WO2014102394A1

WO2014102394A1 - Processes for the preparation of olefins

Info

Publication number: WO2014102394A1
Application number: PCT/EP2013/078172
Authority: WO
Inventors: Ye Mon Chen; Leslie Andrew Chewter; Sivakumar SADASIVAN VIJAYAKUMARI; Jeroen Van Westrenen
Original assignee: Shell Internationale Research Maatschappij B.V.; Shell Oil Company
Priority date: 2012-12-31
Filing date: 2013-12-31
Publication date: 2014-07-03

Abstract

The invention provides a process for the preparation of olefins, which process comprises the steps of: (a) reacting an oxygenate and/or olefinic feed in a reactor in the presence of a molecular sieve catalyst to form a reaction effluent which comprises olefins, water, unreacted oxygenate and catalyst fines; (b) cooling the reaction effluent in a spray tower to a temperature at or lower than the dew point temperature of the reaction effluent by direct injection of an aqueous liquid into the reaction effluent at two or more different levels in the spray tower; (c) recovering from an upper part of the spray tower a first hydrocarbon fraction comprising olefins and from a lower part of the spray tower a first water fraction which comprises hydrocarbons, catalyst particles and dissolved oxygenate; and (d) processing the first hydrocarbon fraction as recovered in step (c) to form a first olefinic product fraction comprising ethylene and/or propylene and a second olefinic product fraction which comprise olefins containing 4 or more carbon atoms.

Description

PROCESSES FOR THE PREPARATION OF OLEFINS

Field of the Invention

The present invention relates to a process for the preparation of olefins such as ethylene and/or propylene from an oxygenate and/or olefinic feed.

Background of the Invention

Processes for the preparation of olefins from

oxygenates (OTO processes) are known in the art. Of particular interest is often the production of light olefins, in particular ethylene and/or propylene. The oxygenate feedstock can for example comprise methanol and/or dimethylether, and an interesting route includes their production from synthesis gas derived from e.g. natural gas or via coal gasification. For example, WO2007 /135052 discloses a process wherein an alcohol and/or ether containing oxygenate feedstock and an olefinic co- feed are reacted in the presence of a zeolite having one- dimensional 10-membered ring channels to prepare an olefinic reaction mixture, and wherein part of the obtained olefinic reaction mixture is recycled as

olefinic co-feed. With a methanol and/or dimethylether containing feedstock, and an olefinic co-feed comprising C4 and/or C5 olefins, an olefinic product rich in light olefins can be obtained.

The reaction effluent obtained from the OTO reactor, comprising the olefins, any unreacted oxygenates and other reaction products such as water, is separated from the bulk of the catalyst, usually by one or more cyclonic separation devices. Since carbonaceous deposits are formed on the catalyst during the OTO process the bulk of the catalyst is continuously regenerated to remove a portion of the carbonaceous deposits. After separation of the bulk of the catalyst, some solids, such as catalyst fines will however still remain in the reaction effluent stream.

After the reaction in the OTO reactor, the reaction effluent must be cooled before it can be treated to provide separate product streams. Normally, the reaction effluent stream is firstly cooled by means of one or more heat exchangers before it is contacted with a cooled aqueous stream in a quench tower. A quench tower

comprises usually at least one set of internals such as packing and/or trays to enable the use of a more compact column. The gaseous stream to be quenched is fed into the quench tower below the internals and one or more aqueous liquid quenching streams are fed into the quench tower above the internals. The gaseous stream to be quenched travels upwards through the quench tower and is brought counter-currently into contact with the quenching

streams. The quenched gaseous stream is removed from the top of the quench tower, and a liquid quenching stream is removed at the bottom of the quench tower, and can be cooled and recycled to the quench tower.

US 7329790 discloses a process for wet scrubbing and recycle of effluent contaminating catalyst particles in an OTO process. In said process two quench tower-type apparatuses are used to contact the reaction effluent stream with liquid in a counter-current fashion in the presence of trays and/or packing. Water removed from the bottom of the quench towers is recycled for re-use in the quench towers .

Such continuous recycling of the aqueous streams, and the catalyst fines contained therein, has however the disadvantage that it will result in solids building up on the internals of the quench towers, causing blockages.

