WO2016109369A1 - An oxygenate to olefins conversion reactor system and process for converting oxygenates to olefins - Google Patents

Abstract

An oxygenate to olefins conversion reactor system comprising: a) a riser reactor having an inlet for feed, an inlet for catalyst and an outlet for products; and b) a cyclone for the separation of products from catalyst comprising: i) an upper zone where the products and catalyst are separated; and ii) a lower zone comprising a bed of accumulated solids wherein a stripping medium is fed into the lower zone to fluidize the catalyst and strip residual products from the catalyst.

Description

AN OXYGENATE TO OLEFINS CONVERSION REACTOR SYSTEM AND PROCESS FOR CONVERTING OXYGENATES TO OLEFINS

Reference to Related Applications

The present application claims the benefit of European Patent Application Serial No. 14200582.6, filed on 30 December 2014.

Field of the Invention

The invention relates to a reactor system for carrying out an oxygenate to olefins conversion reaction. The invention further relates to a process for converting oxygenates to olefins.

Background of the Invention

Oxygenate-to-olefin processes are well described in the art. Typically, oxygenate- to-olefin processes are used to produce predominantly ethylene and propylene. An example of such an oxygenate-to-olefin process is described in US Patent Application Publication No. 2011/112344, which is herein incorporated by reference. The publication describes a process for the preparation of an olefin product comprising ethylene and/or propylene, comprising a step of converting an oxygenate feedstock in an oxygenate-to-olefins conversion system, comprising a reaction zone in which an oxygenate feedstock is contacted with an oxygenate conversion catalyst under oxygenate conversion conditions, to obtain a conversion effluent comprising ethylene and/or propylene.

One preferred embodiment of an oxygenate to olefins conversion reactor system involves the use of a riser reactor. U.S. Patent 7767870 describes a gas-solids reaction system with termination devices to connect a riser with one or more separation devices.

Additional compounds, especially higher molecular weight hydrocarbons are typically produced with the ethylene and propylene in an oxygenate-to-olefins process. A method of improving the yield of lower molecular weight olefins is desired as these olefins, mainly ethylene and propylene, serve as feeds for the production of numerous chemicals. The reactor system design is an important factor in improving the yield of the desired products.

Summary of the Invention

The invention provides an oxygenate to olefins conversion reactor system comprising: a) a riser reactor having an inlet for feed, an inlet for catalyst and an outlet for products; and b) a cyclone for the separation of products from catalyst comprising: i) an upper zone where the products and catalyst are separated; and ii) a lower zone comprising a bed of accumulated solids wherein a stripping medium is fed into the lower zone to fluidize the catalyst and strip residual products from the catalyst.

The invention further provides a process for converting oxygenates to olefins comprising: a) feeding an oxygenate containing feed stream into a riser reactor; b) contacting the feed stream with a fluidized oxygenate conversion catalyst under oxygenate conversion conditions in the riser reactor to produce an effluent stream comprising olefins and solid catalyst particles; c) passing the effluent stream from the riser reactor into a cyclone; d) separating a product stream comprising olefins from the solid catalyst particles in an upper zone of the cyclone; e) fluidizing and stripping the solid catalyst particles in a lower zone of the cyclone by injecting a stripping medium into the lower zone; f) removing a products stream from the upper zone of the cyclone; and g) removing catalyst from the lower zone of the cyclone.

The invention also provides a cyclone for the separation of gas and solids comprising: a) an upper zone where gas and solids are separated; and b) a lower zone comprising a bed of accumulated solids wherein a stripping medium is fed into the lower zone to fluidize the accumulated solids and strip residual gas from the solids.

Brief Description of the Drawings

Figure 1 depicts an embodiment of a riser reactor with a cyclone according to the invention.

