Process for converting oxygenates to olefins using molecular sieve catalysts comprising desirable carbonaceous deposits

A process for converting oxygenates to olefins using molecular screen catalysts containing a desired carbon residue.

Full description

Background of the invention

The present invention relates to methods for selectively converting oxygenates into light olefins, preferably ethylene and propylene, whereby desirable carbonaceous deposits are preserved in the total reaction volume of a catalyst by regenerating only a portion of the total reaction volume of the catalyst entirely and mixing the regenerated portion with the total reaction volume Non-catalytic regenerative.

Light olefins (known as ethylene, propylene, and butylene) are used as feed streams for the production of many chemicals. Light olefins are usually produced by the petroleum cracking process. Due to the limited supply and/or high cost of petroleum sources, the cost of producing olefins from petroleum sources is constantly increasing.

Oxygenates are considered alternative feedstocks for the production of light olefins, such as alcohols, especially methanol, dimethyl ether, and ethanol. Alcohols can be produced

By fermentation, or from synthesis gas derived from natural gas, petroleum liquids, carbonaceous materials, including coal, recycled plastics, municipal wastes, or any organic material . Because of the wide variety of sources, alcohol, alcohol derivatives, and other oxygenates have been provided as an economical, non-petroleum source for olefin production.

The catalysts used to enhance the conversion of oxygenates into olefins are molecular screen catalysts. Because ethylene and propylene are the most desirable products for such a reaction, research has focused on catalysts that are most selective toward ethylene and/or propylene, and on ways to increase the selectivity of catalysts from molecular sieves toward ethylene and/or propylene. It is known that the selectivity of certain molecular sieve catalysts towards ethylene and propylene increases if the level of coke in the total reaction volume of the molecular sieve catalyst is maintained in the range from about 2% by weight to about 30% by weight.

Some people have suggested maintaining this desired level of coke by removing all or part of the total reaction volume of the catalyst, partially regenerating the catalyst thus removed, and returning the partially regenerated catalyst to the reactor. However, partial regeneration may not produce maximum selectivity of the catalyst toward light olefins.

Methods are needed to maintain a desired level of coke in molecular screen catalysts during the conversion of oxygenates to olefins while maintaining maximum catalyst efficiency.

General description of the invention

The present invention provides a method for treating a catalyst from a molecular sieve comprising: contacting a feed stream containing oxygenates with a total reaction volume of a catalyst from a molecular sieve under effective conditions to produce a stream containing C2-C3 olefins, wherein said total reaction volume includes On a desired carbon residue, which makes the mentioned catalyst more selective towards C2-C3 olefins than in the absence of the desired carbon residue.

mentioned; Whereas, when undesirable carbon deposits accumulate that affect the effectiveness of the catalyst, the aforementioned desirable carbon deposits are preserved in the aforementioned molecular sieve catalyst through a process that includes: Separating the stated total reaction volume of the molecular sieve catalyst into a fraction and a remainder; Treating said part with a regeneration medium under effective conditions to remove said undesirable carbon deposits, forming a regenerated part containing from about 0% by weight to a regenerated amount of carbon deposits; and, mixing said regenerated portion with said remaining portion, wherein said regenerated portion of carbonaceous residue includes a sufficient amount, upon said mixing, to produce a total regenerated reaction volume containing said desired carbonic residue.

Brief explanation of fees:

Figure 1: A schematic diagram of a preferred embodiment of a high-speed fluid bed reactor

High velocity fluid bed reactor with recirculation catalyst for use in the present invention.

Detailed description

The conversion of oxygenates into light olefins is catalyzed by several molecular screen catalysts. Because of the high temperatures required during the conversion process, carbon deposits known as “coke” inevitably form on the surface of the molecular sieve catalyst. To avoid a significant reduction in catalyst efficiency, the catalyst should be regenerated by removing coke deposits by burning.

One of the goals during the conversion of oxygenates into olefins is to maximize the production of light olefins, preferably ethylene and propylene, and to reduce the production of methane, ethane, propane, and substances containing more than five C5+ carbon atoms. materials. The present invention uses coke, which must be deposited on the catalyst, to achieve this goal by allowing the desired carbon deposits to accumulate on the molecular screen catalyst while removing unwanted carbon deposits.

