EP2655300A1

EP2655300A1 - Production of fuel additives via simultaneous dehydration and skeletal isomerisation of isobutanol on acid catalysts followed by etherification

Info

Publication number: EP2655300A1
Application number: EP11811537.7A
Authority: EP
Inventors: Cindy Adam; Delphine Minoux; Nikolai Nesterenko
Original assignee: Total Raffinage Marketing SA
Current assignee: TotalEnergies Marketing Services SA
Priority date: 2010-12-21
Filing date: 2011-12-20
Publication date: 2013-10-30
Also published as: FR2969147B1; US20130261345A1; WO2012084950A1; BR112013011269A2; FR2969147A1; CN103261129A

Abstract

The present invention relates to a process for the production of fuel additives in which in a first step isobutanol is subjected to a simultaneous dehydration and skeletal isomerisation to make substantially corresponding olefins, having the same number of carbons and consisting essentially of a mixture of n-butenes and isobutene and in a second step the butene mixture is subjected to etherification, said process comprising: a)introducing in at least one reactor a stream (A) comprising at least 40wt% isobutanol, optionally an inert component, b)contacting said stream with at least one catalyst in said reactor(s) at conditions effective to simultaneously dehydrate and skeletal isomerise at least a portion of the isobutanol to make a mixture of n-butenes and iso-butene, c)removing the inert component if any, recovering from said reactor(s) a stream (B) comprising a mixture of n-butenes and isobutene, d)sending the stream (B) to at least one etherification reactor and contacting stream (B) with at least one catalyst in said etherification reactor(s), in the presence of ethanol and/or methanol, at conditions effective to produce ETBE and/or MTBE respectively, e)recovering from said etherification reactor a stream (E) comprising essentially ETBE and/or MTBE, unreacted butenes, heavies, optionally unreacted ethanol and/or methanol respectively, f)fractionating stream (E) to recover ETBE and/or MTBE.

Description

PRODUCTION OF FUEL ADDITIVES VIA SIMULTANEOUS DEHYDRATION AND SKELETAL ISOMERISATION OF ISOBUTANOL ON ACID CATALYSTS FOLLOWED BY ETHERIFI CATION

[Field of the invention]

The present invention relates to the production of fuel additives from renewable energy source via simultaneous dehydration and skeletal isomerisation of isobutanol to make a corresponding olefin, having substantially the same number of carbons but different skeleton structure, followed by an etherification step. Isobutanol can be obtained by fermentation of carbohydrates or by condensation of lighter alcohols, obtained by fermentation of carbohydrates. Made up of organic matter from living organisms, biomass is the world's leading renewable energy source.

[Background of the invention] Isobutanol (2 -methyl- 1-propanol) has historically found limited applications and its use resembles that of 1-butanol. It has been used as solvent, diluent, wetting agent, cleaner additive and as additive for inks and polymers. Recently, isobutanol has gained interest as fuel or fuel component as it exhibits a high octane number (Blend Octane R+M/2 is 102- 103 (R stands for Research and M for Motor- these numbers being used to calculate octane number)) and a low vapor pressure (RVP- Reid Vapor Pressure- is 3.8-5.2 psi).

Isobutanol is often considered as a byproduct of the industrial production of 1-butanol (Ullmann's encyclopedia of industrial chemistry, 6^th edition, 2002). It is produced from propylene via hydroformylation in the oxo-process (Rh-based catalyst) or via carbonylation in the Reppe-process (Co-based catalyst). Hydroformylation or carbonylation makes n-butanal and iso-butanal in ratios going from 92/8 to 75/25. To obtain isobutanol, the iso-butanal is hydrogenated over a metal catalyst. Isobutanol can also be produced from synthesis gas (mixture of CO, ¾ and CO2) by a process similar to Fischer-Tropsch, resulting in a mixture of higher alcohols, although often a preferential formation of isobutanol occurs (Applied Catalysis A, general, 186, p. 407, 1999 and Chemiker Zeitung, 106, p. 249, 1982). Still another route to obtain isobutanol, is the base-catalysed Guerbet condensation of methanol with ethanol and/ or propanol (J. of Molecular Catalysis A: Chemical 200, 137, 2003 and Applied Biochemistry and Biotechnology, 1 13- 1 16, p. 913, 2004).

Recently, new biochemical routes have been developed to produce selectively isobutanol from carbohydrates. The new strategy uses the highly active amino acid biosynthetic pathway of microorganisms and diverts its 2-keto acid intermediates for alcohol synthesis. 2-Keto acids are intermediates in amino acid biosynthesis pathways. These metabolites can be converted to aldehydes by 2-keto- acid decarboxylases (KDCs) and then to alcohols by alcohol dehydrogenases (ADHs). Two steps are required to produce alcohols by shunting intermediates from amino acid biosynthesis pathways to alcohol production (Nature, 451, p. 86, 2008 and US patent

2008/0261230). Recombinant microorganisms are required to enhance the flux of carbon towards the synthesis of 2-keto-acids. In the valine biosynthesis 2-ketoisovalerate is an intermediate. Glycolyse of carbohydrates results in pyruvate that is converted into acetolactate by acetolactate synthase. 2,4-dihydroxyisovalerate is formed out of acetolactate, catalysed by isomero reductase. A dehydratase converts the 2,4-dihydroxyisovalerate into 2-keto-isovalerate. In the next step, a keto acid decarboxylase makes isobutyraldehyde from 2-keto- isovalerate. The last step is the hydrogenation of isobutyraldehyde by a dehydrogenase into isobutanol.

Of the described routes towards isobutanol above, the Guerbet condensation, the synthesis gas hydrogenation and the 2-keto acid pathway from carbohydrates are routes that can use biomass as primary feedstock. Gasification of biomass results in synthesis gas that can be converted into methanol or directly into isobutanol. Ethanol is already at very large scale produced by fermentation of carbohydrates or via direct fermentation of synthesis gas into ethanol. So methanol and ethanol resourced from biomass can be further condensed to isobutanol. The direct 2-keto acid pathway can produce isobutanol from carbohydrates that are isolated from biomass. Simple carbohydrates can be obtained from plants like sugar cane, sugar beet. More complex carbohydrates can be obtained from plants like maize, wheat and other grain bearing plants. Even more complex carbohydrates can be isolated from substantially any biomass, through unlocking of cellulose and hemicellulose from lignocelluloses. In the mid nineties, many petroleum companies attempted to produce more isobutene for the production of MTBE (methyl ieri-butyl ether). Hence many skeletal isomerisation catalysts for the conversion of n-butenes into iso-butene have been developed (Adv. Catal. 44, p. 505, 1999; Oil & Gas Science and Technology, 54 ( 1) p. 23, 1999 and Applied Catalysis A: General 212, 97, 2001). Among promising catalysts are 10-membered ring zeolites and modified alumina's. The reverse skeletal isomerisation of iso-butene into n-butenes has not been mentioned.

