WO2014151903A1

WO2014151903A1 - Liquid fuel production process and apparatus employing direct and indirect coal liquefaction

Info

Publication number: WO2014151903A1
Application number: PCT/US2014/026643
Authority: WO
Inventors: Richard F. Bauman; Kenneth Lee TRACHTE
Original assignee: Accelergy Corporation
Priority date: 2013-03-14
Filing date: 2014-03-13
Publication date: 2014-09-25
Also published as: US20140262965A1

Abstract

A method and apparatus for producing liquids from coal including diesel and jet fuel blends and BTX in which the coal feed is converted to liquids in a DCL reactor, the produced liquids are upgraded and separated into, naphtha and DCL jet or diesel blend stock streams, the C3/180-350°F stream is converted in a Fischer Tropsch reactor to produce a highly paraffinic diesel or jet fuel blend stock and the DCL and Fischer Tropsch blend stocks are blended in controlled ratios to produce premium diesel or jet fuels meeting applicable specifications.

Description

Liquid Fuel Production Process and Apparatus Employing Direct and Indirect Coal Liquefaction Field of the Invention

The present invention relates to a combination of direct and indirect coal liquefaction processes and apparatus for maximizing production of fully synthetic jet fuel and /or diesel fuel meeting applicable specifications from coal.

Background

European specifications for diesel fuel (Euro 4) require a high Cetane number (CN) (minimum51) with a low specific gravity (0.845 maximum). This fuel can be readily produced from biomass, natural gas or coal via Fischer-Tropsch (F-T) synthesis. The syngas for the F-T synthesis from coal is produced via gasification of the coal or unconverted product from pyrolysis, hydropyrolysis, or DCL. These approaches produce a diesel fuel that is very paraffinic with high CN and low specific gravity. Such diesel fuels, however, have relatively low energy content per gallon, and thus provide lower mileage per gallon when used as a transportation fuel. Additionally, in order to produce a diesel fuel having acceptable cold flow properties for use in many parts of the world, the F-T product needs to be a hydroisomerized. Other important disadvantages of F-T processes for producing diesel fuel from coal are that they have low thermal efficiency and relatively high C02 emissions.

Direct coal liquefaction (DCL) methods have been developed for liquefying coal that have advantages in many applications relative to coal conversion by FT synthesis, including substantially higher thermal efficiency and lower C0₂ emissions. Diesel and jet fuel from DCL have excellent cold flow properties and relatively high energy content as a result of being rich in naphthenes (typically 85+%). However, the specific gravity (0.86 to 0.90) of the diesel fuel produced is significantly above the Euro 4 specification and CN is lower than the Euro 4 specification, even after severe hydrotreating to remove heteroatoms and hydrogenate aromatics to naphthenes. Selectivity to diesel has been also generally lower for DCL than for Fischer-Tropsch plants. The specific gravity of the DCL produced jet fuel is also typically higher than that permitted by the JP-A or JP-8 product specifications.

It has also been proposed to combine F-T and DCL units in a single plant to produce a blended product. Typically, the DCL and F-T processes have been described as being operated completely or partially in parallel, each receiving a coal feed, with the DCL and F-T products being blended to produce the final liquid fuel product. The syngas feed to the F-T process is produced by gasifying either feed coal or bottoms from the DCL process or both. The syngas produced for Fischer-Tropsch in this fashion is rich in carbon monoxide (typically 1:2 hydrogen to carbon monoxide molecular ratio). In order to obtain syngas with the desired 2:1 hydrogen to carbon monoxide ratio, part of feed must be used to generate additional hydrogen prior to F- T synthesis, or the feed must be converted in an iron catalyzed F-T reactor that promotes the water-gas shift reaction. This F-T process typically has a thermal efficiency of 40 to 50% and requires cleanup of the syngas for removal of fines, sulfur, and nitrogen compounds before being used as a feed to the F-T process.

