WO2003033439A2

WO2003033439A2 - A method for producing olefins from methanol

Info

Publication number: WO2003033439A2
Application number: PCT/EP2002/011408
Authority: WO
Inventors: Wilfried Borgmann; Josef Kunkel; Helmut Fritz; Gerhard Lauermann; Roland Walzl; Klaus Müller
Original assignee: Exxonmobil Chemical Patents Inc.
Priority date: 2001-10-16
Filing date: 2002-10-11
Publication date: 2003-04-24
Also published as: WO2003033439A3; DE10150481B4; DE10150481A1; AU2002362888A1

Abstract

A method is described for producing olefins from methanol in a reactor, wherein liquid methanol is evaporated and introduced into the reactor as gaseous methanol for olefin synthesis, while hot olefin-containing reactor exit gas is removed from the reactor. For better utilization of heat, it is proposed to send at least a part of the liquid methanol (1) with the reactor exit gas (15) into a heat exchanger (9), where the liquid methanol (1) is evaporated and the reactor exit gas (15) is simultaneously cooled. The methanol evaporated in the way (14) can be impute to the reactor directly or after addition to another stream (3).

Description

A METHOD FOR PRODUCING OLEFTNS FROM METHANOL

[0001] This application claims benefit of the filing date of German Patent

Application No. 101 50 481.0, filed October 16, 2001, the contents of which are fully incorporated herein by reference.

Field of the Invention

[0002] The invention concerns a method for producing olefins from methanol in a reactor, wherein liquid methanol is evaporated and sent as gaseous methanol to the reactor for olefin synthesis, while hot olefin-containing reactor exit gas is removed from the reactor.

Background of the Invention

[0003] Olefin synthesis from methanol is seen as a highly promising alternative to traditional olefin production from petroleum. Methanol is considered to be a readily storable and manageable intermediate product for utilization of hitherto unused natural gas. Thus, the increasing demand for olefins on the world market could also be served by using very cheap methane. The processes that have been proposed up to now for producing short-chain olefins from methanol operate, for example, catalytically, according to the overall equation 2CH₃OH → C^ + 2H₂O. Here the olefin synthesis takes place in a reactor to which gaseous methanol is supplied. In the reactor the methanol is converted at high temperatures to olefins and byproducts, which leave the reactor as hot reactor exit gas. For fractionation of the olefins and byproducts, the reactor gas first has to be cooled. Since the fresh methanol is usually in the liquid state, it has to be evaporated by supplying heat before it goes to the reactor. To cool the reactor exit gas, and to heat the fresh methanol, water-steam heat exchangers are usually provided. The operation of these water steam cycles, however, results in a significant energy requirement. Accordingly, additional methods of conserving energy are sought. Summary of the Invention

[0004] This invention provides a process for producing olefins from methanol in a way that conserves a significant amount of energy. According to the method, olefins are produced from methanol in a reactor such that the methanol is evaporated for use as feed gas to the reactor. In addition, the methanol is used to cool down the reactor exit gas for fractionation in an economical way.

[0005] In one embodiment, the invention provides a method for producing olefins from methanol in a reactor in which methanol is heat exchanged with hot olefin-containing reactor exit gas to at least partially vaporize the methanol and cool the hot olefin-containing reactor exit gas. The at least partially vaporized methanol is depressurized, and the depressurized methanol is mixed with another methanol stream to form a mixed methanol stream. The mixed methanol stream is contacted with olefin forming catalyst in a reactor to form the hot olefin- containing reactor exit gas that is used to at least partially vaporize the methanol. [0006] In another embodiment, the invention provides a method for producing olefins from methanol in a reactor in which hot olefin-containing reactor exit gas is heat exchanged with water to form steam and a cooled olefin- containing reactor exit gas. The cooled olefin-containing reactor exit gas is heat exchanged with methanol to at least partially vaporize the methanol, and the at least partially vaporized methanol is contacted with an olefin forming catalyst to form the hot olefin-containing reactor exit gas.