One method proposed to maintain the desired carbon deposits on the catalyst is to regenerate only some or all of the total reaction volume of the molecular sieve catalyst. Without limiting the present invention by special theory, it is believed that only partially regenerating part or all of the total reaction volume of a catalyst from a shuttle molecular screen has a serious drawback. It is known that coke produced during the conversion of oxygenates into olefins is deposited on the surface and in the “micropores” of molecular screen catalysts. Reactions that selectively convert oxygenates into ethylene and propylene occur in the micropores of the molecular sieve catalyst. It is relatively difficult for the regeneration medium (usually oxygen) to penetrate the fine pores. For this reason, it is more difficult to remove coke accumulated in the fine pores during the regeneration process. It is highly likely that partial regeneration does not remove coke from the fine pores of the catalyst, resulting in an adverse effect on the selectivity of the catalyst toward ethylene and propylene.

The present invention maintains desirable carbon deposits on the catalyst by removing only a portion of the total reaction volume of the shuttle molecular sieve catalyst and completely regenerating only that portion of the catalyst. It is believed that total regeneration removes coke from the finer pores and from the less selective surface areas of the regenerated catalytic fraction. When the regenerated portion of the catalyst is mixed with the remaining non-regenerated portion of the catalyst, the result is the preservation of desirable carbon residues that occupy less selective surface areas in the non-regenerated portion of the catalyst, and an increase in the available sites for selective conversion of oxygenates to light olefins (space microporous surface) in the regenerated part of the catalyst.

As used in this statement, the term “desirable carbon deposit” is defined as containing

To an amount of at least about 2 wt% carbonaceous residue, ranging in range from about 2% by weight to about 30% by weight carbonaceous residue, based on the weight of the total reaction volume of the shuttle catalyst. The “desired carbon deposits”—even if they contain more than 30% by weight of the total reaction volume of the molecular sieve catalyst—are carbon deposits that essentially clog parts of the catalyst surface that are not selective toward the production of C2-C3 olefins.

C2-C3 olefins.

Any catalysts of small or mesoporous molecular screens and their substantially equivalents may be used in the present invention. Microporous molecular screen catalysts are known as catalysts with a pore diameter of less than about 5 angstroms. Mesoporous molecular screen catalysts are defined as catalysts with a pore diameter in the range of about 5 to 10 angstroms. The term & equivalent is defined to refer to catalysts with a pore size that perform substantially the same function in substantially the same way to achieve a substantially similar result as catalysts with the diameter or pore size mentioned above.

One suitable group of molecular sieve catalysts is zeolite. There are many types of zeolites; Each shows different properties and uses. The structural types of small-pore zeolites suitable for use in the present invention include different levels of effectiveness, but are not necessarily limited to AFT, AEI, ATT, ATO, APC, CAS, BIK, AWW, ATV, EDI, DDR, DAC, CHI, CHA, ERI, GOO, LTA, LOV, LEV, KFI, MON, ROG, RHO, PHI, PAU, and THO, and examples of these structural types carry substitutions, as reported by W.M. Meier and D.H. Olsen DH Olseir, in a book entitled Atlas of Zeolite Structural Types (Buttervorth Heineman, 3rd ed., 1997), which is mentioned in this statement for reference. Preferred zeolite catalysts include, but are not necessarily limited to, 5-ZSM-34, ZSM, erionite, and chabazite.

Silicoaluminophosphates (SAPO) represent a group

Other molecular sieve catalysts useful in the invention. SAPO catalysts have a three-dimensional microporous crystal framework of tetralredral units of PO2+, AlO2, and SiO2. SAPO catalysts suitable for use in the invention include, but are not necessarily limited to, SAPO-17, SAPO-18, and 34-SAPO. 34-SAPO is considered a preferred SAPO catalyst that can be synthesized according to US Patent No. 4,440,871, which is cited herein by reference, and Zeolites Magazine, Volume 17, Page 512-522 (1996), which is cited herein by reference.

SAPO catalysts with added substitutes may also be useful in the present invention. SAPO catalysts bearing these substituents form a class of molecular screens known as “MeAPSO” compounds. Substitutes may include, but are not necessarily limited to, nickel, cobalt, strontium, barium, and calcium.