Summary of the Invention

In accordance with the present invention, there has been provided a combined DCL and cobalt catalyzed F-T coal conversion process and system having high thermal efficiency and high conversion that efficiently produces jet and/or diesel fuel, and that is capable of accepting less expensive coals containing high ash (from 6 to 30% on a moisture free basis) and high inertinite content (greater than 10% on a moisture free basis). The present invention overcomes the above described and other problems with prior hybrid DCL and F-T processes and systems by using the C3/180-400°F, preferably the C3/180-350°F, boiling point fraction, i.e. having boiling point of in the range of between 180 and 350 or 400°F (hereinafter referred to as the C3+ fraction) produced by the DCL reactor as the feed to a partial oxidation (POX) reactor, or more preferably to an Steam Naphtha Reformer (SNR) for production of syngas, as feed to a preferably cobalt catalyzed F-T reactor to produce a highly paraffinic, low specific gravity F-T blend stock. The DCL produced C3+ feed stock to the SNR/F-T unit is free of fines and difficult to remove heteroatoms, and its use allows the direct production of a 2:1 ratio H2/CO syngas that is ideal for cobalt catalyzed hydrocarbon synthesis in an F-T plant. Its use also eliminates the need for the expensive and inefficient production of syngas for F-T via gasification of coal and/or DCL bottoms. Moreover, a cobalt catalyzed F-T plant based on a hydrocarbon feed has a substantially higher thermal efficiency (60+%) than an iron catalyzed F-T plant using a coal feed. The naphtha and lighter product of the F-T plant is preferably used to produce additional syngas, and thereby recycled to extinction.

The preferred DCL system includes a slurry DCL reactor containing molybdenum or iron, preferably molybdenum, microcatalyst and is operated at high conversion with the product boiling above the jet fuel or the diesel fuel boiling range (typically 650F+ or 700F+)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 greater than 1.6:1, preferably greater than 1.7:1, more preferably between 1.8:1 and 3.5:1, still more preferably between 2.0:1 and 3.5:1, and most preferably between 2.0:1 and 3.0:1. 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.

Because the viscosity and density of the coal slurry in the DCL reactor is much higher than is the case with a slurry comprising a hydrotreated donor solvent, the maintenance of a stable slurry and the settling of ash in the reactor is not a concern. Therefore the present process operates with a much lower gas hold-up than is required in prior hydrotreated donor solvent DCL systems. As a result of the low gas hold-up and the preferred use of slurry reactors in series, which operate in a substantially plug flow mode, the reactor volume can be significantly less than that required for a DCL system having the same output capacity operating with a donor solvent and high gas hold-up. The jet or diesel fuel product from the DCL reactor is rich in naphthenes and typically will have a specific gravity higher than that required by the applicable jet or diesel fuel specification.

In accordance with the invention, and DCL jet or diesel fuel product is blended with the low specific gravity F-T blend stock to produce a jet or diesel fuel meeting the required product specifications for the particular application, e.g., jet fuel meeting JP-A or JP-8 product specifications or diesel fuel meeting the Euro 4 CN and specific gravity specifications.

In accordance with an alternative embodiment of the invention, instead of being used to produce gasoline blend stock, the intermediate naphtha product from the DCL reactor boiling in the range of 160/250°F to 160/300°F, which is rich in naphthenes (typically 85+%), constitutes an ideal feed stock for the production of aromatics (benzene, toluene, and xylene, i.e. BTX) for chemicals production.

Description of the Drawings

Figure 1 is a schematic of the overall flow scheme of the illustrated embodiment of the invention.

Figure 2 is a schematic diagram of the flow scheme for the direct coal liquefaction portion of the illustrated embodiment of the invention.