Detailed Description of the Invention

Prior art processes typically use a large gas/solid disengaging zone at the exit of the reactor or use close coupled cyclones which have openings for discharge of effluent into a disengaging zone. The cyclone with a lower zone stripping section quickly and efficiently separates the catalyst from the products. This suppresses undesirable reactions by reducing the amount of time that the products are in contact with the catalyst. The desirable products are formed as intermediates in the process so it is critical that the reaction is stopped as quickly as possible downstream of the reactor. The stripping bed in the cyclone ensures that the separation is made as soon as possible. This is an improvement over other processes that use a stripper, but only strip the catalyst after it has passed through the cyclone dipleg. The use of this stripping cyclone also allows for the use of direct coupled cyclones. The oxygenate to olefins process receives as a feedstock a stream comprising one or more oxygenates. An oxygenate is an organic compound that contains at least one oxygen atom. The oxygenate is preferably one or more alcohols, preferably aliphatic alcohols where the aliphatic moiety has from 1 to 20 carbon atoms, preferably from 1 to 10 carbon atoms, more preferably from 1 to 5 carbon atoms and most preferably from 1 to 4 carbon atoms. The alcohols that can be used as a feed to this process include lower straight and branched chain aliphatic alcohols. In addition, ethers and other oxygen containing organic molecules can be used. Suitable examples of oxygenates include methanol, ethanol, n-propanol, isopropanol, methyl ethyl ether, dimethyl ether, diethyl ether, di- isopropyl ether, formaldehyde, dimethyl carbonate, dimethyl ketone, acetic acid and mixtures thereof. In a preferred embodiment, the feedstock comprises one or more of methanol, ethanol, dimethyl ether, diethyl ether or a combination thereof, more preferably methanol or dimethyl ether and most preferably methanol.

In one embodiment, the oxygenate is obtained as a reaction product of synthesis gas. Synthesis gas can, for example, be generated from fossil fuels, such as from natural gas or oil, or from the gasification of coal. In another embodiment, the oxygenate is obtained from biomaterials, such as through fermentation.

The oxygenate feedstock can be obtained from a pre-reactor, which converts methanol at least partially into dimethylether and water. Water may be removed, by e.g., distillation. In this way, less water is present in the process of converting oxygenates to olefins, which has advantages for the process design and lowers the severity of

hydrothermal conditions to which the catalyst is exposed.

The oxygenate to olefins process, may in certain embodiments, also receive an olefin co-feed. This co-feed may comprise olefins having carbon numbers of from 1 to 8, preferably from 3 to 6 and more preferably 4 or 5. Examples of suitable olefin co-feeds include butene, pentene and hexene.

Preferably, the oxygenate feed comprises one or more oxygenates and olefins, more preferably oxygenates and olefins in an oxygenate:olefin molar ratio in the range of from 1000:1 to 1:1, preferably 100:1 to 1:1. More preferably, in a oxygenate: olefin molar ratio in the range of from 20:1 to 1:1, more preferably in the range of 18:1 to 1:1, still more preferably in the range of 15:1 to 1:1, even still more preferably in the range of 14: 1 to 1:1. It is preferred to convert a C4 olefin, recycled from the oxygenate to olefins conversion reaction together with an oxygenate, to obtain a high yield of ethylene and propylene, therefore preferably at least one mole of oxygenate is provided for every mole of C4 olefin.

The olefin co-feed may also comprise paraffins. These paraffins may serve as diluents or in some cases they may participate in one or more of the reactions taking place in the presence of the catalyst. The paraffins may include alkanes having carbon numbers from 1 to 10, preferably from 3 to 6 and more preferably 4 or 5. The paraffins may be recycled from separation steps occurring downstream of the oxygenate to olefins conversion step.

The oxygenate to olefins process, may in certain embodiments, also receive a diluent co-feed to reduce the concentration of the oxygenates in the feed and suppress side reactions that lead primarily to high molecular weight products. The diluent should generally be non-reactive to the oxygenate feedstock or to the catalyst. Possible diluents include helium, argon, nitrogen, carbon monoxide, carbon dioxide, methane, water and mixtures thereof. The more preferred diluents are water and nitrogen with the most preferred being water.

The diluent may be used in either liquid or vapor form. The diluent may be added to the feedstock before or at the time of entering the reactor or added separately to the reactor or added with the catalyst. In one embodiment, the diluents is added in an amount in the range of from 1 to 90 mole percent, more preferably from 1 to 80 mole percent, more preferably from 5 to 50 mole percent, most preferably from 5 to 40 mole percent.

During the conversion of the oxygenates in the oxygenate to olefins conversion reactor, steam is produced as a by-product, which serves as an in-situ produced diluent. Typically, additional steam is added as diluent. The amount of additional diluent that needs to be added depends on the in-situ water make, which in turn depends on the composition of the oxygenate feed. Where the diluent provided to the reactor is water or steam, the molar ratio of oxygenate to diluent is between 10:1 and 1:20.

The oxygenate feed is contacted with the catalyst at a temperature in the range of from 200 to 1000 °C, preferably of from 300 to 800 °C, more preferably of from 350 to 700 °C, even more preferably of from 450 to 650°C. The feed may be contacted with the catalyst at a temperature in the range of from 530 to 620 °C, or preferably of from 580 to 610 °C. The feed may be contacted with the catalyst at a pressure in the range of from 0.1 kPa (1 mbar) to 5 MPa (50 bar), preferably of from 100 kPa (1 bar) to 1.5 MPa (15 bar), more preferably of from 100 kPa (1 bar) to 300 kPa (3 bar). Reference herein to pressures is to absolute pressures.

