Technical Background

The present invention relates to a process for the manufacture ofC4-olefin-derived chemicals, in particular citral, from renewably-sourced ethanol.

Ethylene is a cornerstone of the modern petrochemical industries. Important ethylene derivatives (at the end of their respective chains) include (meth)acrylic acid, (meth)acrylic esters, isononanols, ethylhexanol, and ethylene glycols. One of the problems faced by the manufacture of chemicals and intermediates from ethylene is that the starting raw materials are from fossil fuels, such as natural gas or crude oil, which are non-renewable feedstocks. Steam cracking, which employs petroleum fractions and natural gas liquids as feedstocks, is the dominant method for large-scale ethylene production worldwide.

Lower olefins, such as isobutylene or propylene, are of significant interest for industrial and chemical applications. Isobutylene, also known as isobutene or 2-methylpropene, is a hydrocarbon of significant interest that is widely used as an intermediate in the production of industrially important products, including para-xylene, jet fuel blendstocks, gasoline oxygenates, isooctane, methacrolein, methyl methacrylate, and butyl rubber. Propylene is a hydrocarbon of significant interest that is widely used as an intermediate in the production of acrylic acid. Historically, lower olefins have been obtained through the catalytic or steam cracking of fossil fuel feedstocks.

Applicants have realized that the production of ethylene and ethylene derivatives compounds 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. Even after the purification processes, the advantage of bioethanol is frequently decreased by small amounts of impurities which it contains. Bioethanol impurities may include oxygen-containing organics, for example other alcohols such as isopropanol, n-propanol, and isobutanol, and/or aldehydes such as acetaldehyde. Bioethanol impurities may further include sulfur-containing impurities, such as inorganic sulfur compounds dialkyl sulfides, dialkyl sulfoxides, alkyl mercaptans, 3-methylthio-1- propanol, and/or sulfur-containing amino acids.

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.

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.

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 [0033] and [0061].

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 alpha-olefins are transformed into internal olefins that are subsequently subjected to metathesis with ethylene.

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. EP 3067340 A discloses a process comprising fermenting a renewable source of carbon for the production of a mixture of alcohols comprising ethanol, isopropanol and 1 -butanol; joint dehydration of the alcohols to produce a mixture of olefins comprising chiefly ethylene, propylene and linear butenes, the linear butenes being a mixture of 1 -butene and 2-butenes (cis- and trans-isomers), besides water and by-products; removal of water, oxygenated compounds and other by-products from the mixture of olefins, to generate a mixture of olefin comprising chiefly ethylene, propylene and linear butenes; and passing the mixture of olefins through an isomerization bed so that 1 -butene is isomerized to 2-butene and subsequently passing the mixture of olefins comprising chiefly ethylene, propylene and 2-butenes through a metathesis bed, for reaction between ethylene and 2-butenes, generating additional propylene.

WO 2009/098268 discloses a process for the dehydration of an alcohol to make an olefin. The alcohol may be ethanol that can be obtained from carbohydrates. For this purpose a stream comprising the ethanol and an inert component is contacted with a catalyst to give ethylene. It is indicated that the ethylene can be used for dimerization to butene and then isomerization to isobutene, dimerization to 1 -butene, which is isomerized to 2- butene and further converted by metathesis with ethylene to propylene, or conversion to ethylene oxide and glycol. Experimental details are provided only for ethanol dehydration. A similar process is disclosed in WO 2011/089235.

WO 2009/098269 discloses a process for conversion of ethanol that can be obtained from carbohydrates to propylene. The ethanol is dehydrated to ethylene which is reacted with olefins having four carbon atoms or more to give propylene. WO 2009/098267 discloses a similar process.

WO 2021/067294 discloses a process for simultaneously dehydrating, dimerizing and metathesizing a C2-C5 alcohol which can be from biobased processes in one reactor to produce a C2-C7 olefin.

WO 2009/070858 discloses an integrated process for the production of ethylenebutylene copolymers. The ethylene is obtained by dehydration of ethanol that is produced by the fermentation of sugars. One method of obtaining 1 -butylene used for the polymerization is indicated to be dimerization of ethylene produced by dehydration of ethanol that is produced by the fermentation of sugars. No details as to the dimerization are given. In embodiments, the invention seeks to advise a reaction scheme that provides renewably-sourced light olefins, such as ethylene, butenes and isobutene, which partially or fully replace 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.