Detailed Description of Preferred Embodiments

Figure 1 of the drawings, is a schematic of the overall flow scheme of the illustrated embodiment of the hybrid DCL/F-T plant 100 of the invention. The coal feed 101 is supplied to the DCL unit 103 that is preferably operated at a high conversion of 80+% on a moisture and ash free (MAF) basis. Optionally, additional coal can be supplied to a hydrogen source 105 for generating hydrogen for the DCL unit 103. In the DCL unit 103, coal is hydrogenated to produce raw liquid products and an effluent stream 107 that consists of ash, unconverted coal, and liquids boiling above 1000°F. The raw liquid products from the DCL unit 103 flow to the upgrader 109. In upgrading, heteroatoms are removed and a major portion of the aromatics present in the raw liquids are converted into naphthenes. Light gases produced in DCL and upgrading may be used to supply a portion of the fuel for the plant. Excess light gases (C2-) may be sent to a steam methane reformer to supply a portion of the hydrogen required by the DCL unit 103. The bottoms stream 107 from the DCL unit 103 is sent to the bottoms disposal unit 111 which may comprise a partial oxidation (POX) gasification for supplying additional hydrogen to the DCL unit 103 or to a Circulating Fluidized Bed (CFB) boiler to produce additional heat and power for the plant.

The liquid stream from the upgrader 109 is sent to the atmospheric fractionator 113 in which it is separated into three output streams; a nominally C3+ stream 115, a naphtha stream 117 and a DCL diesel or jet fuel blend stock stream 119. In accordance with the invention, the C3+ stream 115 is fed to a POX or SNR unit included in unit 121 in which it is converted into syngas, which is converted into a paraffinic hydrocarbon stream via F-T synthesis in unit 121. The amount of the naptha stream 117 being produced will depend on the upper bound of the C3+ fraction and the lower bound of the diesel or jet fuel blend stock stream 119, and in the limit may be zero. The thermal efficiency of the cobalt catalyzed F-T synthesis will be approximately 60 to 62% and the C3+ feed stock is free of fines and difficult to remove heteroatoms. If necessary, the C3+ stream 115 may be further hydrotreated before being fed to unit 121 to remove any remaining trace contaminants that would poison the cobalt F-T catalyst. The output from the F-T unit 121 is fed to the atmospheric fractionator 123 in which it is separated into a naphtha and lighter stream 125 and jet or diesel fuel blend stock (typically 250/500°F or 250/700F, respectively) stream 127. The naphtha and lighter portion (typically 250°F-) 125 is recycled to the input of the F-T unit 121 for being recycled to extinction.

The F-T jet or diesel fuel blend stock stream 127 is then blended with the DCL jet or diesel fuel blend stock stream in a ratio required to produce a jet or diesel fuel that meets the applicable specifications. If the desired product of the plant is a diesel fuel meeting the Euro 4 CN and specific gravity specifications, the ratio of DCL to F-T blend stocks needs to be in the range of about 3 to 1 to 7 to 1. If it is not necessary to meet the Euro 4 specific gravity specification for a particular application, a higher CN can be achieved by adding a

conventional cetane improver, such as the Lubrizol 8090 Cetane Improver, to the blended DCL/F-T diesel fuel. This will also facilitate use of a higher ratio of DCL to F-T blend stocks.

If the desired product of the plant is a jet fuel meeting the JP-A or JP-8 specifications, the ratio of DCL to F-T blend stock needs to be in the range of about 4 to 1 to 15 to 1. If it is beneficial for a particular application, is also possible to provide an additional natural gas feed to the F-T synthesis unit 121 to provide additional F-T blend stock.

DCL

Referring now to the embodiment of a DCL system illustrated in Fig. 2 of the drawings, the coal feed is dried and crushed in a conventional gas swept roller mill 201 to a moisture content of 1 to 4 %. Crushed and dried coal is fed into a mixing tank 203 where it is mixed with a non-donor recycle stream that is preferably constituted by a 600 to 700°F+ recycle stream 204 of the output of the liquefaction reactor to form a slurry strea m. The catalyst precursor in the illustrated embodiment preferably is in the form of an aqueous water solution of phosphomolybdic acid (PMA) in an amount that is equivalent to adding between 50wppm and 2 % molybdenum relative to the dry coal feed. I n the slurry mix tank 203, typica l operating temperature ranges from 300 to 600°F and more preferably between 300 and 500°F. From the slurry mix tank, the catalyst containing slurry is delivered to the slurry pump 205. The selection of the appropriate mixing and temperature conditions is based on experimental work qua ntifying the rheological properties of the specific slurry blend being processed.

