Patent Application
20140296595
The present invention relates to direct coal liquefaction processes for efficiently producing high-value aromatics from coal.
Current and presently proposed methods for the production of high value aromatics (benzene, toluene, and xylene, i.e. BTX) from coal include either the capture of gaseous byproducts from the pyrolysis of coal in an airless environment, or the proposed conversion of coal to methanol followed by the conversion of the methanol to aromatics.
The pyrolysis method is of economic interest only as a byproduct of another process, e.g., producing coke from coal. The methanol to aromatics (MTA) process would involve the gasification of the coal feed, typically by partial oxidation (PDX), to produce syngas as the feed for methanol synthesis, converts coal to methanol and then converts methanol to aromatics by processing over a fixed or fluid bed of catalyst. A published estimate stated that the required investment for a 1 million metric ton per stream day MTA plant in China would be about $4.6 billion. This high cost results in major part because of the need to use expensive PDX for gasifying all of the coal. Available publications indicate that about 3 MT of coal would be required to produce 1 MT of methanol. A separate report states that 1.42 to 1.59 MT of methanol would be required to produce 1 MT of aromatics. If these two numbers are combined, it requires 4.26 to 4.77 MT of coal to produce 1 MT of aromatics. It is estimated that the thermal efficiency of the MTA process is about 40 to 45%.
In accordance with the invention, a highly efficient and lower cost method and system for producing high-value aromatics from coal is provided in which the feed coal is converted by direct coal liquefaction (DCL) to a 1000° F.−, preferably an 800° F.−, more preferably a 600-750° F.− product, most preferably a 600-700° F.− product, at least the 350° F.+ portion of which is then hydrocracked to produce a 350° F.− product stream. The DCL and hydrocracked 350° F.− product streams are then hydroprocessed to remove sulfur, nitrogen and oxygen compounds and fractionated into approximately 160° F.− and 160° F.+ output streams. The approximately 160/350° F. output stream is ideal for aromatics production. It typically contains 85 to 90% naphthenes, 5 to 10% paraffins, plus some single ring aromatics, and the naphthenes are easily converted into BTX in a catalytic reformer. The product from the catalytic reformer can be processed in a solvent extraction unit to produce a pure aromatics product and a paraffinic raffinate. Alternatively, after converting the naphthenes and heavier paraffins to aromatics in the catalytic reformer, the entire product can be passed over a bed of cracking catalyst in the same unit. This bed isomerizes lower value ethyl benzene to more valuable para-xylene and benzene, and cracks the remaining paraffins into an approximately 160° F.− product that can be separated from the aromatics in the distillation tower, thereby producing a higher value aromatic product and eliminating the need for solvent extraction.
The preferred DCL system includes a slurry DCL reactor containing a molybdenum or iron, preferably molybdenum, microcatalyst and is operated at high conversion with the product boiling above the 600-700° F. range preferably being recycled and mixed with the DCL feed coal as a non-donor stream in a ratio of non-donor stream to coal at the input to the reactor (on a moisture free weight basis) of between 1.6 and 3.5:1. By “non-donor” is meant that the recycle stream has not been processed in a hydrotreater to partially hydrogenate multi-ring aromatic compounds in the stream in order to produce compounds that can donate hydrogen during liquefaction.
In order to provide the additional hydrogen required for the DCL and hydrocracking processes, the bottoms from the DCL reactor can be gasified in a PDX reactor. As an alternative, the additional hydrogen may be provided by processing 160° F.− product of the upgrading and the catalytic reformer by liquid PDX or steam naphtha reforming (SNR). If natural gas is available, it can be used as the feed to the liquid PDX or SNR instead of the 160° F.− product.
Referring now to
The approximately 160/350° F. stream 119 from the fractionator 115 is typically made up of 85% to 90% naphthenes, 5 to 10% paraffins, and some single ring aromatics. This stream is fed to the catalytic reformer 121 where the naphthenes are converted into aromatics. The output of the catalytic reformer 121 is fed to the solvent extractor unit 123 where it is separated into a pure aromatics stream 125 and a lower boiling point paraffinic raffinate stream 127. The aromatics stream 125 can be separated into its benzene, toluene and xylene components by distillation.
