethy l-3-hydroxyhexanal .

[0164] Aldol condensations can occur under a variety of conditions under weak acidic or strong basic conditions and in the presence of various catalysts. The reaction can typically be carried out in liquid phase using an aqueous caustic catalyst at a temperature of about 80 to 140 °C. In another embodiment, the reaction can be carried out in gaseous phase by contacting the aldehyde in the vapor phase with a particulate catalyst comprising at least one basic alkali metal compound on an inert substrate at a temperature above 175 °C. Further details are provided in WO 2000 / 031011 . The hydrogenation of the obtained 2-ethyl-3-hydroxyhexanal to 2-ethylhexanol can be carried out analogously to the above-described hydrogenation of n-butyraldehyde.

[0165] Oxidation of Propylene to Produce Acrylic Acid

[0166] In an embodiment, the sequence of chemical conversion according to step d) additionally comprises 5) and E):

[0167] 5) oxidation reaction of the propylene to produce acrylic acid;

[0168] E) esterification reaction of the acrylic acid with one of n-butanol obtained in p) or 2-ethylhexanol obtained in y) to produce an acrylic ester.

[0169] Acrylic acid is an important basic chemical. Owing to its very reactive double bond and the acid function, it is suitable in particular for use as monomer for preparing polymers. Of the amount of acrylic acid monomer produced, the major part is esterified before polymerization, for example to form acrylate adhesives, dispersions or coatings. Only the smaller part of the acrylic acid monomer produced is polymerized directly, for example to form water-absorbent resins. Whereas, in general, the direct polymerization of acrylic acid requires high purity monomer, the acrylic acid for conversion into acrylate before polymerization does not have to be so pure.

[0170] It is common knowledge that acrylic acid can be produced by heterogeneously catalyzed gas phase oxidation of propylene with molecular oxygen over solid catalysts at temperatures between 200° to 400° C. in two stages via acrolein (cf. for example DE-A 19 62 431 , DE-A 29 43 707, DE-C 1 205 502, EP-A 257 565, EP-A 253 409, DE-B 22 51 364, EPA 117 146, GB-C 1 450 986 and EP-A 293 224). The catalysts used are oxidic multicomponent catalysts based for example on oxides of the elements molybdenum, chromium, vanadium or tellurium Five most commonly used catalyst systems for acrolein production are cuprous oxides, uranium antimony oxides, tin antimony oxides, bismuth molybdate oxides and multicomponent bismuth molybdate based oxides. The most efficient catalysts for partial oxidation of propylene to acrolein consist of multi-component metal oxides systems. In almost every multi-component catalyst system, bismuth molybdate serves as the main ingredient. The following components are most commonly used as catalyst additives in molybdate bismuth oxide based catalysts: iron, cobalt, nickel, tungsten, potassium and phosphorous. Typical catalyst supports are inert porous solids, such as SiO2, AI2O3, MgO, TiC>2, ZrO2, aluminosilicates, zeolites, activated carbon, and ceramics

[0171] The oxidation of propylene to acrylic acid can be carried out in one stage or two stages. Catalysts used for the heterogeneously catalyzed reaction are as a rule multimetal oxide materials which generally contain heavy metal molybdates as main component and compounds of various elements as promoters. The oxidation of propylene takes place in a first step to give acrolein and in a second step to give acrylic acid. Since the two oxidation steps may differ in their kinetics, uniform process conditions and a single catalyst do not as a rule lead to optimum selectivity. Recently, two-stage processes with optimum adaptation of catalyst and process variables have therefore preferably been developed. In general, propylene is oxidized to acrolein in the presence of molecular oxygen in the first stage in an exothermic reaction in a fixed-bed tubular reactor. The reaction products are passed directly into the second reactor and are further oxidized to acrylic acid. The reaction gases obtained in the second stage can be condensed and the acrylic acid can be isolated therefrom by extraction and / or distillation.

[0172] The oxidation of propylene to acrolein and / or acrylic acid is highly exothermic. The tubes of the fixed-bed tubular reactor which are filled with the heterogeneous catalyst are therefore surrounded by a cooling medium, as a rule a salt melt, such as a eutectic mixture of KNO3 and NaNC . The heat of reaction is released through the wall of the catalyst-filled tubes to the salt bath.