Detailed Description of the Invention

To this effect, the present invention relates to a process for the manufacture of a chemical of interest selected from C4-olefin-derived chemicals, said process comprising the steps of: a) subjecting a feedstock comprising a renewably-sourced ethanol to dehydration to produce a renewably-sourced ethylene stream; b) subjecting the renewably-sourced ethylene stream to an olefin-interconversion, to obtain one or more renewably-sourced C4-olefins, selected from n-butenes and isobutene; the olefin-interconversion comprising (i) and, where required, (ii):

(i) ethylene dimerization to obtain n-butenes;

(ii) isomerization of n-butenes obtained according to (i) to obtain isobutene; and c) subjecting the renewably-sourced C4-olefin to a chemical conversion or sequence of chemical conversions to obtain the chemical of interest.

In particular, the present invention relates to a process for the manufacture of citral, said process comprising the steps of: a) subjecting a feedstock comprising a renewably-sourced ethanol to dehydration to produce a renewably-sourced ethylene stream; b) subjecting the renewably-sourced ethylene stream to an olefin-interconversion, to obtain renewably-sourced isobutene; the olefin-interconversion comprising (i) and (ii):

(i) ethylene dimerization to obtain n-butenes; and

(ii) isomerization of n-butenes obtained according to (i) to obtain isobutene; and c) subjecting the renewably-sourced isobutene to a sequence of chemical conversions to obtain citral, comprising: a) reacting the isobutene with formaldehyde to produce isoprenol;

P) isomerizing isoprenol obtained according to (i) to prenol; y) subjecting isoprenol obtained according to (i) to an oxidation reaction to produce isoprenal;

5) isomerizing the isoprenal to prenal; and

E) reacting the prenol with the prenal to produce citral.

The process of the invention is preferably a continuous process in the sense that at least the upstream steps of a reaction route leading to a chemical of interest, including at least step a) and any of steps b-(i) through (ii) as far as they are involved in the reaction route, are carried out continuously. In a still more preferred embodiment all steps of a reaction route leading to a chemical of interest are carried out continuously. This does not preclude the presence of buffer volumes between subsequent reaction steps in a reaction route.

The present invention is based on the idea of eliminating impurities that are inherently present in renewably-sourced ethanol during the ethylene manufacturing process itself. Hence, the renewably-sourced ethylene, or the renewably-sourced C4-olefin, in particular n-butenes and isobutene, produced therefrom, can be blended or used interchangeably with a fossil-derived intermediate of the same chemical structure without necessitating adjustments in downstream processes.

It is envisaged that the renewably-sourced olefins involved in the process according to the invention may be blended with complementary olefins from other sources. This can ensure the efficient utilization of downstream processes, e.g., for transitional periods when supply of renewably-sourced olefin is limited. These complementary olefins, including complementary ethylene, complementary isobutene and complementary n-butenes, may be fossil-based, partially renewably-sourced or renewably-sourced by another production route.

Hence 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 isobutene with complementary isobutene prior to step c), the complementary isobutene not being obtained from renewably-sourced n-butenes in accordance with steps a) and b); 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).

Examples for complementary ethylenes are ethylenes obtained by steam cracking of fossil based feeds, like naphtha, natural gas or crude oil. Examples for complementary isobutenes are isobutenes 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.

In another aspect, the invention also relates to a process for enhancing the environmental sustainability of citral by blending or replacing a fossil-derived isobutene with a renewably-sourced isobutene to obtain a sustainability-enhanced isobutene and subjecting the sustainability-enhanced isobutene to a sequence of chemical conversions to obtain citral, wherein the renewably-sourced isobutene is obtained by a) subjecting a feedstock comprising a renewably-sourced ethanol to dehydration to produce the renewably-sourced ethylene stream; and b) subjecting the renewably-sourced ethylene stream to an olefin-interconversion, to obtain the renewably-sourced isobutene; the olefin-interconversion comprising (i) and (ii):

(i) ethylene-dimerization to obtain n-butenes; and

(ii) isomerization of n-butenes obtained according to (i) to obtain isobutene; the sequence of chemical conversions comprising: a) reacting the isobutene with formaldehyde to produce isoprenol;

P) isomerizing the isoprenol obtained in a) to prenol; y) subjecting the isoprenol obtained in a) to an oxidation reaction to produce isoprenal;

5) isomerizing the isoprenal to prenal; and

E) reacting the prenol with the prenal to produce citral.