As attractive alternative to the use of the extraction unit 123 is that described in U.S. Pat. No. 5,472,593 in which, after converting the naphthenes and heavier paraffins to aromatics in the catalytic reformer, the entire product is passed over a bed of catalyst in the same unit comprising a medium-pore molecular sieve having a pore size of from 5 to 6.5 .ANG., a refractory inorganic oxide, a platinum-group metal component and a lead or bismuth metal attenuator. This bed converts the lower value ethyl benzene to more valuable para-xylene and benzene, and cracks the remaining paraffins into a 160° F.− product that can be separated from the aromatics by distillation, thereby producing a higher value aromatic product and eliminating the need for solvent extraction. This alternative does, however, require additional hydrogen for cracking. The disclosure of U.S. Pat. No. 5,472,593 is hereby incorporated by reference in its entirety.
CO2 produced by the PDX unit 111, and optionally, by the DCL reactor system 103 and/or other components of the liquefaction and upgrading system, is fed to the algae production system 129, which includes a photo-bio reactor (PBR) in which the CO2 is used to produce preferably blue-green algae through photosynthesis. The DCL reactor system 103 and especially the upgrading system 109 also produce NH3, which can be fed to the algae production system 129 as a nutrient. The algae from the algae production system 129 is preferably used to produce a biofertilizer 131. Methods and apparatus suitable for use in the present invention for producing algae and biofertilizer are disclosed in U.S. patent application Ser. No. 13/316,546 that was filed on Dec. 11, 2011, the disclosure of which is hereby Incorporated by reference in its entirety.
Referring now to the embodiment of a DCL system illustrated in
Most of the remaining moisture in the coal is driven off in the mixing tank due to the hot atmospheric fractionator bottoms feeding to the mixing tanks. Residual moisture and any entrained volatiles are condensed out as sour water (not shown in
The coal slurry and hydrogen mixture is fed to the input of the first stage of the series-connected liquefaction reactors 209, 211 and 213 at between 600 to 700° F. (316 to 371° C.) and 2,000 to 3,000 psig (138 to 206 kg/cm2 g). The reactors 209, 211 and 213 are simple up-flow tubular vessels, the total length of the three reactors being 40 to 200 feet. The temperature rises from one reactor stage to the next as a result of the highly exothermic coal liquefaction reactions. In order to maintain the maximum temperature in each stage below about 800 to 900° F. (427 to 482° C.), a portion of the hydrogen based treat gas is preferably injected between reactor stages. The hydrogen partial pressure in each stage is preferably maintained at a minimum of about 1,000 to 2,000 psig (69 to 138 kg/cm2 g).
The effluent from the last stage of liquefaction reactor is separated into a gas stream and a liquid/solid stream, and the liquid/solid stream let down in pressure, in the separation and cooling system 215. The gas stream is cooled to condense out the liquid vapors of H2O, naphtha, distillate, and solvent. The remaining gas is then processed to remove H2S, NH3 and CO2
Most of the processed gas is then sent to a hydrogen recovery system, not shown, for further processing by conventional means to recover the hydrogen contained therein, which is then recycled to be mixed with the coal slurry. The remaining portion of the processed gas is purged to prevent buildup of light ends in the recycle loop. Hydrogen recovered therefrom can be used in the downstream hydro-processing upgrading system.
The depressurized liquid/solid stream and the hydrocarbons condensed during the gas cooling are sent to the atmospheric fractionator 219 where they are separated into light ends and in the preferred embodiment, a 600 to 700° F.− fraction, and a 600 to 700° F.+ fraction. The light ends are processed to recover hydrogen and C1-C2 hydrocarbons that can be used for fuel gas and other purposes. The 600 to 700° F.− fraction is sent to upgrading for aromatics production. Alternatively, the fractionator 219 could be arranged to produce 1000° F.− and 1000° F.+ fractions or 800° F.− and 800° F.+ fractions in which the 1000° F.− or 800° F.− fraction would be sent to the upgrading step.
In the preferred embodiment, a portion of the 600 to 700° F.+ (316 to 371° C.+) is recycled to the slurry mix tank. The remaining 600 to 700° F.+ fraction produced from the atmospheric fractionator 219 is fed to the vacuum fractionator 221 wherein it is separated into a 1000° F.− fraction and a 1000° F.+ fraction. The 1000° F.− fraction is added to the 600 to 700° F.+ stream being recycled to the slurry mix tank 203.
In the preferred embodiment illustrated in
Catalysts useful in DCL processes also include those disclosed in U.S. Pat. Nos. 4,077,867, 4,196,072 and 4,561,964, the disclosures of which are hereby incorporated by reference in their entirety. Other DCL reactor systems suitable for use in the process of the invention are disclosed in U.S. Pat. Nos. 4,485,008, 4,637,870, 5,200,063, 5,338,441, and 5,389,230, and U.S. patent application Ser. No. 13/657,087,the disclosures of which are hereby incorporated by reference in their entirety.