[0173] Particularly preferred multimetal oxide materials have the formula I or II

[0174] [X1aX2bOx]p[X3cX4dX5eXWgX2hOy]q(I)

[0175] MOl2BiiX8kFeiX9mX10nOz (II) where

[0176] X1is bismuth, tellurium, antimony, tin and / or copper, preferably bismuth, X2is molybdenum and / or tungsten, X3is an alkali metal, thallium and / or samarium, preferably potassium,

[0177] X4is an alkaline earth metal, nickel, cobalt, copper, manganese, zinc, tin, cadmium and / or mercury, preferably nickel and / or cobalt,

[0178] X5is iron, chromium, cerium and / or vanadium, preferably iron,

[0179] X6is phosphorus, arsenic, boron and / or antimony,

[0180] X7is a rare earth metal, titanium, zirconium, niobium, tantalum, rhenium, ruthenium, rhodium, silver, gold, aluminum, gallium, indium, silicon, germanium, lead, thorium and / or uranium, preferably silicon, aluminum, titanium and / or zirconium, a is from 0.01 to 8, b is from 0.1 to 30, c is from 0 to 4, d is from 0 to 20, e is from 0 to 20, f is from 0 to 6, g is from 0 to 15, h is from 8 to 16, x and y are numbers which are determined by the valency and frequency of the elements other than oxygen in I, p and q are numbers whose ratio p / q is from 0.1 to 10,

[0181] X8is cobalt and / or nickel, preferably cobalt,

[0182] X9is silicon and / or aluminum, preferably silicon,

[0183] X10is an alkali metal, preferably potassium, sodium, cesium and / or rubidium, in particular potassium, i is from 0.1 to 2, k is from 2 to 10, I is from 0.5 to 10, m is from 0 to 10, n is from 0 to 0 5, z is a number which is determined by the valency and frequency of the elements other than oxygen in II.

[0184] Multimetal oxide materials of the formula I are known per se from EP 0 000 835 and EP 0 575 897, and multimetal oxide materials of the formula II are known per se from DE 198 55 913.

[0185] Briefly, a process for preparing acrylic acid typically comprises the steps of:

[0186] (a) catalytic gas phase oxidation of propylene and / or acrolein to acrylic acid to obtain a gaseous reaction product comprising acrylic acid;

[0187] (b) solvent absorption of the reaction product;

[0188] (c) distillation of the solvent loaded with reaction product in a column to obtain a crude acrylic acid and the solvent,

[0189] (d) purification of the crude acrylic acid by crystallization.

[0190] Step (a) affords not pure acrylic acid, but a gaseous mixture which in addition to acrylic acid can substantially include unconverted acrolein and / or propylene, water vapor, carbon monoxide, carbon dioxide, nitrogen, oxygen, acetic acid, propionic acid, formaldehyde, further aldehydes and maleic anhydride.

[0191] The remaining, unabsorbed reaction gas of step (a) is further cooled down so that the condensable part of the low-boiling co-components thereof, especially water, formaldehyde and acetic acid, may be separated off by condensation. This condensate is known as acid water. The remaining gas stream, hereinafter called recycle gas, consists predominantly of nitrogen, carbon oxides and unconverted starting materials. Preferably, the recycle gas is partly recirculated into the reaction stages as diluting gas.

[0192] The oxidation of propylene to acrolein, as well as the oxidation of acrolein to acrylic acid, proceed with less than 100% selectivity and are accompanied by the combustion of propylene or acrolein over the catalyst, which gives carbon monoxide and carbon dioxide, herein collectively referred to as COX. It should be appreciated that emission of the carbon dioxide side product does not contribute to the carbon footprint of this process, as the starting propylene is carbon neutral.

[0193] Hydrogenation of COXfrom Oxidation Reaction

[0194] In an embodiment of this aspect, the propylene oxidation reaction progresses under formation of COXas a side product, and the process further comprises subjecting said COXto hydrogenation to produce at least one of synthesis gas, methanol, formaldehyde and formic acid. While the carbon monoxide and / or carbon dioxide may be sequestered by, e.g., underground storage, it may be beneficial to subject said COXto hydrogenation to produce at least one of synthesis gas, methanol, formaldehyde and formic acid. The synthesis gas, methanol, formaldehyde and / or formic acid thus produced may then be certified as carbon negative, and can at least partly displace their fossil-based counterparts and reduce the carbon footprint of chemical conversion processes making use of synthesis gas, methanol, formaldehyde and / or formic acid.

[0195] Esterification of n-Butanol or 2-Ethylhexanol

[0196] In an embodiment, the sequence of chemical conversion according to step d) further comprises esterification of the n-butanol, the isobutanol and / or the 2-ethylhexanol with carboxylic acids. Suitable carboxylic acids include saturated and non-saturated Ci-Ci6-carboxylic acid, in particular (meth)acrylic acid. When the carboxylic acid is a saturated carboxylic acid, a polymerization inhibitor is not required.

[0197] Esters of C4-C10 carboxylic acids, such as phthalic acid and adipic acid, and n-butanol and / or 2-ethylhexanol are widely used as plasticizers in plastics, such as cellulose acetates, polyurethanes, PVC, polyacrylates, etc. They may be prepared by reacting the acid component or an anhydride thereof with the alcohol component in the presence of an esterification catalyst. The reaction is an equilibrium reaction. The equilibrium may be shifted to the product side, i.e. the ester side, by continuous removal of the water produced as by-product from the reaction. 2-Ethylhexanol has a region in which it is not miscible with water; hence it is possible to distill off continuously from the reaction mixture a mixture of the water of the reaction and 2-ethylhexanol, and, after phase separation, to return the organic phase to the esterification, while the aqueous phase is removed from the system.