The dehydration reactions of alcohols to produce alkenes have been known for a long time (J. Catal. 7, p. 163, 1967 and J. Am.

Chem. Soc. 83, p. 2847, 1961). Many available solid acid catalysts can be used for alcohol dehydration (Stud. Surf. Sci. Catal. 51, p. 260, 1989). However, γ-aluminas are the most commonly used, especially for the longer chain alcohols (with more than three carbon atoms) . This is because catalysts with stronger acidity, such as the silica-aluminas, molecular sieves, zeolites or resin catalysts can promote double-bond shift, skeletal isomerization and other olefin inter-conversion reactions. The primary product of the acid-catalysed dehydration of isobutanol is iso-butene:

CH₃-CH-CH₂-OH CH₃-C=CH₂ + H₂0

I I

CH₃ CH₃

The dehydration of alcohols with four or more carbons over solid acid catalysts is expected to be accompanied by the double-bond shift reaction of the alkene product. This is because the two reactions occur readily and at comparable rates (Carboniogenic Activity of Zeolites, Elsevier Scientific Publishing Company, Amsterdam ( 1977) p. 169). The primary product, iso-butene is very reactive in presence of acid catalyst because of the presence of a double bond linked to a tertiary carbon. This allows easy protonation, as the tertiary structure of the resulting carbocation is the most favourable one among the possible carbocation structures (tertiary > secondary > primary carbocations) . The resulting t-butyl-cation undergoes easy oligo/ polymerisation or other electrophilic substitution on aromatics or aliphatics or electrophilic addition reactions. The rearrangement of t-butyl-cation is not a straightforward reaction as, without willing to be bound to any theory, it involves an intermediate formation of secondary or primary butyl- cation and hence the probability of secondary reactions (substitutions or additions) is very high and would reduce the selectivity for the desired product.

Dehydration of butanols has been described on alumina-type catalysts (Applied Catalysis A, General, 214, p. 251, 2001). Both double-bond shift and skeletal isomerisation has been obtained at very low space velocity (or very long reaction time) corresponding to a GHSV (Gas Hourly Space Velocity = ratio of feed rate (gram/h) to weight of catalyst (ml)) of less than 1 gram. ml ¹. h ¹. The US7473812 patent discloses a process to remove iso-butene from a butenes mixture by a process for coproducing butene oligomers and tert-butyl ethers by partly oligomerizing the iso-butene over an acidic catalyst to give butene oligomers and subsequently etherifying the remaining isobutene with an alcohol under acidic catalysis to give tert-butyl ethers.

US446991 1 discloses a process for isobutene oligomerization in the presence of a fixed bed cation exchange resin at a temperature in the range from 30° to 60°C and a LHSV of from 2.5 to 12 tr¹.

US5895830 describes an enhanced dimer selectivity of a butene oligomerization process using SPA (supported phosphoric acid) catalyst, by diluting the butene feed with a heavy saturate stream comprising paraffins having a carbon number of at least 8.

US5877372 discloses dimerization of isobutene in the presence of isooctane diluent and tert-butyl alcohol (at least 1 wt-% and preferably 5 to 15 wt-%), over a sulfonic acid type ion exchange resin such as Amberlyst A- 15, Dowex 50 or the like, at temperatures in the range 10° to 200° C. and pressures in the range of 50 to 500 psig. It is suggested that tert-butyl alcohol improves the selectivity of dimer formation and reduces the formation of trimer and higher oligomers.

US6689927 describes a low temperature butene oligomerization process having improved selectivity for dimerization and improved selectivity for the preferred 2,4,4-trimethylpentene isomer, caused by carrying out oligomerization in the presence of an SPA catalyst at a temperature below 1 12°C in the presence of a saturated hydrocarbon diluent having a carbon number of at least 6.

WO2007/ 14399 and WO2008/016428 discloses the conversion of isobutanol derived from fermentation broth into butenes, which can be converted into isoalkanes, alkyl-substituted aromatics, isooctanols, and octyl ethers, and may be used in transportation fuel. US 2008/0220488 describes a process for the production of isooctenes as fuel additives, using dry 1-butanol derived from fermentation broth. From examples, dehydration of alcohol and dimerization are performed in one-step using acidic catalysts.

It has now been discovered that the dehydration of isobutanol and the skeletal isomerisation of the iso-butyl moiety of isobutanol can be carried out simultaneously and that the resulting mixture of iso- butene and n-butenes, can be efficiently used in the etherification to produce MTBE/ETBE. Optionally, the residual butenes (essentially n- butenes) can be valorized into heavies (gasoline, middle distillates) by either alkylation (Polyalkylate formation) or oligomerization.

By way of example it has been discovered that for the simultaneous dehydration and skeletal isomerisation of isobutanol, crystalline silicates of the group FER, MWW, EUO, MFS, ZSM-48, MTT, MFI, MEL or TON having Si/Al higher than 10,

or a dealuminated crystalline silicate of the group FER, MWW, EUO, MFS, ZSM-48, MTT, MFI, MEL or TON having Si/Al higher than 10,

or a phosphorus modified crystalline silicate of the group FER, MWW, EUO, MFS, ZSM-48, MTT, MFI, MEL or TON having Si/Al higher than 10,

or molecular sieves of the type silicoaluminophosphate of the group AEL

or silicated, zirconated, titanated or fluorinated alumina's, have many advantages.

Said dehydration can be made with a WHSV (Weight Hourly Space Velocity = ratio of feed flow rate (gram/h) to catalyst weight) of at least 1 h ¹ , at a temperature from 200 to 600°C and using a isobutanol-diluent composition from 30 to 100wt% isobutanol at a total operating pressure from 0.05 to 1.0 MPa.

By way of example, in the dehydration/ isomerisation of isobutanol on a ferrierite having a Si/Al ratio from 10 to 90 and with a WHSV of at least 2 h ¹ to make n-butenes beside iso-butene, the isobutanol conversion is at least 98% and often 99%, advantageously the butenes (iso and n-butenes) yield is at least 90%, the n-butenes selectivity is between 5% and the thermodynamic equilibrium at the given reaction conditions.

The isobutanol conversion is the ratio (isobutanol introduced in the reactor - isobutanol leaving the reactor) / (isobutanol introduced in the reactor).