The preferred DCL Process combines several elements that contribute to maximum BTX Product production and maximum thermal efficiency. These include, very importantly, the recycle of a non-donor 600 to 700° F.+ stream, preferably including atmospheric fractionator bottoms, to maintain a ratio of the recycle stream to coal at the input to the reactors 209, 211, 213 that is between 1.6:1 and 3.5:1 on a moisture free weight basis; the use of a microcatalyst in the form of finely divided molybdenum; and the use of a much lower treat gas rate than in previous systems. Also, the use of bottoms recycle, and multiple slurry reactors in series contribute to the benefits of the process.
Use of a microcatalyst, which is either a compound of molybdenum or iron, more preferably molybdenum, and added at 100 to 1,000 wppm, more preferably 100 to 500 wppm, and most preferably 100 to 300 wppm, eliminates several disadvantages to the use of a donor solvent such as required by prior DCL systems. First, energy is lost during preparation of the donor solvent. Second, energy is required to preheat the donor solvent in the solvent hydrotreater and hydrogen must be compressed and circulated around the hydrotreater. Thirdly, the heat release during partial hydrogenation of the donor solvent is lost during cooling prior to separation of hydrogen for recycle. In comparison, all of the heat release occurs in the liquefaction reactors during operation with a 600 to 700° F.+ recycle stream, which minimizes the preheat requirement prior to liquefaction. These factors contribute to the higher thermal efficiency of the microcatalytic coal liquefaction process. Moreover, the use of a microcatalyst and the consequent elimination of the need for a donor solvent also eliminates the need for an expensive solvent hydrotreater to generate the donor solvent, thereby substantially reducing the capital and operating cost of the system. It also permits the use of coals having substantially higher ash contents, from 6 to 20 wt % or more on a moisture free basis, and the recycle of a substantially higher portion of bottoms than were possible with donor solvent systems. Examples of microcatalysts and their method of preparation are described in U.S. Pat. No. 4,226,742, the contents of which are hereby incorporated by reference in their entirety.
The 600 to 700° F.+ fraction recycled from the atmospheric fractionator 219 and the 1000° F.− fraction from the vacuum fractionator 221 as the non-donor stream being recycled to the slurry mix tank 203 provides preheat for the coal and solvent in the slurry mix tank 203. This raises the temperature in the mix tank to 300° F. to 500° F., more preferably 350° F. to 500° F., and most preferably about 400 to 500° F. This further reduces the energy requirement for preheating the slurry prior to liquefaction. A significant portion of the of the microcatalyst is entrained in the 600 to 700° F.+ fraction recycled from the atmospheric tower 219, so that recycling a larger portion of such fraction increases the catalyst concentration in the DCL reactors 209, 211, 213, thereby decreasing the requirement for the addition of fresh catalyst precursor and increasing the conversion efficiency of the DCL process.
Use of the non-donor 600° F. to 700° F.+ stream, more preferably 630° F. to 670° F+, and most preferably a 650° F+, process derived recycle solvent in the DCL process reduces cracking, relative to donor solvent, and produces a 650° F.− product with a greater fraction of diesel and less light gases and naphtha. The 650° F.− product can be selectively upgraded to finished products in fixed bed upgrading reactors.
The much lower treat gas rate of 600 to 900 NL per kg of slurry has a significant impact on thermal efficiency, plant investment, and operating cost. The required recycle treat gas rate for the DCL process of the invention is up to three times lower than the preferred gas rate in the NEDOL program (without taking into account the treat gas rate to the solvent hydrotreater, which makes the difference even larger). This has an important impact on power requirements for the compressor and fuel requirements for slurry preheat furnace 207 and solvent hydrotreater preheat.
The use of two to four, more preferably three slurry reactors in series approaches a plug flow reactor and hence has as little as two thirds of the required volume of one or two ebullated bed reactors such as used in some prior DCL systems. Since all of the heat is released in the three liquefaction reactors, the temperature profile can be also maintained to maximize selectivity to liquids. Operation of the initial reactor at a somewhat lower temperature has been reported in previous patents as a route to increase conversion and liquid yields.
An exemplary process for upgrading the liquid product of the DCL reactors 209, 211, 213 is disclosed in U.S. Pat. No. 5,198,099, the disclosure of which is hereby incorporated by reference in its entirety. Other processes and systems suitable for upgrading the liquid products are commercially available from vendors such as UOP, Axens, Criterion and others.
Referring now to
Referring now to
The embodiment of the invention illustrated in
The embodiment of the invention illustrated in
The embodiment of
| Number | Date | Country | |
|---|---|---|---|
| 61807423 | Apr 2013 | US |