[0198] Esterification of the acrylic acid with an alcohol yields an acrylic ester. In particular, esterification of acrylic acid with n-butanol obtained in p) yields n-butyl acrylate, while esterification of acrylic acid with 2-ethylhexanol obtained in y) yield 2-ethylhexyl acrylate.

[0199] Acrylic esters are generally known and are important, for example, as reactive monoethylenically unsaturated monomers for the preparation of aqueous polymer dispersions by the free radical aqueous emulsion polymerization method, which dispersions are used, for example, as adhesives.

[0200] The acrylic acid can be esterified in a conventional manner to produce the desired acrylic acid ester using the corresponding alkanol, i.e., n-butanol or 2-ethylhexanol. Processes for the preparation of alkyl acrylates by reacting acrylic acid with alkanols in the homogeneous liquid phase at elevated temperatures and in the presence of catalysts are equilibrium reactions in which the conversion of the acrylic acid and of the alkanol to the corresponding ester is limited by the equilibrium constant. Consequently, for an economical procedure, the unconverted starting materials have to be separated from the resulting ester and recycled to the reaction zone.

[0201] Conveniently, the reaction zone may consist of a cascade of reaction regions, connected in series, and the discharge stream of one reaction region forms a feed stream of a subsequent reaction region and the concentration of the esterification catalyst increases along the reaction cascade. Acrylic acid, the alkanol and the catalyst are fed continuously to the reaction zone. An azeotropic mixture comprising the alkyl acrylate, water and optionally starting alkanol is separated off by rectification via the top of a rectification zone mounted on the reaction zone. The azeotropic mixture is separated into an organic phase containing the alkyl acrylate and an aqueous phase, with a part of the organic phase being recycled to the reaction zone. The alkyl acrylate is isolated from the excess organic phase. The latter is usually carried out by separation steps involving rectification (cf. for example DE 19536178).

[0202] The temperature in the reaction zone depends on the type of alcohol used and is suitably in the range of 70 to 160 °C, preferably 100 to 140 °C. The total residence time of the reactants in the reaction zone is as a rule from 0.25 to 15 h, frequently from 1 to 7 h, or from 2 to 5 h.

[0203] Suitable acidic esterification catalysts include acidic ion exchange resins and strong mineral acids, e.g. sulfuric acid, or organic sulfonic acids, such as methanesulfonic acid, benzenesulfonic acid, dodecanesulfonic acid or para-toluenesulfonic acid, or a mixture of some or all of the abovementioned acids. Sulfuric acid is particularly suitable for carrying out the novel process. This applies in particular to the preparation of n-butyl acrylate.

[0204] The content of acidic esterification catalyst in the reaction zone is expediently from 0.1 to 20wt.-%, frequently from 0.5 to 5 wt.-%, based on the reaction mixture contained therein.

[0205] To prevent undesired formation of polymer initiated by free radicals, a polymerization inhibitor is typically used during esterification Examples of suitable polymerization inhibitors are hydroquinone, 4-methoxyphenol, and phenothiazine, which may be used singly or in admixture with each other. It is usual to add from about 0.01 to 0.1 wt.-% of polymerization inhibitor to the esterification mixture and mixtures containing the methacrylic ester.

[0206] The invention is further illustrated by the following examples.

[0207] Examples

[0208] The examples are based on simulations performed via the flow sheet simulation platform Aspen Plus V14.0.

[0209] Example 1

[0210] This example is a plasma gasification process comprising a plasma fixed bed gasifier, converting a first feed stream of refuse derived fuel into a clean synthesis gas. The plasma reactor is fed with the feed stream at a mass flow ratio of 1.00 t / (t clean syngas) at 1 bar(abs) and 25 °C. The feed stream comprises 40 wt.-% carbon, 8 wt.-% hydrogen, 40 wt.-% oxygen, 1 wt.-% sulfur and 1 wt.-% nitrogen, rest percentage is ash Additionally, gasification agents are injected into the plasma reactor, enabling a gasification reaction at a temperature of 1350°C. The gasification agent is steam with a mass flow ratio of 0.10 t / (t clean syngas) at 5.4 bar(abs) and 180°C, forming a plasma in the reactor with 3.26 MWh / (t clean syngas) of electricity. The resulting raw synthesis gas stream is sent to the gasifier outlet. The resulting high-temperature synthesis gas stream, comprising CO, H2O, CO2, H2, CH4, H2S, N2, NH3, is washed with water and dried in order to reduce the amount of ash and tars. The raw synthesis gas exits the wash a d drying step at a temperature of 25 °C. The acid gas removal, enabled by an amine scrubbing, is used to separate acids like H2S and CO2 with a mass flow ratio of 0.002 t / (t clean syngas) from the raw synthesis gas. The obtained clean synthesis gas stream comprises a molar ratio of H2 / CO = 1 .47 at a mass flow rate of 0.09 1 H2 / (t clean syngas) and 0.86 t CO / (t clean syngas).