A wide range of WHS V for the feedstock may be used. WHS V is defined as the mass of the feed (excluding diluents) per hour per mass of catalyst. The WHSV should preferably be in the range of from l hr^ to SOOO hr^"1.

The process takes place in a reactor and the catalyst may be present in the form of a fixed bed, a moving bed, a fluidized bed, a dense fluidized bed, a fast or turbulent fluidized bed, a circulating fluidized bed. In addition, riser reactors, hybrid reactors or other reactor types known to those skilled in the art may be used. In another embodiment, more than one of these reactor types may be used in series. In one embodiment, the reactor is a riser reactor. The advantage of a riser reactor is that it allows for very accurate control of the contact time of the feed with the catalyst, as riser reactors exhibit a flow of catalyst and reactants through the reactor that approaches plug flow.

The feedstocks described above are converted primarily into olefins. The olefins produced from the feedstock typically have from 2 to 30 carbon atoms, preferably from 2 to 8 carbon atoms, more preferably from 2 to 6 carbon atoms, most preferably ethylene and/or propylene. In addition to these olefins, diolefins having from 4 to 18 carbon atoms, conjugated or nonconjugated dienes, polyenes, vinyl monomers and cyclic olefins may be produced in the reaction.

In a preferred embodiment, the feedstock, preferably one or more oxygenates, is converted in the presence of a molecular sieve catalyst into olefins having from 2 to 6 carbon atoms. Preferably the oxygenate is methanol, and the olefins are ethylene and/or propylene.

The products from the reactor are typically separated and/or purified to prepare separate product streams in a recovery system. Such systems typically comprise one or more separation, fractionation or distillation towers, columns, and splitters and other associated equipment, for example, various condensers, heat exchangers, refrigeration systems or chill trains, compressors, knock-out drums or pots, pumps and the like.

The recovery system may include a demethanizer, a deethanizer, a depropanizer, a wash tower often referred to as a caustic wash tower and/or quench tower, absorbers, adsorbers, membranes, an ethylene-ethane splitter, a propylene-propane splitter, a butene- butane splitter and the like. Typically in the recovery system, additional products, by-products and/or contaminants may be formed along with the preferred olefin products. The preferred products, ethylene and propylene are preferably separated and purified for use in derivative processes such as polymerization processes.

In addition to the propylene and ethylene, the products may comprise C4+ olefins, paraffins and aromatics that may be further reacted, recycled or otherwise further treated to increase the yield of the desired products and/or other valuable products. C4+ olefins may be recycled to the oxygenate to olefins conversion reaction or fed to a separate reactor for cracking. The paraffins may also be cracked in a separate reactor, and/or removed from the system to be used elsewhere or possibly as fuel.

Although less desired, the product will typically comprise some aromatic compounds such as benzene, toluene and xylenes. Although it is not the primary aim of the process, xylenes can be seen as a valuable product. Xylenes may be formed in the OTO process by the alkylation of benzene and, in particular, toluene with oxygenates such as methanol. Therefore, in a preferred embodiment, a separate fraction comprising aromatics, in particular benzene, toluene and xylenes is separated from the gaseous product and at least in part recycled to the oxygenate to olefins conversion reactor as part of the oxygenate feed. Preferably, part or all of the xylenes in the fraction comprising aromatics are withdrawn from the process as a product prior to recycling the fraction comprising aromatics to the oxygenate to olefins conversion reactor.

The C4+ olefins and paraffins formed in the oxygenate to olefins conversion reactor may be further reacted in an additional reactor containing the same or a different molecular sieve catalyst. In this additional reactor, the C4+ feed is converted over the molecular sieve catalyst at a temperature in the range of from 500 to 700 °C. The additional reactor is also referred to as an OCP reactor and the process that takes place in this reactor is referred to as an olefin cracking process. In contact with the molecular sieve catalyst, at least part of the olefins in the C4+ feed is converted to a product, which includes at least ethylene and/or propylene and preferably both. In addition to ethylene and/or propylene, the gaseous product may comprise higher olefins, i.e. C4+ olefins, and paraffins. The gaseous product is retrieved from the second reactor as part of a second reactor effluent stream.