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 ethylene and C4-olefins, in particular n-butenes and isobutene, can be fully or partially substituted by respective renewably-sourced ethylene and C4-olefins. Hereby, one obtains a respective chemical of interest, the carbon atoms of which are fully or partially based on a renewable-sourced carbon (so-called “green” carbon).

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. Oxidation reactions of a substrate to a desired product, for example, are almost invariably accompanied to a certain extent by an over-oxidation of the substrate to form carbon oxides, in particular carbon dioxide. By a full or partial replacement of fossil ethylene and C4-olefins by their renewably-sourced counterparts, the fossil-based carbon dioxide emissions of the entire production site can be reduced because respective emissions resulting from yield losses along the value chain are at least partially based on green carbon. The resulting carbon dioxide emissions therefore do not contribute to the green house emission of the production site.

In addition, in non-oxidative reactions 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 full or partial replacement of fossil ethylene and C4- olefins by their renewably-sourced counterparts at the beginning of the processing chain reduces the emission of fossil-based carbon dioxide resulting from the combustion of downstream side-products.

Hence, it is envisaged that direct and indirect benefits are associated with the process of the invention with regard to any chemical of interest that is manufactured via the process according to the present invention.

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.

The expression “chemical of interest” collectively refers to any desired compound appearing in a value chain starting out from and including ethylene. Thus, the expression includes any intermediates and final products. In certain cases, a chemical compound can be an intermediate and final product at the same time. For example, isobutene can be the final product of a value chain and yet can be an intermediate when it is further processed, if desired.

All patent and literature documents addressed in the following are incorporated herein by reference in their entirety.

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

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.

Biomass feedstock can originate from several sources. Bioethanol production may be based on food crop feedstocks such as corn and sugar cane, sugarcane bagasse, cassava (first generation biofeedstock).

Another source of biomass feedstock is lignocellulosic materials from agricultural crops (second-generation biofeedstock). Potential feedstocks include agricultural residue byproducts 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.

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

(1 ) Enzymatic saccharification or hydrolysis of starch into sugars

(2) Microbial fermentation of sugars

(3) Purification by distillation to give hydrous ethanol

(4) Dehydration (water removal) to produce anhydrous ethanol

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.

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

As a first step, the invention 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.

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.

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.

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.

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 undesireable 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.

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.

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.

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 AhOa-MgO/SiC^, 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.

Dimerization of Ethylene

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.

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

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 ^TM process is based on this technology and leads to the industrial production of olefins.

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.

As the bivalent nickel compound, nickel carboxylates of general formula (R ¹ COO)2Ni are preferably used, where R ¹ is 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.

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.

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 R ² COOH in which R ² is 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 CX _P H3 _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. 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 Bronsted 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 Bronsted acid to the nickel compound is 0.25/1 to 10/1 , and preferably 0.25/1 to 5/1.

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 AIR _n X3- _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.gCh.i; and dibromoethyl aluminum enriched with aluminum bromide, the mixture having a formula AIEto.gBr ₂ .i.

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.

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.

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 .

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

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

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.

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.

The higher 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 steamcracking to produce renewably-sourced light olefins, dienes and aromatics.

Hence, in an embodiment, 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;

- 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

- optionally subjecting the stream consisting essentially of heavier olefins to hydrogenation so as to obtain renewably-sourced naphtha.

Skeletal Isomerization of n-Butenes

In one aspect, the process of the invention involves isomerization of n-butenes obtained according to (i) to obtain isobutene in accordance with step b)-(ii).