[0211] Example 2

[0212] This example is a plasma gasification process comprising a plasma fixed bed gasifier, converting a first feed stream of refuse derived fuel into a clean synthesis gas. The plasma reactor is fed with the feed stream at a mass flow ratio of 1.07 t / (t clean syngas) at 1 bar(abs) and 25 °C. The feed stream comprises 40 wt.-% carbon, 8 wt.-% hydrogen, 40 wt.-% oxygen, 1 wt.-% sulfur and 1 wt.-% nitrogen, rest percentage is ash. Additionally, gasification agents are injected into the plasma reactor, enabling a gasification reaction at a temperature of 1350°C. The gasification agent is steam with a mass flow ratio of 1.07 t / (t clean syngas) at 5.4 bar(abs) and 180°C, forming a plasma in the reactor with 5.06 MWh / (t clean syngas) of electricity. The resulting raw synthesis gas stream is sent to the gasifier outlet. The resulting high-temperature synthesis gas stream, comprising CO, H2O, CO2, H2, CH4, H2S, N2, NH3, is washed with water and dried in order to reduce the amount of ash and tars. The raw synthesis gas exits the wash and drying step at a temperature of 25 °C The acid gas removal, enabled by an amine scrubbing, is used to separate acids like H2S and CO2 with a mass flow ratio of 0.31 t / (t clean syngas) from the raw synthesis gas. The obtained clean synthesis gas stream comprises a molar ratio of H2 / CO = 2.02 at a mass flow rate of 0.12 t H2 / (t clean syngas) and 0.80 1 CO / (t clean syngas).

[0213] Example 3

[0214] This example is a plasma gasification process comprising a plasma fixed bed gasifier, converting a first feed stream of refuse derived fuel into a clean synthesis gas. The plasma reactor is fed with the feed stream at a mass flow ratio of 0.95 t / (t clean syngas) at 1 bar(abs) and 25 °C. The feed stream comprises 65 wt.-% carbon, 9 wt.-% hydrogen, 21 wt.-% oxygen, 0 wt.-% sulfur and 0 wt.-% nitrogen, rest percentage is ash. Additionally, gasification agents are injected into the plasma reactor, enabling a gasification reaction at a temperature of 1350°C. The gasification agent is steam with a mass flow ratio of 0.10 t / (t clean syngas) at 5.4 bar(abs) and 180°C, forming a plasma in the reactor with 2.95 MWh / (t clean syngas) of electricity. The resulting raw synthesis gas stream is sent to the gasifier outlet. The resulting high-temperature synthesis gas stream, comprising CO, H2O, CO2, H2, CH4, H2S, N2, NH3, is washed with water and dried in order to reduce the amount of ash and tars. The raw synthesis gas exits the wash a d drying step at a temperature of 25 °C The acid gas removal, enabled by an amine scrubbing, is used to separate acids like H2S and CO2 with a mass flow ratio of 0.001 t / (t clean syngas) from the raw synthesis gas. The obtained clean synthesis gas stream comprises a molar ratio of H2 / CO = 2.69 at a mass flow rate of 0.10 t H2 / (t clean syngas) and 0.50 1 CO / (t clean syngas).

[0215] Example 4

[0216] This example is a plasma gasification process comprising a plasma fixed bed gasifier, converting a first feed stream of refuse derived fuel into a clean synthesis gas. The plasma reactor is fed with the feed stream at a mass flow ratio of 1.43 t / (t clean syngas) at 1 bar(abs) and 25 °C. The feed stream comprises 30 wt.-% carbon, 6 wt.-% hydrogen, 20 wt.-% oxygen, 2 wt.-% sulfur and 2 wt.-% nitrogen, rest percentage is ash. Additionally, gasification agents are injected into the plasma reactor, enabling a gasification reaction at a temperature of 1350°C. The gasification agent is steam with a mass flow ratio of 0.14 t / (t clean syngas) at 5.4 bar(abs) and 180°C, forming a plasma in the reactor with 3.09 MWh / (t clean syngas) of electricity. The resulting raw synthesis gas stream is sent to the gasifier outlet. The resulting high-temperature synthesis gas stream, comprising CO, H2O, CO2, H2, CH4, H2S, N2, NH3, is washed with water and dried in order to reduce the amount of ash and tars. The raw synthesis gas exits the wash and drying step at a temperature of 25 °C The acid gas removal, enabled by an amine scrubbing, is used to separate acids like H2S and CO2 with a mass flow ratio of 0.001 t / (t clean syngas) from the raw synthesis gas. The obtained clean synthesis gas stream comprises a molar ratio of H2 / CO = 1 .92 at a mass flow rate of 0 10 t H2 / (t clean syngas) and 0.72 1 CO / (t clean syngas).