The olefin feed is contacted with the catalyst at a temperature in the range of from 500 to 700 °C, preferably of from 550 to 650°C, more preferably of from 550 to 620°C, even more preferably of from 580 to 610°C; and a pressure in the range of from 0.1 kPa (1 mbara) to 5 MPa (50 bara), preferably of from 100 kPa (1 bara) to 1.5 MPa (15 bara), more preferably of from 100 kPa ( 1 bara) to 300 kPa (3 bara). Reference herein to pressures is to absolute pressures.

In one embodiment, the C4+ olefins are separated into at least two fractions: a C4 olefin fraction and a C5+ olefin fraction. In this embodiment, the C4 olefins are recycled to the oxygenate to olefins conversion reactor and the C5+ olefins are fed to the OCP reactor. The cracking behavior of C4 olefins and C5 olefins is believed to be different when contacted with a molecular sieve catalyst, in particular above 500 °C.

The cracking of C4 olefins is an indirect process which involves a primary oligomerisation process to a C8, C12 or higher olefin followed by cracking of the oligomers to lower molecular weight hydrocarbons including ethylene and propylene, but also, amongst other things, to C5 to C7 olefins, and by-products such as C2 to C6 paraffins, cyclic hydrocarbons and aromatics. In addition, the cracking of C4 olefins is prone to coke formation, which places a restriction on the obtainable conversion of the C4 olefins. Generally, paraffins, cyclics and aromatics are not formed by cracking. They are formed by hydrogen transfer reactions and cyclisation reactions. This is more likely in larger molecules. Hence the C4 olefin cracking process, which as mentioned above includes intermediate oligomerisation, is more prone to by-product formation than direct cracking of C5 olefins. The conversion of the C4 olefins is typically a function of the temperature and space time (often expressed as the weight hourly space velocity). With increasing temperature and decreasing weight hourly space velocity (WHSV) conversion of the C4 olefins in the feed to the OCP increases. Initially, the ethylene and propylene yields increase, but, at higher conversions, yield decreases at the cost of a higher byproduct make and, in particular, a higher coke make, limiting significantly the maximum yield obtainable.

Contrary to C4 olefins, C5 olefin cracking is ideally a relatively straight forward- process whereby the C5 olefin cracks into a C2 and a C3 olefin, in particular above 500°C. This cracking reaction can be run at high conversions, up to 100%, while maintaining, at least compared to C4 olefins, high ethylene and propylene yields with a significantly lower by-product and coke make. Although, C5+ olefins can also oligomerise, this process competes with the more beneficial cracking to ethylene and propylene. In a preferred embodiment of the process according to the present invention, instead of cracking the C4 olefins in the OCP reactor, the C4 olefins are recycled to the oxygenate toi olefins conversion reactor. Again without wishing to be bound by any particular theory, it is believed that in the oxygenate to olefins conversion reactor the C4 olefins are alkylated with, for instance, methanol to C5 and/or C6 olefins. These C5 and/or C6 olefins may subsequently be converted into at least ethylene and/or propylene. The main by-products from this oxygenate to olefins conversion reaction are again C4 and C5 olefins, which can be recycled to the oxygenate to olefins conversion reactor and olefin cracking reactor, respectively.

Therefore, preferably, where the gaseous products further include C4 olefins, at least part of the C4 olefins are provided to (i) the oxygenate to olefins conversion reactor together with or as part of the oxygenate feed, and/or (ii) the olefin cracking reactor as part of the olefin feed, more preferably at least part of the C4 olefins is provided to the oxygenate to olefins conversion reactor together with or as part of the oxygenate feed.

Preferably, where the gaseous products further include C5 olefins, at least part of the C5 olefins are provided to the olefin cracking reactor as part of the olefin feed.

Preferably, the olefin feed to the olefin cracking reactor comprises C4+ olefins, preferably C5+ olefins, more preferably C5 olefins.

In a preferred embodiment, the oxygenate to olefins conversion reactor and the optional OCP reactor are operated as riser reactors where the catalyst and feedstock are fed at the base of the riser and an effluent stream with entrained catalyst exits the top of the riser. In this embodiment, gas/solid separators are necessary to separate the entrained catalyst from the reactor effluent. The gas/solid separator may be any separator suitable for separating gases from solids. Preferably, the gas/solid separator comprises one or more centrifugal separation units, preferably cyclone units.

The cyclone comprises an upper zone where the products and catalyst are separated due to the centrifugal forces in the cyclone. The cyclone also comprises a lower zone that contains a bed of accumulated catalyst and possibly other solids. A stripping medium is fed into the lower zone to fluidize the catalyst in the bed and strip residual products from the catalyst.

