EP1973645A1

EP1973645A1 - Producing alkylene glycols in microchannel apparatus

Info

Publication number: EP1973645A1
Application number: EP06830757A
Authority: EP
Inventors: Bernardus Franciscus Josef Marie Ramakers
Original assignee: Shell Internationale Research Maatschappij BV
Current assignee: Shell Internationale Research Maatschappij BV
Priority date: 2005-12-22
Filing date: 2006-12-20
Publication date: 2008-10-01
Also published as: TW200735958A; EP1981632A1; CN101384352A; CA2634409A1; EA200870100A1; TW200740519A; EP1979087A1; CN101384350A; EP1976625A1; JP2009520943A; CN101384349A; JP2009520767A; WO2007071744A1; TW200803973A; WO2007071739A1; WO2007071737A1; WO2007071741A1; JP2009520764A; EP1976625B1; DE602006007180D1

Abstract

Process for preparing alkylene glycol by reacting the corresponding oxide and water by flowing these through a microchannel reactor and transferring the heat to a heat transfer medium. The reactor contains a catalyst suitablefor the hydrolysis. Process for preparing a mono-alkylene glycol by reacting the alkylene oxide and water under a first set of temperature/pressure conditions to achieve vapour phase conversion, altering the conditions to a second set of conditions and removing the glycols deposited on the surface of the catalyst, reestablishing the first set of conditions to repeat the first step and recovering the glycol from the vapour phase mixture .

Description

PRODUCING ALKYLENE GLYCOLS IN MICROCHANNEL APPARATUS

Field of the Invention

The present invention relates to improvements in process operations involving particularly hydrocarbons. The process improvements envisaged find especial application in the production of olefin oxide from olefin and oxygen and in its optional further conversion. Background of the Invention

When operating on a commercial scale, process operations have to meet a number of important design criteria. In the modern day environment, process design has to take account of environmental legislation and keep to health and safety standards. Processes that utilise or produce dangerous chemicals pose particular problems and often, in order to minimise risks of explosion or reaction runaway, such process operations have to be run at conditions that are not optimal; this increases the running costs of a plant (the operational expenditure or OPEX) . Such processes may also have to utilise more equipment than is necessary just to perform the process; this leads to an increase in building costs (the capital expenditure or CAPEX) .

There is an on-going need to provide process operations that can reduce CAPEX and OPEX costs and particularly without increasing the risk of damage to the plant and danger to the public and/or to the process plant workers . Summary of the Invention

The present invention provides for the utilisation of microchannel apparatus in process operations. Such apparatus have previously been proposed for use in certain specific fields of application but have not previously been proposed to provide the combination of reduced CAPEX and/or OPEX with maintained or reduced plant safety risks. In one aspect the present invention provides a process for the preparation of an alkylene glycol by the reaction of a corresponding alkylene oxide and water, which process comprises a) flowing the alkylene oxide and water through a microchannel reactor, optionally in the presence of a catalyst, wherein the oxide and water undergo an exothermic reaction to form the corresponding alkylene glycol, b) transferring heat from the microchannel reactor to a heat transfer medium, and c) recovering the alkylene glycol product from the microchannel reactor.

In another aspect the present invention provides a process for the preparation of a mono-alkylene glycol by the reaction of a corresponding alkylene oxide and water, which process comprises a) reacting the alkylene oxide and water in a first reactor under a first set of conditions and in the presence of a catalyst so as to achieve vapour phase conversion to the mono-alkylene glycol, b) altering the conditions in the first reactor to a second set of conditions whereby glycols deposited on the surface of the catalyst are removed, c) re-establishing the first set of conditions in the first reactor in order to repeat step a) , and d) recovering the mono-alkylene glycol from a vapour phase mixture produced in step a) and/or step b) . Brief Description of the Drawings

FIG. 1 shows a schematic drawing of a microchannel reactor and its main constituents.

FIG. 2 shows a schematic drawing of a typical example of a repeating unit which comprises process microchannels and heat exchange channels and its operation when in use in the practice of the invention. A microchannel apparatus or reactor utilised in this invention may comprise a plurality of such repeating units.

FIG. 3 shows a schematic drawing of glycol production unit. Detailed Description of the Invention

The present invention provides, in a number of aspects, processes that utilise microchannel apparatus.

In a number of these processes the microchannel apparatus may house a chemical reaction and optionally may also contain catalytic components; in other processes the microchannel apparatus are utilised for physical operations. Hereinafter a discussion of such apparatus is given and reference is made generally to 'microchannel reactors' ; this term will be understood to encompass microchannel apparatus whether utilised for physical processes or for chemical reaction processes, with or without a catalytic component.

MicroChannel reactors suitable for use in this invention and their operation have been described in WO-A-2004/099113, WO-A-01/12312 , WO-01/54812, US-A-6440895, US-A-6284217 , US-A-6451864 , US-A-6491880, US-A-6666909, US-A-6811829, US-A-6851171, US-A-6494614 , US-A-6228434 and US-A-6192596. Methods by which the microchannel reactor may be manufactured, loaded with catalyst, and operated, as described in these references, - A -

may generally be applicable in the practice of the present invention.

With reference to FIG. 1, microchannel reactor 100 may be comprised of a header 102, a plurality of process microchannels 104, and a footer 108. The header 102 provides a passageway for fluid to flow into the process microchannels 104. The footer 108 provides a passageway for fluid to flow from the process microchannels 104.

The number of process microchannels contained in a microchannel reactor may be very large. For example, the number may be up to 10⁵, or even up to 10⁶ or up to 2 x 10⁶. Normally, the number of process microchannels may be at least 10 or at least 100, or even at least 1000.

The process microchannels are typically arranged parallel, for example they may form an array of planar microchannels. Each of the process microchannels may have at least one internal dimension of height or width of up to 15 mm, for example from 0.05 to 10 mm, in particular from 0.1 to 5 mm, more in particular from 0.5 to 2 mm. The other internal dimension of height or width may be, for example, from 0.1 to 100 cm, in particular from 0.2 to 75 cm, more in particular from 0.3 to 50 cm. The length of each of the process microchannels may be, for example, from 1 to 500 cm, in particular from 2 to 300 cm, more in particular from 3 to 200 cm, or from 5 to 100 cm.