[0217] Example 5

[0218] This example is a two-step gasification process comprising a fluidized bed reactor, followed by an entrained flow gasifier, converting a first feed stream of refuse derived fuel into a clean synthesis gas. The fluidized bed reactor is fed with the feed stream at a mass flow ratio of 1.40 t / (t clean syngas) at 4 bar(abs) and 35 °C. The feed stream comprises 40 wt.-% carbon, 8 wt.-% hydrogen, 40 wt.-% oxygen, 1 wt -% sulfur and 1 wt.-% nitrogen, rest percentage is ash. Additionally, gasification agents are injected into the fluidized bed reactor, enabling a gasification reaction at a temperature of 800°C. The gasification agents are steam with a mass flow ratio of 0.7 t / (t clean syngas) at 5.4 bar(abs) and 180°C, as well as oxygen with a mass flow ratio of 0.46 t / (t clean syngas) at 25 °C and 6 bar(abs). The resulting raw synthesis gas stream is sent to the top of the entrained flow gasifier whereby being kept at a temperature of 800 °C as a second stage and injected via the burner. No second fuel is added. Here, additional oxygen at a mass flow ratio of 0.22 t / (t clean syngas) at 15 bar(abs) and 100°C is used to increase the temperature up to 1350 °C at the gasifier outlet and convert remaining hydrocarbons, tars, and soot into synthesis gas. The resulting high- temperature synthesis gas stream, comprising CO, H O, CO , H2, CH4, H2S, N , NH3, is washed with water and dried in order to reduce the amount of ash and tars. The raw synthesis gas exits the wash and drying step at a temperature of 25 °C. The acid gas removal, enabled by an amine scrubbing, is used to separate acids like H2S and CO2 with a mass flow ratio of 0.65 t / (t clean syngas) from the raw synthesis gas. The obtained clean synthesis gas stream comprises a molar ratio of H2 / CO = 1.22 at a mass flow rate of 0.08 1 H2 / (t clean syngas) and 0.89 1 CO / (t clean syngas).

[0219] Example 6

[0220] This example is a two-step gasification process comprising a fluidized bed reactor, followed by an entrained flow gasifier, converting a first feed stream of refuse derived fuel into a clean synthesis gas. The fluidized bed reactor is fed with the feed stream at a mass flow ratio of 1.13 t / (t clean syngas) at 4 bar(abs) and 35 °C. The feed stream comprises 40 wt.-% carbon, 8 wt.-% hydrogen, 40 wt.-% oxygen, 1 wt.-% sulfur and 1 wt.-% nitrogen, rest percentage is ash. Additionally, gasification agents are injected into the fluidized bed reactor, enabling a gasification reaction at a temperature of 800°C. The gasification agents are steam with a mass flow ratio of 0.56 t / (t clean syngas) at 5.4 bar(abs) and 180°C, as well as oxygen with a mass flow ratio of 0.37 t / (t clean syngas) at 25 °C and 6 bar(abs). The resulting raw synthesis gas stream is sent to the top of the entrained flow gasifier whereby being kept at a temperature of 800 °C as a second stage and injected via the burner. A bio oil is added as a second fuel to the entrained flow gasifier with a mass ratio of 0.11 t / (t clean syngas) at 100 °C and 20 bar comprising 75 wt.-% carbon, 7 wt.-% hydrogen and 18 wt.-% oxygen. Also, additional oxygen at a mass flow ratio of 0.27 t / (t clean syngas) at 15 bar(abs) and 100°C is used to increase the temperature up to 1350 °C at the gasifier outlet and convert remaining hydrocarbons, tars, and soot into synthesis gas. The resulting high-temperature synthesis gas stream, comprising CO, H2O, CO2, H2, CH4, H2S, N2, N H3, is washed with water and dried in order to reduce the amount of ash and tars. The raw synthesis gas exits the wash and drying step at a temperature of 25 °C. The acid gas removal, enabled by an amine scrubbing, is used to separate acids like H2S and CO2 with a mass flow ratio of 0.55 t / (t clean syngas) from the raw synthesis gas. The obtained clean synthesis gas stream comprises a molar ratio of H2 / CO = 1.13 at a mass flow rate of 0.07 1 H2 / (t clean syngas) and 0 90 1 CO / (t clean syngas).