[0007] In yet another embodiment, there is provided a method for producing olefins from methanol in a reactor in which a first methanol stream is heat exchanged with hot olefin-containing reactor exit gas to at least partially vaporize the methanol in the first methanol stream and cool the hot olefin- containing reactor exit gas. A second methanol stream is heat exchanged with steam to at least partially vaporize the methanol in the second methanol stream, and the vaporized methanol from the first and second methanol streams are mixed together to form a mixed methanol stream. The mixed methanol stream is contacted with olefin forming catalyst in a reactor to form the hot olefin- containing reactor exit gas that is used to at least partially vaporize the methanol. Brief Description of the Drawing

[0008] The attached Figure shows an example of but one type of flow scheme of the invention in which methanol is evaporated and sent as gaseous methanol to a reactor for olefin synthesis, while hot olefin-containing reactor exit gas is removed from the reactor and cooled.

Detailed Description of the Invention

[0009] The problem of conserving energy is solved in accordance with the invention by providing heat exchange between at least a part of the liquid methanol and hot reactor exit gas. The heat exchange is such that the liquid methanol becomes evaporated, and the reactor exit gas is cooled at the same time. The evaporated methanol is sent directly to the reactor, or after adding it to another stream.

[0010] According to this invention, fresh methanol is evaporated and hot reactor exit gas is cooled at the same time. In one embodiment, the reactor exit gas is cooled while producing vapor (i.e., steam) at one or more pressures. The fresh methanol can be heated with the thus produced vapor, and in this way the methanol is evaporated. In this embodiment, which calls for a process vapor cycle in which the heating of the methanol and cooling of the reactor exit gas are connected is, however, limited in that it is desirable to limit the pressure of the vaporized methanol from about 2.5 bar absolute to about 5 bar absolute. This means that the reactor exit gas should only be cooled to about 180°C, before it goes, for example, to a wet quenching step, before going on to fractionation. In the quenching step, the maximum heat level is from about 80°C to about 100°C, which is unsuitable for evaporating the methanol. In the end, additional heat would be removed by outside cooling.

[0011] In another embodiment of the invention, at least a part of the liquid methanol is brought into heat exchange with the reactor exit gas without intermediate connection of a heat exchange cycle. In this embodiment, it is desirable to use a heat exchanger with vertical tubes that can be cleaned during operation by injecting water. In the heat exchanger, the methanol to be evaporated is brought into indirect heat exchange directly with the reactor exit gas. By omitting an intermediately connected heat exchange cycle, unnecessary heat losses can be avoided. Through the economy of the heat/cold circulation, and the corresponding reduction of the necessary driving temperature difference, heat is obtained at a favorable level. In a 1000 KTA plant this is, for example, about 21 megawatts, which, without the methanol evaporation step in accordance with the invention, must be generated on one side of a steam generator, and has to be eliminated at the cold end at high cost.

[0012] According to one particularly preferred embodiment of the invention, a methanol pressure is established in the heat exchange between the methanol and the reactor exit gas at which the water dew point of 113 °C at 2137 bar is not exceeded on the reactor exit gas side. In this way, no water can condense on the reactor exit gas side. Avoiding water condensation in the gas cooler brings process engineering advantages.

[0013] Expediently, the reactor exit gas is cooled in several steps; in particular, in three steps by means of heat exchangers designed as gas coolers. The methanol is advantageously brought into heat exchange with the reactor exit gas in the last step of gas cooling.

[0014] In principle, all of the fresh methanol provided for input to the reactor can be evaporated by heat exchange with the reactor exit gas. However, it is also possible for only a part of the methanol to be evaporated in this way. The remaining amount of the methanol can be mixed in with the evaporated methanol before going into the reactor, or it can be partly or fully evaporated by other means, such as by using steam. Preferably, the methanol evaporated in accordance with the invention is depressurized into a stream of unevaporated methanol to form a mixed methanol stream. The mixed methanol stream is then sent to the reactor.

[0015] The invention offers a number of advantages. For example, the methanol can be evaporated at the pressure at which the reactor is operated. As another example, the reactor exit gas can be substantially cooled before it is sent to a wet quenching step. Moreover, water condensation in the gas cooler can be avoided prior to the wet quenching step. It is also possible for other process streams to be cooled through methanol evaporation. In this way the utilization of the heat can be increased even further.

[0016] In one embodiment of the invention, the reactor exit gas is obtained by contacting methanol with an olefin forming catalyst in a reactor. Preferably, the catalyst is a molecular sieve catalyst.