The n-butenes yield is the ratio, on carbon basis, (n-butenes leaving the reactor) / (isobutanol introduced in the reactor) .

The n-butenes selectivity is the ratio, on carbon basis, (n- butenes leaving the reactor) / (isobutanol converted in the reactor) .

The simultaneous dehydration/ isomerisation of isobutanol results in a mixture of n-butenes (but- l-ene and but-2-ene) and iso- butene (2-methyl propene). According to the present invention, often a composition close to thermodynamic equilibrium is obtained while maintaining the high yield of total butenes. The thermodynamic equilibrium for n-butenes varies between 50 and 65% and for iso- butene between 35 and 50% depending on operating conditions. An important advantage of the present invention is that the composition of obtained products resembles the composition of a raffinate I C4 cut obtained from a steam naphtha cracker. Raffinate I is obtained by removing 1,3-butadiene from the raw C4 cut produced on a steam naphtha cracker. Typical compositions of raffinate I are: 35-45wt% isobutene, 3- 15wt% butanes and the remaining 52-40wt% n-butenes.

Said product from the simultaneous dehydration/ isomerisation of isobutanol according to the invention can readily replace the use of raffinate I in existing petrochemical plants. The result is that capital investment can be minimised and that the derivatives from such iso- butene/n-butenes mixture can hence be produced from renewable resources instead of fossil resources simply by substituting fossil raffinate I by the product of the present invention.

[Brief summary of the invention]

The present invention relates to a process for the production of fuel additives (ETBE/ MTBE, gasoline, middle distillates) for example from renewable energy sources in which in a first step isobutanol is subjected to a simultaneous dehydration and skeletal isomerisation to make substantially corresponding olefins, having the same number of carbons and consisting essentially of a mixture of n- butenes and iso-butene and in a second step, the mixture of butenes are subjected to an etherification, said process comprising :

a) introducing in at least one reactor a stream (A) comprising at least 40wt% of isobutanol, optionally an inert component,

b) contacting said stream (A) with at least one catalyst in said reactor(s) at conditions effective to simultaneously dehydrate and skeletal isomerise at least a portion of the isobutanol to make a mixture of n-butenes and iso-butene,

c) removing the inert component if any, recovering from said reactor(s) a stream (B) comprising a mixture of n-butenes and iso- butene,

d) sending the stream (B) to at least one etherification reactor and contacting stream (B) with at least one catalyst in said etherification reactor(s), in the presence of ethanol and/ or methanol at conditions effective to produce ETBE and/or MTBE,

e) recovering from said etherification reactor(s) a stream (E) comprising essentially ETBE and/or MTBE, unreacted butenes (mainly n-butenes), optionally heavies, optionally unreacted ethanol and/ or methanol,

f) fractionating stream (E) to recover ETBE and/ or MTBE and optionally recycling unreacted butenes and unreacted ethanol and / or methanol to the etherification reactor.

Advantageously, stream (A) used in step a), and/ or methanol and/or ethanol used in step d) are issued from renewable energy sources.

Optionally, subjecting stream (A) is subjected to a purification treatment before step b) . This purification treatment is intended to remove contaminants, such as oxygenates, for example by usual separation processes (distillation, adsorption,...)

In a first embodiment the WHSV of the isobutanol is at least

1 h ¹.

In a second embodiment the temperature of the simultaneous dehydration and skeletal isomerisation of isobutanol ranges from

200°C to 600°C.

Advantageously the dehydration/ isomerization catalyst is a crystalline silicate of the group FER, MWW, EUO, MFS, ZSM-48, MTT, MFI, MEL or TON having Si/Al higher than 10,

or a silicoaluminaphosphate molecular sieve of the group AEL, or a silicated, zirconated or titanated or fluorinated alumina.

It would not depart from the scope of the invention if the isobutanol feedstock comprises one or more of the other C4 alcohols such as 2-butanol, tertiary-butanol and n-butanol.

In the process of the invention, isobutanol is a major component in the feedstock (stream A). Notably, stream (A) may contain other C4 alcohols.

This means the ratio of isobutanol to all other components of stream (A), for example to the C4 alcohols in the feedstock is of 40wt% or more. More advantageously the previous ratio is 70% or more and preferably 80% or more.

If the proportion of isobutanol is too low, the invention is of low interest, as the dehydration should then better be done without occurrence of skeletal isomerisation and there are a lot of catalysts in the prior art capable to dehydrate isobutanol, 2-butanol and n-butanol to produce the corresponding butenes. Dehydration of tertiary-butanol to isobutene followed by a skeletal isomerisation of at least a part of the tertiary-butanol is described in WO 20051 10951.

Optionally, stream (E) recovered from step e) is fractionated to recover unreacted butenes and/ or unreacted ethanol and/ or methanol.

Then, at least a part of said recovered unreacted butenes and/or at least a part of said recovered unreacted ethanol and/ or methanol may be recycled to the etherification reactor(s) .

Alternatively or in combination, at least a part of said recovered unreacted butenes may be sent to a purification zone before being sent to at least one oligomerization reactor and / or to at least one alkylation reactor to produce heavies, ie hycrocarbonaceous compounds of more than 4 carbon atoms, such as iso-octenes, middle distillates or Polyalkylate. Advantageously, said unreacted butenes may be subjected to a purification step before being sent to oligomerization reactor(s) and/ or alkylation reactor(s).

In a specific embodiment, the residual butene (mainly n-butene) coming out from the etherification reactor(s) is selectively transformed in an easy separable product (iso-octenes, middle distillates) via oligomerization or alkylation.

[Detailed description of the invention]

As regards the stream (A), the isobutanol may be subjected to simultaneous dehydration and skeletal isomerisation alone or in mixture with an inert medium. The inert component is any component provided it is substantially not converted on the catalyst. Because the dehydration step is endothermic the inert component can be used as energy vector. The inert component allows reducing the partial pressure of the isobutanol and other reaction intermediates and will hence reduce secondary reactions like oligo/ polymerisation. The inert component may be selected among water, nitrogen, hydrogen, CO2 and saturated hydrocarbons. It may be such that some inert components are already present in the isobutanol because they were used or co- produced during the production of isobutanol. Examples of inert components that may already be present in the isobutanol are water and CO2. The inert component may be selected among the saturated hydrocarbons having up to 10 carbon atoms, naphtenes.