As the isomerization is an equilibrium reaction, the reaction mixture invariably contains unreacted n-butenes. Expediently, step b)-(ii) comprise:

- contacting the n-butenes with a skeletal isomerization catalyst in an isomerization zone to produce a mixture of n-butenes and isobutene;

- recovering from the mixture a stream consisting essentially of n-butenes and a stream consisting essentially of isobutene; and

- recycling the stream consisting essentially of n-butenes into the isomerization zone. A suitable recovery scheme utilizes the reaction of isobutene with alkanol to produce alkyl tertiary butyl ether. The etherification reaction is selective with respect to isobutene, while n-butenes are unreactive in the reaction. The reaction therefore can be utilized as a method to separate n-butenes and isobutene.

Hence, isobutene may be recovered from the mixture of n-butenes and isobutene by the following steps:

(a) reacting the mixture of n-butenes and isobutene with isobutanol in the presence of an acidic ion exchange resin in an etherification unit to form a mixture of isobutyl tert-butyl ether (IBTBE) and unconverted n-butenes;

(b) distilling the reaction mixture in a first distillation unit to obtain a top product stream consisting essentially of n-butenes, and a bottom product comprising IBTBE;

(c) feeding the bottom product to a ether cleavage unit to decompose the IBTBE to obtain isobutene and isobutanol;

(d) distilling the mixture of isobutene and isobutanol produced in step (c) in a second distillation unit to obtain a top product stream consisting essentially of isobutene, and a bottom product comprising isobutanol; and

(e) recycling the bottom product of step (d) to step (a).

Skeletal isomerization generally requires acidic catalysts. Known skeletal isomerization catalysts include aluminas and halogenated aluminas, particularly F- or Cl-promoted aluminas.

Certain zeolites have been shown to be highly effective in skeletal isomerization of normal olefins. Such zeolites include those selected from the group consisting of zeolites having the framework structure of ZSM-22, ZSM-23, and ZSM-35.

Examples for high selectivity, high stability catalysts are chlorinated Y-AI2O3, ferrierite SAPO-11 (silico-alumino phosphate molecular sieve) and MeAPO-11 (Me = Co, Mn, Mg) (molecular sieve). A particularly preferred catalyst is ferrierite. The typical elemental composition of ferrierite zeolite is Na ₂ Mg ₂ [AleSi3o072]-18H ₂ 0 as, for example, disclosed in US 6323384.

Spent catalysts can be regenerated by heating in an oxygen-containing gas, such as air, at temperatures ranging from about 200° C to about 700° C.

Skeletal isomerization of n-butenes to isobutene is an equilibrium controlled process where equilibrium conversion decreases with increasing temperature.

The skeletal isomerization is carried out by contacting the feed with the catalyst, using any suitable contacting techniques, at temperatures at which skeletal isomerization of the feed of n-butenes occurs. The feed is preferably maintained in the vapor phase during contacting. The reactor temperature is preferably in the range of about 300° to about 650° C, more preferably about 400° to about 580° C. The weight hourly space velocity (WHSV) is not narrowly critical but will generally be within the range of about 0.1 to about 40 hr ¹ , preferably from about 1 to about 20 hr ¹ . Any convenient pressure can be used, with the lowest practical pressure preferred in order to minimize side reactions such as polymerization. Preferred pressures are within the range of about 0.1 to about 10 atmospheres, more preferably about 1 to about 4 atmospheres.

The equilibrium may not be achieved in the case of a single contact of the feed with the catalyst. However, in a particular variant of the process, the product stream leaving the catalyst bed can be divided up, and only one part is directly conveyed to the working-up process, while the other part is again conducted over the catalyst bed.

Several commercial processes for n-butene isomerization are known. In one embodiment the n-butene feedstock is vaporized, in general by heat exchange with reactor effluent, and further heated to reaction temperature. In the reactor, vapor reacts with up to 44% of n-butenes converted to isobutylene with greater than 86% selectivity. Typically, two reactors are cyclically operated: one in reaction mode, the other in regeneration mode. Reactor effluent is cooled, compressed and fractionated. Heavy ends are separated and removed as bottoms from the overhead isobutylene product.

Operating conditions, process and catalyst modifications are, for example, disclosed in US 6,111 ,160 and US 6,323,384. Typical operating conditions are: 340-360°C reaction temperature, WHSV of 2 H ¹ , an olefin partial pressure of 1-2 bar, and a total pressure of 1-3 bar. Reaction of Isobutene with Formaldehyde to Produce Isoprenol

A reaction of isobutene with formaldehyde produces isoprenol.