The structural types of mesoporous molecular sieves useful in the present invention,

It is not necessarily limited to FER, HEU, MTT, EUO, MTW, MEL, MFL, TON, AEL, AFO, and examples of these structural types have alternatives, as described in a book entitled Atlas of Zeolite Types, which was previously mentioned in this statement for reference. For reference.

The process of converting oxygenates into olefins uses an organic raw material (feedstock), which preferably contains “oxygenates”. As used in this statement, the term “oxygenates” is defined to include, but not necessarily be limited to, aliphatic alcohols, ethers, carbonyl compounds (aldehydes, ketones, carboxylic acids, carbonates, and the like). ), as well as compounds containing heteroatoms, such as halides, mercaptans, sulfides, amines, and mixtures thereof. It is preferable for the aliphatic moiety to contain carbon atoms in the range from about 1 to 1, and preferably in the range from about 1 to 4. Representative oxygenates include, but are not necessarily limited to, lower straight cliain or branched aliphatic alcohols, their corresponding unsaturated parts, and their nitrogenous, halogenated lralogeir and sulfur analogues. Examples of suitable vehicles include, but are not necessarily limited to: methanol; ethanol; p-propanol n-propanol; isopropanol; alcolrols C4-C10 C4-C10; methyl ethyl ether; Dimethyl ether &dimethyl ether; diethyl ether etlrer; di-isopropyl ether; methyl mercaptan; methyl sulfide; methyl amine; ethyl mercaptan; di-ethyl sulfide; di-ethyl airline; ethyl chloride; Formaldehyde

formaldehyde; di-methyl carbonate; di-methyl ketone; acetic acid leid; acetic; n-alkyl amines; p-alkyl halides 11-alkyl halides; N-alkyl sulfides contain 11-alkyl groups with about 3 to 10 carbon atoms; And mixtures thereof. As used in this statement, the term oxygenate refers only to the organic matter used as a feed stream. The total feed stream charged to the reaction zone may contain other compounds, such as diluents.

Preferably, the oxygenated feedstock in the vapor phase is brought into contact in a reaction zone with a molecular screen catalyst defined at effective process conditions for producing the desired olefins, i.e. effective conditions of temperature, pressure, weight hourly space velocity (WHSV) and Weight Hourly Space Velocity. and, optionally, an effective amount of diluent. Alternatively, the process may be conducted in a liquid phase or a vapor-liquid mixture. When the process is carried out in the liquid phase or in the mixed phase of vapor and liquid, different conversion rates and selectivity of the feedstock towards the product may result depending on the catalyst and reaction conditions.

The temperature used in the conversion process may vary over a wide range depending on

At least partly, on the chosen catalyst. Although it is not required at a specific temperature, the best results can be obtained if the process is performed at temperatures in the range from about 200 to 700 C (Celsius degree C), and preferably in the range from about 250 to

About 600 m, and most preferably in the range from about 300 m to about 500 m. And in a way

In general, low temperatures result in low reaction rates, and the formation of the desired light olefin products may be significantly slow. However, at high temperatures, the process may not form an optimal amount of light olefins products and the coking rate may become very high.

Light olefins, although not necessarily at optimal amounts, will form over a wide range of pressures, including but not limited to autogenous pressures and pressures in the range from about 0.1 kilopascal (kPa) to about 100 MPa (kPa). (megapascal (MPa). Pressures vary

Preferred in the range from about 6.9 kPa to about 34 MPa, and most preferable in the range from about 48 kPa to about 0.34 MPa. The above pressures exclude the diluent, if present, and are due to the partial pressure of the feedstock relating to oxygenates and/or mixtures thereof. Pressures outside the indicated ranges may be used and are not excluded from the scope of the invention. The lower and upper limits of pressure may adversely affect selectivity, conversion, coking rate, and/or reaction speed; However, it may still be possible to form light olefins such as ethylene.

The process should continue for a sufficient period of time to produce the desired olefin products. The reaction time may vary from tenths of a second to several hours. The reaction time is determined to a large extent by the reaction temperature, pressure, catalyst chosen, gravimetric velocity per hour, the phase in which the process is conducted (liquid or vapor), and the design features chosen for the process.