Advantageously it is a saturated hydrocarbon or a mixture of saturated hydrocarbons having from 3 to 7 carbon atoms, more advantageously having from 4 to 6 carbon atoms and is preferably pentane. An example of inert component can be any individual saturated compound, a synthetic mixture of the individual saturated compounds as well as some equilibrated refinery streams like straight naphtha, butanes etc. Advantageously the inert component is a saturated hydrocarbon having from 3 to 6 carbon atoms and is preferably pentane. The weight proportions of respectively isobutanol and inert component are, for example, 30- 100/70-0 (the total being 100). The stream (A) can be liquid or gaseous.

As regards the reactor for the simultaneous dehydration/ isomerisation, it can be a fixed bed reactor, a moving bed reactor or a fluidized bed reactor. A typical fluid bed reactor is one of the FCC type used for fluidized-bed catalytic cracking in the oil refinery. A typical moving bed reactor is of the continuous catalytic reforming type. The simultaneous dehydration/ isomerisation may be performed continuously in a fixed bed reactor configuration using a pair of parallel "swing" reactors. This enables the dehydration process to be performed continuously in two parallel "swing" reactors wherein when one reactor is operating, the other reactor is undergoing catalyst regeneration. The catalyst of the present invention also can be regenerated several times.

The simultaneous dehydration/ isomerisation may be performed continuously in a moving bed reactor in which the catalyst circulates from a reaction zone to a regeneration zone and backwards with a residence time of the catalyst in the reaction zone of at least 12 hours. In each zone the catalyst behaves substantially like in a fixed bed reactor, but the catalyst moves slowly, by gravity or pneumatically through the respective zone. The use of a moving bed reaction allows accomplishing a continuous operation with no switching of the feedstock and regeneration gas from one reactor to another one. The reaction zone receives continuously the feedstock while the regeneration zone receives continuously the regeneration gas.

The simultaneous dehydration/ isomerisation may be performed continuously in a fluidised bed reactor in which the catalyst circulates from a reaction zone to a regeneration zone and backwards with a residence time of the catalyst in the reaction zone of less than 12 hours. In each zone the catalyst is in a fluidised state and exhibit such a shape and size that it remains fluidised in the flow of the feedstock and reaction products or regeneration gas. The use of a fluidised bed reactor allows regenerating very rapidly deactivated catalyst by regeneration in the regeneration zone.

As regards the pressure for the simultaneous dehydration/ isomerisation, it can be any pressure but it is more easy and economical to operate at moderate pressure. By way of example the pressure of the reactor ranges from 0.5 to 10 bars absolute (50 kPa to 1 MPa), advantageously from 0.5 to 5 bars absolute (50 kPa to 0.5 MPa), more advantageously from 1.2 to 5 bars absolute (0. 12 MPa to 0.5 MPa) and preferably from 1.2 to 4 bars absolute (0. 12 MPa to 0.4 MPa). Advantageously the partial pressure of the isobutanol is from 0. 1 to 4 bars absolute (0.01 MPa to 0.4 MPa), more advantageously from

0.5 to 3.5 bars absolute (0.05 MPa to 0.35 MPa).

As regards the temperature for the simultaneous dehydration/ isomerisation, it ranges from 200°C to 600°C, advantageously from 250°C to 500°C, more advantageously from 300°C to 450°C.

These reaction temperatures refer substantially to average catalyst bed temperature. The isobutanol dehydration is an endothermic reaction and requires the input of reaction heat in order to maintain catalyst activity sufficiently high and shift the dehydration thermodynamic equilibrium to sufficiently high conversion levels.

In case of fluidised bed reactors: (i) for stationary fluidised beds without catalyst circulation, the reaction temperature is substantially homogeneous throughout the catalyst bed; (ii) in case of circulating fluidised beds where catalyst circulates between a converting reaction section and a catalyst regeneration section, depending on the degree of catalyst backmixing the temperature in the catalyst bed approaches homogeneous conditions (a lot of backmixing) or approaches plug flow conditions (nearly no backmixing) and hence a decreasing temperature profile will install as the conversion proceeds.

In case of fixed bed or moving bed reactors, a decreasing temperature profile will install as the conversion of the isobutanol proceeds. In order to compensate for temperature drop and consequently decreasing catalyst activity or approach to thermodynamic equilibrium, reaction heat can be introduced by using several catalyst beds in series with interheating of the reactor effluent from the first bed to higher temperatures and introducing the heated effluent in a second catalyst bed, etc. When fixed bed reactors are used, a multi-tubular reactor can be used where the catalyst is loaded in small-diameter tubes that are installed in a reactor shell. At the shell side, a heating medium is introduced that provides the required reaction heat by heat-transfer through the wall of the reactor tubes to the catalyst.

As regards the WHSV of the isobutanol for the simultaneous dehydration/ isomerisation, it ranges advantageously from 1 to 30 tr ^l, more advantageously from 2 to 21 h ¹, preferably from 5 to 15 h ¹, more preferably from 7 to 12 h ¹.

As regards the stream (B) from the simultaneous dehydration/ isomerisation, it comprises essentially water, olefin, the inert component (if any) and unconverted isobutanol. Said unconverted isobutanol is supposed to be as less as possible. The olefin is recovered by usual fractionation means. Advantageously the inert component, if any, is recycled in the stream (A) as well as the unconverted isobutanol, if any. Unconverted isobutanol, if any, is recycled to the reactor in the stream (A) . As regards the catalyst for the simultaneous dehydration/ isomerisation, advantageously it is a crystalline silicate of the group FER(ferrierite, FU-9, ZSM-35), MWW (MCM-22, PSH-3, ITQ- 1, MCM- 49), EUO(ZSM-50, EU- 1), MFS(ZSM-57), ZSM-48, MTT(ZSM-23) , MFI (ZSM-5), MEL (ZSM- 1 1) or TON (ZSM-22, Theta- 1, NU- 10),

or a dealuminated crystalline silicate of the group FER (ferrierite, FU-9, ZSM-35), MWW (MCM-22, PSH-3, ITQ- 1, MCM-49), EUO (ZSM- 50, EU- 1), MFS (ZSM-57), ZSM-48, MTT (ZSM-23) , MFI (ZSM-5), MEL (ZSM- 1 1) or TON (ZSM-22, Theta- 1, NU- 10),

or a phosphorus modified crystalline silicate of the group FER

(ferrierite, FU-9, ZSM-35), MWW (MCM-22, PSH-3, ITQ- 1, MCM-49), EUO (ZSM-50, EU- 1), MFS (ZSM-57), ZSM-48, MTT (ZSM-23) , MFI (ZSM-5), MEL (ZSM- 1 1) or TON (ZSM-22, Theta- 1, NU- 10),

or a silicoaluminophosphate molecular sieve of the group AEL (SAPO- 1 1),

or a silicated, zirconated or titanated or fluorinated alumina.