The microchannel reactor 100 additionally comprises heat exchange channels (not shown in FIG. 1) which are in heat exchange contact with the process microchannels 104. The heat exchange channels may be microchannels. The microchannel reactor is adapted such that heat exchange fluid can flow from heat exchange header 110 through the heat exchange channels to heat exchange footer 112. The heat exchange channels may be aligned to provide a flow in a co-current, counter-current or, in some aspects, preferably cross-current direction, relative to a flow in the process microchannels 104. The cross-current direction is as indicated by arrows 114 and 116. Each of the heat exchange channels may have at least one internal dimension of height or width of up to 15 mm, for example from 0.05 to 10 mm, in particular from 0.1 to 5 mm, more in particular from 0.5 to 2 mm. The other internal dimension of height or width may be, for example, from 0.1 to 100 cm, in particular from 0.2 to 75 cm, more in particular from 0.3 to 50 cm. The length of each of the heat exchange channels may be, for example, from 1 to 500 cm, in particular from 2 to 300 cm, more in particular from 3 to 200 cm, or from 5 to 100 cm.

The separation between each process microchannel 104 and the next adjacent heat exchange channel may be in the range of from 0.05 mm to 5 mm, in particular from 0.2 to 2 mm. In some embodiments of this invention, there is provided for first heat exchange channels and second heat exchange channels, or first heat exchange channels, second heat exchange channels and third heat exchange channels, or even up to fifth heat exchange channels, or even further heat exchange channels. Thus, in such cases, there is a plurality of sets of heat exchange channels, and accordingly there may be a plurality of heat exchange headers 110 and heat exchange footers 112, whereby each set of heat exchange channels may be adapted to receive heat exchange fluid from a heat exchange header 110 and to deliver heat exchange fluid into a heat exchange footer 112. The header 102, footer 108, heat exchange header 110, heat exchange footer 112, process microchannels 104 and heat exchange channels may independently be made of any construction material which provides sufficient strength, optionally dimensional stability, and heat transfer characteristics to permit operation of the processes in accordance with this invention. Suitable construction materials include, for example, steel (for example stainless steel and carbon steel) , monel, titanium, copper, glass and polymer compositions. The kind of heat exchange fluid is not material to the present invention and the heat exchange fluid may be selected from a large variety. Suitable heat exchange fluids include steam, water, air and oils. In embodiments of the invention which include a plurality of sets of heat exchange channels, such sets of heat exchange channels may operate with different heat exchange fluids or with heat exchange fluids having different temperatures . A microchannel reactor of use in the invention may comprise a plurality of repeating units each comprising one or more process microchannels and one or more heat exchange channels. Reference is now made to FIG. 2, which shows a typical repeating unit and its operation. Process microchannels 210 have an upstream end 220 and a downstream end 230 and may comprise of a first section 240 which may optionally, for certain aspects of the present invention, contain a catalyst (not shown) . First section 240 may be in heat exchange contact with first heat exchange channel 250, allowing heat exchange between first section 240 of process microchannel 210 and first heat exchange channel 250. The repeating unit may comprise first feed channel 260 which leads into first section 240 through one or more first orifices 280. Typically one or more first orifices 280 may be positioned downstream relative to another first orifice 280. During operation, feed may enter into first section 240 of process microchannel 210 through an opening in upstream end 220 and/or through first feed channel 260 and one or more first orifices 280.

Process microchannels 210 may comprise a second section 340 which may or may not be adapted to contain a catalyst. Second section 340 is positioned down stream of first section 240. Second section 340 may be in heat exchange contact with second heat exchange channel 350, allowing heat exchange between second section 340 of process microchannel 210 and second heat exchange channel 350. In some embodiments second section 340 is adapted to quench product obtained in and received from first section 240 by heat exchange with a heat exchange fluid in second heat exchange channel 350. Quenching if required may be achieved in stages by the presence of a plurality of second heat exchange channels 350, for example two or three or four. Such a plurality of second heat exchange channels 350 may be adapted to contain heat exchange fluids having different temperatures, in particular such that in downstream direction of second section 340 heat exchange takes place with a second heat exchange channel 350 containing a heat exchange fluid having a lower temperature. The repeating unit may comprise second feed channel 360 which leads into second section 340 through one or more second orifices 380. During operation, feed may enter into second section 340 from upstream in process microchannel 210 and through second feed channel 360 and one or more second orifices 380.

The first and second feed channels 260 or 360 in combination with first and second orifices 280 or 380 whereby one or more first or second orifices 280 or 380 are positioned downstream to another first or second orifice 280 or 380, respectively, allow for replenishment of a reactant. Replenishment of a reactant can be utilised in some embodiments of this invention. Process microchannels 210 may comprise an intermediate section 440, which is positioned downstream of first section 240 and upstream of second section 340. Intermediate section 440 may be in heat exchange contact with third heat exchange channel 450, allowing heat exchange between intermediate section 440 of the process microchannel 210 and third heat exchange channel 450.

In some embodiments, process microchannel 210 may comprise a third section (not drawn) downstream of second section 340, and optionally a second intermediate section (not drawn) downstream of second section 340 and upstream of the third section. The third section may be in heat exchange contact with a fourth heat exchange channel (not drawn) , allowing heat exchange between the third section of the process microchannel 210 and fourth heat exchange channel. The second intermediate section may be in heat exchange contact with a fifth heat exchange channel (not drawn) , allowing heat exchange between the second intermediate section of the process microchannel 210 and fifth heat exchange channel. The repeating unit may comprise a third feed channel (not drawn) which ends into the third section through one or more third orifices (not drawn) . Typically one or more third orifices may be positioned downstream relative to another third orifice. During operation, feed may enter into the third section from upstream in process microchannel 210 and through the third feed channel and the one or more third orifices.

