WO2000050369A1

WO2000050369A1 - Alcohol production

Info

Publication number: WO2000050369A1
Application number: PCT/GB2000/000485
Authority: WO
Inventors: Gordon John Haining; Mark Julian Howard; Nigel Keith Potter
Original assignee: Bp Chemicals Limited
Priority date: 1999-02-25
Filing date: 2000-02-14
Publication date: 2000-08-31
Also published as: GB9904239D0; AU2454800A

Abstract

The invention relates to a process for the hydration of an olefin, said process comprising: introducing a feed comprising the olefin and water into a reactor containing an olefin hydration catalyst, and reacting the olefin with the water in the presence of said catalyst in the vapour phase, characterised in that the amount of olefin in the feed to the reactor is from 10-60 mol %.

Description

ALCOHOL PRODUCTION

The present invention relates to a process for the hydration of an olefin to its corresponding alcohol using a catalyst composition comprising a heteropolyacid catalyst on a support.

The hydration of olefins, such as ethylene or propylene, to their corresponding alcohols is well known. Such hydration reactions are carried out in the vapour phase using phosphoric acid deposited on a siliceous support as a catalyst. Examples of such processes are described in GB-A-1570650, US-A-4808559, GB-A-1371905, US-A- 4038211, US-A-4012452, GB-A-1476534, GB-A-1306141, US-A-3996338 and CAN- A-844004. US-A-2173187 also describes a process for the hydration of olefins. This process employs a heteropolyacid whose complex anion includes one element from group VIA of the Periodic Table as a catalyst. It is stated that the catalyst can be used with or without a support. When used, the support is preferably silica gel. However, other siliceous supports such as silicic acid, Japanese acid clay, bentonite, kieselguhr, or asbestos are also said to be suitable.

EP-A-0704240 describes another vapour phase process for producing alcohols by olefin hydration. In particular, the reference describes an olefin hydration process which employs a heteropolyacid catalyst supported on a siliceous support in the form of extrudates or pellets. The support is an amorphous synthetic silica support formed from non-porous fumed silica powder. The document provides experimental evidence that such heteropolyacid catalysts are superior in activity to conventional supported phosphoric acid catalysts. In all the examples in EP-A-0704240, the olefin to be hydrated is substantially pure (>99%v/v). Although olefin feedstocks of lesser purity may be employed, such feedstocks have to be purified, for example, by distillation in order to bring their purity up to the required levels. Under the process conditions, only a small percentage (eg 5 to 10 mol %) of the olefin in the feedstock is hydrated. Thus, alcohol only forms a small proportion of the reactor effluent. A significant amount of unreacted olefin is also present, and this is compressed and recycled to the reactor. This recycled olefin is supplemented with fresh olefin, such that the feed stream entering the reactor inlet comprises recycle gas, fresh olefin and water. The olefin concentration at the reactor inlet is typically above 60 mol %, for example, 80 mol %.

In order to maintain the olefin concentration at the reactor inlet at this desired value, the reactor effluent, and fresh olefin are generally purified using a series of complex purification steps. These steps are necessary to prevent the build-up of inert gases in the reactor feed. As the build-up of inert gas would occur at the expense of olefin in the reactor, this was believed to be detrimental to the productivity of the process.

To compensate for such a decrease in productivity, the recycle flow rate of the process can be increased. Such a solution, however, is undesirable, as the compressors and pipework required to cope with the increased flow rates would have to be significantly larger, adding to both equipment and production costs.

We have now found that, with the correct choice of catalyst, the presence of inert gases in the reactor feed is not as detrimental to the process productivity as conventionally believed. Thus, the hardware requirements previously believed to be necessary to operate using feeds with lower olefin concentrations have now been found to be unnecessary. In fact, it has been found that by reducing the amount of olefin in the reactor feed, olefin hydration can be operated more efficiently, and at significantly lower costs.

Accordingly, the present invention is a process for the hydration of an olefin, said process comprising: introducing a feed comprising the olefin and water into a reactor containing a heteropolyacid catalyst, and reacting the olefin with the water in the presence of said catalyst in the vapour phase, characterised in that the amount of olefin in the feed to the reactor is from 10 - 60 mol %, preferably, 20 - 50 mol %, more preferably, 20 - 40 mol %, for example, 25 - 35 mol %. Olefins that may be used in the process of the present invention include ethylene and propylene. Preferably, either ethylene or propylene is employed as the olefin.