The production of isoprenol (3-methyl-3-butene-1-ol) is well-known. For example, it can be produced by a Prins reaction between isobutene and formaldehyde in liquid phase at temperatures of 220 to 280 °C and a pressure of 230 to 270 bar with or without a catalyst. The reaction mixture may be fractionated to obtain isoprenol. Further details are provided in W02008037693 or Ullmann’s Encyclopedia of Industrial Chemistry. Wiley-VCH Verlag GmbH & Co. KGaA., 15. Juni 2000, S. 86, doi:10.1002/14356007.a14_627.

Isomerization of Isoprenol to Prenol

Isomerization of isoprenol yields prenol.

The isomerization is carried out in the presence of hydrogen and a catalyst. A preferred catalyst is a fixed bed catalyst containing palladium and selenium or tellurium or a mixture of selenium and tellurium supported on silicium dioxide. The isomerization is carried out at a temperature of 50 to 150 °C to produce a reaction mixture of prenol and isoprenol. Prenol may be separated by distillation, the isoprenol can be recycled. Further details are provided in W02008037693.

Oxidation of Isoprenol to Isoprenal

An oxidation reaction of the isoprenol produces isoprenal.

The selective oxidation of primary alcohols such as isoprenol to the corresponding aldehyde is a well-known reaction. Isoprenol can be oxidized to isoprenal by oxidative dehydrogenation by means of an oxygen-containing gas under catalysis, for example a supported copper, silver and/or gold catalyst, preferably a silver catalyst. The oxidation is carried out at a reaction temperature of 300 to 500 °C and results in a mixture of isoprenal (3-methyl-3-butenal) and prenal (3-methyl-2-butenal) and unreacted isoprenol. The mixture contains an excess of isoprenal, for example in a wt-ratio of 2:1 to 5:1 of isoprenal to prenal. Further details are provided in W02008037693.

Isomerization of Isoprenal to Prenal

Isomerization of isoprenal yields prenal.

Isoprenal is subjected to isomerization in order to produce additional prenal, wherein the isoprenal is preferably used in the form of the mixture obtained from the oxidation of isoprenol to isoprenal. The isomerization is carried out in the presence of an isomerization catalyst, preferably sodium acetate, at a temperature of 100 to 200 °C to produce a mixture of prenal and isoprenol. The mixture is fractionated to produce a prenal stream and an isoprenol stream. The isoprenol can be recycled. Further details are provided in W02008037693.

Reaction of Prenal with Prenol to Produce Citral

Conversion of prenol and prenal produces citral.

Prenol and prenal are first subjected to an acetalization to produce the diprenol acetal of prenal, 3-methyl-2-butenal-diprenylacetal. The acetalization is carried out under vacuum and acidic catalysis, for example a mineral acid, such as nitric acid or sulfuric acid, at a temperature up to 100 - 120 °C. The water formed during acetalization is continuously removed.

The acetal is then subjected to thermal cleavage in the presence of an acidic catalyst such as phosphoric acid, at a temperature of 150 - 170 °C to obtain cis/trans-prenyl-(3- methyl-butadienyl)ether. Under the reaction conditions this ether undergoes a Claisen and Cope rearrangement to give citral.

Further details are provided in W02008037693.

Additional Embodiments

The present invention also relates to processes for the manufacture of a chemical of interest selected from C4-olefin-derived chemicals according to the following embodiments.

1 . Process for the manufacture of a chemical of interest selected from C4-olefin-derived chemicals, said process comprising the steps of: a) subjecting a feedstock comprising a renewably-sourced ethanol to dehydration to produce a renewably-sourced ethylene stream; b) subjecting the renewably-sourced ethylene stream to an olefin-interconversion, to obtain one or more renewably-sourced C4-olefins, selected from n-butenes and isobutene; the olefin-interconversion comprising (i) and, where required, (ii):

(i) ethylene dimerization to obtain n-butenes;

(ii) isomerization of n-butenes obtained according to (i) to obtain isobutene; and c) subjecting the renewably-sourced C4-olefin to a chemical conversion or sequence of chemical conversions to obtain the chemical of interest.