In order to overcome this disadvantage, processes in the prior art have used further cyclonic separation devices in order to increase separation of solids from the reactor effluent. Alternatively, in order to

compensate for the loss of tray efficiency the length of the quench tower needs to be increased when more trays are used at similar tray spacing. A two-stage quench tower system, in which solids are removed in a first stage is for instance described in WO 03/104170. Such a design would, however, significantly increase the capital expenditure required for an OTO process.

In view of the above, it would be desirable to provide a process in which the separation of solid materials, especially catalyst fines, is established in a simple and more effective manner and blockages in the OTO process quench system can be avoided.

Summary of the Invention

It has now been found that this can be established when an OTO reaction effluent is cooled in a particular spray tower.

Accordingly, the present invention provides a process for the preparation of olefins, which process comprises the steps of:

(a) reacting an oxygenate and/or olefinic feed in a reactor in the presence of a molecular sieve catalyst to form a reaction effluent which comprises olefins, water, unreacted oxygenate and catalyst fines;

(b) cooling the reaction effluent in a spray tower to a temperature at or lower than the dew point temperature of the reaction effluent by direct injection of an aqueous liquid into the reaction effluent at two or more

different levels in the spray tower;

(c) recovering from an upper part of the spray tower a first hydrocarbon fraction comprising olefins and from a lower part of the spray tower a first water fraction which comprises hydrocarbons, catalyst fines and

dissolved oxygenate; and

(d) processing the first hydrocarbon fraction as

recovered in step (c) to form a first olefinic product fraction comprising ethylene and/or propylene and a second olefinic product fraction which comprise olefins containing 4 or more carbon atoms .

Detailed Description of the Invention

In step (a), an oxygenate and/or olefinic feed is reacted in a reactor in the presence of a molecular sieve catalyst to form a mixture which comprises olefins and at least partially coked catalyst. The reactor in step (a) can be an OTO reaction zone wherein the oxygenate feed is contacted with an oxygenate conversion catalyst under oxygenate conversion conditions, to obtain a conversion effluent comprising lower olefins. Reference herein to an oxygenate feed is to an oxygenate-comprising feed. In the OTO reaction zone, at least part of the feed is converted into a product containing one or more olefins, preferably including lower olefins, in particular ethylene and typically propylene.

The oxygenate used in the process according to the invention is preferably an oxygenate which comprises at least one oxygen-bonded alkyl group. The alkyl group preferably is a C1-C5 alkyl group, more preferably C1-C4 alkyl group, i.e. comprises 1 to 5, respectively, 4 carbon atoms; more preferably the alkyl group comprises 1 or 2 carbon atoms and most preferably one carbon atom. Examples of oxygenates that can be used in the oxygenate feed include alcohols and ethers. Examples of preferred oxygenates include alcohols, such as methanol, ethanol, propanol; and dialkyl ethers, such as dimethylether, diethylether, methylethylether . Preferably, the oxygenate is methanol or dimethylether, or a mixture thereof. More preferably, the oxygenate comprises methanol or

dimethylether .

Preferably the oxygenate feed comprises at least 50 wt% of oxygenate, in particular methanol and/or

dimethylether, based on total hydrocarbons, more

preferably at least 70 wt%.

The oxygenate feed can comprise an amount of diluent, such as nitrogen and water, preferably in the form of steam. In one embodiment, the molar ratio of oxygenate to diluent is between 10:1 and 1:10, preferably between 4:1 and 1:2, in particular when the oxygenate is methanol and the diluent is water (steam) .

A variety of OTO processes is known for converting oxygenates such as for instance methanol or dimethylether to an olefin-containing product, as already referred to above. One such process is described in WO 2006/020083. Processes integrating the production of oxygenates from synthesis gas and their conversion to light olefins are described in US20070203380A1 and US20070155999A1.

Catalysts suitable for converting the oxygenate feed in accordance with the present invention include

molecular sieve-catalysts. The molecular sieve catalyst suitably comprises one or more zeolite catalysts and/or one or more SAPO catalysts. Molecular sieve catalysts typically also include binder materials, matrix material and optionally fillers. Suitable matrix materials include clays, such as kaolin. Suitable binder materials include silica, alumina, silica-alumina, titania and zirconia, wherein silica is preferred due to its low acidity.