A feature of the present invention is that the amount of olefin in the reactor feed is less than the amount conventionally employed for olefin hydration. Thus, the present process may be carried out at overall reaction pressures comparable to those conventionally employed, using olefin feed streams of reduced purity. An advantage of carrying out the process in this manner is that cheaper, less pure feed stocks may be employed.

As with the conventional process, the conversion of olefin to alcohol is not 100% efficient. Thus, the alcohol may only form a small proportion (e.g. ca 3-4 mol%) of the reactor effluent, the remainder of which comprises unreacted olefin. The alcohol may be recovered from the reactor effluent, for example, by distillation or as an aqueous solution. The remainder of the effluent is recycled.

Because the amount of olefin in the feed to the reactor is less than the amount conventionally employed, a significant portion of the reactor effluent may be compressed and returned directly to the reactor, without purification. The remainder of the effluent, however, is generally purified before it is returned to the reactor.

A number of purification stages may be employed. For example, the stream may be introduced into a heavy ends removal unit, which separates heavy end impurities from the stream. Such a unit may operate under reflux. Preferably, a refluxing fluid, such as a fresh olefin, is introduced into the heavy end removal unit to aid the separation of heavy end impurities from the rest of the olefin stream. In a preferred embodiment, the heavy end removal unit is in the form of a vent gas fractionator.

The purified stream, now substantially free of heavy end impurities, may be purified further. A stream comprising propylene, for example, may comprise saturated impurities such as propane. Such saturated impurities may be removed using a distillation column. In removing such saturated impurities, a fraction of the olefin may be removed together with the saturated impurities. Such olefin/alkane mixtures may be sold or otherwise used, for example, as a fuel. Advantageously, the distillation column may also be employed to remove light end impurities such as methane from the stream.

Thus purified, the olefin-containing stream may then be returned to the reactor. Preferably, this stream is supplemented with a fresh olefin feed. A significant advantage of the present invention is that it is not necessary for the fresh olefin feed to be as pure as that employed in the conventional process. This is because the concentration of olefin at the reactor inlet need only be between 10 and 60 mol %, in contrast to concentrations of 80 mol % generally employed in conventional processes.

In view of the above, the fresh olefin feed employed in the present process may be of refinery grade. Suitable refinery grade streams include mixed hydrocarbon feeds produced, for example, by fluid catalytic cracking. Typically, such streams having a purity of between 50 and 95 w/w%, preferably between 60 and 80 w/w%, and more preferably, between 65 and 75 w/w%. Although such refinery grade olefins may be employed in conventional olefin hydration processes, they have to be purified to at least 98.5 w/w %, in order to ensure that the desired reactor inlet concentrations are achieved. The purification stages required to achieve such high levels of purity add complexity and cost to the overall process, as large distillation units are required to cope with the volume of fresh olefin to be purified.

Instead of using an olefin feed of lesser purity, the reduced partial pressures of olefin required by the present invention may be achieved by using feeds of conventionally high purity, but by running the process at a lower than normal overall reaction pressure. Thus, whereas the hydration of ethylene is conventionally carried out at pressures of 4500 to 8000 kPa, ethylene hydration according to the present invention may be carried out at 4500 kPa and below, for example, 300 to 4000 kPa, preferably, 2000 to 4000 kPa. Similarly, whereas the hydration of propylene may be carried out between

3000 to 5000 kPa, propylene hydration according to the present invention may be carried out at 3000kPa and below, for example, 300 - 3000 kPa, for example, 1500 - 3000 kPa. The main advantage of carrying out the process at a reduced overall pressure is that the process may be carried out more economically. The hydration reaction is catalysed by a heteropolyacid catalyst. The term