2. Process according to embodiment 1 , wherein the renewably-sourced C4-olefin is n-butenes, wherein step c) comprises an oligomerization reaction of the n-butenes to produce at least one of dibutenes and tributenes.

3. Process according to embodiment 2, wherein step c) further comprises a hydroformylation reaction of the dibutenes to produce isononanal, and, optionally, a hydrogenation reaction of the isononanal to produce isononanol.

4. Process according to embodiment 2, wherein step c) further comprises a hydroformylation reaction of the tributenes to produce tridecanals and, optionally, a hydrogenation reaction of the tridecanals to produce tridecanols.

5. Process according to embodiment 1 , wherein the renewably-sourced C4-olefin is n-butenes, wherein step c) comprises a hydroformylation reaction of the n-butenes to produce valeraldehyde.

6. Process according to embodiment 5, wherein step c) further comprises a condensation reaction of the valeraldehyde to produce 2-propyl-2-heptenal and optionally, subjecting the 2-propyl-2-heptenal to a hydrogenation reaction to produce 2-propylheptanoL

7. Process according to embodiment 1 , wherein the renewably-sourced C4-olefin is isobutene, and step c) comprises an oxidation reaction to produce an intermediate selected from methacrolein and methacrylic acid.

8. Process according to embodiment 7, wherein step c) further comprises an esterification reaction of methacrylic acid to produce a methacrylic ester.

9. Process according to embodiment 1 , wherein the renewably-sourced C4-olefin is isobutene, and step c) comprises a polymerization reaction of isobutene, optionally together with one or more comonomers, to produce polyisobutene.

Dimerization and Trimerization of Butenes

In an embodiment, the renewably-sourced C4-olefin is n-butenes, wherein step c) comprises an oligomerization reaction of the n-butenes to produce at least one of dibutenes and tributenes. The oligomerization of straight-chain C4-olefins to their dimers, trimers and tetramers may be carried out by means of a fixed-bed catalyst, at superatmospheric pressure and at room temperature or elevated temperatures. The oligomerization is preferably carried out under supercritical conditions with respect to the starting material. For the oligomerization of n-butenes, reaction temperatures from 20 to 280 °C are preferred, more preferably above 160 °C and in particular from 180 to 210 °C.

The reaction pressure is, in general, from 20 to 300 bar, in particular from 60 to 80 bar. In one embodiment, the oligomerization is carried out at a pressure from 60 to 300 bar, in particular from 60 to 80 bar, and at a temperature from 160 to 280 °C, in particular from 180 to 210 °C. In another embodiment, the oligomerization is carried out at a pressure from 10 to 30 bar, in particular from 15 to 25 bar, and at a temperature from 20 to 140 °C, in particular from 40 to 120 °C.

The catalyst used typically contains, as active components, after deduction of the loss on ignition following heating at 900 °C, from 10 to 70% by weight of nickel oxide, calculated as NiO, from 5 to 30% by weight of titanium dioxide or zirconium dioxide, from 0 to 20% by weight of alumina, from 20 to 40% by weight of silica and from 0.01 to 1 wt.-% of an alkali metal oxide, wherein the contents of the individual components in the catalyst add up to 100 wt.-%.

Further details with regard to the oligomerization process and the catalyst used are provided in WO 95/014647.

Hydroformylation of Dibutenes and/or Tributenes

In an embodiment, step c) further comprises a hydroformylation reaction of the dibutenes to produce isononanal, and, optionally, a hydrogenation reaction of the isononanal to produce isononanol.

In another embodiment of this aspect, step c) further comprises a hydroformylation reaction of the tributenes to produce tridecanals, and, optionally, a hydrogenation reaction of the tridecanals to produce tridecanols.

Hydroformylation or the oxo process is an important large-scale industrial process for preparing aldehydes from olefins, carbon monoxide and hydrogen. These aldehydes can optionally 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.

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. However, Co-catalysts are preferred for hydroformylation of dibutenes and tributenes. If alcohols having a very low degree of branching are desired, the hydroformylation reaction may preferably be carried out using unmodified cobalt catalysts. 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. Further details may be taken from US 9,115,069, WO 2001/014297, WO 2021/197953 and US 6,015,928.