Each of the feed channels may be a microchannel. They may have at least one internal dimension of height or width of up to 15 mm, for example from 0.05 to 10 mm, in particular from 0.1 to 5 mm, more in particular from 0.5 to 2 mm. The other internal dimension of height or width may be, for example, from 0.1 to 100 cm, in particular from 0.2 to 75 cm, more in particular from 0.3 to 50 cm. The length of each of the feed channels may be, for example, from 1 to 250 cm, in particular from 2 to 150 cm, more in particular from 3 to 100 cm, or from 5 to 50 cm.

The length of each of the sections of the process microchannels may be selected independently of each other, in accordance with, for example, the heat exchange capacity needed or the quantity of catalyst which may be contained in the section. The lengths of the sections may independently be at least 1 cm, or at least 2 cm, or at least 5 cm. The lengths of the sections may independently be at most 250 cm, or at most 150 cm, or at most 100 cm, or at most 50 cm. Other dimensions of the sections are defined by the corresponding dimensions of process microchannel 210. The microchannel reactor of this invention may be manufactured using known techniques, for example conventional machining, laser cutting, molding, stamping and etching and combinations thereof. The microchannel reactor of this invention may be manufactured by forming sheets with features removed which allow passages. A stack of such sheets may be assembled to form an integrated device, by using known techniques, for example diffusion bonding, laser welding, cold welding, diffusion brazing, and combinations thereof. The microchannel reactor of this invention comprises appropriate headers, footers, valves, conduit lines, and other features to control input of reactants, output of product, and flow of heat exchange fluids. These are not shown in the drawings, but they can be readily provided by those skilled in the art. Also, there may be further heat exchange equipment (not shown in the drawings) for temperature control of feed, in particular for heating feed or feed components, before it enters the process microchannels, or for temperature control of product, in particular for cooling product, after it has left the process microchannels. Such further heat exchange equipment may be integral with the microchannel reactor, but more typically it will be separate equipment. These are not shown in the drawings, but they can be readily provided by those skilled in the art.

Where catalyst is present, it may be in any suitable form to be accommodated in one or more of the process microchannels. Such catalyst may be installed by any known technique in the designated section of the process microchannels. The catalyst may be in solid form and form a packed bed in the designated section of the process microchannels and/or may form a coating on at least a portion of the wall of the designated section of the process microchannels. Alternatively the catalyst may be in the form of a coating on inserts which may be positioned in the designated section of the microchannel apparatus. Coatings may be prepared by any suitable deposition method such as wash coating or vapour deposition. Where a catalyst is comprised of several catalytically effective components, deposition may be achieved by deposition of a first catalytic component, e.g. a metal or metal component, on at least a portion of the wall of the designated section of the process microchannels with the deposition of one or more additional catalyst components on at least the same wall prior to, together with, or subsequent to that of the first component. In some embodiments the catalyst may be homogeneous and not in solid form in which case the catalyst may be fed to the designated section of the process microchannels together with one or more components of the relevant feed or process stream and may pass through the microchannels along with the reaction mixture or process stream.

The thermal conversion of ethylene oxide and water to ethylene glycol is well known and commercially practised world-wide, see for example the description in Ullmann' s Encyclopedia of Industrial Chemistry, Volume A 10, pages 104 & 105. The thermal process requires a high molar excess of water, as much as a 20-fold molar excess, to yield the most desired product of mono-ethylene glycol. Catalytic conversions that are selective to mono- ethylene glycol and that do not require such high excess of water are also of interest. Catalytic processes for converting alkylene oxides directly to alkylene glycols in general have been investigated and catalysts capable of promoting a higher selectivity to monoalkylene glycol product at reduced water levels are known, (e.g.

EP-A-015649, EP-A-0160330, WO 95/20559 and US-A-6124508 ) .

All of these conversion reactions are highly exothermic .

The present invention provides a process for the preparation of an alkylene glycol by the reaction of a corresponding alkylene oxide and water, which process comprises a) flowing the alkylene oxide and water through a microchannel reactor, wherein the oxide and water undergo an exothermic reaction to form the corresponding alkylene glycol, b) transferring heat from the microchannel reactor to a heat transfer medium, and c) recovering the alkylene glycol product from the microchannel reactor.

Utilising a microchannel reactor provides the advantages of a high removal rate of the heat of reaction, and a much greater temperature control of the full conversion process.

The microchannel reactor can also incorporate a catalyst system that permits the reduction of the high water excess. Such a catalyst system may be a homogeneous catalyst that is mixed with the reactants either before entry to the reactor or within the reactor, or it may be a heterogeneous system present as a solid catalyst or as a coating, preferably a wash-coating, on the walls of one or more, and desirably all, of the process microchannels present in the reactor.

Catalysts that may be employed in the present process are known in the art. Suitable catalysts are acid catalysts and basic catalysts.

Homogeneous catalysts include acidic catalysts which are liquid under the conditions of the reaction. Suitably such catalysts are mineral acids, such as sulphuric acid and phosphoric acid, and such catalysts as known from JP- A-56-092228. Homogeneous metalate catalysts are also very suitable; such catalysts comprise a salt selected from vanadates, molybdates and tungstates. Suitable examples are described in US-A-4 , 551, 566, EP-A-156447, and EP-A- 156448.

Less preferred are heterogenous catalysts. Ones that may be mentioned are acidic catalysts such as strongly acidic ion exchange resins, such as those comprising sulphonic acid groups on a styrene/divinylbenzene copolymer matrix, and silicas and oxides of metals selected from Groups 3 to 6 of the Periodic Table of Elements, for example zirconium oxide and titanium oxide. As basic catalysts there may be mentioned those comprising an ion exchange resin (IER) as a solid support, in particular the strongly basic (anionic) IER' s wherein the basic groups are quaternary ammonium or quaternary phosphonium on a styrene/divinylbenzene copolymer matrix. Also suitable as heterogeneous catalysts are metalates, such as vanadates, molybdates and tungstates, contained on a solid support such as an ion exchange resin or a hydrotalcite clay as described in EP-A-156449 and EP-A-318099.