[0017] Although the use of methanol to produce the olefin stream is preferred, other oxygenate components can be used as a feed. Such oxygenates comprise at least one organic compound which contains at least one oxygen atom, such as aliphatic alcohols, ethers, carbonyl compounds (aldehydes, ketones, carboxylic acids, carbonates, esters and the like). When the oxygenate is an alcohol, the alcohol includes an aliphatic moiety having from 1 to 10 carbon atoms, more preferably from 1 to 4 carbon atoms. Representative alcohols include but are not necessarily limited to lower straight and branched chain aliphatic alcohols and their unsaturated counterparts; Examples of suitable oxygenate compounds include, but are not limited to: methanol; ethanol; n-propanol; isopropanol; C₄ - C₂₀ alcohols; methyl ethyl ether; dimethyl ether; diethyl ether; di-isopropyl ether; formaldehyde; dimethyl carbonate; dimethyl ketone; acetic acid; and mixtures thereof. Dimethyl ether, or a mixture of dimethyl ether and methanol, are also preferred feeds.

[0018] Molecular sieves capable of converting an oxygenate such as methanol to an olefin compound include zeolites as well as non-zeolites, and are of the large, medium or small pore type. Small pore molecular sieves are preferred in one embodiment of this invention, however. As defined herein, small pore molecular sieves have a pore size of less than about 5.0 angstroms. Generally, suitable catalysts have a pore size ranging from about 3.5 to about 5.0 angstroms, preferably from about 4.0 to about 5.0 angstroms, and most preferably from about 4.3 to about 5.0 angstroms.

[0019] Zeolite materials, both natural and synthetic, have been demonstrated to have catalytic properties for various types of hydrocarbon conversion processes. In addition, zeolite materials have been used as adsorbents, catalyst carriers for various types of hydrocarbon conversion processes, and other applications. Zeolites are complex crystalline alumino silicates which form a network of AlO₂ ^" and SiO₂ tetrahedra linked by shared oxygen atoms. The negativity of the tetrahedra is balanced by the inclusion of cations such as alkali or alkaline earth metal ions. In the manufacture of some zeolites, non-metallic cations, such as tetramethylammonium (TMA) or tetrapropylammonium (TPA), are present during synthesis. The interstitial spaces or channels formed by the crystalline network enable zeolites to be used as molecular sieves in separation processes, as catalyst for chemical reactions, and as catalyst carriers in a wide variety of hydrocarbon conversion processes.

[0020] Zeolites include materials containing silica and optionally alumina, and materials in which the silica and alumina portions have been replaced in whole or in part with other oxides. For example, germanium oxide, tin oxide, and mixtures thereof can replace the silica portion. Boron oxide, iron oxide, gallium oxide, indium oxide, and mixtures thereof can replace the alumina portion. Unless otherwise specified, the terms "zeolite" and "zeolite material" as used herein, shall mean not only materials containing silicon atoms and, optionally, aluminum atoms in the crystalline lattice structure thereof, but also materials which contain suitable replacement atoms for such silicon and aluminum atoms. [0021] One type of olefin forming catalyst capable of producing large quantities of ethylene and propylene is a silicoaluminophosphate (SAPO) molecular sieve. Silicoaluminophosphate molecular sieves are generally classified as being microporous materials having 8, 10, or 12 membered ring structures. These ring structures can have an average pore size ranging from about 3.5 to about 15 angstroms. Preferred are the small pore SAPO molecular sieves having an average pore size of less than about 5 angstroms, preferably an average pore size ranging from about 3.5 to about 5 angstroms, more preferably from about 3.5 to about 4.2 angstroms. These pore sizes are typical of molecular sieves having 8 membered rings.

[0022] According to one embodiment, substituted SAPOs can also be used in oxygenate to olefin reaction processes. These compounds are generally known as MeAPSOs or metal-containing silicoaluminophosphates. The metal can be alkali metal ions (Group IA), alkaline earth metal ions (Group DA), rare earth ions (Group DIB, including the lanthanoid elements: lanthanum, cerium, praseodymium, neodymium, samarium, europium, gadolinium, terbium, dysprosium, holmium, erbium, thulium, ytterbium and lutetium; and scandium or yttrium) and the additional transition cations of Groups IVB, VB, VTB, VDB, VHIB, and IB.