"heteropolyacid" as used herein and throughout the specification in the context of the catalyst is meant to include the free acids. The heteropolyacids used to prepare the hydration catalysts of the present invention therefore include inter alia the free acids and co-ordination type partial acid salts thereof in which the anion is a complex, high molecular weight entity. Typically, the anion comprises 2-18 oxygen-linked polyvalent metal atoms, which are called peripheral atoms. These peripheral atoms surround one or more central atoms in a symmetrical manner. The peripheral atoms are usually one or more of molybdenum, tungsten, vanadium, niobium, tantalum and other metals. The central atoms are usually silicon or phosphorus but can comprise any one of a large variety of atoms from Groups I-NIII in the Periodic Table of elements. These include, for instance, cupric ions; divalent beryllium, zinc, cobalt or nickel ions; trivalent boron, aluminium, gallium, iron, cerium, arsenic, antimony, phosphorus, bismuth, chromium or rhodium ions; tetravalent silicon, germanium, tin, titanium, zirconium, vanadium, sulphur, tellurium, manganese nickel, platinum, thorium, hafnium, cerium ions and other rare earth ions; pentavalent phosphorus, arsenic, vanadium, antimony ions; hexavalent tellurium ions; and heptavalent iodine ions. Such heteropolyacids are also known as "polyoxoanions", "polyoxometallates" or "metal oxide clusters". The structures of some of the well known anions are named after the original researchers in this field and are known e.g. as Keggin, Wells-Dawson and Anderson-Evans-PerlofF structures.

Heteropolyacids usually have a high molecular weight eg in the range from 700- 8500 and include dimeric complexes. They have a relatively high solubility in polar solvents such as water or other oxygenated solvents, especially if they are free acids and in the case of several salts, and their solubility can be controlled by choosing the appropriate counter-ions. Specific examples of heteropolyacids that may be used as the catalysts in the present invention include: 12-tungstophosphoric acid - H₃[PWι₂O ₀].xH₂0 12-molybdophosphoric acid - H₃[PMθι₂O₄₀].xH₂O

12-tungstosilicic acid - H [SiWι₂O ₀].xH₂O

12-molybdosilicic acid - H₄[SiMθι₂O₄o].xH₂O

Caesium hydrogen tungstosilicate - Cs₃H[SiW₁₂O o].xH₂O

The heteropolyacid catalyst may be used as a free acid or as a partial acid salt thereof. Preferably, such acids are supported on a siliceous support. The siliceous support is suitably in the form of granules, beads, globules, agglomerates, extrudates or pellets. The siliceous support used can be derived from an amorphous, non-porous synthetic silica especially fumed silica, such as those produced by flame hydrolysis of SiCl . Specific examples of such siliceous supports include Support 350 made by pelletisation of AEROSIL® 200 (both ex Degussa). This pelletisation procedure is suitably carried out by the process described in US Patent 5,086,031 (see especially the Examples) and is incorporated herein by reference. Such a process of pelletisation or extrusion does not involve any steam treatment steps and the porosity of the support is derived from the interstices formed during the pelletisation or extrusion step of the non- porous silica The silica support is suitably in the form of pellets or beads or are globular in shape having an average particle diameter of 2 to 10 mm, preferably 4 to 6 mm. The siliceous support suitably has a pore volume in the range from 0.3-1.2 ml/g, preferably from 0.6-1.0 ml/g. The support suitably has a crush strength of at least 2 Kg force, suitably at least 5 Kg force, preferably at least 6 Kg and more preferably at least 7 Kg. The crush strengths quoted are based on an average of that determined for each set of 50 beads/globules on a CHATTILLON tester which measures the minimum force necessary to crush a particle between parallel plates. The bulk density of the support is suitably at least 380 g/1, preferably at least 440 g/1.

The support suitably has an average pore radius (prior to use) of 10 to 500 A preferably an average pore radius of 30 to 100 . In order to achieve optimum performance, the siliceous support is suitably free of extraneous metals or elements which might adversely affect the catalytic activity of the system. The siliceous support suitably has at least 99% w/w purity, i.e. the impurities are less than 1% w/w, preferably less than 0.60% w/w and more preferably less than 0.30% w/w. Other silica supports are the Grace 57 and 1371 grades of silica gel. In particular,

Grace 57 grade silica has a bulk density of about 0.4 g/ml and an average pore volume of 1.15 ml/g. Grace silica grade No. 1371 has an average bulk density of about 0.39 g/ml, an average pore volume of about 1.15 ml/g and an average particle size ranging from about 0.1-3.5 mm. These supports can be used as received. The impregnated support is suitably prepared by dissolving the heteropolyacid, which is preferably a tungstosilicic acid, in eg distilled water, and then adding the support to the aqueous solution so formed. The support is suitably left to soak in the acid solution for a duration of several hours, with periodic manual stirring, after which time it is suitably filtered using a Buchner funnel in order to remove any excess acid.