This will be further described in reference to Figure 1. Figure 1 depicts a riser reactor 10 used for the conversion of oxygenates to olefins. The feed and catalyst are preferably fed at the bottom of the riser (not shown). The products and catalyst pass through the reactor outlet 20 and then into the cyclone 30. The cyclone may be directly coupled to the riser reactor as shown in Figure 1 where the outlet of the riser 20 is directly connected to the cyclone inlet 40, with no openings or other outlets for products or catalyst.

The cyclone 30 is divided into two zones. The upper zone 60 is the separation zone where the catalyst and products are separated by the centrifugal forces of the cyclone. The lower zone 70 contains a bed of accumulated catalyst that is stripped and fluidized by stripping medium 80 that enters towards the bottom of lower zone 70. The catalyst passes through dipleg 90 and out of the cyclone. The catalyst may be directed back towards the riser reactor or alternatively to a regenerator. The products separated in the cyclone pass through the gas outlet pipe 50. In one embodiment, the products are then passed into a secondary cyclone 110. The secondary cyclone may be a typical cyclone or it may also contain a lower stripping zone. Vortex stabilizer 100 acts to hold the vortex in the middle of the cyclone.

The reactor effluent is preferably cooled in the gas/solid separator to terminate the conversion process and prevent the formation of by-products outside the reactors. The cooling may be achieved by use of a water quench.

A number of gas/solids separators may be used, but the one connected to the reactor outlet is the primary separator. Any additional gas/solids separators that are connected in series behind the primary gas/solids separator are secondary separators. There may be one or more primary separators and one or more secondary separators.

In a preferred embodiment, the primary gas/solids separator inlet is directly coupled to the riser reactor outlet such that all of the reactor effluent directly enters the gas/solids separator. This provides for a quick separation of the solids from the effluent gas to stop the reaction and prevent the formation of undesirable products and/or byproducts. There are no gaps or openings between the riser reactor outlet and the gas/solids separator inlet.

A transitional element may be used to connect the riser reactor outlet to the gas/solids separator inlet. The transitional element may be a pipe or other conduit, but it does not contain any openings or gaps that would allow the effluent to exit the system.

In another embodiment, secondary gas/solids separators may be directly coupled to the gas outlet of the primary gas/solids separator. The inlet of the secondary gas/solids separator may be coupled to the gas outlet of the primary gas/solids separator using a transitional element. Once the catalyst is separated from the effluent, the catalyst may be returned to the reaction zone from which it came, to another reaction zone or to a regeneration zone. Further, the catalyst that has been separated in the gas/solid separator may be combined with catalyst from other gas/solid separators before it is sent to a reaction zone or to the regeneration zone.

During conversion of the oxygenates to olefins, carbonaceous deposits known as "coke" are formed on the surface of and/or within the molecular sieve catalysts. To avoid a significant reduction in activity of the catalyst, the catalyst must be regenerated by burning off the coke deposits.

In one embodiment, a portion of the coked molecular sieve catalyst is withdrawn from the reactor and introduced into a regeneration system. The regeneration system comprises a regenerator where the coked catalyst is contacted with a regeneration medium, preferably an oxygen-containing gas, under regeneration temperature, pressure and residence time conditions.

Examples of suitable regeneration media include oxygen, O₃, SO₃, N₂O, NO, NO₂, N₂O₅, air, air enriched with oxygen, air diluted with nitrogen or carbon dioxide, oxygen and water, carbon monoxide and/or hydrogen. The regeneration conditions are those capable of burning at least a portion of the coke from the coked catalyst, preferably to a coke level of less than 75% of the coke level on the catalyst entering the regenerator. More preferably the coke level is reduced to less than 50% of the coke level on the catalyst entering the regenerator and most preferably the coke level is reduced to less than 30% of the coke level on the catalyst entering the regenerator. Complete removal of the coke is not necessary as this may result in degradation of the catalyst.

The regeneration temperature is in the range of from 200 °C to 1500 °C, preferably from 300 °C to 1000 °C, more preferably from 450 °C to 700 °C and most preferably from 500 °C to 700 °C. In a preferred embodiment, the catalyst is regenerated at a temperature in the range of from 550 to 650 °C.

The preferred residence time of the coked molecular sieve catalyst in the regenerator is in the range of from 1 minute to several hours, most preferably 1 minute to 100 minutes. The preferred volume of oxygen in the regeneration medium is from 0.01 mole percent to 10 mole percent based on the total volume of the regeneration medium.