Suitable ion exchange resins utilised may be based on vinylpyridine, polysiloxanes . Other solid supports having electropositive complexing sites of an inorganic nature may also be utilised, such as carbon, silica, silica-alumina, zeolites, glass and clays such as hydrotalcite. Further, immobilised complexing macrocycles such as crown ethers, etc. can be used as well as a solid support .

Such heterogeneous catalyst may be based on a strongly basic quaternary ammonium resin or a quaternary phosphonium resin, for example an anion exchange resin comprising a trimethylbenzyl ammonium group. Examples of commercially available anion exchange resins on which the catalyst may be based include LEWATIT M 500 WS (LEWATIT is a trademark) , DUOLITE A 368 (DUOLITE is a trademark) and AMBERJET 4200 (AMBERJET is a trademark) , DOWEX MSA-I (DOWEX is a trademark), MARATHON-A and MARATHON-MSA (MARATHON is a trademark) (all based on polystyrene resins, cross-linked with divinyl benzene) and Reillex HPQ (based on a polyvinylpyridine resin, cross-linked with divinyl benzene) . The anion exchange resin in the fixed bed of solid catalyst may comprise more than one anion which may be selected from the group of bicarbonate, bisulfite, metalate and carboxylate anions . When the anion is a carboxylate anion, it maybe a polycarboxylic acid anion having in its chain molecule one or more carboxyl groups and one or more carboxylate groups, the individual carboxyl and/or carboxylate groups being separated from each other in the chain molecule by a separating group consisting of at least one atom. The polycarboxylic acid anion is suitably a citric acid derivative, more preferably a mono-anion of citric acid. A suitable solid catalyst is a catalyst based on a quaternary ammonium resin, preferably a resin comprising a trimethylbenzyl ammonium group, and wherein the anion is a bicarbonate anion.

The alkylene oxides used as starting materials in the process of the present invention, have their conventional definition, i.e. they are compounds having a vicinal oxide (epoxy) group in their molecules.

Preferred alkylene oxides are alkylene oxides of the general formula :-

1 2 3 4

R -CR -CR -R

\ / 0 wherein each of R¹ to R⁴ independently represents a hydrogen atom or an optionally substituted alkyl group having from 1 to 6 carbon atoms. Any alkyl group, represented by R¹, R², R³ and/or R⁴, preferably has from 1 to 3 carbon atoms. Optional substituents on the alkyl groups include hydroxy groups. Preferably, R¹, R², and R³ represent hydrogen atoms and R⁴ represents a non- substituted Ci-C₃-alkyl group and, more preferably, R¹, R², R³ and R⁴ all represent hydrogen atoms. Examples of alkylene oxides which may conveniently be employed include ethylene oxide, propylene oxide, 1,2- epoxybutane, 2, 3-epoxybutane and glycidol. The alkylene oxide is preferably ethylene oxide or propylene oxide; ethylene glycol and propylene glycol being alkylene glycols of particular commercial importance. Most preferably the alkylene oxide of the present invention is ethylene oxide or propylene oxide and the alkylene glycol is ethylene glycol or propylene glycol. When the conversion is a thermal conversion, the temperature may be in the range of from 100 to 300⁰C, in particular from 150 to 250⁰C. When the conversion is a catalytic conversion, the temperature may be in the range of from 30 to 200⁰C, in particular from 50 to 15O⁰C. The molar ratio of water to the alkylene oxide may be in the range of from 5 to 50, in particular from 10 to 30. The pressure may be in the range of from 500 to 3500 kPa, as measured at the second feed channel, described hereinbefore . In certain embodiments of the present invention it may be beneficial to add carbon dioxide to the (catalytic) reactor to establish advantageous conditions for the hydrolysis. Such carbon dioxide may conveniently be added directly to the reactor or it may be added to the alkylene oxide feed. If carbon dioxide is to be added, the amount of carbon dioxide added may be varied to obtain optimum performance in relation to other reaction parameters, in particular the type of catalyst employed. However the amount added will preferably be less than 0.1 % wt, more preferably less than 0.01 % wt, based on a total amount of reactants in the second reactor . With reference to FIG. 2, as an example, alkylene oxide may be fed into first section 240 via feed channels 260 and/or 220, and water may be co-fed through the same channels or fed into the microchannel system via feed channel 360 of second section 340. In a specific embodiment when the oxide is fed through channel 240, then, water may be fed through channels 220 and mixed with the oxide in the microchannels . To remove heat evolved during the reaction, coolant may flow via heat exchange channels 250 and/or 350 depending on the site of reaction and the channel through which the water feed is fed. If alkylene oxide and water are co-fed to the first section, then where catalyst is present, any additional component useful for the reaction, such as carbon dioxide, may be fed to the reactants via the second section 340.

The use of a microchannel reactor permits a greater control of the exothermic reaction than has hitherto been possible which reduces the need for excess volumes of water to act as a heat sink.

The hydration of ethylene oxide to mono-ethylene glycol (MEG) is normally carried out in the liquid phase in, for example, a pipe or tube reactor. As noted above previously proposals to use catalysts in such conversions have been made. Additionally it has been proposed to react EO and water in the vapour phase since this can be beneficial in terms of process integration and separation. Regarding the latter in particular the removal of MEG from a gas stream is possibly easier than from a dilute aqueous stream as in conventional plants.