About the crystalline silicate of FER structure (ferrierite, FU-9, ZSM-35) it can be the lamellar precursor which becomes FER by calcinations.

The Si/Al ratio of the crystalline silicate is advantageously higher than 10.

The crystalline silicate is such as the Si/Al ratio ranges more advantageously from 10 to 500, preferably from 12 to 250, more preferably from 15 to 150.

The acidity of the catalyst can be determined by the amount of residual ammonia on the catalyst following contact of the catalyst with ammonia which adsorbs on the acid sites of the catalyst with subsequent ammonium desorption at elevated temperature measured by differential thermogravimetric analysis or analysis of ammonia concentration in the desorbed gases.

The crystalline silicate can be subjected to various treatments before use in the dehydration including, ion exchange, modification with metals (in a not restrictive manner alkali, alkali-earth, transition, or rare earth elements), external surface passivation, modification with P-compounds, steaming, acid treatment or other dealumination methods, or combination thereof.

In a specific embodiment the crystalline silicate is steamed to remove aluminium from the crystalline silicate framework. The steam treatment is conducted at elevated temperature, preferably in the range of from 425 to 870°C, more preferably in the range of from 540 to 815°C and at atmospheric pressure and at a water partial pressure of from 13 to 200kPa. Preferably, the steam treatment is conducted in an atmosphere comprising from 5 to 100 vol% steam. The steam atmosphere preferably contains from 5 to 100 vol% steam with from 0 to 95 vol% of an inert gas, preferably nitrogen. The steam treatment is preferably carried out for a period of from 1 to 200 hours, more preferably from 4 hours to 10 hours. As stated above, the steam treatment tends to reduce the amount of tetrahedral aluminium in the crystalline silicate framework, by forming alumina. In a more specific embodiment the crystalline silicate is dealuminated by heating the catalyst in steam to remove aluminium from the crystalline silicate framework and extracting aluminium from the catalyst by contacting the catalyst with a complexing agent for aluminium to remove from pores of the framework alumina deposited therein during the steaming step thereby to increase the silicon/ aluminium atomic ratio of the catalyst. Advantageously, the commercially available crystalline silicate is modified by a steaming process which reduces the tetrahedral aluminium in the crystalline silicate framework and converts the aluminium atoms into octahedral aluminium in the form of amorphous alumina. Although in the steaming step aluminium atoms are chemically removed from the crystalline silicate framework structure to form alumina particles, those particles cause partial obstruction of the pores or channels in the framework. This could inhibit the dehydration process of the present invention. Accordingly, following the steaming step, the crystalline silicate is subjected to an extraction step wherein amorphous alumina is removed from the pores and the micropore volume is, at least partially, recovered. The physical removal, by a leaching step, of the amorphous alumina from the pores by the formation of a water-soluble aluminium complex yields the overall effect of de-alumination of the crystalline silicate. In this way by removing aluminium from the crystalline silicate framework and then removing alumina formed therefrom from the pores, the process aims at achieving a substantially homogeneous de-alumination throughout the whole pore surfaces of the catalyst. This reduces the acidity of the catalyst. The reduction of acidity ideally occurs substantially homogeneously throughout the pores defined in the crystalline silicate framework. Following the steam treatment, the extraction process is performed in order to de-aluminate the catalyst by leaching. The aluminium is preferably extracted from the crystalline silicate by a complexing agent which tends to form a soluble complex with alumina. The complexing agent is preferably in an aqueous solution thereof. The complexing agent may comprise an organic acid such as citric acid, formic acid, oxalic acid, tartaric acid, malonic acid, succinic acid, glutaric acid, adipic acid, maleic acid, phthalic acid, isophthalic acid, fumaric acid, nitrilotriacetic acid, hydroxyethylenediaminetriacetic acid, ethylenediaminetetracetic acid, trichloroacetic acid trifluoroacetic acid or a salt of such an acid (e.g. the sodium salt) or a mixture of two or more of such acids or salts. The complexing agent may comprise an inorganic acid such as nitric acid, halogenic acids, sulphuric acid, phosphoric acid or salts of such acids or a mixture of such acids. The complexing agent may also comprise a mixture of such organic and inorganic acids or their corresponding salts. The complexing agent for aluminium preferably forms a water-soluble complex with aluminium, and in particular removes alumina which is formed during the steam treatment step from the crystalline silicate.

Following the aluminium leaching step, the crystalline silicate may be subsequently washed, for example with distilled water, and then dried, preferably at an elevated temperature, for example around 1 10°C.

Additionally, if during the preparation of the catalysts used in the present invention alkaline or alkaline earth metals have been used, the molecular sieve might be subjected to an ion-exchange step.

Conventionally, ion-exchange is done in aqueous solutions using ammonium salts or inorganic acids.

Following the de-alumination step, the catalyst is thereafter calcined, for example at a temperature of from 400 to 800°C at atmospheric pressure for a period of from 1 to 10 hours.

Another suitable catalyst for the present process is the silicoaluminophosphate molecular sieves of the AEL group with typical example the SAPO- 1 1 molecular sieve. The SAPO- 1 1 molecular sieve is based on the ALPO- 1 1, having essentially an Al/P ratio of 1 atom/ atom. During the synthesis silicon precursor is added and insertion of silicon in the ALPO framework results in an acid site at the surface of the micropores of the 10-membered ring sieve. The silicon content ranges from 0.1 to 10 atom% (Al + P + Si is 100). In another specific embodiment the crystalline silicate or silicoaluminophosphate molecular sieve is mixed with a binder, preferably an inorganic binder, and shaped to a desired shape, e.g. pellets. The binder is selected so as to be resistant to the temperature and other conditions employed in the dehydration process of the invention. The binder is an inorganic material selected from clays, silica, metal silicates, metal oxides such as ZrC and/or metals, or gels including mixtures of silica and metal oxides. If the binder which is used in conjunction with the crystalline silicate is itself catalytically active, this may alter the conversion and/or the selectivity of the catalyst. Inactive materials for the binder may suitably serve as diluents to control the amount of conversion so that products can be obtained economically and orderly without employing other means for controlling the reaction rate. It is desirable to provide a catalyst having a good crush strength. This is because in commercial use, it is desirable to prevent the catalyst from breaking down into powder-like materials. Such clay or oxide binders have been employed normally only for the purpose of improving the crush strength of the catalyst. A particularly preferred binder for a catalyst used in the present invention comprises silica. The relative proportions of the finely divided crystalline silicate material and the inorganic oxide matrix of the binder can vary widely. Typically, the binder content ranges from 5 to 95% by weight, more typically from 20 to 75% by weight, based on the weight of the composite catalyst. Such a mixture of the crystalline silicate and an inorganic oxide binder is referred to as a formulated crystalline silicate. In mixing the catalyst with a binder, the catalyst may be formulated into pellets, extruded into other shapes, or formed into spheres or a spray-dried powder. Typically, the binder and the crystalline silicate are mixed together by a mixing process. In such a process, the binder, for example silica, in the form of a gel is mixed with the crystalline silicate material and the resultant mixture is extruded into the desired shape, for example cylindrical or multi-lobe bars. Spherical shapes can be made in rotating granulators or by oil- drop technique. Small spheres can further be made by spray-drying a catalyst-binder suspension. Thereafter, the formulated crystalline silicate is calcined in air or an inert gas, typically at a temperature of from 200 to 900°C for a period of from 1 to 48 hours.