[0221] Example 7

[0222] This example is a two-step gasification process comprising a fluidized bed reactor, followed by an entrained flow gasifier, converting a first feed stream of refuse derived fuel into a clean synthesis gas. The fluidized bed reactor is fed with the feed stream at a mass flow ratio of 1.09 t / (t clean syngas) at 4 bar(abs) and 35 °C. The feed stream comprises 40 wt.-% carbon, 8 wt.-% hydrogen, 40 wt.-% oxygen, 1 wt.-% sulfur and 1 wt.-% nitrogen, rest percentage is ash. Additionally, gasification agents are injected into the fluidized bed reactor, enabling a gasification reaction at a temperature of 800°C. The gasification agents are steam with a mass flow ratio of 0.55 t / (t clean syngas) at 5.4 bar(abs) and 180°C, as well as oxygen with a mass flow ratio of 0.36 t / (t clean syngas) at 25 °C and 6 bar(abs). The resulting raw synthesis gas stream is sent to the top of the entrained flow gasifier whereby being kept at a temperature of 800 °C as a second stage and injected via the burner. A bio oil is added as a second fuel to the entrained flow gasifier with a mass ratio of 0.11 t / (t clean syngas) at 100 °C and 20 bar comprising 85 wt.-% carbon, 8 wt.-% hydrogen and 7 wt.-% oxygen. Also, additional oxygen at a mass flow ratio of 0.27 t / (t clean syngas) at 15 bar(abs) and 100°C is used to increase the temperature up to 1350 °C at the gasifier outlet and convert remaining hydrocarbons, tars, and soot into synthesis gas. The resulting high-temperature synthesis gas stream, comprising CO, H2O, CO2, H2, CH4, H2S, N2, NH3, is washed with water and dried in order to reduce the amount of ash and tars. The raw synthesis gas exits the wash and drying step at a temperature of 25 °C. The acid gas removal, enabled by an amine scrubbing, is used to separate acids like H2S and CO2 with a mass flow ratio of 0.52 t / (t clean syngas) from the raw synthesis gas. The obtained clean synthesis gas stream comprises a molar ratio of H2 / CO = 1.13 at a mass flow rate of 0.07 1 H2 / (t clean syngas) and 0 90 1 CO / (t clean syngas).

[0223] Example 8

[0224] This example is a two-step gasification process comprising a fluidized bed reactor, followed by an entrained flow gasifier, converting a first feed stream of refuse derived fuel into a clean synthesis gas. The fluidized bed reactor is fed with the feed stream at a mass flow ratio of 1.23 t / (t clean syngas) at 4 bar(abs) and 35 °C. The feed stream comprises 40 wt.-% carbon, 8 wt.-% hydrogen, 40 wt.-% oxygen, 1 wt.-% sulfur and 1 wt.-% nitrogen, rest percentage is ash. Additionally, gasification agents are injected into the fluidized bed reactor, enabling a gasification reaction at a temperature of 800°C. The gasification agents are steam with a mass flow ratio of 0.62 t / (t clean syngas) at 5.4 bar(abs) and 180°C, as well as oxygen with a mass flow ratio of 0.41 t / (t clean syngas) at 25 °C and 6 bar(abs). The resulting raw synthesis gas stream is sent to the top of the entrained flow gasifier whereby being kept at a temperature of 800 °C as a second stage and injected via the burner. A bio oil is added as a second fuel to the entrained flow gasifier with a mass ratio of 0.12 t / (t clean syngas) at 100 °C and 20 bar comprising 50 wt.-% carbon, 5 wt.-% hydrogen and 45 wt.-% oxygen. Also, additional oxygen at a mass flow ratio of 0.26 t / (t clean syngas) at 15 bar(abs) and 100°C is used to increase the temperature up to 1350 °C at the gasifier outlet and convert remaining hydrocarbons, tars, and soot into synthesis gas. The resulting high-temperature synthesis gas stream, comprising CO, H2O, CO2, H2, CH4, H2S, N2, NHg, is washed with water and dried in order to reduce the amount of ash and tars. The raw synthesis gas exits the wash and drying step at a temperature of 25 °C. The acid gas removal, enabled by an amine scrubbing, is used to separate acids like H2S and CO2 with a mass flow ratio of 0.63 t / (t clean syngas) from the raw synthesis gas. The obtained clean synthesis gas stream comprises a molar ratio of H2 / CO = 1.13 at a mass flow rate of 0.07 t H2 / (t clean syngas) and 0.90 1 CO / (t clean syngas).

[0225] Example 9

[0226] This example is a one-step gasification process comprising an entrained flow gasifier, converting a first feed stream of bio oil into a clean synthesis gas. The entrained flow reactor is fed with the feed stream at a mass flow ratio of 0.57 t / (t clean syngas) at 50 bar(abs) and 100 °C. The feed stream comprises 75 wt.-% carbon, 7 wt.-% hydrogen and 18 wt.-% oxygen. Additionally, gasification agents are injected into the reactor, enabling a gasification reaction at a temperature of 1350°C. The gasification agents are steam with a mass flow ratio of 0.11 t / (t clean syngas) at 70 bar(abs) and 400°C, as well as oxygen with a mass flow ratio of 0.45 t / (t clean syngas) at 25 °C and 50 bar(abs). The resulting raw synthesis gas stream is sent to the gasifier outlet. The resulting high-temperature synthesis gas stream, comprising CO, H2O, CO2, H2, CH4, H2S, N2, NH3, is washed with water and dried in order to reduce the amount of ash and tars. The raw synthesis gas exits the wash and drying step at a temperature of 25 °C. The acid gas removal, enabled by an amine scrubbing, is used to separate acids like H2S and CO2 with a mass flow ratio of 0.08 t / (t clean syngas) from the raw synthesis gas. The obtained clean synthesis gas stream comprises a molar ratio of H2 / CO = 0.66 at a mass flow rate of 0.05 1 H2 / G clean syngas) and 0.95 1 CO / (t clean syngas).