Heterogeneous hydrolysis catalysts can also be utilized in vapour phase hydration, where mono-ethylene glycol will be formed as the main product. However inevitably the EO present will also react with the formed glycols to form higher molecular weight glycols, for example EO with mono-ethylene glycol will form di- ethylene glycol, with di-ethylene glycol EO will form tri-ethylene glycol, and so forth. The major problem with the use of heterogeneous catalysts for vapour phase reactions is that the higher glycols have a high boiling point and thus are liquid at the typical reaction temperature and pressure applied. Thus the catalyst surface will be covered with glycols quickly growing in molecular weight leading to deactivation of the catalyst. MEG may also be trapped as liquid on the catalyst surface. Also the reaction of EO with glycols on the surface of the catalyst will result in a reduced selectivity of EO to mono-ethylene glycol product. The use of certain highly selective heterogeneous catalysts in the vapour phase hydration of ethylene oxide has been proposed in the literature, most recently in EP- A-318099 and EP-A-529726 which describe the use of specific hydrotalcites, which are anionic clays, both for vapour and liquid phase hydration.

However even with a high degree of selectivity to mono-ethylene glycol, the problem of deactivation by deposition of other glycols produced as by-products in vapour phase hydration will still exist. In the present invention it is proposed to operate a gas phase reactor in a ^λswing' mode. In mode 1, the reaction mode, the temperature and pressure are optimised to achieve the desired production of mono-ethylene glycol. A heterogeneous catalyst which is highly selective to the production of MEG is preferably used.

Preferably the temperature is maintained in the range of from 200 to 35O⁰C, most suitably 200 to 275⁰C, and the pressure in the range of from 100 to 1000 kPa. Suitably this process is performed without excessive amounts of water. The amount of water is preferably in the range of from 1 to 35 moles per mole of alkylene oxide, more preferably from 1 to 20 moles, most preferably from 1 to 10 moles, per mole of alkylene oxide. The heterogeneous catalysts preferred for use in such a process are based on support materials selected from members of the family of clays; aluminas, for example α- and γ~ alumina; zirconias; silicas; and hydrotalcites (anionic clays) . Such support materials suitably have a metal component, for example a metal ion or a metal oxide deposited thereon to enhance activity and/or selectivity, but can be utilised above. Any metal or metal component can be incorporated into the catalyst. Most suitably such a metal component may be selected from one or more metals of Groups IA, HA, IHA, IVA, VA, VIA, VIIA, VIIIA, IB, HB, and IHB of the Periodic Table (using the IUPAC notation) . Very suitable metals include sodium, cesium, molybdenum, nickel, cobalt, zinc, aluminium, lanthanum, rhenium, tungsten, and vanadium.

Suitable catalyst components may also include anionic groups such as hydroxide ions, carbonate ions, sulphate ions and phosphate ions .

Where the support material is a hydrotalcite, such materials are anionic clays and consist of positively charged layers of oxides and/or hydroxides, for example in conjunction with a mixture of Mg²⁺ and Al³⁺ cations, separated by a layer containing water and charge compensating anions, for examples hydroxides or carbonates.

Examples of suitable catalyst systems are:

MoO₄/ZrO_x (OH) ₄-_2X; Cs/α-Al₂O₃; Co/Mo/α-Al₂0₃; Zn/Al/CO₃ hydrotalcite; Co/Mo/Si0₂; Mo/Co/Zn/Al hydrotalcite; hydrotalcite/Na-citrate; Co/Zn/Al-hydrotalcite; Mo- Co/Zn/Al-hydrotalcite; SiO₂ granules; 12wt% La/ Oc-Al₂O₃;

SO₄/ZrO_x(OH)₄__2x; SO₄/ZrO₂; ZrO_x (OH) ₄-_2x; ZrO₂; P0₄/Zr0_x (OH) ₄_

_2x; PO₄/ZrO₂; ReO₄/ZrO_x (OH) ₄__2x; ReO₄/ZrO₂; WO₄/ZrO_x (OH) ₄-_2x;

WO₄/ZrO₂; Mo0₄/Zr0_x (OH) ₄__2x; Mo0₄/Zr0₂; Ni/V hydrotalcite; Ni/V hydrotalcite-coated Al/5Mg; α-Al₂O₃; Co/α-Al₂0₃ dried; Co/α-Al₂0₃ calcined; Cs/α-Al₂O₃ dried; Cs/α-Al₂O₃ calcined; Co/Mo/α-Al₂0₃ dried; Co/Mo/α-Al₂0₃ calcined;

ZnAlCO₃ hydrotalcite; and CoMoSiO₂. In the preceding list x, where it appears, is a number from 0 to 2. Such catalysts are either available commercially or may be easily prepared by techniques well known to the person skilled in the art.

Hydrotalcite-type catalysts of the type proposed in

EP-A-529726 are very suitable. These are hydrotalcite- type catalysts of the general formula

M_xQ_y (OH) 2x+3y-nzA_z ^n" ■ aH₂0 wherein M is at least one divalent metal cation; Q is at least one trivalent metal cation; A is at least one component having a valence n- selected from a metalate anion, selected from vanadate (suitably metavanadate, orthovanadate, pyrovanadate, and hydrogen pyrovanadate) , tungstate, niobate, tantalite and perrhenate, and a large organic anion spacer; and a is a positive number. M, Q and A are present such that x/y is greater than or equal to 1, z > 0, and 2x+3y-nz is a positive number. The composition has a layered structure where A is located in anionic sites of the composition.