[0023] Preferably, the Me represents atoms such as Zn, Mg, Mn, Co, Ni,

Ga, Fe₅ Ti, Zr, Ge, Sn, and Cr. These atoms can be inserted into the tetrahedral framework through a [MeO₂] tetrahedral unit. The [MeO₂] tetrahedral unit carries a net electric charge depending on the valence state of the metal substituent. When the metal component has a valence state of +2, +3, +4, +5, or +6, the net electric charge is between -2 and +2. Incorporation of the metal component is typically accomplished adding the metal component during synthesis of the molecular sieve. However, post-synthesis ion exchange can also be used. In post synthesis exchange, the metal component will introduce cations into ion-exchange positions at an open surface of the molecular sieve, not into the framework itself. [0024] Suitable silicoaluminophosphate molecular sieves include SAPO-5,

SAPO-8, SAPO-11, SAPO-16, SAPO-17, SAPO-18, SAPO-20, SAPO-31, SAPO- 34, SAPO-35, SAPO-36, SAPO-37, SAPO-40, SAPO-41, SAPO-42, SAPO-44, SAPO-47, SAPO-56, the metal containing forms thereof, and mixtures thereof Preferred are SAPO-18, SAPO-34, SAPO-35, SAPO-44, and SAPO-47, particularly SAPO-18 and SAPO-34, including the metal containing forms thereof, and mixtures thereof. As used herein, the term mixture is synonymous with combination and is considered a composition of matter having two or more components in varying proportions, regardless of their physical state. [0025] An aluminophosphate (ALPO) molecular sieve can also be included in the catalyst composition. Aluminophosphate molecular sieves are crystalline microporous oxides which can have an AlPO₄ framework. They can have additional elements within the framework, typically have uniform pore dimensions ranging from about 3 angstroms to about 10 angstroms, and are capable of making size selective separations of molecular species. More than two dozen structure types have been reported, including zeolite topological analogues. A more detailed description of the background and synthesis of aluminophosphates is found in U.S. Pat. No. 4,310,440, which is incorporated herein by reference in its entirety. Preferred ALPO structures are ALPO-5, ALPO-11, ALPO-18, ALPO-31, ALPO-34, ALPO-36, ALPO-37, and ALPO-46. [0026] The ALPOs can also include a metal substituent in its framework.

Preferably, the metal is selected from the group consisting of magnesium, manganese, zinc, cobalt, and mixtures thereof. These materials preferably exhibit adsorption, ion-exchange and/or catalytic properties similar to aluminosilicate, aluminophosphate and silica aluminophosphate molecular sieve compositions. Members of this class and their preparation are described in U.S. Pat. No. 4,567,029, incorporated herein by reference in its entirety. [0027] The metal containing ALPOs have a three-dimensional microporous crystal framework structure of MO₂, AlO₂ and PO₂ tetrahedral units. These as manufactured structures (which contain template prior to calcination) can be represented by empirical chemical composition, on an anhydrous basis, as: mR: (M_xAl_yP_z)O₂ wherein "R" represents at least one organic templating agent present in the intracrystalline pore system; "m" represents the moles of "R" present per mole of (M_xAlyP_z)O₂ and has a value of from zero to 0.3, the maximum value in each case depending upon the molecular dimensions of the templating agent and the available void volume of the pore system of the particular metal aluminophosphate involved, "x", "y", and "z" represent the mole fractions of the metal "M', (i.e. magnesium, manganese, zinc and cobalt), aluminum and phosphorus, respectively, present as tetrahedral oxides. [0028] The metal containing ALPOs are sometimes referred to by the acronym as MeAPO. Also in those cases where the metal "Me" in the composition is magnesium, the acronym MAPO is applied to the composition. Similarly ZAPO, MnAPO and CoAPO are applied to the compositions which contain zinc, manganese and cobalt respectively. To identify the various structural species which make up each of the subgeneric classes MAPO, ZAPO, CoAPO and MnAPO, each species is assigned a number and is identified, for example, as ZAPO-5, MAPO-11, CoAPO-34 and so forth. [0029] The silicoaluminophosphate molecular sieve is typically admixed

(i.e., blended) with other materials. When blended, the resulting composition is typically referred to as a SAPO catalyst, with the catalyst comprising the SAPO molecular sieve.