In one embodiment, regeneration promoters, typically metal containing compounds such as platinum and palladium are added to the regenerator directly or indirectly, for example with the coked catalyst composition. In another embodiment, a fresh molecular sieve catalyst is added to the regenerator.

In an embodiment, a portion of the regenerated molecular sieve catalyst from the regenerator is returned to the reactor, directly to the reaction zone or indirectly by pre- contacting with the feedstock.

The burning of coke is an exothermic reaction and in certain embodiments, the temperature in the regeneration system is controlled to prevent it from rising too high. Various known techniques for cooling the system and/or the regenerated catalyst may be employed including feeding a cooled gas to the regenerator, or passing the regenerated catalyst through a catalyst cooler. A portion of the cooled regenerated catalyst may be returned to the regenerator while another portion is returned to the reactor.

In certain embodiments, there is not sufficient coke on the catalyst to raise the temperature of the catalyst to desired levels. In one embodiment, a liquid or gaseous fuel may be fed to the regenerator where it will combust and provide additional heat to the catalyst.

Catalysts suitable for use in the conversion of oxygenates to olefins may be made from practically any small or medium pore molecular sieve. One example of a suitable type of molecular sieve is a zeolite. Suitable zeolites include, but are not limited to AEI, AEL, AFT, AFO, APC, ATN, ATT, ATV, AWW, BIK, CAS, CHA, CHI, DAC, DDR, EDI, ERI, EUO, FER, GOO, HEU, KFI, LEV, LOV, LTA, MFI, MEL, MON, MTT, MTW, PAU, PHI, RHO, ROG, THO, TON and substituted forms of these types. 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, and 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. Preferred zeolites for this process include ZSM-5, ZSM-22 and ZSM-23.

A suitable molecular sieve catalyst may have a silica-to-alumina ratio (SAR) of less than 280, preferably less than 200 and more preferably less than 100. The SAR may be in the range of from 10 to 280, preferably from 15 to 200 and more preferably from 20 to 100.

A preferred MFI-type zeolite for the oxygenate to olefins conversion catalyst has a silica-to-alumina ratio, SAR, of at least 60, preferably at least 80. More preferred MFI- type zeolite has a silica-to-alumina ratio, SAR, in the range of 60 to 150, preferably in the range of 80 to 100.

The zeolite-comprising catalyst may comprise more than one zeolite. In that case it is preferred that the catalyst comprises at least 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, and a one-dimensional zeolite having 10-membered ring channels, such as of the MTT and/or TON type.

It is preferred that zeolites in the hydrogen form are used in the zeolite-comprising catalyst, e.g., HZSM-5, HZSM-11, and HZSM-22, HZSM-23. Preferably at least 50wt%, more preferably at least 90wt%, still more preferably at least 95wt% and most preferably 100wt% of the total amount of zeolite used is in the hydrogen form. It is well known in the art how to produce such zeolites in the hydrogen form.

Another example of suitable molecular sieves are siliocoaluminophosphates (SAPOs). SAPOs have a three dimensional microporous crystal framework of P02+, A102-, and Si02 tetrahedral units. Suitable SAPOs include 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 lanthanides of the Periodic Table of Elements. Preferred SAPOs for this process include SAPO-34, SAPO-17 and SAPO-18. Preferred substituent metals for the MeAlPO include Co, Cr, Cu, Fe, Ga, Ge, Mg, Mn, Ni, Sn, Ti, Zn and Zr.

The molecular sieves described above are formulated into molecular sieve catalyst compositions for use in the oxygenates to olefins conversion reaction and the olefin cracking step. The molecular sieves are formulated into catalysts by combining the molecular sieve with a binder and/or matrix material and/or filler and forming the composition into particles by techniques such as spray-drying, pelletizing, or extrusion. The molecular sieve may be further processed before being combined with the binder and/or matrix. For example, the molecular sieve may be milled and/or calcined.

Suitable binders for use in these molecular sieve catalyst compositions include various types of aluminas, aluminophosphates, silicas and/or other inorganic oxide sol. The binder acts like glue binding the molecular sieves and other materials together, particularly after thermal treatment. Various compounds may be added to stabilize the binder to allow processing.

Matrix materials are usually effective at among other benefits, increasing the density of the catalyst composition and increasing catalyst strength (crush strength and/or attrition resistance). Suitable matrix materials include one or more of the following: rare earth metals, metal oxides including titania, zirconia, magnesia, thoria, beryllia, quartz, silica or sols, and mixtures thereof, for example, silica-magnesia, silica- zirconia, silica- titania, and silica-alumina. In one embodiment, matrix materials are natural clays, for example, kaolin. A preferred matrix material is kaolin.