The wet catalyst thus formed is then suitably placed in an oven at elevated temperature for several hours to dry, after which time it is allowed to cool to ambient temperature in a desiccator. The catalyst can also be dried suitably by using a flow of heated gas such as e.g. nitrogen or air. The catalyst loading in g/litre was determined by deducting the weight of the support used from the weight of the catalyst on drying.

Alternatively, the support may be impregnated with the catalyst using the incipient wetness technique and dried by using a flow of heated gas such as e.g. nitrogen or air. This supported catalyst (measured by weight) can then be used in the process of the present invention. The amount of heteropolyacid deposited/impregnated on the support for use as catalyst in the reaction is suitably in the range of from 5 to 60% by weight, preferably from 20 to 50% by weight, more preferably from 20-45 % by weight, for example, 20% to 35% by weight (corresponding to a loading in the range of about 100- 215 g/litre) based on the total weight of the heteropolyacid and the support.

The process is suitably carried out using the following reaction conditions: a. the mole ratio of water to olefin passing through the reactor is suitably in the range from 0.1-3.0, preferably 0.1-1.0 b. the gas hourly space velocity (GHSN) of the water/olefm mixture is suitably from 0.010 to 0.25 g/min/cm³ of the catalyst system, preferably from 0.02-0.10 g/min/cπ of the catalyst composition. c. the heteropolyacid catalyst concentration is from 5 to 60% w/w based on the total weight of the catalyst system, preferably from 20 to 35 % w/w of the total weight of the catalyst composition. The supported heteropolyacid catalysts may also be further modified by the addition of phosphoric acid or other mineral acids thereto.

The hydration reaction is carried out by placing the catalyst composition in a reactor, sealing the reactor and then heating the catalyst composition to the reaction temperature. The catalyst composition may be heated to a temperature of from 160-300° C, preferably from 165-180°C for propylene hydration, and preferably, 185 to 240°C for ethylene hydration. Preferably, the catalyst composition is heated under a purge of dry nitrogen. For the hydration of propylene, the overall reaction pressure may be between 300 to 5000 kPa, preferably, 1500 to 4000 kPa. Propylene partial pressures in the feed to the reactor may be between 20 and 2700 kPa, preferably 100 and 2500kPa. For the hydration of ethylene, the overall reaction pressure may be between 300 and 8000 kPa, preferably, between 2000 and 4500 kPa. Ethylene partial pressures in the feed to the reactor may be between 20 and 4350 kPa, preferably, between 100 and 2500 kPa.

According to a further aspect, the present invention is a process for the hydration of propylene with steam in the vapour phase and in the presence of a catalyst comprising a heteropolyacid on a siliceous support in an olefin hydration reactor, characterised in that the amount of propylene in the hydrocarbon feed to the reactor is from 20-40 mol %.

These and other aspects of the present invention will now be described, by way of example, with reference to the accompanying drawings, in which:

Figure 1 is a schematic view of an apparatus for carrying out a conventional propylene hydration process, Figure 2 is a schematic view of an apparatus for carrying out a propylene hydration process in accordance with a first embodiment of the present invention, and

Figures 3 and 4 are graphs showing the results of Example 2.

Referring to Figure 1, there is depicted an apparatus comprising a reactor 10 which is coupled to a main gas scrubber 12. The main gas scrubber 12 is connected to a vent gas fractionator 14, which in turn, is coupled to a propylene splitter 16. Connected to the propylene splitter 16 is a propylene fractionator 18. In an alternative embodiment (not shown), this fractionator 18 is effectively the integral with the splitter 16.

In operation, water and propylene are introduced into the reactor 10 at the reactor inlet 20. The concentration of propylene at the reactor inlet 20 is 80 mol %. The ensuing reaction produces a product stream comprising iso-propyl alcohol. The product stream also comprises significant amounts of unreacted propylene.

The product stream is cooled using a reactor feed/effluent exchanger system (not shown). After cooling, the stream is removed from the base of the reactor 10, and introduced into the main gas scrubber 12. Water is introduced into the scrubber via inlet 22 to dissolve the iso-propyl alcohol. The resulting solution (approximately 12 % iso- propyl alcohol) is recovered from the base of the scrubber 12 via outlet 24.