Molecular sieve catalysts preferably have a

molecular framework of one, preferably two or more corner-sharing [Ti0₄] tetrahedral units, more preferably, two or more [Si0₄], [A10₄] and/or [P0₄] tetrahedral units. These silicon, aluminum and/or phosphorous based

molecular sieves and metal containing silicon, aluminum and/or phosphorous based molecular sieves have been described in detail in numerous publications including for example, U.S. Pat. No. 4,567,029. In a preferred embodiment, the molecular sieve catalysts have 8-, 10- or 12-ring structures and an average pore size in the range of from about 3 A to 15 A.

Suitable molecular sieve catalysts are

silicoaluminophosphates (SAPO) , such as SAPO-17, -18, 34, -35, -44, but also SAPO-5, -8, -11, -20, -31, -36, 37, - 40, -41, -42, -47 and -56; aluminophosphates (A1PO) and metal substituted (silico) aluminophosphates (MeAlPO) , wherein the Me in MeAlPO refers to a substituted metal atom, including metal selected from one of Group IA, IIA, IB, IIIB, IVB, VB, VIB, VIIB, VIIIB and Lanthanide ' s of the Periodic Table of Elements, preferably Me is selected from one of the group consisting of Co, Cr, Cu,Fe, Ga, Ge, Mg, Mn, Ni, Sn, Ti, Zn and Zr.

Preferably, the conversion of the oxygenate feed may be accomplished by the use of an aluminosilicate- comprising catalyst, in particular a zeolite-comprising catalyst. In a zeolite-comprising catalyst the amount of zeolite is suitably from 20 to 50 wt%, preferably from 35 to 45 wt%, based on total catalyst composition.

Suitable catalysts include those containing a zeolite of the ZSM group, in particular of the MFI type, such as ZSM-5, the MTT type, such as ZSM-23, the TON type, such as ZSM-22, the MEL type, such as ZSM-11, or the FER type. Other suitable zeolites are for example zeolites of the STF-type, such as SSZ-35, the SFF type, such as SSZ-44 and the EU-2 type, such as ZSM-48.

Aluminosilicates-comprising catalyst, and in particular zeolite-comprising catalyst are preferred when an olefinic co-feed is fed to the oxygenate conversion zone together with oxygenate, for increased production of ethylene and propylene.

Preferred catalysts comprise a more-dimensional zeolite, in particular of the MFI type, more in

particular ZSM-5, or of the MEL type, such as zeolite ZSM-11. Such zeolites are particularly suitable for converting olefins, including iso-olefins, to ethylene and/or propylene. The zeolite having more-dimensional channels has intersecting channels in at least two directions. So, for example, the channel structure is formed of substantially parallel channels in a first direction, and substantially parallel channels in a second direction, wherein channels in the first and second directions intersect. Intersections with a further channel type are also possible. Preferably, the channels in at least one of the directions are 10-membered ring channels. A preferred MFI-type zeolite has a Silica-to-

Alumina ratio (SAR) of at least 60, preferably at least 80.

Particular preferred catalysts include catalysts comprising one or more zeolites having one-dimensional 10-membered ring channels, i.e. one-dimensional 10- membered ring channels, which are not intersected by other channels. Preferably, the zeolite is a zeolite of the MFI-type, the MEL-type, the MTT-type, the TON-type or any mixture thereof.

Preferred examples are zeolites of the MTT and/or TON type.

Preferably, the catalyst comprises at least 40wt%, preferably at least 50%wt of such zeolites based on total zeolites in the catalyst.

In a particularly preferred embodiment the catalyst comprises in addition to one or more one-dimensional zeolites having 10-membered ring channels, such as of the

MTT and/or TON type, a more-dimensional zeolite, in particular of the MFI type, more in particular ZSM-5, or of the MEL type, such as zeolite ZSM-11.

The catalyst may comprise phosphorus as such or in a compound, i.e. phosphorus other than any phosphorus included in the framework of the molecular sieve. It is preferred that an MEL or MFI-type zeolite comprising catalyst additionally comprises phosphorus. The

phosphorus may be introduced by pre-treating the MEL or MFI-type zeolites prior to formulating the catalyst and/or by post-treating the formulated catalyst