[0030] Materials which can be blended with the molecular sieve can be various inert or catalytically active materials, or various binder materials. These materials include compositions such as kaolin and other clays, various forms of rare earth metals, metal oxides, other non-zeolite catalyst components, zeolite catalyst components, alumina or alumina sol, titania, zirconia, magnesia, thoria, beryllia, quartz, silica or silica or silica sol, and mixtures thereof. These components are also effective in reducing, inter alia, overall catalyst cost, acting as a thermal sink to assist in heat shielding the catalyst during regeneration, densifying the catalyst and increasing catalyst strength. It is particularly desirable that the inert materials that are used in the catalyst to act as a thermal sink have a heat capacity of from about 0.05 to about 1 cal/g-°C, more preferably from about 0.1 to about 0.8 cal/g-°C, most preferably from about 0.1 to about 0.5 cal/g-°C. [0031] Additional molecular sieve materials can be included as a part of the SAPO catalyst composition or they can be used as separate molecular sieve catalysts in admixture with the SAPO catalyst if desired. Structural types of small pore molecular sieves that are suitable for use in this invention include AEI, AFT, APC, ATN, ATT, ATV, AWW, BIK, CAS, CHA, CHI, DAC, DDR, EDI, ERL GOO, KFI, LEV, LOV, LTA, MON, PAU, PHI, RHO, ROG, THO, and substituted forms thereof. Structural types of medium pore molecular sieves that are suitable for use in this invention include MFL MEL, MTW, EUO, MTT, HEU, FER, AFO, AEL, TON, and substituted forms thereof. These small and medium pore molecular sieves are described in greater detail in the Atlas of Zeolite Structural Types, W.M. Meier and D.H. Olsen, Butterworth Heineman, 3rd ed., 1997, the detailed description of which is explicitly incorporated herein by reference. Preferred molecular sieves which can be combined with a silicoaluminophosphate catalyst include ZSM-5, ZSM-34, erionite, and chabazite. [0032] The catalyst composition, according to an embodiment, preferably comprises from about 1% to about 99 %, more preferably from about 5 % to about 90 %, and most preferably from about 10% to about 80%, by weight of molecular sieve. It is also preferred that the catalyst composition have a particle size of from about 20 angstroms to about 3,000 angstroms, more preferably from about 30 angstroms to about 200 angstroms, most preferably from about 50 angstroms to about 150 angstroms.

[0033] The catalyst can be subjected to a variety of treatments to achieve the desired physical and chemical characteristics. Such treatments include, but are not necessarily limited to hydrothermal treatment, calcination, acid treatment, base treatment, milling, ball milling, grinding, spray drying, and combinations thereof. [0034] A molecular sieve catalyst particularly useful in making ethylene and propylene is a catalyst which contains a combination of SAPO-34, and SAPO- 18 or ALPO-18 molecular sieve. In a particular embodiment, the molecular sieve is a crystalline intergrowth of SAPO-34, and SAPO-18 or ALPO-18. [0035] To convert methanol or other oxygenate to olefin, conventional reactor systems can be used, including fixed bed, fluid bed or moving bed systems. Preferred reactors of one embodiment are co-current riser reactors and short contact time, countercurrent free-fall reactors. Desirably, the reactor is one in which an oxygenate feedstock can be contacted with a molecular sieve catalyst at a weight hourly space velocity (WHSV) of at least about 1 hr^"1, preferably in the range of from about 1 hr^"1 to 1000 hr^"1, more preferably in the range of from about 20 hr^"1 to about 1000 hr^"1, and most preferably in the range of from about 50 hr^"1 to about 500 hr^"1. WHSV is defined herein as the weight of oxygenate, and reactive hydrocarbon which may optionally be in the feed, per hour per weight of the molecular sieve in the reactor. Because the catalyst or the feedstock may contain other materials which act as inerts or diluents, the WHSV is calculated on the weight basis of the oxygenate feed, and any reactive hydrocarbon which may be present with the oxygenate feed, and the molecular sieve contained in the reactor.