The remainder of the product stream, now substantially free of iso-propyl alcohol, is removed from the top of the scrubber 12. A portion of this stream 26 is recycled to the reactor 10 via a recycle gas compressor 28. Water is introduced into this stream portion 26 via line 27. The remaining portion 30 of the stream is purged at a rate of 4536-5443 kg/hr (10-12 klb/hr), [This is for a 300tpd IPA unit] into the vent gas fractionator 14. This purged stream 30 may comprise as much as 95 mol % propylene.

In the fractionator 14, the purged stream 30 is mixed with fresh propylene 35, which is introduced into the fractionator 14 as reflux. The fresh propylene feedstream employed is of refinery grade, having a purity of 72 w/w%. The addition of propylene 35 aids the removal of heavy ends 32 from the fresh propylene/purge stream mixture. The fresh propylene added in this manner accounts for approximately 30% of the fresh propylene added to the apparatus. The remaining 70% of propylene is added to the propylene splitter, as will be described in further detail below.

Once the heavy ends 32 are removed from the purged stream/fresh propylene mixture 33, the mixture is removed from the top of the fractionator 14 and introduced into the propylene splitter 16, as an overhead stream 33.

In the splitter 16, the stream 33 is mixed with further fresh propylene, which is introduced into the splitter via line 34. The splitter 16 purifies the resulting propylene mixture by removing light end impurities as an overhead stream 36. Other impurities such as propane are removed from the propylene using the second fractionator 18, which is coupled to the base of the splitter 16.

Purified propylene (98.5 mol %) is recovered via stream 38. This is mixed with the water and recycled propylene in stream 26, such that the concentration of propylene at the reactor inlet is 80 mol %.

Reference is now made to Figure 2, which depicts an apparatus for carrying out a propylene hydration process of an embodiment of the present invention. The apparatus of Figure 2 is similar to that of Figure 1 and corresponding parts have been numbered with like numerals increased by 100. The apparatus of Figure 2 is designed to perform the olefin hydration process at comparable overall pressures as the process described in relation to Figure 1. However, the olefin feed introduced into the reactor need not be as pure as the olefin employed in the process of Figure 1. Thus, unlike the apparatus of Figure 1, the second fractionator 18 is absent from the apparatus of Figure 2. A further difference is that fresh propylene 134 is not introduced into the propylene splitter 116, but directly into feedstream 138. Thus, 70 to 80 % of the fresh propylene added is introduced directly into the feedstream 138. The remainder of the fresh propylene added to the apparatus is used as reflux in the fractionator 114.

In operation, water and propylene are introduced into the reactor 110 at the reactor inlet 120. The concentration of propylene at the reactor inlet 120 is 10 - 60 mol %, for example, 30 mol %. The ensuing reaction produces a product stream comprising iso- propyl alcohol, as well as significant amounts of unreacted water and propylene.

Iso-propyl alcohol is isolated from the product stream using the main gas scrubber 112 in the manner described with reference to Figure 1. The remaining stream is then removed from the top of the scrubber 112. As in Figure 1, a portion 126 of this stream is recycled to the reactor 110 via a recycle gas compressor 128, whilst the remainder of the stream 130 is purged into the vent gas fractionator 114. Unlike the purge stream of Figure 1, the concentration of propylene in the purge stream 130 is typically as low as 25 mol %. Also, the number and size of the purification devices required to purify the unreacted propylene in the purge stream for the purposes of recycle are substantially reduced. For example, the size of the vent stream to the vent gas fractionator 114 of Figure 2 is smaller (by approximately 20% by vol) than the corresponding fractionator 14 of Figure 1.

In the vent gas fractionator 114, the purge stream 130 is mixed with fresh propylene 135, which aids the removal of heavy ends 132 from the propylene mixture. Once the heavy ends 132 are removed, the remainder 133 of the stream is removed from the top of the fractionator 114 and introduced into the propylene splitter 116 for the last stage of purification.

The splitter 116 purifies the resulting propylene mixture by removing light end impurities as an overhead stream 136. Other impurities such as propane are removed from the base of the splitter 116. The purified stream is then recycled to the reactor via line 138.

Table 1 below summarises the conditions used when operating the processes described with reference to Figures 1 and 2. Table 1

* - The refinery grade propylene feed had 72 % w/w propylene.