comprising the MEL or MFI-type zeolites. Preferably, the catalyst comprising MEL or MFI-type zeolites comprises phosphorus as such or in a compound in an elemental amount of from 0.05 - 10 wt% based on the weight of the formulated catalyst. A particularly preferred catalyst comprises phosphorus-treated MEL or MFI-type zeolites having SAR of in the range of from 60 to 150, more preferably of from 80 to 100. An even more particularly preferred catalyst comprises phosphorus-treated ZSM-5 having SAR of in the range of from 60 to 150, more preferably of from 80 to 100. It is preferred that the molecular sieves in the hydrogen form are used in the oxygenate conversion catalyst, e.g., HZSM-22, HZSM-23, and HZSM-48, HZSM-5. Preferably at least 50% w/w, more preferably at least 90% w/w, still more preferably at least 95% w/w and most preferably 100% of the total amount of molecular sieve used is in the hydrogen form. It is well known in the art how to produce such molecular sieves in the hydrogen form.

The catalyst particles used in the process of the present invention can have any shape known to the skilled person to be suitable for this purpose, for it can be present in the form of spray dried catalyst particles, spheres, tablets, rings, extrudates, etc. Extruded catalysts can be applied in various shapes, such as, cylinders and trilobes. Spherical particles are normally obtained by spray drying. Preferably the average particle size is in the range of 1-500 ym, preferably 50-100 ym.

The reaction conditions of the oxygenate conversion in step (a) include a reaction temperature from 350 to

750 °C, preferably from 450 to 750 °C, more preferably from 450 to 700°C, even more preferably 500 to 650°C; and a pressure of from 1-15 bara, preferably from 1-4 bara, more preferably from 1.1-3 bara, and even more preferably from 1.3-2 bara.

Suitably, the oxygenate-comprising feed is preheated to a temperature in the range of from 120 to 550°C, preferably 250 to 500°C prior to contacting with the molecular sieve catalyst in step (a) .

Preferably, in addition to the oxygenate, an

olefinic co-feed is provided along with and/or as part of the oxygenate feed. Reference herein to an olefinic co- feed is to an olefin-comprising co-feed. The olefinic co- feed preferably comprises C4 and higher olefins, more preferably C4 and C5 olefins. Preferably, the olefinic co-feed comprises at least 25 wt%, more preferably at least 50 wt%, of C4 olefins, and at least a total of 70 wt% of C4 hydrocarbon species. The olefinic co-feed can also comprise propylene.

The reaction in step (a) may suitably be operated in a fluidized bed or moving bed, e.g. a dense, turbulent or fast fluidized bed or a riser reactor or a downward reactor system, and also in a fixed bed reactor or a tubular reactor. A fluidized bed or moving bed, e.g. a turbulent fluidized bed, fast fluidized bed or a riser reactor system are preferred. These could be arranged in a single or multiple reactors arranged in parallel or in series.

The superficial velocity of the gas components in a dense fluidized bed will generally be from 0 to 1 m/s; the superficial velocity of the gas components in a turbulent fluidized bed will generally be from 1 to 3 m/s; the superficial velocity of the gas components in a fast fluidized bed will generally be from 3 to 5 m/s; and the superficial velocity of the gas components in a riser reactor will generally be from 5 to about 25 m/s.

It will be understood that dense, turbulent and fast fluidized beds will include a dense lower reaction zone with densities generally above 300 kg/m³. Moreover, when working with a fluidized bed several possible

configurations can be used: (a) co-current flow meaning that the gas (going upward) and the catalyst travels through the bed in the same direction, and (b)

countercurrent , meaning that the catalyst is fed at the top of the bed and travels through the bed in opposite direction with respect to the gas, whereby the catalyst leaves the vessel at the bottom. In a conventional riser reactor system the catalyst and the vapors will travel co-currently .

More preferably, a fluidized bed, in particular a turbulent fluidized bed system is used. Suitably, in such a moving bed reactor the oxygenate feed is contacted with the molecular sieve catalyst at a weight hourly space velocity of at least 1 hr^-1, suitably from 1 to 1000 hr^-1, preferably from 1 to 500 hr^-1, more preferably 1 to 250 hr^-1, even more preferably from 1 to 100 hr^-1, and most preferably from 1 to 50 hr^-1.

The reactor in step (a) can also be an OCP (Olefin Cracking Process) reaction zone wherein the olefinic feed is contacted with an olefin conversion catalyst under olefin conversion conditions, to obtain a conversion effluent comprising lower olefins.