In catalysts of the above general formula in which A is a large organic anion spacer, the selectivity to MEG is increased. Therefore preferably A is a large organic spacer and may be any organic acid containing from one to 20 carbon atoms, provided its steric bulk is large. Such organic acid or its alkali salt must be somewhat soluble in a solvent, and may have one or more carboxylic acid functional groups, and may have one or more sulphonic acid functional groups. Large organic anion spacers containing carboxylic acid functional groups are preferred, since these functional groups are readily removed by heating. Preferred large organic anion spacers include terephthalate, benzoate, cyclohexanecarboxylate, sebacate, glutarate and acetate. Preferably, the large organic anion spacer is selected from the group consisting of terephthalate and benzoate. Terephthalate is the most preferred large organic anion spacer. Mixtures of large organic anion spacers may also be used. Preferably x/y is in the range of from 1 to 12, more preferably 1 to 6, and most preferably 1 to 4. Suitable divalent cations M broadly include elements selected from the transition elements and Groups HA, IVA and VA of the Periodic Table (IUPAC version), as well as certain rare earth elements. Specific examples of divalent metal cations are magnesium, calcium, titanium, vanadium, chromium, manganese, iron, cobalt, nickel, palladium, platinum, copper, zinc, cadmium, mercury, tin, lead and mixtures thereof. Divalent metal cations which are particularly suitable are magnesium, nickel, cobalt, zinc, calcium, iron, titanium and copper. Suitable trivalent metal cations Q broadly include elements selected from the transition elements and Groups IHA and VA of the Periodic Table as well as certain rare earth elements and actinide elements. Specific examples of trivalent metal cations are aluminium, antimony, titanium, scandium, bismuth, vanadium, yttrium, chromium, iron, manganese, cobalt, ruthenium, nickel, gold, gallium, thallium, cerium, lanthanum and mixtures thereof. Trivalent metal cations which are particularly suitable are aluminum, iron, chromium, and lanthanum. The foregoing lists of suitable divalent and trivalent metal cations are meant to be illustrative but not exclusive. Those skilled in the art will recognize that other cations can be used, provided that the types of cations and relative amounts (x/y ratio) result in a hydrotalcite-type catalyst, M is nickel, Q is aluminum, E is metavanadate and x/y is in the range of 1 to 6. Another preferred hydrotalcite-type catalyst is formed when M is nickel, Q is aluminum, E is niobate and x/y is in the range of 1 to 6. Hydrotalcite-type materials in which M is nickel and Q is aluminum are known as takovites .

Such hydrotalcite catalsyst may be prepared by the procedures described in EP-A-529726. The reactor is changed to mode 2 (step b) once glycols form on the surface of the catalyst. This can be after reaction time of from 1 seconds to 10 hours, preferably from 10 seconds to 1 hour, depending on the reaction conditions (temperature and pressure) . In mode 2, the desorption or evaporation mode, the temperature, the pressure, or both temperature and pressure, are adjusted to desorb or evaporate the glycols from the surface of the catalyst, while the gas stream is fed to a second reactor. Essentially the temperature has to be increased and/or the pressure decreased to such conditions as are necessary to cause the glycols to desorb and/or to evaporate. Preferably if temperature alone is altered then the temperature is changed to be in the range of from 250 to 400⁰C. If the pressure alone is altered then the pressure is preferably changed to be in the range of from 1 Pa to 500 kPa . If both temperature and pressure conditions are altered, then the conditions are preferably changed to a temperature in the range of from 300 to 350⁰C and a pressure in the range of from 1 Pa to 300 kPa. The rate of change of the temperature and pressure conditions can be optimised to achieve maximum economical benefit.

It may be beneficial also to utilise a sweep gas in mode 2. Such a gas may be introduced into the reactor, when in mode 2, in order to sweep or carry the desorbed glycol (s) out of the reactor and onto a separation section or unit. Suitably such a sweep gas would be an inert gas, such as steam or preferably nitrogen. The product MEG may be deposited on the catalyst surface with the other glycols or preferably remains in the vapour phase. Where the MEG is predominantly produced and maintained in the vapour phase, the reactor mode switching is still necessary to prevent deactivation of the catalyst by the higher glycols deposited thereon.

The vapour phase mixture of unconverted EO and water and mono-ethylene glycol product coming out of the reactor zone that is operated in mode 1, is suitably led through a downstream zone that is operated at a lower temperature and/or higher pressure, where the mono- ethylene glycol can be separated from the gaseous mixture by condensation. If desired, a water/mono-ethylene glycol mixture can be separated by condensation from the vapour phase by further lowering the temperature or by further increasing the pressure. This separation or condensation zone may also operated in a 'swing' mode. After having condensed mono-ethylene glycol or mono-ethylene glycol and a part of the water from the reactor product vapour stream, this stream may be directed to a second, similar separator or condensation zone. The condensed product is removed from the first zone. During the removal, the temperature or pressure of the separation zone may, or may not, be changed from the conditions during condensation. Where MEG is trapped on the catalyst surface, then it may similarly be recovered from the vapour mixture produced in mode 2 by evaporation.

Quick changes in temperature are possible but somewhat difficult in the conventional large vapour phase reactors that are normally used in the process industry, particularly those utilizing heterogeneous catalyst, because of the large gas and catalyst volumes, large heat transfer medium volumes, the steel mass (heat sink) and heat transfer limitations. Therefore the reaction, as well as the separation, is advantageously performed in the process microchannels of one or more microchannel reactors, which enables fast and accurate temperature plus pressure change and control. The downstream condensation zone(s) may also advantageously be one or more process microchannels of one or more further microchannel apparatus.

The present invention accordingly provides a process for the preparation of a mono-alkylene glycol by the reaction of a corresponding alkylene oxide and water, which process comprises a) reacting the alkylene oxide and water in a first reactor under a first set of conditions and in the presence of a catalyst so as to achieve vapour phase conversion to the mono-alkylene glycol, b) altering the conditions in the first reactor to a second set of conditions whereby glycols deposited on the surface of the catalyst are removed, c) re-establishing the first set of conditions in the first reactor in order to repeat step a) , and d) recovering the mono-alkylene glycol from a vapour phase mixture produced in step a) and/or step b) .

In a preferred embodiment the process is operated using two reactors whereby simultaneously with operating step b) , the gaseous alkylene oxide and water feeds are switched to a second reactor which is operating under the first set of conditions. When the first set of conditions are re-established in the first reactor under step c) of the process of the invention, the feeds are switched back to the first reactor and the conditions of the second reactor are changed to the second set of conditions.

The vapour product stream from the first reactor thus comprises the mono-alkylene glycol and possibly heavier components, the latter may be for example di- and tri-ethylene glycol. However, because of the preferential deposition of the heavier glycols onto the surface of the catalyst in the first reactor, the amount of these 'heavies' in the product stream of step a) will be low. Greater amounts of heavier glycols may be present in the product mixture from step b) . In both cases the heavier glycols where present can optionally be removed via a distillation column (a 'topping and tailing column') where the pure mono-ethylene glycol is withdrawn as a side-stream and these heavies are drawn off as a separate stream and utilised, or incinerated as waste.