The operating conditions of the present invention were such that the IP A production rate was the same for both Figures 1 and 2. The advantages of the embodiment of Figure 2 are as follows:

1. The required amount of propylene feed capacity is greatly reduced since: a. the splitter is no longer treating the whole of the propylene feed b. the process can operate with a lower propylene product purity c. the size of the reactor vent is significantly reduced. 2. No additional heat exchange area is required in the olefin hydration section. This is because the recycle gas, propylene and process water (including added demineralised water (DMW)) flows are at least as favourable as, if not better, than that used in conventional processes.

3. The propylene loss in the splitting section is substantially halved. 4. At 15% propylene conversion per pass, the recycle gas flow is approximately 90% of that needed in conventional processes.

These advantages were confirmed by simulating the olefin hydration process using

ASPEN PLUS Release 9.2 supplied by AspenTech UK Ltd (hereinafter "ASPEN").

ASPEN is a sequential modular simulation program design to model chemical processes using the underlying physical relationships of mass and energy balances, equilibrium relationships and reaction kinetics. If all these advantages are implemented in a plant, it would result in almost halving the capital expenditure required in the propylene splitting / recovery section of the plant compared with the conventional processes.

Another feature of the embodiment of Figure 2 is that the size of the vent stream to the vent gas fractionator 114 of Figure 2 is substantially less than that of Figure 1 (eg. by as much as 20% when using a heteropolyacid on a siliceous support). This was confirmed using ASPEN' s distillation model on two different mixtures of propylene/propane feeds, one at 95 mol % propylene and the other at 25 mol % propylene. In both cases, 99.5% recovery of propylene was specified. It was found that the relationship between reflux ratio and the number of trays for the two feedstocks tested indicated that treating a 25 mol % propylene vent requires approximately half the number of trays as the old/existing system, thereby enabling the elimination of at least one of the columns.

Table 2 below illustrates the improvements achievable when a refinery grade propylene is fed directly into the reactor loop operating at a propylene conversion of 12- 15% per pass. As can be seen from the results, the demands conventionally placed on the propylene splitter are relieved. Thus, it is no longer necessary to (a) process the whole of the propylene feed and/or (b) employ propylene of such high purity as used hitherto. In fact, maintaining the desired recovery of propylene from the purge in itself forces up the propylene product purity to approximately 85 mol %. In short: i. The number of separation stages required to recover propylene are halved, ii. The loss of propylene during the recovery stages is halved, iii. The reactor loop purge is 25-30% smaller than used hitherto.

TABLE 2

Examples

Example 1; Catalyst Preparation 3.392kg (8.5 litres) of Grace 57 silica gel support were placed in a large polyethylene lined container. 8 litres of a 37.33 %wt/vol aqueous solution of 12- tungstosilicic acid (HUSi W₁₂O₄₀.xH₂0, ex. JMN, in distilled water ) were then slowly added. The support was then left to soak in the acid solution for 24 hours at atmospheric pressure and ambient temperature. After soaking, the support was filtered using a Buchner funnel to remove any excess acid solution present. The resulting catalyst was then dried in a standard fan assisted laboratory drying oven for 24 hours at 100°C, followed by a further 24 hours at 120°C .

After drying, and cooling back to ambient temperature, the bulk density of the catalyst was found to be 550.5 g/1. Given that the bulk density of the support material (Grace 57 silica gel) was 399 g/1, this indicates that the catalyst has an "acid loading" of ca. 151.5 g/1. EXAMPLE 2: ACTIVITY TESTING OF CATALYST

Eight litres of the catalyst produced in Example 1 were tested for ethylene hydration ability in an 8 litre pilot plant. In this example, the olefin hydration process was carried out using an ethylene feed of a purity comparable to that used in conventional ethylene hydration reactions. The process, however, was operated at a reduced overall pressure.

The 8 litre pilot plant employed is a continuous flow pilot plant, designed to accurately simulate the reaction section of a gas phase ethylene hydration plant. The pilot plant also has facilities for recycling all unreacted ethylene, and the majority of the by-products produced.