[0227] Example 10

[0228] This example is a one-step gasification process comprising an entrained flow gasifier, converting a first feed stream of pyrolysis oil into a clean synthesis gas. The entrained flow reactor is fed with the feed stream at a mass flow ratio of 0.51 t / (t clean syngas) at 50 bar(abs) and 100 °C. The feed stream comprises 86 wt.-% carbon, 6 wt.-% hydrogen, 6 wt.-% oxygen and 2 wt.-% sulfur. Additionally, gasification agents are injected into the reactor, enabling a gasification reaction at a temperature of 1350°C. The gasification agents are steam with a mass flow ratio of 0.10 t / (t clean syngas) at 70 bar(abs) and 400°C, as well as oxygen with a mass flow ratio of 0.41 t / (t clean syngas) at 25 °C and 50 bar(abs). The resulting raw synthesis gas stream is sent to the gasifier outlet. The resulting high-temperature synthesis gas stream, comprising CO, H2O, CO2, H2, CH4, H2S, N2, NH3, is washed with water and dried in order to reduce the amount of ash and tars. The raw synthesis gas exits the wash and drying step at a temperature of 25 °C. The acid gas removal, enabled by an amine scrubbing, is used to separate acids like H2S and CO2 with a mass flow ratio of 0.01 t / (t clean syngas) from the raw synthesis gas. The obtained clean synthesis gas stream comprises a molar ratio of H2 / CO = 0.60 at a mass flow rate of 0.04 1 H2 / (t clean syngas) and 0.90 1 CO / (t clean syngas). Example 11

[0229] This example is a one-step gasification process comprising an entrained flow gasifier, converting a first feed stream of pyrolysis oil into a clean synthesis gas. The entrained flow reactor is fed with the feed stream at a mass flow ratio of 0.51 t / (t clean syngas) at 50 bar(abs) and 100 °C. The feed stream comprises 92 wt.-% carbon, 7 wt.-% hydrogen, 1 wt.-% oxygen and 0 wt.-% sulfur. Additionally, gasification agents are injected into the reactor, enabling a gasification reaction at a temperature of 1350°C. The gasification agents are steam with a mass flow ratio of 0.10 t / (t clean syngas) at 70 bar(abs) and 400°C, as well as oxygen with a mass flow ratio of 0.41 t / (t clean syngas) at 25 °C and 50 bar(abs). The resulting raw synthesis gas stream is sent to the gasifier outlet. The resulting high-temperature synthesis gas stream, comprising CO, H2O, CO2, H2, CH4, H2S, N2, NH3, is washed with water and dried in order to reduce the amount of ash and tars. The raw synthesis gas exits the wash and drying step at a temperature of 25 °C The acid gas removal, enabled by an amine scrubbing, is used to separate acids like H2S and CO2 with a mass flow ratio of 0.01 t / (t clean syngas) from the raw synthesis gas. The obtained clean synthesis gas stream comprises a molar ratio of H2 / CO = 0.73 at a mass flow rate of 0.04 1 H2 / (t clean syngas) and 0.85 1 CO / (t clean syngas).

[0230] Example 12

[0231] This example is a one-step gasification process comprising an entrained flow gasifier, converting a first feed stream of pyrolysis oil into a clean synthesis gas. The entrained flow reactor is fed with the feed stream at a mass flow ratio of 0.63 t / (t clean syngas) at 50 bar(abs) and 100 °C. The feed stream comprises 70 wt.-% carbon, 3 wt.-% hydrogen, 23 wt.-% oxygen and 4 wt.-% sulfur. Additionally, gasification agents are injected into the reactor, enabling a gasification reaction at a temperature of 1350°C. The gasification agents are steam with a mass flow ratio of 0.13 t / (t clean syngas) at 70 bar(abs) and 400°C, as well as oxygen with a mass flow ratio of 0.44 t / (t clean syngas) at 25 °C and 50 bar(abs). The resulting raw synthesis gas stream is sent to the gasifier outlet. The resulting high-temperature synthesis gas stream, comprising CO, H2O, CO2, H2, CH4, H2S, N2, N H3, is washed with water and dried in order to reduce the amount of ash and tars. The raw synthesis gas exits the wash and drying step at a temperature of 25 °C. The acid gas removal, enabled by an amine scrubbing, is used to separate acids like H2S and CO2 with a mass flow ratio of 0.14 t / (t clean syngas) from the raw synthesis gas. The obtained clean synthesis gas stream comprises a molar ratio of H2 / CO = 0.36 at a mass flow rate of 0.021 H2 / (t clean syngas) and 0.951 CO / (t clean syngas).