A wide range of gravimetric hourly spatial velocity (WHSV) of the feedstock can be used in the present invention. WHSV is defined as the weight of the feed stream (excluding diluent) per hour per weight of total reaction volume of molecular sieve catalyst (excluding inerts and/or fillers). In general, WHSV should be in the range from about 0.01 hour-1 (hr-1) to about 500 hour-1, varying in the range from

About 0.5 hours-1 to about 300 hours, the most preferable in the range from

About 0.1 h to about 200 h-1. One or more diluents may be fed to the reaction zone with oxygenates, such that the total feed mixture includes an amount of diluent in the range from about 1 mol% to about 99 mol%. Diluents that may be used in the process include, but are not necessarily limited to, helium, argon, nitrogen, carbon monoxide, carbon dioxide, hydrogen, water, paraffins, and other hydrocarbons (e.g. Methane), aromatic compounds, and mixtures thereof. Preferred diluents include water and nitrogen

A preferred embodiment of a reaction system of the present invention is a circulating fluid bed reactor with continuous regeneration, analogous to a modern fluid catalytic cracker. Fixed beds are considered impractical for the process because converting oxygenates to olefins is a highly exothermic process, requiring multiple stages with intercoolers or other cooling devices. The reaction also leads to a high pressure drop due to the production of low-pressure, low-density gas.

Because the catalyst must be regenerated continuously, the reactor should allow for easy removal of part of the catalyst to the regenerator, exposing the catalyst to a regeneration medium, preferably a gas containing oxygen, most preferably air, to remove the coke from the catalyst by burning, resulting in To restore the effectiveness of the catalyst. The conditions of temperature, partial pressure of oxygen, and residence time in the regenerator should be chosen to obtain a coke content in the regenerated catalyst of less than about 0.5% by weight. At least part of the regenerated catalyst should be returned to the reactor.

It is necessary to design the reactor so as to maintain a relatively high average level of coke in the reactor—at an amount greater than about 1.5% by weight, preferably in the range from about 2% by weight to about 30% by weight, and most preferably in the range from about 2% by weight

To about 20% by weight. If the reactor is a high-speed fluidized bed reactor (sometimes referred to as a riser reactor), a portion of the catalyst that exits the top of the reactor should be returned to the reactor inlet. This differs from a typical fluid catalytic cracker (FCC) riser reactor (fcc) in which all or most of the catalyst exiting the top of the reactor is sent to the regenerator. Recycling the shuttle catalyst directly into the reactor, without renewing the shuttle catalyst, allows the average level of coke in the reactor to be increased to a preferred level. By adjusting the flow rate of the shuttle catalyst between the regenerator and the reactor, a preferred level of coking, or carbon deposits, can be maintained.

Desirable&.

If the fluidized bed reactor is designed with low gas velocities, less than about 2 meters/second (mSec), cyclones can be used to return the catalyst particles to the fluidized bed reaction zone. Such reactors generally have high solids recycling rates Within the fluidized bed, which allows the coke level on the catalyst to be increased to a preferred level. The desired carbon deposit is maintained by withdrawing the catalyst from the bed, regenerating the catalyst in the manner described above, and then returning at least a portion of this regenerated catalyst to the reactor.

In general, a preferred embodiment of a riser reactor used in the present invention is shown in Figure 1, referred to as 10. The methanol feed stream 12 is at least partially evaporated in a preheater (not shown). The methanol feed stream is mixed with regenerative catalyst 8 2 and shuttle catalyst 2 2 at the bottom of the riser reactor 14. An inert gas and/or water vapor can be used to dilute the methanol, raise catalytic streams 22 and 28, and keep the pressure equipment lines free of catalyst. This inert gas and/or water vapor is mixed with methanol and the catalyst in reactor 41. The reaction is exothermic, and the preferred reaction temperature, which is in the range from about 300°C to about 500°C, is maintained by removing heat. The heat can be removed by any suitable means, but does not have to be confined, cooling the reactor with a catalyst cooler (not shown), feeding a certain amount of methanol in liquid form, cooling the catalyst feed stream to the reactor, or any combination of these methods. .