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.

The proportions of carbon monoxide, hydrogen, and dibutenes/tributenes in the hydroformylation reaction medium can be selected within a wide range. In some embodiments, based on the total amount of CO, hydrogen, and ethylene, CO is from about 1 to 50 mol-%, preferably about 1 to 35 mol-%; H ₂ is from about 1 to 98 mol-%, preferably about 10 to 90 mol-%; and ethylene is from about 0.1 to 35 mol-%, preferably about 1 to 35 mol-%.

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.

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.

The ratio of H2 to 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. In some modes of practice, syngas in the practice of the present invention is anhydrous.

The aldehydes obtained can optionally be hydrogenated with hydrogen in the same operation or subsequently in a separate hydrogenation step, to produce the corresponding alcohols. The aldehydes can be hydrogenated with hydrogen in the same reaction step, or subsequently in a separate hydrogenation step, to produce the corresponding alcohols. The hydrogenation of the aldehydes to their corresponding alcohols is a well-known reaction and can be conducted by any suitable known process.

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, 5 ^th edition, vol. A1 , 1984.

Hydroformylation of n-Butenes to Produce Valeraldehyde

In an embodiment, the renewably-sourced C4-olefin is n-butenes, wherein step c) comprises a hydroformylation reaction of the n-butenes to produce valeraldehyde.

The hydroformylation may generally be carried out in the manner customary and known to those skilled in the art. In the hydroformylation, valeraldehyde (n-pentanal) is prepared under transition metal catalysis from 1 -butene with addition of synthesis gas (CO:H2 of from 3:1 to 1 :3, preferably from 1.5:1 to 1 :1.15). The catalysts used for the hydroformylation reaction are generally rhodium complexes having phosphorus ligands. The phosphorus ligands are typically a mono- or diphosphine, preferably a triarylphosphine, more preferably triphenylphosphine. The hydroformylation is carried out typically at temperatures of from 50 to 150° C., preferably from 70 to 120° C, and pressures of from 5 to 50 bar, preferably from 10 to 30 bar.

Condensation of Valeraldehyde to Produce 2-Propyl-2-Heptenal

In an embodiment, step c) further comprises a condensation reaction of the valeraldehyde to produce 2-propyl-2-heptenal, and, optionally, subjecting the 2-propyl- 2-heptenal to a hydrogenation reaction to produce 2-propylheptanol.

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 P-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 valeraldehyde is reacted in a self-aldol condensation to obtain 2-propyl-2-heptenaL

Aldol condensations can occur under a variety of conditions under weak acidic or strong basic conditions and in the presence of various catalysts. In the present case a catalyst system comprising a secondary amine, preferably a di-Ci-Ce-alkylamine or a 4-, 5- or 6-membered cyclic mono or diamine, and an organic acid with up to 10 carbon atoms, preferably a mono or dicarboxylic acid, is preferably used. Examples for amines are dimethylamine, diethylamine, methyl ethyl amine, ethyl butyl amine, di-n-butylamine, di-2-ethylhexylamine, diisooctylamine, diphenylamine, dicyclohexylamine, piperidine, piperazine or morpholine or a combination thereof. Examples for organic acids are formic, acetic, propionic, malic, malonic, glutaric, tartaric, adipic, succinic, hydroxy succinic, maleic, 2-ethylhexanoic or salicylic acid or combinations thereof.

The secondary amine is preferably used in a molar ratio of 0.005:1 to 0.1 :1 to the valeraldehyde. The organic acid may be used in a molar ratio of 0.002:1 to 0.05:1 to valeraldehyde.

The reaction temperature may be in the range from 70 °C to 120 °C, preferably 80 °C to 100 °C, and the pressure may be from 100 kPa to 300 kPa, preferably 150 to 250 kPa.

Optionally, 2-propyl-2-heptenal may be subjected to a hydrogenation reaction to produce 2-propylheptanol, analogous to the hydrogenation of aldehydes derived from dibutenes and/or tributenes described above. Oxidation of Isobutene to Produce Methacrolein and Methacrylic Acid

In another aspect of the invention, the renewably-sourced C4-olefin is isobutene, and step c) comprises an oxidation reaction to produce an intermediate selected from methacrolein and methacrylic acid.