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. Residua l moisture and any entrained volatiles are condensed out as sour water (not shown in Fig. 2). The coa l in the slurry leaving the mixing tank 203 has about 0.1 to 1.0% moisture. The slurry formed by the coa l and the recycle stream 204, which consists of the 600 to 700 to 1,000°F stream from the vacuum fractionator 221, and the 600 to 700°F+ strea m fraction from the atmospheric fractionator 219, is pumped from the mixing tank 203 and the pressure is raised to about 2,000 to 3,000 psig (138 to 206 kg/cm² g) by the slurry pum ping system 205. The resulting high pressure slurry may be preheated in a heat exchanger (not shown), mixed with a treat gas consisting of recycled and ma keup treat gas containing over 80% hydrogen, and then further heated in furnace 207. 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 a re simple up-flow tubular vessels, the tota l 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 ⁰ F (427 to 482 ⁰ 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/cm²g).

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 is let down in pressure in the separation and cooling system 215. The gas stream is cooled to condense out the liquid vapors of H20, naphtha, distillate, and solvent. The remaining gas is then processed to remove H₂S and C0₂.

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 a re sent to the atmospheric fractionator 219 where they are separated into light ends, naphtha, distillate, and 650 to 700°F+ fractions. The light ends are processed to recover hydrogen and Ci-C₄ hydroca rbons that can be used for fuel gas and other purposes. The naphtha is hydrotreated to saturate olefins and other reactive hydrocarbon compounds. The 160°F+ fraction of the naphtha can be hydrotreated and catalytically reformed to produce gasoline blend stock. The distillate fraction can be upgraded to produce products such as diesel and jet fuel. A portion of the 650 to 700°F+ (343 to 371°C+) is recycled to the slurry mix tank 203. The remaining 650 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 650 to 700°F+ strea m being recycled to the slurry mix tank 203. The 1000° F+ bottoms fraction from the vacuum fractionator 221 can be processed in a pa rtial oxidation system, a Circulating Fluid Bed boiler, a cement plant, or sold as a feed for asphalt paving or electrode manufacture. G.E., Shell, and others offer com mercial processes for gasification (partial oxidation) of the 1000°F+ bottoms and Circulating Fluid Bed boiler manufactures such as Foster-Wheeler and Alstom offer technology for combusting the 1000°F+ bottoms. Hydrogen for liquefaction and upgrading can also be produced by Steam Methane Reforming of a stream such as natural gas, sha le gas, or coal mine methane. This technology is utilized worldwide in refineries and offered by many com mercial vendors such as Ha ldor-Topsoe.

Catalysts useful in DCL processes also include those disclosed in U .S. Patents 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 a re disclosed in U.S. Patents Nos. 4,485,008, 4,637,870, 5,200,063, 5,338,441, and 5,389,230, U.S. Patent Application No. 13/657,087 and U.S. Provisional Application No. 61/752,428, the disclosures of which are hereby

incorporated by reference in their entirety. Examples of microcatalysts and their method of preparation are described in U.S. Patent No. 4,226,742, the contents of which are hereby incorporated by reference in their entirety.

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. An exemplary process for upgrading the liquid product of the DCL reactors 209, 211, 213 is disclosed in U .S. Patent number 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. The diesel product from DCL after upgrading will have a CN of approximately 42 and 47 depending upon cut points of the product and aromatics content. Specific gravity of the product will also vary between 0.83 and 0.90. Hence, making Euro 4 diesel, directly from DCL alone, is not

possible.