The effluent from reactor 16, which contains products, shuttle catalyst, diluents and undenatured feed compounds, should flow into disengaging zone 18. In separation zone 18, the shuttle catalyst separates the gaseous materials by gravity and/or separators Vortex. A portion of the shuttle catalyst 2 2 is returned to the reactor inlet. The portion of the shuttle catalyst 22 that should be regenerated is first sent to stripping zone 9 2, where water vapor or an inert gas is used to extract adsorbed hydrocarbons from the catalyst. The stripped spent shuttle catalyst 23 should flow to the regenerator 24. The portion of the catalyst sent to the regenerator 24 should be brought into contact with a regeneration medium, preferably

A gas containing oxygen 0 3, at temperatures, pressures, and residence times capable of removing coke from the catalyst by burning and reducing its concentration to a level of less than about 0.5% by weight. The preferred temperature in the regenerator ranges from about 0.55°C to about 700°C, and the preferred oxygen concentration in the gas discharged from the regenerator ranges from about 0.1% by volume to about 5% by volume, and the preferred residence time in the range ranges from about a minute. One to about 100 minutes.

The process of removing coke by burning is an exothermic process. The temperature may be maintained at a suitable level using any acceptable method, including but not limited to feeding coolant, cooling the catalyst in the regenerator using a catalytic cooler 26, or a combination of these methods.

The regenerated catalyst 8 2 is sent to reactor 4 1, where it is mixed with the recycled shuttle catalyst 2 2 and the methanol feed stream 12. The regenerated catalyst 8 2 can be lifted into reactor 14 via inert gas, water vapor or methanol vapor (not shown in the drawing). The process should be repeated continuously or almost continuously. The hot gases resulting from the reactor should be cooled, the by-product water is condensed and collected, and the desired resulting olefin gases are extracted for further processing.

To determine the level of coke in the reactor and regenerator, small samples can be withdrawn periodically from different points in the recycling system to measure the carbon content. Accordingly, the interaction variables can be adjusted.

The following examples illustrate but do not limit the present invention.

Example 1

Charge the continuous circulation fluidized bed reactor with 3200 g of catalyst, spray dried using a mixture of 34-SAPO powder (obtained from Universal for oily products (UOP, Des Plaines, IL) with materials Binders of alumina and clay whose average particle size ranges from 90 to 100 microns. In three different tests, methanol was charged at a rate of 900 grams/hour and evaporated in a preheater and mixed with the evaporator feed stream

Catalyst capacity ranging from 20,000 to 25,000 g/hour, enriched into a reactor with an internal diameter of 1.02 centimeter (cm) (0.4 inch) and a length of 6.71 m (22 feet)> and used about 268.21 liters/hour (7 standard feet per hour scf/hr) of nitrogen to raise the catalyst, keep the pressure tools free of catalyst particles, and mix the nitrogen with methanol and the catalyst in the reactor, and the temperature in the reactor was maintained at 450°C using electric heaters. The effluent from the reactor flowed to the stripper, where the catalyst was removed from the resulting gas. The catalyst is contacted with nitrogen at the bottom of the catalyst to extract volatile hydrocarbons from the catalyst. The stripped catalyst was sent to a regenerator, where it came into contact with a mixture of nitrogen and air. The temperature in the regenerator was maintained at 620°C using electric heaters, and the air rate may be varied to adjust the coke level on the regenerator catalyst. The catalyst was returned from the regenerator to the reactor, where it was mixed with the methanol feed stream. The process itself was repeated continuously. The hot reactor product gases were cooled, and the water byproduct was condensed and collected. Hydrocarbon gases were separated from the water and analyzed by gas chromatography (GC). The flue gas from the regenerator was analyzed to determine the content of oxygen, carbon monoxide, and carbon dioxide. The rate was measured using a dry test meter. It was withdrawn. Small samples of the catalyst from both the sifter and the strands are periodically measured to measure the carbon content. Based on these measurements, the production rate of products, including coke, is calculated.

In Test 1, the air rate was adjusted so that nearly all of the carbon content on the catalyst was removed during each pass through the regenerator. The carbon content of the catalyst discharged from the reactor was by weight, and the regenerator removed all of this carbon except for 0.2% by weight. The selectivity towards ethylene reached 10.8% by weight, the selectivity towards heavy products was 3409% by weight, and towards coke 14.3% by weight.