In addition, the mixing of the catalyst with the binder may be carried out either before or after the steaming and extraction steps. Another family of suitable catalysts for the simultaneous dehydration and skeletal isomerisation are alumina's that have been modified by surface treatment with silicon, zirconium, titanium or fluor. Alumina's are generally characterised by a rather broad acid strength distribution and having both Lewis-type and Bronsted-type acid sites. The presence of a broad acid strength distribution makes the catalysis of several reactions, requiring each a different acid strength, possible. This often results in low selectivity for the desired product. Deposition of silicon, zirconium, titanium or fluor on the surface of alumina allows rendering the catalyst significantly more selective. For the preparation of the alumina based catalyst, suitable commercial alumina's can be used, preferably eta or gamma alumina, having a surface area of 10 to 500 m²/gram and an alkali content of less than 0.5%. The catalyst used in the present invention is for example prepared by adding 0.05 to 10% of silicon, zirconium or titanium. The addition of these metals can be done during the preparation of the alumina or can be added to the existing alumina, eventually already activated. Addition of the metal during the preparation of the alumina can be done by dissolving the metal precursor together with the aluminium precursor before precipitation of the final alumina or by addition of the metal precursor to the aluminium hydroxide gel. A preferred method is adding metal precursors to the shaped alumina. Metal precursors are dissolved in a suitable solvent, either aqueous or organic, and contacted with the alumina by incipient wetness impregnation or by wet impregnation or by contacting with an excess of solute during a given time, followed by removing the excess solute. The alumina can also be contacted with vapour of the metal precursor. Suitable metal precursors are halides of silicon, zirconium or titanium, oxyhalides of zirconium or titanium; alcoxides of silicon, zirconium or titanium; oxalates or citrates of zirconium or titanium or mixtures of the above. The solvent is selected according to the solubility of the metal precursor. The contacting can be done at temperature of 0°C to 500°C, most preferred from 10°C to 200°C. After the contacting, the alumina is eventually washed, dried and finally calcined in other to enhance the surface reaction between the silicon, zirconium or titanium and the alumina and the removal of the metal precursor ligands. The use of silicated, zirconated or titanated or fluorinated alumina's for the simultaneous dehydration and skeletal isomerisation of isobutanol is preferably done in the presence of water. The weight ratio of water to isobutanol ranges from 1/25 to 3/ 1. Fluorinated alumina is known in itself and can be made according to the prior art.

As regards the use of the product issued from the simultaneous dehydration/ isomerisation, the mixture of n-butenes and iso-butene can replace the use of raffinate I in the refinery or petrochemical plants. The main applications of n-butenes and isobutene are described below.

1,3-butadiene can be isolated from a raw C4 stream (often containing 40% 1,3-butadiene, 25%isobutene, 30% n-butenes, 5% butanes) issued from a cracker.

1,3-butadiene can be used to produce:

- styrene-butadiene rubber (SBR),

polybutadiene (PB),

nitrile-butadiene rubber (NBR), styrene-butadiene-styrene (SBS/SEBS), acrylonitrile-butadiene-styrene (ABS),

caprolactam,

1,4-butanediol/THF,

- telomers, etc.

Raffinate I can also be isolated from a raw C4 stream issued from a cracker. Such raffinate I generally contains 40% isobutene, 50% n-butenes, 10% butanes.

Isobutanol dehydration and isomerisation leads to a mixture of n-butenes and isobutene (generally 40% isobutene and 60% n- butenes).

Raffinate I, products of isobutanol dehydration and isomerisation, and C4 issued from FCC stream can be used to product:

( 1) MTBE/ETBE, TBA or di/tri-isobutylene can be used to produce:

(a) high purity isobutene by decomposition of MTBE/ETBE or TBA, such isobutene can be used to produce:

butyl-rubber

poly-isobutene

- methylmethacrylate

isoprene

hydrocarbons resins

t-butyl-amine

alkyl-phenols

- t-butyl-mercaptan

(b) gasoline or jet fuels

(2) Raffinate II (generally containing 80-45% n-butenes, <2% isobutenes, > 18-53% butanes) can be used to produce:

- sec-butanol (used to produce MEK)

- alkylate

- polygasoline

- oxo-alcohols - propylene

- high purity butene- 1 via superfractionation, which can be used to produce LLDPE/HDPE, poly-butene- 1 or n-butyl-mercaptan.

The most typical application of such mixture is the conversion of the contained iso-butene into ethers (MTBE and ETBE), into t- butylalcohol (TBA) or oligomers (e.g. di/tri-iso-butenes), all being gasoline components. The higher oligomers of iso-butene can be used for jet fuel/kerosene applications.

In a specific embodiment, the mixture of butenes produced from simultaneous dehydration/ isomerisation step does not contain butane (or only very small amounts), thus allowing complete upgrading of the olefinic fraction through accurate process enchainment.

The n-butenes, having not reacted during the production of ethers have applications in the production of Alkylate (addition of isobutane to butenes), Polygasoline (oligomerization of butenes).

As regards the etherification step, the production of butyl alkylethers from the acid catalyzed reaction of C₄-alkene and alcohol (such as methanol, or ethanol) is well known in the art.

The etherification reaction is for example described by Stiiwe A. Et al. (Synthesis of MTBE and TAME and related reactions, Section 3. 1 1, in handbook of heterogeneous Catalysis,n Volume 4, (Ertl G., Knozinger H., Weitkamp J., (eds), 1997, VCH Verlagsgesellschaft mbH,

Weinheim, Germany) for the production of methyl-t-butyl ether.