Suitably, the olefinic feed comprises C4+ olefins that will be converted to ethylene and/or propylene by contacting such a feed with a zeolite-comprising

catalyst. Preferably, the olefinic feed is contacted with the zeolite-comprising catalyst in step (a) at a reaction temperature of 350 to 1000 °C, preferably from 375 to 750 °C, more preferably 450 to 700°C, even more

preferably 500 to 650°C; and a pressure from 1 bara to 50 bara, preferably from 1-15 bara. Optionally, the olefinic feed also contains a diluent. Examples of suitable diluents include, but are not limited to, diluents such as water or steam, nitrogen, paraffins and methane. Under these conditions, at least part of the olefins in the olefinic feed are converted to further ethylene and/or propylene .

In an OCP suitable aluminosilicate catalysts are used. Aluminosilicate catalysts, and in particular zeolite catalysts, have the additional advantage that in addition to the conversion of methanol or ethanol, these catalysts also induce the conversion of olefins to ethylene and/or propylene. Therefore, aluminosilicate catalysts, and in particular zeolite catalysts, are particularly suitable for use as the catalyst in an OCP. Particular preferred catalyst for the OCP reaction are catalysts comprising at least one zeolite selected from MFI, MEL, TON and MTT type zeolites, more preferably at least one of ZSM-5, ZSM-11, ZSM-22 and ZSM-23 zeolites.

Also an OCP may suitably be operated in a fluidized bed or moving bed, e.g. a fast fluidized bed or a riser reactor or a downward reactor system, and also in a fixed bed reactor or a tubular reactor. A fluidized bed or moving bed, e.g. a fast fluidized bed or a riser reactor system are preferred.

The olefins and at least partially coked catalyst as obtained in step (a) will be separated. The separation can be carried out by one or more cyclone separators. Such one or more cyclone separators may be located inside, partly inside and partly outside, or outside the reactor used in step (a) . Such cyclone separators are well known in the art. Cyclone separators are preferred, but also methods for separating the catalyst from the olefins can be used that apply plates, caps, elbows, and the like.

The preferred molar ratio of oxygenate in the oxygenate feed to olefin in the olefinic co-feed provided to the OTO reactor in step (a) depends on the specific oxygenate used and the number of reactive oxygen-bonded alkyl groups therein. Preferably the molar ratio of oxygenate to olefin in the total feed lies in the range of 20:1 to 1:10, more preferably in the range of 18:1 to 1:5, still more preferably in the range of 15:1 to 1:3, even still more preferably in the range of 12:1 to 1:3.

In step (b) , the reactant effluent as obtained in step (a) is cooled to a temperature at or lower than the dew point temperature of the reaction effluent by direct injection of an aqueous liquid at two or more different levels in a spray tower. The direct injection of the aqueous liquid rapidly cools the reaction effluent, enabling solids to be separated along with the liquid. The solids-containing liquid phase, which also contains catalyst fines, can then be separated and removed from the process.

The aqueous liquid is preferably water. It may suitably be fresh water, but preferably is a recycled stream from a later stage of the process for the

preparation of an olefinic product.

Suitably, the aqueous liquid is at a temperature in the range of from 10-65°C, preferably in the range of from 20-50°C, more preferably in the range of from 25- 40°C.

In the context of the present application the term irect injection' indicates that the aqueous liquid is provided straight into the reaction effluent stream by a spray fitting suitable for dispersing the water into fine droplets, effectively forcing the liquid as fine droplets into the gas stream. The injection of the aqueous stream is such that the aqueous stream travels through the spray tower counter-current to the direction of flow of the reaction effluent stream. The injection itself may occur at a tangential angle to the flow of the reaction

effluent stream. The aqueous liquid is injected into the reaction effluent stream by means of a spray fitting. The spray fitting may be any means known in the art capable of rapidly introducing a large quantity of water directly into a gaseous stream, but suitably comprises a spray device such as a spray nozzle. The aqueous stream is injected into the reaction effluent stream such that the mass ratio of the aqueous stream to be injected to the total contents of the reaction effluent stream is

preferably at least 0.4:1, more preferably at least 1:1, most preferably at least 3:1. Preferably, the mass ratio of the aqueous stream to be injected to the total

contents of the reaction effluent stream is at most 20:1, more preferably at most 15:1, even more preferably at most 10:1, most preferably at most 8:1.