Fresh ethylene is fed to the plant as a gas from a high pressure ethylene compressor. Steam is generated from liquid water fed (by diaphragm metering pump) into a "drip-feed" vapouriser, and the feeds are combined with recycled ethylene for passage through the (normally) 8L of catalyst bed. The catalyst is held in a copper lined tubular reactor, which is also fitted with a central multipoint thermocouple for accurately measuring catalyst temperatures at various (fixed) depths down the catalyst bed. The gaseous reactor effluent is cooled to ambient using a simple shell & tube type heat exchanger, & the mixture of liquid and gaseous products are separated in a high pressure gas/liquid separator. The gaseous product, which still contains significant levels of ethanol, is then further processed by passing it through a wash tower, where the majority of the water soluble components are scrubbed out. The liquid effluent from the wash tower is then mixed with the liquid effluent from the gas/liquid separator to form the main product stream. This stream is collected and analysed (by gas chromatography) on a regular basis to provide catalyst activity and selectivity data. The scrubbed gas from the wash tower is fed to a recycle compressor for return to the reactor. The recycle gas flow rate is carefully controlled using a Coriolis meter to provide a similar contact time through the catalyst bed as is encountered in the commercial ethanol plants. An on-line gas chromatograph also analyses this recycle stream every 15 minutes in order to determine the recycle gas composition. Tests were carried out at 2068.5 - 3792 kPa (300 - 550 psig) total reaction pressure; ethylene partial pressures from 1655 - 3034 kPa (240 - 440 psig); reactor inlet temps from 190 - 205°C; [water]: [ethylene] feed mole ratio from 0.25 - 0.30; typical ethylene GHSN=1275 hr ^); typical steam GHSV=370 hr ^_1).

The catalyst was kept on stream for 2 months, during which time 178 test periods were carried out. The results are shown graphically in Figures 3 and 4.

It can be seen from Figure 3 that the ethanol STYs obtained using the catalyst of Example 1 is comparable to those obtained using a commercially available H₃PO₄ catalyst operating at 240°C and an ethylene partial pressure of 5378 kPa (780 psi). Thus, by using the catalyst of Example 1 acceptable ethanol STYs may be achieved at substantially lower reaction temperatures & ethylene partial pressures. Example 3: Effect of Acid loading on Catalysts The performance of the catalyst can be enhanced or retarded to some degree by varying the amounts of heteropoly acid supported on the silica carrier.

Two further catalysts have been prepared using the method given in Example 1, but with different concentrations of HP A solution (33.9%wt/vol, & 41.5%wt/vol).

Using these catalysts in pilot plant experiments identical to those described in Example 2, the commercially viable ethanol production rate (109 STY) was achieved at 205oC & 201oC respectively.

This shows that the catalyst has a reasonably broad range of working loadings, but that higher loadings (than 151.5 g/L) tend to boost activity.

Claims

Claims:

1. A process for the hydration of an olefin, said process comprising: introducing a feed comprising the olefin and water into a reactor containing a heteropolyacid catalyst, and reacting the olefin with the water in the presence of said catalyst in the vapour phase, characterised in that the amount of olefin in the feed to the reactor is from 10 to 60 mol %.

2. A process as claimed in clam 1, wherein the olefin is propylene or ethylene.

3. A process as claimed in any preceding claim, wherein the olefin in the feed comprises fresh olefin and recycled olefin.

4. A process as claimed in claim 3, wherein the fresh olefin comprises refinery grade olefin, which is introduced into the reactor without purification.

5. A process as claimed in claim 3 or 4, wherein some of the recycled olefin returned to the reactor without purification, whilst the remainder of the recycled olefin is purified before it is returned to he reactor.

6. A process as claimed in claim 5, wherein the amount of olefin in the feed to the reactor is controlled by controlling the concentration of fresh olefin to the process, and/or controlling the degree to which the unreacted olefin is purified.

7. A process as claimed in claim 5, wherein the amount of olefin in the feed to the reactor is controlled by controlling the overall pressure at which the process is carried out.

8. A process as claimed in claim 7, wherein the olefin is ethylene and the ethylene hydration reaction is carried out at 4500 kPa or below.

9. A process as claimed in claim 8, wherein the olefin is propylene and the propylene hydration reaction is carried out at 3000kPa or below.

10. A process as claimed in any preceding claim, wherein the feed to the reactor is from 20 to 50 mol % olefin.

11. A process as claimed in any preceding claim wherein the olefin hydration catalyst is a heteropolyacid catalyst supported on a siliceous support.