In Test 2, the rate of air flowing to the regenerator was reduced so that each catalyst was only partially regenerated in each pass. The carbon content on the rotating catalyst increased, until a steady state was finally reached such that carbon was removed at the same rate at which it was deposited. And at

At this point, the carbon content of the catalyst flowing to the regenerator was 5.5%, and the carbon content of the catalyst draining from the regenerator was 4.9%. Selectivity towards ethylene improved to 26.7%, and selectivity towards undesired heavy products decreased to 17.9%. The coke production rate remained relatively constant at 13.6% and the methanol conversion reached 91.3%, which indicates a decrease in the effectiveness of the catalyst due to the coke deposited on it.

In Test 3, the methanol feed stream was stopped and the rotary catalyst was left to regenerate completely (until the carbon content reached 0.15% by weight). Then the air flow to the regenerator was stopped, the methanol feed stream was resupplied, and the coke was left to accumulate on the catalyst without renewal for approximately 5 hours. The carbon content on the rotating catalyst after 5 hours reached 5.8%. At this point, the mobility toward ethylene improved to 35%, the mobility toward heavy products decreased further to 13.4%, and the mobility toward coke decreased to 4.1%. The coke production rate was calculated from measurements of carbon accumulation on the catalyst during the last hour of startup. The conversion at this point was 89.9%, indicating that the catalyst was approximately as effective as the catalyst used in Test 2.

After Test 3 was conducted, the reactor was returned to the same operating conditions used in Test 1 and then Test 2. Methanol conversion ratios and product production rates were essentially the same as the original production rates in Tests 1 and 2 after a total of 150 hours of continuous operation, indicating that the results were not affected simply by suspending the catalyst.

Column 4 represents product selectivity calculated for a commercial reactor using the present invention,

Based on the data obtained from tests 1 and 3. Column 4 assumes that 10% of the methanol is converted on a newly regenerated catalyst (selectivities according to Test 1), and the remaining 90% of the methanol is converted on a shuttle catalyst (selectivities according to Test 3). The calculated selectivities were slightly worse than the results of Test 3, but still significantly better than using only a partially regenerated catalyst as in Test 2:

<img file="SA1197B1_D0001.tif" />

''Selectivities&amp;* were calculated based on the absence of water.

Previous results showed that shuttle catalysts (containing approximately 5% of coke on the catalyst) achieved higher selectivities towards ethylene and propylene than catalysts containing less than 1% of coke. The results also showed that the completely renewed catalyst that was left to accumulate coke achieved higher selectivities than the partially renewed catalyst to reduce the coke content to the same level. These results are consistent with the theory that partial regeneration of the catalyst selectively removes coke, which hinders undesired surface reactions that lead to the formation of propane and materials containing more than 5 carbon atoms. The interactions that occur

It leads to the formation of ethylene and propylene selectively in small pores, and removing coke accumulated in these pores is more difficult than removing coke on the outer surface areas (macropore). When coking of the 'pure' catalyst is possible, coke deposits on the macroporous surfaces faster than the fine pores, and this slows down the non-selective surface reactions, thereby improving the selectivity of the 'coke' catalyst compared to the pure catalyst. However, when more coke accumulates, the catalyst becomes unselective. Finally effective. Efficiency is restored by regenerating the catalyst with air, but it is important to burn the catalyst completely to remove as much carbon as possible, and then allow it to decompose again in the reactor. Partial regeneration of the catalyst, as described in US Patent No. 4,873,390 under Lewis's name, was not nearly as effective in maintaining etliylene and propylene selectivity in the reactor.

Based on what was mentioned above, it was concluded that better selectivity towards light olefins can be achieved in the process of converting oxygenates to olefins if a desired amount of coke is maintained over the total reaction volume of the molecular sieve catalyst by completely regenerating only part of the catalyst. Returning at least a portion of the regenerated portion to the total reaction volume.

Those familiar with the technology will realize that it is possible to make several modifications to this invention without deviating from the principle and scope of the current invention. The embodiment described herein is described by way of illustration only and shall not limit the invention defined in the following claims.