All commercialized MTBE/ ETBE / TAME processes use similar operating conditions although different reactor technologies are applied (EP0455029, DE2706465, DE-OS 291 1077).The etherification reaction proceeds in the liquid phase with a liquid hourly space velocity (LHSV) of 1 to 2 h ¹ at mild conditions (temperature from 40 to 200°C, preferably from 50 to 130°C, pressure ranging from 0. 1 to 20.7 MPa, preferably from 6 to 25bar, pressure greater than 8bar allowing to ensure that the liquid phase is maintained in the presence of a solid acid catalyst. Suitable acid catalysts include, but are not limited to, acidic ion-exchange resins.

The reaction of isobutene with alcohol (ethanol and/ or methanol) is conducted in presence of a small excess of alcohol relative to that required for the stoechiometric reaction of the isoolefin contained in the hydrocarbon feed. Operation with small excess of alcohol allows displacing the equilibrium toward the production of ether to favour higher per-pass conversion, thus limiting secondary reactions such as oligomerization .

The reaction temperature is kept low and can be adjusted over a fairly broad range: indeed, the etherification rate increases with increasing temperature but higher conversion to ethers is favoured at lower temperatures due to the thermodynamic equilibrium.

Furthermore, if higher temperatures are possible, excessive temperatures are not recommended because resin fouling from polymers can occur. Around 130°C, sulfonic ion-exchange resins become unstable. Operation in the lower temperature range ensures stable operation and long catalyst life.

The etherification reaction is exothermic. The removal of the heat of reaction from the reactors and the balance between conversion and reaction rate has greatly influenced the development of etherification technology. In general, two reactors are used in series - the first reactor maintained at around 70°C achieves about 80% of the reaction, and the second reactor operated at 50 to 60°C allows to maximize the olefin conversion to ethers (in the case of isobutene, the conversion reached is above 97%) .

Regarding the etherification technology, the first commercial MTBE plant (started in 1973 in Ravenna, Italy operated by

Anic/Snamprogretti), used a tubular reactor as primary reactor followed by a fixed bed reactor. At that time, four companies Huels (Huels, Belgian 856,401 - Jan.2, 1978), Snamprogretti (Ancillotti F. et al., US 4,071,567 or US 3,979,461), Arco (Oil Gas J. 76, 26, June 26, 1978, 62) and Sun (Oil Gas J., June 16, 1975, 50-2) have developed MTBE processes and are offering them for license. For example, PEP Report n° 131 August 1979 and references therein described in a detailed manner the MTBE manufacture from isobutylene contained in cracked gas mixture.

Today, most of the etherification plants use catalytic distillation (CR&L, EP0008860, 1979 or CDTech, Hydrocarbon Processing, November 1992), which was used for the first time in 1981, in Houston

(Texas).

The etherification catalysts applied in etherification technology are strongly acidic ion exchange resins. Variation of this catalyst, called trifunctional catalyst (TFC) is also reported, which incorporated palladium on the etherification catalyst to hydrogenate the small amounts of dialkenes present in the C5 ex-stream cracked feed stream, thereby improving catalyst life and product color (B.Schleppinghoff, S.Pritchard, Erdoel and Kohle-Erdgas-Petrochimie; 41, April 1988 or K.Rock et al., Hydrocarbon Technology International, 1994, 42-46).

Figure 1 shows a non limitative process enchainment with etherification implementation illustrating the invention. After the simultaneous dehydration/ skeletal isomerisation the water product is separated and the mixture of butenes, eventually containing some heavies (C4+) is etherified by mixing with methanol and/or ethanol.

The etherification effluent contains ethers, not-reacted n-butenes and minor amounts of not-converted iso-butene. The substantially not- reacted butenes mixtures are optionally recovered and sent to either an oligomerization reactor or an alkylation reactor to produce heavies (PolyAlkylate, iso-octenes or middle distillates). [Examples]

Experimental:

The stainless-steel reactor tube has an internal diameter of 10mm. 10ml of catalyst, as pellets of 35-45 mesh, is loaded in the tubular reactor. The void spaces before and after the catalyst are filled with SiC granulates of 2mm.The temperature profile is monitored with the aid of a thermocouple well placed inside the reactor. The reactor temperature is increased at a rate of 60^°C/h to 550°C under air, kept 2 hours at 550°C and then purged by nitrogen. The nitrogen is then replaced by the feed at the indicated operating conditions.

The catalytic tests are performed down-flow, at 1.5 and 2.0 bara, in a temperature range of 280-380°C and with a weight hour space velocity (WHSV) varying from 7 to 211Γ¹.

Analysis of the products is performed by using an on-line gas chromatography.

Example 1 (according to the invention) The catalyst used here is a crystalline silicate of the FER structure. The H-FER has a Si/Al of 33 under powder form. The catalyst is calcinated with air at 550°C during 4 hours before formulation in pellets of 35-45 mesh.

An isobutanol/ water mixture having the 95/5 wt% composition has been processed on the catalyst under 2 bara, at temperatures between 350 and 375°C, and with an isobutanol space velocity from 7 to 211Γ¹.

In this set of operating conditions, isobutanol conversion is almost complete, with a butenes selectivity of over 95%wt, and an iso- butene selectivity of around 41-43%. Low amounts of C₄ ⁺ compounds are formed. Table 1 resumes the data of this example. Table 1: example 1

Comparative example 2

The catalyst is cylinder- shaped gamma-alumina from Sasol® formulated. The catalyst has a specific surface are of 182 m²/g and a porous volume of 0.481 ml/g. The impurities present on the alumina in small amount are summarized below:

0.25 %wt Si, 0.02 %wt P, 0.02 %wt fe, 29 ppm Na.

An isobutanol/ water mixture having the 95/5 wt% composition has been processed on the catalyst under 2 bara, at temperatures between 350 and 380°C, and with an isobutanol space velocity from 7 to 121T¹.

In this set of operating conditions, isobutanol conversion is almost complete, with a butenes selectivity of over 98%wt, and an iso- butene selectivity of around 90-94%. Thus very low amounts of n- butenes are produced over this catalyst. Low amounts of Cs⁺ compounds are formed. Table 2 resumes the data of this example.