Claims

Claims1. Process for the manufacture of an alcohol selected from n-butanol and 2-ethy lhexanol , said process comprising the steps of: a) providing a gasifier feedstock; gasifying the gasifier feedstock by partial oxidation to form a raw product gas stream comprising a plurality of gases comprising methane, hydrogen and carbon monoxide; and recovering syngas from the raw product gas stream; b) obtaining propylene from renewably-sourced ethanol by subjecting renewably-sourced ethanol to dehydration to produce a renewably-sourced ethylene stream; and subjecting the renewably-sourced ethylene stream to an olefin-interconversion, to obtain propylene; the olefin-interconversion comprising (i) and (ii):(i) ethylene dimerization to obtain n-butenes; and(ii) metathesis reaction between n-butenes obtained according to (i) and ethylene to obtain propylene; c) mixing the propylene obtained from renewably-sourced ethanol with propylene not obtained from renewably-sourced ethanol to form mixed propylene; d) subjecting the mixed propylene to a sequence of chemical conversions to obtain n-butanol or 2-ethylhexanol, the sequence of chemical conversions comprising a) and p), or a) and y): a) hydroformylation of the propylene in the presence of syngas obtained in a) to obtain n- butyraldehyde; p) hydrogenation of the n-butyraldehyde obtained in a) in the presence of hydrogen to produce n-butanol; y) condensation of the n-butyraldehyde obtained in a) to produce 2-ethyl-3-hydroxyhexanal, and subjecting the 2-ethyl-3-hydroxyhexanal to a hydrogenation reaction in the presence of hydrogen to produce 2-ethylhexanol.

2. The process according to claim 1 , additionally comprising recovering hydrogen from the raw product gas stream, and directing the hydrogen at least partially to p) and / or y).

3. The process according to claim 1 or 2, wherein recovering syngas from the raw product gas stream comprises subjecting the raw product gas stream to a washing step to remove particulate solids and to a drying step to remove water.

4. The process according to claim 3, wherein recovering syngas from the raw product gas stream further comprises sweetening the raw product gas stream to form a sweetened product gas stream.

5. The process according to claim 3 or 4, wherein recovering syngas from the raw product gas stream further comprises adjusting the hydrogen to carbon monoxide molar ratio in the raw product gas stream or the sweetened gas stream, respectively, to obtain an adjusted gas stream comprising hydrogen and carbon monoxide at a molar ratio different from the molar ratio in the raw product gas stream or the sweetened gas stream, respectively; separating the adjusted gas stream to provide a CO-comprising gas stream, a ^-comprising gas stream and a CH4-comprising gas stream.

6. The process according to any one of the preceding claims, wherein the gasifier feedstock comprises a liquid gasifier feedstock and / or a solid gasifier feedstock.

7. The process according to claim 6, wherein the gasifier feedstock comprises a liquid gasifier feedstock, and gasifying the liquid gasifier feedstock comprises subjecting the liquid gasifier feedstock to partial oxidation in an entrained flow reactor to obtain the raw product gas stream.

8. The process according to claim 7, wherein the entrained flow reactor is operated at a temperature above 400 °C and at a pressure of 1 bar(abs) or more.

9. The process according to claim 6, wherein the gasifier feedstock comprises a solid gasifier feedstock, and gasifying the solid gasifier feedstock comprises subjecting the solid gasifier feedstock to partial oxidation in a fluidized bed reactor so as to obtain an intermediate gasification product, and directing the intermediate gasification product to gasification in an entrained flow reactor to obtain the raw product gas stream.

10. The process according to claim 9, wherein the fluidized bed reactor is operated at a temperature in the range of from 350 to 1000 °C and at a pressure in the range of from 1 to 200 bar(abs), and the entrained flow reactor is operated at a temperature above 400 °C and at a pressure of 1 bar(abs) or more, preferably at a temperature in the range of from 1000 to 1700 °C and at a pressure in the range of from 1 to 200 bar(abs).

11. The process according to claim 9 or 10, additionally comprising directing a liquid gasifier feedstock into the entrained flow reactor.

12. The process according to any one of claims 1 to 6, wherein gasifying the gasifier feedstock comprises partial oxidation in a plasma reactor.

13. The process according to any one of the preceding claims, wherein step b)-(i) comprises:- contacting the renewably-sourced ethylene stream with a dimerization catalyst in a dimerization zone;- operating said dimerization zone at conditions effective to produce an effluent consisting essentially of n-butenes, a stream consisting essentially of heavier olefins, and optionally an unconverted ethylene stream;- fractioning the effluent to recover a stream consisting essentially of n-butenes, a stream consisting essentially of heavier olefins, and an optional ethylene stream; and- optionally directing the stream consisting essentially of heavier olefins to step a) as a gasifier feedstock or subjecting the stream consisting essentially of heavier olefins to hydrogenation so as to obtain renewably-sourced naphtha.

14. The process according to any one of the preceding claims, wherein the n-butenes are a mixed stream including 1-butene and 2-butenes, and wherein b)-(ii) comprises passing the mixed stream through a metathesis / isomerization zone comprising both a metathesis catalyst and an isomerization catalyst.

15. The process according to any one of the preceding claims, wherein the sequence of chemical conversions additionally comprises 5) and E):5) oxidation reaction of the propylene to produce acrylic acid;E) esterification reaction of the acrylic acid with one of n-butanol obtained in P) or 2-ethyl hexanol obtained in y) to produce an acrylic ester.