[0036] Preferably, the oxygenate feed is contacted with the catalyst when the oxygenate is in a vapor phase. Alternately, the process may be carried out in a liquid or a mixed vapor/liquid phase. When the process is carried out in a liquid phase or a mixed vapor/liquid phase, different conversions and selectivities of feed-to-product may result depending upon the catalyst and reaction conditions. [0037] The process can generally be carried out at a wide range of temperatures. An effective operating temperature range can be from about 200°C to about 700°C, preferably from about 300°C to about 600°C, more preferably from about 350°C to about 550°C. At the lower end of the temperature range, the formation of the desired olefin products may become markedly slow with a relatively high content of oxygenated olefin by-products being found in the olefin product. However, the selectivity to ethylene and propylene at reduced temperatures may be increased. At the upper end of the temperature range, the process may not form an optimum amount of ethylene and propylene product, but the conversion of oxygenate feed will generally be high. [0038] Operating pressure also may vary over a wide range, including autogenous pressures. Effective pressures include, but are not necessarily limited to, a total pressure of at least 1 psia (7 kPa), preferably at least about 5 psia (34 kPa). The process is particularly effective at higher total pressures, including a total pressure of at least about 20 psia (138 kPa). Preferably, the total pressure is at least about 25 psia (172 kPa), more preferably at least about 30 psia (207 kPa). For practical design purposes it is desirable to use methanol as the primary oxygenate feed component, and operate the reactor at a pressure of not greater than about 500 psia (3445 kPa), preferably not greater than about 400 psia (2756 kPa), most preferably not greater than about 300 psia (2067 kPa). [0039] Undesirable by-products can be avoided by operating at an appropriate gas superficial velocity. As the gas superficial velocity increases the conversion decreases avoiding undesirable by-products. As used herein, the term, "gas superficial velocity" is defined as the combined volumetric flow rate of vaporized feedstock, which includes diluent when present in the feedstock, as well as conversion products, divided by the cross-sectional area of the reaction zone. Because the oxygenate is converted to a product having significant quantities of ethylene and propylene while flowing through the reaction zone, the gas superficial velocity may vary at different locations within the reaction zone. The degree of variation depends on the total number of moles of gas present and the cross section of a particular location in the reaction zone, temperature, pressure and other relevant reaction parameters. [0040] In one embodiment, the gas superficial velocity is maintained at a rate of greater than 1 meter per second (m/s) at least one point in the reaction zone. In another embodiment, it is desirable that the gas superficial velocity is greater than about 2 m/s at least one point in the reaction zone. More desirably, the gas superficial velocity is greater than about 2.5 m/s at least one point in the reaction zone. Even more desirably, the gas superficial velocity is greater than about 4 m/s at least one point in the reaction zone. Most desirably, the gas superficial velocity is greater than about 8 m/s at least one point in the reaction zone.

[0041] According to yet another embodiment of the invention, the gas superficial velocity is maintained relatively constant in the reaction zone such that the gas superficial velocity is maintained at a rate greater than 1 m/s at all points in the reaction zone. It is also desirable that the gas superficial velocity be greater than about 2 m/s at all points in the reaction zone. More desirably, the gas superficial velocity is greater than about 2.5 m/s at all points in the reaction zone. Even more desirably, the gas superficial velocity is greater than about 4 m/s at all points in the reaction zone. Most desirably, the gas superficial velocity is greater than about 8 m/s at all points in the reaction zone.

[0042] The amount of ethylene and propylene produced in the oxygenate to olefin process can be increased by reducing the conversion of the oxygenates in the oxygenate to olefins reaction. However, reducing the conversion of feed oxygenates in the oxygenate conversion reaction tends to increase the amount of oxygenated hydrocarbons, particularly including dimethyl ether, that are present in the olefin product. Thus, control of the conversion of feed to the oxygenate reaction process can be important.

[0043] According to one embodiment, the conversion of the primary oxygenate, e.g., methanol, is from 90 wt % to 98 wt %. According to another embodiment the conversion of methanol is from 92 wt % to 98 wt %, preferably

[0044] According to another embodiment, the conversion of methanol is above 98 wt % to less than 100 wt %. According to another embodiment, the conversion of methanol is from 98.1 wt % to less than 100 wt %; preferably from 98.2 wt % to 99.8 wt %. According to another embodiment, the conversion of methanol is from 98.2 wt % to less than 99.5 wt ; preferably from 98.2 wt % to 99 wt %.