In one embodiment, the molecular sieve, binder and matrix material are combined in the presence of a liquid to form a molecular sieve catalyst slurry. The amount of binder is in the range of from 2 to 40 wt%, preferably in the range of from 10 to 35 wt%, more preferably in the range of from 15 to 30 wt%, based on the total weight of the molecular sieve, binder and matrix material, excluding liquid (after calcination).

After forming the slurry, the slurry may be mixed, preferably with rigorous mixing to form a substantially homogeneous mixture. Suitable liquids include one or more of water, alcohols, ketones, aldehydes and/or esters. Water is the preferred liquid. In one embodiment, the mixture is colloid-milled for a period of time sufficient to produce the desired texture, particle size or particle size distribution.

The molecular sieve, matrix and optional binder can be in the same or different liquids and are combined in any order together, simultaneously, sequentially or a combination thereof. In a preferred embodiment, water is the only liquid used.

In a preferred embodiment, the slurry is mixed or milled to achieve a uniform slurry of sub-particles that is then fed to a forming unit. A slurry of the zeolite may be prepared and then milled before combining with the binder and/or matrix. In a preferred embodiment, the forming unit is a spray dryer. The forming unit is typically operated at a temperature high enough to remove most of the liquid from the slurry and from the resulting molecular sieve catalyst composition. In a preferred embodiment, the particles are then exposed to ion-exchange using an ammonium nitrate or other appropriate solution.

In one embodiment, the ion exchange is carried out before the phosphorous impregnation. The ammonium nitrate is used to ion exchange the zeolite to remove alkali ions. The zeolite can be impregnated with phosphorous using phosphoric acid followed by a thermal treatment to H+ form. In another embodiment, the ion exchange is carried out after the phosphorous impregnation. In this embodiment, alkali phosphates or phosphoric acid may be used to impregnate the zeolite with phosphorous, and then the ammonium nitrate and heat treatment are used to ion exchange and convert the zeolite to the H+ form.

Alternatively to spray drying the catalyst may be formed into spheres, tablets, rings, extrudates or any other shape known to one of ordinary skill in the art. The catalyst may be extruded into various shapes, including cylinders and trilobes.

The average particle size is in the range of from 1-200 μιη, preferably from 50-100 μιη. If extrudates are formed, then the average size is in the range of from 1 mm to 10 mm, preferably from 2 mm to 7 mm.

The catalyst may further 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 a MEL or MFI-type zeolite comprising catalyst additionally comprises phosphorus.

The molecular sieve catalyst is prepared by first forming a molecular sieve catalyst precursor as described above, optionally impregnating the catalyst with a phosphorous containing compound and then calcining the catalyst precursor to form the catalyst. The phosphorous impregnation may be carried out by any method known to one of skill in the art.

The phosphorus-containing compound preferably comprises a phosphorus species such as P0₄ ^3~, Ρ-(0Ο¾)3, or P2O5, especially P0₄ ^3". Preferably the phosphorus-containing compound comprises a compound selected from the group consisting of ammonium phosphate, ammonium dihydrogen phosphate, dimethylphosphate, metaphosphoric acid and trimethyl phosphite and phosphoric acid, especially phosphoric acid. The phosphorus containing compound is preferably not a Group II metal phosphate. Group II metal species include magnesium, calcium, strontium and barium; especially calcium.

In one embodiment, phosphorus can be deposited on the catalyst by impregnation using acidic solutions containing phosphoric acid (]¾Ρ0₄). The concentration of the solution can be adjusted to impregnate the desired amount of phosphorus on the precursor. The catalyst precursor may then be dried.

The catalyst precursor, containing phosphorous (either in the framework or impregnated) is calcined to form the catalyst. The calcination of the catalyst is important to determining the performance of the catalyst in the oxygenate to olefins process. The calcination may be carried out in any type of calciner known to one of ordinary skill in the art. The calcination may be carried out in a tray calciner, a rotary calciner, or a batch oven optionally in the presence of an inert gas and/or oxygen and/or steam

The calcination may be carried out at a temperature in the range of from 400 °C to 1000 °C, preferably in a range of from 450 °C to 800 °C, more preferably in a range of from 500 °C to 700 ⁰ C. Calcination time is typically dependent on the degree of hardening of the molecular sieve catalyst composition and the temperature and ranges from about 15 minutes to about 2 hours.