The nature of the two sets of conditions may vary, however generally the conditions will be such that a direct change from the first set of conditions to the second set of conditions will cause the evaporation of glycol deposited on the catalyst.

Most preferably the first reactor, and second reactor where present, is a microchannel apparatus such as described herein. This provides the additional advantage of good control of the conditions to be changed. The first reactor may be in a first set of process microchannels operating under the first set of conditions, and then the conditions can be changed by use of a heat transfer medium, flowing through heat exchange channels, to change, for example, the temperature conditions to provide evaporation of the glycol. Two microchannel apparatus may be provided working in tandem with the glycol-containing feed being switched between the two microchannel apparatus so that a continuous operation can occur, with a first microchannel performing the vapour phase hydration while the second is in evaporation mode, and then switching the feed so that the second apparatus performs the reaaction while the first performs the evaporation. Reference is made to the publications US-Bl-

6,508,862 and WO 2005/032693 which describe microchannel apparatus used in temperature swing sorption for fluids. The apparatus and control mechanisms may be readily adapted to operating the process of the present invention by those skilled in the art.

This aspect of the process of the invention may additionally be utilised in conjunction with the use of microchannel apparatus to perform conversion of alkylene oxide to glycol as described above. The present invention will now be illustrated by the following Examples. EXAMPLES Example 1

In a plant 400,000 mt/a ethylene oxide is produced in an combined EO and glycols plant. 200,000 mt/a of this ethylene oxide is fed to the integrated glycol production unit. In the glycol production unit of FIG. 3, ethylene oxide via line 1 is subsequently mixed with fresh water (fed via line 2) and recycle water (10) in vessel 3, preheated in heat exchangers 4, and is reacted without catalysis with water to form mono-ethylene glycol in reactor 5. Since EO not only reacts with water to mono- ethylene glycol but simultaneously with glycols, not only mono-ethylene glycol is formed but also the byproducts di-ethylene glycol, tri-ethylene glycol and even higher glycols are formed. The amount and ratio of these glycols is heavily determined by the concentration of these glycols inside the reactor. High concentration of water favours a high yield of mono-ethylene glycol. On the other hand at a low concentration of water a lot of di- ethylene glycol, tri-ethylene glycol and the heavier glycols are formed which is in most cases undesired. In this example, the water to EO ratio of the reactor feed is adjusted to be 10:1 to achieve a ratio of 10:1:0.1 mono-ethylene glycol: di-ethylene glycol : tri-ethylene glycol in the reactor outlet stream 18.

The water is not only used as feedstock to form glycols and as dilution agent to control the ratio of glycols, but also acts as a heat sink to control the outlet temperature of the reactor outlet stream, since the reactions in the glycol reactor are strongly exothermic. Since the product glycol is produced in an abundance of water the mixture needs to be dehydrated before separation and purification of the glycol mixture can be achieved. Dehydration is typically carried out in a train of concentrator and dehydrator columns 6. The water streams from the top of these columns are combined (10) and recycled to the reactor feed. The water-free bottom stream 11 of dehydrator 6 is fed to the glycol purification section and the glycol mixture is separated into its four product steams mono-ethylene glycol (12), di-ethylene glycol (14), tri-ethylene glycol (16) and heavier glycols (17) . It is evident that the need for large quantities of water will lead to a lot of equipment needed for dehydration, and the dehydration itself will demand a lot of energy use in the form of steam used in the concentrator and dehydrator column reboilers. By making use of the present invention, the reaction of EO to mono- ethylene glycol is performed inside the process microchannels of a microchannel reactor. The temperature can be easily controlled because of the excellent heat transfer, and a large amount of water for heat sink is not needed anymore. The reaction to MEG can be catalyzed to suppress the formation of di-ethylene glycol, tri- ethylene glycol and other heavy glycols. A catalyst may be present in one or more process microchannels. Thus the number of dehydrator columns can be reduced and energy for dehydration can be saved. By using a catalyst the selectivity to mono-ethylene glycol can additionally be increased, enabling reduction of the size of glycol purification equipment. Example 2

A Co/Zn/Al hydrotalcite-type catalyst was prepared as follows: 24g of Co (NO₃) ₂ ■ 6H₂O was dissolved in 200 ml demi-water, 93.8g of Al (NO₃) ₃.9H₂O was dissolved in 300 ml demi-water and 124.2g of Zn (NO₃) ₂.6H₂O in 300 ml demiwater. These three solutions were mixed forming solution A and stored in a drip-flask. Then 7Og NaOH was dissolved in 200 ml demi-water and 53g Na₂CO₃ in 250 ml demiwater. The latter was heated to 50⁰C until clear. Both Na solutions were subsequently mixed in a 2 litre round bottom and stirred for 0.5 hour while cooling to

<5°C. This is solution B. In the next step solution A was added slowly (ca. 8 ml/min totalling 1.5 hours) to B while keeping the temperature below 5⁰C. A thick pink gel was formed. After mixing of A and B the resulting slurry was heated to 60°C and stirred for another 1.5 hours.

Then the heater was turned off and stirring was continued for the night. The next day the slurry was filtered and washed 3 times with demi-water. Half of the filter cake was dried at 120⁰C, the other half was calcined at 425°C for 12 hrs in air. The target composition was

Co₂ZnI₀Al₆. (CO₃) x .yH₂0.

Example 3

The microchannel reactor will be assembled in accordance with methods known from WO-A-2004/099113, and references cited therein.

A microchannel reactor will comprise process microchannels, heat exchange microchannels, and feed channels . The process microchannel section will comprise a hydrolysis catalyst comprising cobalt, zinc and alumina as described above.