Table 2: example 2

FEED i-ButOH /H20 (95/5)%wt

P (bara) 2 2 2 2

T (°C) 380.0 350.0 350.0 325.0

WHSV (H-l) 12.4 7.4 12.4 7.4

Conversion (%

wt CH2) 99.98 99.96 99.93 99.85

Oxygenates (%wt CH2) - average

Other

Oxygenates 0.0 0.0 0.0 0.0

Other alcohol 0.0 0. 1 0. 1 0. 1

Selectivity on C-basis (%wt CH2) - average

Paraffins 0.3 0.3 0. 1 0.3

C2= 0.3 0.2 0.2 0. 1

C3= 0.2 0. 1 0.0 0.0

C4= 98.2 98.6 99.1 98.6

C5+ olef 0.7 0.5 0. 1 0.3

Dienes 0. 1 0.0 0.0 0. 1

Aromatics 0.0 0.0 0.0 0.0

Unknown 0. 1 0. 1 0.3 0.4

C4= distribution

(%wt)

iC4= 90.2 92.5 92.7 94.0 t-2-C4= 3.0 1.8 1.4 1.2

C-2-C4- 3.9 3.2 3.3 2.7

1-C4= 2.9 2.5 2.6 2. 1 n-C4= 9.8 7.5 7.3 6.0 Example 3 (according to the invention)

The catalyst is a phosphorous modified zeolite (P-ZSM5), prepared according to the following recipe. A sample of zeolite ZSM-5 (Si/Al= 13) in H-form was steamed at 550°C for 6h in 100% H₂0. The steamed solid was subjected to a contact with an aqueous solution of H3PO4 (85%wt) for 2h under reflux condition (4ml/ lg zeolite). 69.9 g of CaC03 was then introduced. Then the solution was dried by evaporation for 3 days at 80°C. 750g of the dried sample was extruded with 401,5g of Bindzil and 0.01wt% of extrusion additives. The extruded solid was dried at 1 10°C for 16h and calcinated at 600°C for lOh.

An isobutanol/ water mixture having the 95/5 wt% composition has been processed on the catalyst under 1.5 bara, at temperatures between 280 and 350°C, and with an isobutanol space velocity of about 7h ^l.

In this set of operating conditions, isobutanol conversion is almost complete, with a butenes selectivity of over 90%wt, and an iso- butene selectivity of about 66-67 %. Thus, nearly 90% or more butenes are produced, of which a significant amount are skeletal isomerised into n-butenes. The heavies production is limited to 10% or less. Table 3 resumes the data of this example.

Table 3: example 3

FEED : i-ButOH/H20 (95/5)%wt

P (bara) 1.5 1.5

T (°C) 300 280

WHSV (H-l) 7.4 7.4

Conversion (% wt

CH2) 100.0 83.5

Oxygenates (%wt CH2) - Average

Other alcohols 0.01 0.00

Other Oxygenates 0.03 0.08

Selectivity on C-basis (%wt CH2) - Average

Paraffins C1-C4 0. 1 0. 1

C2= 0.0 0.0

C3= 0.5 0.3

C4= 89.9 93.9 i-Butene 60.3 61.9

1-Butene 5.0 6. 1

2-Butene 24.6 26.0

C5+ olefines 4.8 2.7

C5+ paraffines 1.9 1. 1

Dienes 0.5 0.4

Aromatics 0.5 0.2

Unknown 1.6 1. 1

C4= distribution - Average i-Butene 67. 1 65.9 n-butenes 32.9 34. 1

1-Butene 5.5 6.5

2-Butene 27.4 27.7

Claims

1 Process for the production of fuel additives in which in a first step isobutanol is subjected to a simultaneous dehydration and skeletal isomerisation to make substantially corresponding olefins, having the same number of carbons and consisting essentially of a mixture of n-butenes and iso-butene and in a second step the butene mixture is subjected to etherification, said process comprising:

a) introducing in at least one reactor a stream (A) comprising at least 40wt% isobutanol, optionally an inert component,

b) contacting said stream with at least one catalyst in said reactor(s) at conditions effective to simultaneously dehydrate and skeletal isomerise at least a portion of the isobutanol to make a mixture of n-butenes and iso-butene,

d) sending the stream (B) to at least one etherification reactor and contacting stream (B) with at least one catalyst in said etherification reactor(s), in the presence of ethanol and/ or methanol, at conditions effective to produce ETBE and/or MTBE respectively,

e) recovering from said etherification reactor a stream (E) comprising essentially ETBE and/or MTBE, unreacted butenes, heavies, optionally unreacted ethanol and / or methanol respectively, f) fractionating stream (E) to recover ETBE and/or MTBE.

2. Process according to claim 1, wherein the stream (A) comprises at least 70wt% of isobutanol, preferably at least 80% or more.

3. Process according to claim 1 or 2, wherein the stream (A) comprises one or more C4 alcohols other than isobutanol.

4. Process according to any one of claims 1 to 3, wherein the stream (A) used in step a), and/ or methanol and/ or ethanol used in step d) are issued from renewable energy sources.

5. Process according to any one of claims 1 to 4, wherein the stream (A) is subjected to a purification treatment before step b).

6. Process according to any one of claims 1 to 5, wherein the WHSV of the isobutanol is at least 1 h ¹.

7. Process according to any one of claims 1 to 6, wherein the temperature of the simultaneous dehydration and skeletal isomerisation of isobutanol ranges from 200°C to 600°C.

8. Process according to any one of the preceding claims, wherein the temperature of the simultaneous dehydration and skeletal isomerisation of isobutanol ranges from 250°C to 500°C.

9. Process according to any one of the preceding claims, wherein the temperature of the simultaneous dehydration and skeletal isomerisation of isobutanol ranges from 300°C to 450°C.

10. Process according to any one of the preceding claims, wherein the catalyst for the simultaneous dehydration and skeletal isomerisation is a crystalline silicate of the group FER, MWW, EUO, MFS, ZSM-48, MTT, MFI, MEL or TON having Si/Al higher than 10, or a dealuminated crystalline silicate of the group FER, MWW, EUO, MFS, ZSM-48, MTT, MFI, MEL or TON having Si/Al higher than 10,

1 1. Process according to any one of the preceding claims, wherein the pressure of the reactor(s) of the simultaneous dehydration and skeletal isomerisation of isobutanol ranges from 0.5 to 10 bars absolute.

12. Process according to any one of the preceding claims, wherein stream (E) recovered from step e) is fractionated to recover unreacted butenes and/or unreacted ethanol and/or methanol.

13. Process according to claim 12, wherein at least a part of said recovered unreacted butenes and/ or at least a part of said recovered unreacted ethanol and/ or methanol are recycled to the etherification reactor(s) .

14. Process according to claim 12 or 13, wherein at least a part of said recovered unreacted butenes are sent to a purification zone before being sent to at least one oligomerization reactor and/ or to at least one alkylation reactor to produce heavies.

15. Process according to claim 14, wherein said unreacted butenes are subjected to a purification step before being sent to oligomerization reactor(s) and / or alkylation reactor(s) .