In accordance with the present invention the direct injection with the aqueous liquid takes place at two or more different levels in the spray tower. Preferably, in step (b) the reaction effluent is cooled by direct injection of the aqueous liquid at 2-5 different levels in the spray tower. The distance between the different and adjacent levels at which the reaction effluent is cooled in the spray tower may vary but is preferably in the range of from 40 cm to 1 m.

Preferably, the aqueous liquid is directly injected into the reaction effluent by means of a number of spraying devices that are arranged symmetrically with respect to each other. The spray devices, preferably spray nozzles, can for instance be circumferentially arranged. Suitably, at the levels at which the direct injection occurs the spray devices are in the form of an array of spray devices fixed on the inner wall of the spray tower, whereby the spray devices are suitably distributed in an even pattern over the array, and the array covers the cross-section of the spray tower at the level concerned. The distances between spray devices at the respective levels may be the same or may be

different. Suitably, the distance between the spray devices at the different levels is in the range of from 10 to 30 cm.

The arrays of spray devices to be used may include rectangular or triangular arrays of spray devices.

Suitably, in step (b) the effluent as obtained in step (a) is introduced into a lower part of the spray tower, the effluent is contacted in the spray tower with droplets of the aqueous liquid that are generated at the two or more different levels in the spray tower, and the first hydrocarbon fraction is recovered in step (c) from an upper part of the spray tower.

The upper part of the spray tower is suitably equipped with a mist eliminator to separate the first hydrocarbon fraction more effectively from catalyst fines. The mist eliminator may be washed out with a liquid stream such as a water stream to remove catalyst fines that are entrapped in the mist eliminator device.

Apart from the spray devices the spray tower to be used in accordance with the present invention does not contain any further process equipment such as trays and/or packing.

In step (c) , a first hydrocarbon fraction comprising olefins is recovered from an upper part of the spray tower and a first water fraction which comprises

hydrocarbons, catalyst fines and dissolved oxygenate is recovered from a lower part of the spray tower.

Preferably, the first water fraction as recovered in step (c) has a temperature in the range of from 30-100 °C, more preferably in the range of from 40-95°C, most preferably in the range of from 50-90°C.

In step (d) , the first hydrocarbon fraction as obtained in step (c) is processed to form a first

olefinic product fraction comprising ethylene and/or propylene and a second olefinic product fraction which comprises olefins containing 4 or more carbon atoms.

The processing in step (d) suitably comprises a fractionating process which is carried out in a

fractionation unit.

Preferably, in step (d) at least part of the first hydrocarbon fraction as recovered in step (c) is

subjected in a quenching tower to a quench treatment with an aqueous liquid to form a second hydrocarbon fraction comprising olefins and a second water fraction comprising hydrocarbons, and at least part of the second hydrocarbon fraction is fractionated in a fractionating unit to obtain the first olefinic product fraction comprising ethylene and/or propylene and the second olefinic product fraction which comprises olefins containing 4 or more carbon atoms .

Preferably, at least part of the second olefinic product fraction so obtained is recycled to the reactor in step (a) .

The first hydrocarbon fraction as recovered in step (c) is suitably fed into the quench tower at a point below the internals. The term internals as used herein preferably refers to packing and/or trays.

Gaseous material in the first hydrocarbon fraction will pass upwards through the quench tower and contact an aqueous liquid which is fed into the quench tower at one or more points above at least one set of internals.

Preferably, the aqueous liquid is fed into the quench tower at one or more points above each set of internals. After contacting the aqueous liquid, the second hydrocarbon fraction is removed from the top of the quench tower.

The second water fraction which contains all other materials present in the quench tower, including

oxygenates, liquid hydrocarbons and aqueous materials will pass to the bottom of the quench tower. These materials may be recycled with cooling for use as the aqueous liquid in the quench tower. Preferably, at least a portion of these materials are separated to provide a liquid hydrocarbon stream and an oxygenates containing liquid stream as well as a stream to be recycled as the aqueous liquid to be used in the quench tower.

Due to the lack of solids present after step (b) , blockages in the quench tower due to the deposition and build up of solids are avoided and the lifetime of the quench tower can, thus, be greatly extended. An added advantage of the present invention is that, as any solids present have been removed in step (b) , the second

hydrocarbon fraction can be fed into the quench tower through an inlet device allowing an even vapour

distribution across the vessel cross section. For example a vane type feed inlet device, e.g. a Schoepentoeter™, may be used without risk of solid build up.