[0045] In this invention, weight percent conversion is calculated on a water free basis unless otherwise specified. Weight percent conversion on a water free basis is calculated as: 100 x (weight oxygenate fed on a water free basis - weight oxygenated hydrocarbon in the product on a water free basis). The water free basis of oxygenate is calculated by subtracting out the water portion of the oxygenate in the feed and product, and excluding water formed in the product. For example, the weight flow rate of methanol on an oxygenate free basis is calculated by multiplying the weight flow rate of methanol by 14/32 to remove the water component of the methanol. As another example, the rate flow rate of dimethyl ether on an oxygenate free basis is calculated by multiplying the weight flow rate of dimethyl ether by 28/46 to remove the water component of the dimethyl ether. If there is a mixture of oxygenates in the feed or product, trace oxygenates are not included. When methanol and/or dimethyl ether is used as the feed, only methanol and dimethyl ether are used to calculate conversion on a water free basis.

[0046] In this invention, selectivity is also calculated on a water free basis unless otherwise specified. Selectivity is calculated as: 100 x wt % component / (100 - wt % water - wt % methanol - wt % dimethyl ether) when methanol and/or dimethyl ether is used as the feed.

[0047] An example of the invention is shown in the Figure. According to the Figure, fresh liquid methanol is supplied from a supply tank (not shown in the Figure) and sent through a pipe (1) and through a pipe (3) to sections of the plant for evaporation of the methanol. The methanol flowing in pipe (3) is heated in two steps in heat exchangers (4) and (5) by process steam. The process steam is supplied to heat exchangers (5) and (4) through a pipe (20). Condensate arising from condensation of the process steam following heat exchange to vaporize the methanol is removed through a pipe (6). The vaporized methanol is then sent through a pipe (2) to a olefin synthesis reactor (not shown in the Figure). [0048] The liquid methanol flowing in pipe (1) is sent through pipes (7) and (8) to heat exchanger (9). In heat exchanger (9), the methanol is brought into heat exchange with hot reactor exit gas, which is withdrawn from the olefin synthesis reactor (not shown) and sent through a pipe (15) to the heat exchanger (9). The hot reactor exit gas is cooled prior to contact with heat exchanger (9) in gas coolers (11) and (10). Water is sent to the gas coolers (11) and (10), and steam is formed by the heat exchange with the reactor exit gas. The steam exits the gas coolers (11) and (10) through lines (18) and (19), respectively. A high pressure steam can be formed in gas cooler (11) and a low pressure steam can be formed in gas cooler (10). The reactor exit gas, cooled in the heat exchangers (11), (10) and (9), is removed via a pipe (16). The cooled reactor exit gas can be sent to a quenching step (not shown in the figure), before being further sent to fractionation.

[0049] The methanol evaporated by heat exchange with the reactor exit gas in heat exchanger (9) is sent through a pipe (14), and mixed with the methanol stream flowing in the pipe (2). Liquid methanol can be removed through a pipe (17), and sent to an appropriate use.

[0050] Having now fully described this invention, it will be appreciated by those skilled in the art that the invention can be performed within a wide range of parameters within what is claimed, without departing from the spirit and scope of the invention.

Claims

What is claimed is;

1. A method for producing olefins from methanol in a reactor, comprising: heat exchanging methanol with hot olefin-containing reactor exit gas to at least partially vaporize the methanol and cool the hot olefin-containing reactor exit gas; depressurizing the at least partially vaporized methanol; mixing the depressurized methanol with another methanol stream to form a mixed methanol stream; and contacting the mixed methanol stream with olefin forming catalyst in a reactor to form the hot olefin-containing reactor exit gas that is used to at least partially vaporize the methanol.

2. A method for producing olefins from methanol in a reactor, comprising: heat exchanging hot olefin-containing reactor exit gas with water to form steam and a cooled olefin-containing reactor exit gas; heat exchanging the cooled olefin-containing reactor exit gas with methanol to at least partially vaporize the methanol; and contacting the at least partially vaporized methanol with an olefin forming catalyst to form the hot olefin-containing reactor exit gas.

3. A method for producing olefins from methanol in a reactor, comprising: heat exchanging a first methanol stream with hot olefin-containing reactor exit gas to at least partially vaporize the methanol in the first methanol stream and cool the hot olefin-containing reactor exit gas; heat exchanging a second methanol stream with steam to at least partially vaporize the methanol in the second methanol stream; mixing together the vaporized methanol from the first and second methanol streams to form a mixed methanol stream; and contacting the mixed methanol stream with olefin forming catalyst in a reactor to form the hot olefin-containing reactor exit gas that is used to at least partially vaporize the methanol.