Suitable oxidation catalysts for oxidizing isobutene to methacrolein are mixed metal oxide catalysts well known in the art.

A catalyst suitable for industrial production of methacrylic acid from isobutylene in high yield has not yet been found. Accordingly, it is industrially advantageous to conduct the reaction in two steps using a catalyst for production of methacrolein from isobutylene and a catalyst for production of methacrylic acid from methacrolein.

However, when methacrylic acid is produced in such a two step oxidation, it is relatively difficult to attain high selectivity of methacrylic acid in the second oxidation step if the conversion of methacrolein is high. That is, the selectivity of methacrylic acid formation decreases with increasing conversion of methacrolein. Accordingly, it becomes advantageous to conduct the second oxidation step under conditions such that the conversion of methacrolein is relatively low and then to recover the unreacted methacrolein from the reaction products and recycle them back into the second oxidation step. The unreacted methacrolein is separated from the reaction products in the second oxidation step while the object product, methacrylic acid is recovered. This is accomplished by cooling the reaction products in the second oxidation step to liquefy them, and/or further contacting them with water to separate the gaseous components of oxygen, nitrogen, carbon dioxide gas and the like. As a result, a liquid phase containing methacrolein and/or methacrylic acid, is obtained and these are separated by distillation.

Esterification of Methacrylic Acid to Produce a Methacrylic Ester

In an embodiment, step c) further comprises an esterification reaction of methacrylic acid to produce a methacrylic ester.

(Meth)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. The methacrylic acid can be esterified in a conventional manner to produce the desired methacrylic acid ester using the corresponding alkanol such as methanol, ethanol, n-propanol, isopropanol, n-butanol and 2-ethylhexanoL

Processes for the preparation of alkyl methacrylates by reacting methacrylic 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 methacrylic 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.

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. Methacrylic acid, the alkanol and the catalyst are fed continuously to the reaction zone. An azeotropic mixture comprising the alkyl (meth)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 (meth)acrylate and an aqueous phase, with a part of the organic phase being recycled to the reaction zone. The alkyl (meth)acrylate is isolated from the excess organic phase. The latter is usually carried out by separation steps involving rectification (cf. for example DE 19536178).

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.

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.

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

Polymerization of Isobutene

In an embodiment, the renewable C4-olefin is isobutene, and step c) comprises a polymerization reaction of isobutene, optionally together with one or more comonomers, to produce polyisobutene.

The polymerization of isobutene by means of various initiators has been disclosed, for example in "High Polymers", volume XXIV (H. Wiley & Sons, Inc. New York, 1971), pages 713 f.

In a preferred process as described in US 4,152,499, isobutene is polymerized with boron trifluoride as the initiator, while the polymerization is carried out at from -50 °C to 30 °C. Typically, 1 to 20 mmoles of boron trifluoride are used per mole of isobutene. The mean polymerization time is generally confined to from 1 to 10 minutes.

The polymerization can also be accelerated by the conventional method used for cationic polymerization, i.e. by using co-catalysts, e.g. water or alcohols. The amount of such compounds is usually from 0.2 to at most 1 .0 mole %, based on the amount of the boron trifluoride.

The polymerization may be carried out by introducing gaseous boron trifluoride, with or without addition of the co-catalyst, batchwise, semi-continuously or completely continuously, in the conventional manner, into isobutene cooled to from -50 °C to +30 °C, whilst cooling the mixture efficiently and mixing it thoroughly. The polymerization may then be stopped, again in the conventional manner, by adding alcohols, e.g., methanol, or aqueous or alcoholic alkali metal hydroxide solutions. The catalyst residues may then be filtered off or adsorbed on adsorbents, e.g. aluminum oxide, or extracted with water or alcohol. Solvents, monomers and low molecular weight oligomers are advantageously removed by flash distillation.

To manufacture a petroleum additive, the isobutene polymer is reacted with the stoichiometric amount of maleic anhydride, or a slight excess thereof, in the conventional manner at from 170 to 250 °C. These adducts, in turn, are converted in the conventional manner into lubricating oil additives by reacting them with amines, in particular polyamines such as diethylenetriamine and triethylenetetramine.