The calcination temperatures described above are temperatures that are reached for at least a portion of the calcination time. For example, in a rotary calciner, there may be separate temperature zones that the catalyst passes through. For example, the first zone may be at a temperature in the range of from 100 to 300 °C. At least one of the zones is at the temperatures specified above. In a stationary calciner, the temperature is increased from ambient to the calcination temperatures above and so the temperature is not at the calcination temperature for the entire time.

In a preferred embodiment, the calcination is carried out in air at a temperature of from 500 °C to 600 °C. The calcination is carried out for a period of time from 30 minutes to 15 hours, preferably from 1 hour to 10 hours, more preferably from 1 hour to 5 hours.

The calcination is carried out on a bed of catalyst. For example, if the calcination is carried out in a tray calciner, then the catalyst precursor added to the tray forms a bed which is typically kept stationary during the calcination. If the calcination is carried out in a rotary calciner, then the catalyst added to the rotary drum forms a bed that although not stationary does maintain some form and shape as it passes through the calciner.

Claims

C L A I M S

1. An oxygenate to olefins conversion reactor system comprising:

a. a riser reactor having an inlet for feed, an inlet for catalyst and an outlet for products; and

b. a cyclone for the separation of products from catalyst comprising:

i. an upper zone where the products and catalyst are separated; and ii. a lower zone comprising a bed of accumulated solids wherein a stripping medium is fed into the lower zone to fluidize the catalyst and strip residual products from the catalyst.

2. The system of claim 1 wherein the cyclone is directly coupled to the riser reactor outlet such that all of the reactor effluent directly enters the cyclone.

3. The system of any of claims 1-2 wherein the feed and/or catalyst are fed into the riser reactor by more than one inlets.

4. The system of any of claims 1-3 further comprising a distribution device for

feeding the stripping medium into the lower zone and wherein the upper zone has an inlet for the products and entrained catalyst and an outlet for the products; the lower zone has an outlet for removal of the catalyst from the cyclone; and the upper zone has a larger diameter than the lower zone.

5. The system of any of claims 1-4 further comprising a secondary cyclone that is in fluid communication with the products outlet of the cyclone and has an outlet for catalyst that is in fluid communication with the lower zone of the cyclone.

6. The system of any of claims 1-4 further comprising a secondary cyclone that is in fluid communication with the products outlet of the cyclone and has an outlet for catalyst that is in fluid communication with the upper zone of the cyclone.

7. The system of any of claims 5-6 comprising a plurality of secondary cyclones which are in fluid communication with the products outlet of the cyclone.

8. The system of any of claims 1-7 wherein the cyclones are located inside another vessel.

9. A process for converting oxygenates to olefins comprising:

a. feeding an oxygenate containing feed stream into a riser reactor;

b. contacting the feed stream with a fluidized oxygenate conversion catalyst under oxygenate conversion conditions in the riser reactor to produce an effluent stream comprising olefins and solid catalyst particles; c. passing the effluent stream from the riser reactor into a cyclone;

d. separating a product stream comprising olefins from the solid catalyst particles in an upper zone of the cyclone;

e. fluidizing and stripping the solid catalyst particles in a lower zone of the cyclone by injecting a stripping medium into the lower zone;

f. removing a products stream from the upper zone of the cyclone; and g. removing catalyst from the lower zone of the cyclone.

10. The process of claim 9 wherein the oxygenate comprises methanol and/or dimethyl ether.

11. The process of any of claims 9-10 wherein the oxygenate conversion catalyst is a molecular sieve catalyst.

12. The process of any of claims 9-11 wherein step (c) comprises passing the effluent through a transitional element connecting the riser reactor with the cyclone.

13. The process of any of claims 9-12 wherein the stripping medium comprises steam, nitrogen, methane or another suitable gas.

14. A cyclone for the separation of gas and solids comprising:

a. an upper zone where gas and solids are separated; and

b. a lower zone comprising a bed of accumulated solids

wherein a stripping medium is fed into the lower zone to fluidize the accumulated solids and strip residual gas from the solids.

15. The cyclone of claim 14 wherein the upper zone has an inlet for the gas and solids mixture and an outlet for the gas.

16. The cyclone of any of claims 14-15 wherein the upper zone has a larger diameter than the lower zone.

17. The cyclone of any of claims 14-16 wherein the lower zone has an outlet for

removal of solids from the cyclone.

18. The cyclone of any of claims 14-17 further comprising a distribution device for feeding the stripping medium into the lower zone.

19. The cyclone of any of claims 14-18 further comprising a support member located in the upper zone.