The oxygenates containing second water fraction can be further separated to provide an oxygenate recovery stream for recycle as an oxygenate co-feed in the reactor in step (a) and a further aqueous recycle stream that can be cooled and used as the aqueous liquid in step (b) .

Preferably, at least part of the first water

fraction as recovered in step (c) is passed to a solids recovery unit in which catalyst particles are separated from the first water fraction. The first water fraction comprises at least a portion of the aqueous liquid as well as any condensed materials and solids present. A portion of the aqueous liquid may have evaporated in the process and form part of the first hydrocarbon fraction. The solids recovery unit may be a separation vessel, such as a decanter or a knock out drum. Suitably, the first water fraction contains the majority of the solids present in the reaction effluent. Preferably, at least 80wt% of the solids present in the reaction effluent are removed in the first water fraction. More preferably at least 90wt%, even more preferably at least 95 wt~6 , even more

preferably at least 98 wt%, even more preferably at least 99wt%, most preferably substantially all solids present in the reaction effluent are removed in the first water fraction .

Preferably, at least part of the second water fraction as obtained in the quench tower is passed to a settler unit to separate the second water fraction into a hydrocarbon-enriched fraction and a third water fraction which is depleted in hydrocarbons. Suitably, at least part of the third water fraction which is depleted in hydrocarbons is passed to the reactor in step (a) and/or at least part of the third water fraction which is depleted in hydrocarbons is passed in the form of steam to the reactor in step (a) .

Suitably, the reaction effluent as obtained in step (a) is cooled by means of an indirect heat exchange to a temperature greater than the dew point temperature of the reaction effluent before the reaction effluent is

subjected to step (b) .

In such an initial cooling step, the reaction effluent as obtained in step (a) is cooled by means of an indirect heat exchange to recover the heat contained in the reactor effluent. Typically, the reaction effluent stream is indirectly contacted with a liquid stream, which is at a lower temperature, in a heat exchanger. The liquid stream is suitably a process stream, for instance a reactor feed, or a water stream. The heat exchanger may be of any type known in the art, for instance a transfer line exchanger (TLE) and/or a feed/effluent exchanger. Preferably, a shell and tube type heat exchanger is used. In the initial cooling step the reaction effluent is suitably cooled to a temperature greater than the dew point of the reaction effluent. Preferably, the reaction effluent is cooled to a temperature in the range of from 110-370 °C, preferably in the range of from 140-270°C, more preferably in the range of from 150-250°C.

An embodiment of the present invention will now be described by way of example only and with reference to the accompanying non-limiting figure.

Figure 1 exemplifies an embodiment of the present invention. An oxygenate feedstock 101 is fed into the reactor 105. An oxygenate co-feed 102 may also be

supplied by an oxygenate recovery stream. A diluent 103 may also be provided to the reaction zone. Preferably, an olefinic co-feed 104 is also provided to the reactor. The oxygenate co-feed, diluent and olefinic co-feed may be supplied to the reactor separately or one or more of these streams may be combined with the oxygenate

feedstock or together before being fed to the reactor. In the OTO reactor 105, reaction is carried out in the presence of a catalyst at a temperature in the range of from 350 to 1000°C. Following reaction, the gaseous product is separated from the bulk of the catalyst, e.g. by cyclonic separation devices, to produce a reaction effluent 106. The reaction effluent is cooled in one or more heat exchangers 107 situated in series, to provide a cooled reaction effluent 108 at a temperature greater than the dew point temperature of the reaction effluent stream and preferably in the range of from 150 to 250 °C.

An aqueous liquid 110 is directly injected into the cooled reaction effluent 108 at three different levels in a spray tower 109 (e.g. by means of spray nozzles), rapidly cooling the reaction effluent 108 by mixing it with a large quantity of water to a temperature at or below the dew point temperature of the reaction effluent stream.

From the spray tower 109 a first hydrocarbon fraction 111 is recovered from an upper part of the spray tower 109 and a first water fraction 112 is recovered from a lower part of the spray tower 109. The first water fraction 112 is passed to a solids recovery unit 113, and the first hydrocarbon fraction 111 is then fed into a quench tower 114 containing internal and/or packing 115 and 116.

At least one aqueous liquid stream 117, 118 is fed into the quench tower above the internals and/or packing.