The process microchannel reactor will be filled with a hydrolysis catalyst that will be prepared by milling and sieving a hydrotalcite-type catalyst. The catalyst will be firstly conditioned under N₂ and H₂O₅ for at least

1 hour at reaction temperature before adding the reaction gas mixture.

The process section will be heated at 275°C by heat exchange with the heat exchange fluid flowing in the first heat exchange microchannel, while water is fed through an opening positioned at the upstream end of the process microchannels. This process section will be maintained at 50OkPa. Ethylene oxide gas will be fed through a second set of feed channels upstream of the process microchannels.

The molar ratio of ethylene oxide to water will be 1:10. As an alternative, ethylene oxide and water (molar ratio 1 to 10) will be fed into the microchannel process section using one feed channel upstream of the process section . The product mixture exiting the process section, containing the desired mono-ethylene glycol will be further processed and/or purified by a conventional method. Example 4

A microchannel reactor will comprise process microchannels, heat exchange microchannels, and feed channels .

The process microchannel section will comprise a hydrolysis catalyst comprising cobalt, zinc and alumina as described above. The process microchannel reactor will be filled with a hydrolysis catalyst that will be prepared by milling and sieving a hydrotalcite type catalyst. The catalyst will be firstly conditioned under N₂ and H₂O₉ for at least 1 hour at reaction temperature before adding the reaction gas mixture.

Two such microchannel reactors will be operated in swing mode in parallel, in which simultaneously one reactor is operated with EO/water feed to produce glycol and the other reactor is operated at higher temperature and lower pressure to evaporate condensed higher glycols from the catalyst surface.

The process section will be heated at 275⁰C by heat exchange with the heat exchange fluid flowing in the first heat exchange microchannel, while water is fed through an opening positioned at the upstream end of the process microchannels. This process section will be maintained at 50OkPa. Ethylene oxide gas will be fed through a second set of feed channels upstream of the process microchannels . The molar ratio of ethylene oxide to water will be 1:10.

As an alternative ethylene oxide and water (molar ratio 1 to 10) will be fed into the microchannel process section using one feed channel upstream of the process section .

Simultaneously the second microchannel reactor will be operated at 35O⁰C and 200 kPa without feeding ethylene oxide/water.

Conditions and feed of both parallel reactors will be changed every 30 seconds.

The product mixture exiting the process section, containing the desired mono-ethylene glycol may be further processed and/or purified by a suitable method. Example 5

The microchannel reactor will be assembled in accordance with methods known from WO-A-2004/099113, and references cited therein. A microchannel reactor will comprise process microchannels, heat exchange microchannels, and feed channels .

The process microchannel section will comprise a hydrolysis catalyst comprising cobalt, zinc and alumina as described above.

The process microchannel reactor will be filled with a hydrolysis catalyst that will be prepared by milling and sieving a hydrotalcite type catalyst. The catalyst will be firstly conditioned under N₂ and H₂O_g for at least 1 hour at reaction temperature before adding the reaction gas mixture.

The process section will be heated at 275°C by heat exchange with the heat exchange fluid flowing in the first heat exchange microchannel, while water is fed through an opening positioned at the upstream end of the process microchannels . This process section will be maintained at 50OkPa.

Ethylene oxide gas will be fed through a second set of feed channels upstream of the process microchannels. The molar ratio of ethylene oxide to water will be 1:10. As an alternative, ethylene oxide and water (molar ratio 1 to 10) will be fed into the microchannel process section using one feed channel upstream of the process section.

The vapour phase product mixture exiting the process section, containing unreacted ethylene oxide, water, and the desired mono-ethylene glycol will be further processed in a second set of parallel microchannel reactors operating in swing mode. One reactor will be fed with the product mixture from the process section and will operate at a lower temperature of 120⁰C to enable condensation of the monoethylene glycol, while the unreacted ethylene oxide and water will be recycled back to the process microchannel reactor. The other parallel reactor will operate at an elevated temperature of 200⁰C to vaporize condensed monoethylene glycol for further processing and purification. Conditions and feed of both parallel reactors will be changed every 60 seconds.

Claims

C L A I M S

1. A process for the preparation of an alkylene glycol by the reaction of a corresponding alkylene oxide and water, which process comprises a) flowing the alkylene oxide and water through a microchannel reactor, wherein the oxide and water undergo an exothermic reaction to form the corresponding alkylene glycol, b) transferring heat from the microchannel reactor to a heat transfer medium, and c) recovering the alkylene glycol product from the microchannel reactor.

2. A process as claimed in claim 1, wherein the microchannel reactor contains a catalyst suitable for the catalytic hydrolysis of alkylene oxide.

3. A process as claimed in claim 2, wherein the catalyst is a homogeneous catalyst present in the reaction mixture.

4. A process as claimed in claim 2, wherein the catalyst is a heterogeneous catalyst present as a solid catalyst in, or as a coating on the walls of, one or more process microchannels present in the microchannel reactor.

5. A process for the preparation of a mono-alkylene glycol by the reaction of a corresponding alkylene oxide and water, which process comprises a) reacting the alkylene oxide and water in a first reactor under a first set of conditions and in the presence of a catalyst so as to achieve vapour phase conversion to the mono- alkylene glycol, b) altering the conditions in the first reactor to a second set of conditions whereby glycols deposited on the surface of the catalyst are removed, c) re-establishing the first set of conditions in the first reactor in order to repeat step a) , and d) recovering the mono-alkylene glycol from the vapour phase mixture produced in step a) and/or step b) .

6. A process as claimed in claim 5, wherein the first reactor is a microchannel apparatus.

7. A process as claimed in claim 5 or claim 6, wherein step a) is carried out using the process as claimed in any one of claims 1 to 4.

8. A process as claimed in any one of claims 5 to 7, wherein the conditions of step b) are altered by either or both of the following: i) change of temperature ii) change of pressure.

9. A process as claimed in any one of claims 5 to 8, wherein the alkylene oxide is ethylene oxide and the mono-alkylene glycol prepared is mono-ethylene glycol .