An aqueous stream 119 is removed from the bottom of the quench tower 114 and cooled in air or water coolers 120 and 121 to provide the aqueous liquid streams 117, 118. In the embodiment of the invention exemplified in Figure

1, the second water fraction 120 is removed from the bottom of the quench tower 114 and optionally passed to the separation vessel 123. Liquid hydrocarbon stream 124 may be separated as a waste stream or for use as fuel. An oxygenate containing liquid stream 125 is separated and passed to a separation unit 126 to produce an oxygenate recovered stream 127, which may be used as oxygenate co- feed 102 and an aqueous stream 128, which after cooling in an air or water cooler 129, can be re-used as the aqueous liquid 110.

The second hydrocarbon fraction 130 comprising the olefinic product is removed from the top of the quench tower.

Claims

C L A I M S

1. Process for the preparation of olefins, which process comprises the steps of:

different levels in the spray tower;

(c) recovering from an upper part of the spray tower a first hydrocarbon fraction comprising olefins and from a lower part of the spray tower a first water fraction which comprises hydrocarbons, catalyst particles and dissolved oxygenate; and

(d) processing the first hydrocarbon fraction as

2. Process according to claim 1, wherein in step (d) the processing comprises a fractionating process which is carried out in a fractionating unit.

3. Process according to claim 1, wherein in step (d) at least part of the first hydrocarbon fraction as recovered in step (c) is subjected in a quenching tower to a quench treatment with an aqueous liquid to form a second

hydrocarbon fraction comprising olefins and a second water fraction comprising hydrocarbons, and at least part of the second hydrocarbon fraction is fractionated in a fractionating unit into the first olefinic product fraction comprising ethylene and/or propylene and the second olefinic product fraction which comprises olefins containing 4 or more carbon atoms .

4. Process according to any one of claims 1-3, wherein at least part of the second olefinic product fraction is recycled to the reactor in step (a) .

5. Process according to any one of claims 1-4, wherein at least part of the first water fraction as recovered in step (c) is passed to a solids recovery unit in which catalyst particles are separated from the first water fraction .

6. Process according to any one of claims 1-5, wherein the reaction effluent is cooled by means of an indirect heat exchange to a temperature greater than the dew point temperature of the reaction effluent before the reaction effluent is subjected to step (b) .

7. Process according to any one of claims 3-6, wherein at least part of the second water fraction is passed to a settler unit to separate the second water fraction into a hydrocarbon-enriched fraction and a third water fraction which is depleted in hydrocarbons.

8. Process according to claim 7, wherein at least part of the third water fraction which is depleted in

hydrocarbons is passed to the reactor in step (a) and/or at least part of the third water fraction which is depleted in hydrocarbons is passed in the form of steam to the reactor in step (a) .

9. Process according to any one of claims 1-8, wherein in step (b) the reaction effluent is cooled by direct injection of the aqueous liquid at 2-5 different levels in the spray tower.

10. Process according to any one of claims 1-9, wherein in step (b) the distance between the different and adjacent levels at which the reaction effluent is cooled in the spray tower is in the range of from 40 cm to 1 m.

11. Process according to any one of claims 1-10, wherein in step (b) at the different levels in the spray tower the aqueous liquid is directly injected into the reaction effluent by means of a number of spraying devices that are arranged symmetrically with respect to each other.

12. Process according to any one of claims 1-11, wherein in step (b) the effluent as obtained in step (a) is introduced into a lower part of the spray tower, the effluent is contacted in the spray tower with droplets of the aqueous liquid that are generated at the two or more different levels in the spray tower, and the first hydrocarbon fraction is recovered in step (c) from an upper part of the spray tower.

13. Process according to any one of claims 1-12, wherein the molecular sieve catalyst in step (a) comprises a zeolite of the MFI type or the MEL type.

14. Process according to any one of claims 1-13, wherein the oxygenate feed in step (a) comprises methanol and/or dimethylether .

15. Process according to any one of claims 1-14, wherein the reaction in step (a) is conducted at a temperature from 350 to 750 °C, preferably from 450 to 750 °C, more preferably from 450 to 700°C, even more preferably 500 to 650°C; and a pressure of from 1 to 15 bara, preferably from 1 to 4 bara, more preferably from 1.1 to 3 bara, and even more preferably from 1.3 to 2 bara.