Fischer-Tropsch

Reactors, catalysts and conditions for performing cobalt catalyzed F-T synthesis are well known to those of skill in the a rt and are described in numerous patents and other publications, for example, in U.S. Patents Nos. 7,198,845, 6,942,839, 6,315,891, 5,981608 and RE39,073, the contents of which are hereby incorporated by reference in their entirety. I n the F-T unit, the C3+ stream produced by the DCL reactors 209, 211, and 213 is first converted into syngas having a 2:1 ratio of hydrogen to ca rbon, typically by steam naphtha reforming. Commercial steam naphtha reforming processes are available from Haldor Topsoe and ICI. Alternatively, the C3+ strea m can be converted into syngas having a 2:1 ratio of hydrogen to carbon by POX. The resulting syngas is then converted into a

hydrocarbon stream by F-T synthesis in the unit 121. The F-T synthesis can be performed in fixed bed, moving bed, fluid bed, ebullating bed or slurry reactors using a cobalt catalyst under operating conditions that are selected based on the desired product suite and other factors. Typica l coba lt catalyzed F-T synthesis products include normal paraffins, generally represented by the formula nCH₂. The productivity and selectivity for a given product strea m is determined by reaction conditions including, but not limited to, reactor type, temperature, pressure, space rate and catalyst type. Commercial cobalt catalyzed F-T processes are available from Shell, Sasol, ExxonMobil, Synfuels China, and others.

The normal paraffins produced in F-T may be isomerized in a

hydroisomerization reactor to impart the required characteristics for use as jet or diesel fuel blend stock. The hydroisomerization can be accomplished using a shape selective intermediate pore size molecular sieve. Hydroisomerization catalysts useful for this purpose comprise a shape selective intermediate pore size molecula r sieve and optionally a cata lytica lly active metal hydrogenation component on a refractory oxide support. The phrase "intermediate pore size," as used herein means an effective pore aperture in the range of from about 4.0 to about 7.1 Angstrom when the porous inorganic oxide is in the ca lcined form. The shape selective intermediate pore size molecular sieves used in the practice of the present invention are generally 1-D 10-, 11- or 12-ring molecular sieves. Preferred molecular sieves are of the 1-D 10-ring variety, where 10-(or 11-or 12-) ring molecular sieves have 10 (or 11 or 12)

tetrahedrally-coordinated atoms (T-atoms) joined by oxygens. In the 1-D molecular sieve, the 10-ring (or larger) pores are parallel with each other, and do not

interconnect. The classification of intrazeolite channels as 1-D, 2-D and 3-D is set forth by R. M . Barrer in Zeolites, Science and Technology, edited by F. R. Rodrigues, L. D. Rollman and C. Naccache, NATO ASI Series, 1984.

Preferred shape selective intermediate pore size molecular sieves used for hydroisomerization are based upon aluminum phosphates, such as SAPO-11, SAPO-31, and SAPO-41. SAPO-11 and SAPO-31 are more preferred, with SAPO-11 being most preferred. SM-3 is a particularly preferred shape selective intermediate pore size SAPO, which has a crysta lline structure falling within that of the SAPO-11 molecula r sieves. The preparation of SM-3 and its unique characteristics are described in U.S. Pat. Nos. 4,943,424 and 5,158,665. Also preferred shape selective intermediate pore size molecular sieves used for hydroisomerization a re zeolites, such as ZSM-22, ZSM-23, ZSM-35, ZSM-48, ZSM-57, SSZ-32, offretite, and ferrierite. SSZ-32 and ZSM-23 are more preferred.

A preferred intermediate pore size molecular sieve is cha racterized by selected crystallographic free dia meters of the channels, selected crystallite size (corresponding to selected channel length), and selected acidity. Desirable crystallographic free diameters of the channels of the molecula r sieves are in the range of from about 4.0 to about 7.1 Angstrom, having a maximum crysta llographic free dia meter of not more than 7.1 and a minimum crysta llographic free diameter of not less than 3.9 Angstrom . Preferably the maximum crysta llographic free diameter is not more than 7.1 and the minimum crysta llographic free dia meter is not less than 4.0 Angstrom. Most preferably the maximum crysta llographic free diameter is not more than 6.5 and the minimum crystallographic free dia meter is not less than 4.0 Angstrom .

A particula rly preferred intermediate pore size molecular sieve, which is useful in the present process is described in U.S. Pat. Nos. 5,135,638 and 5,282,958, the contents of which are hereby incorporated by reference in their entirety. Such a preferred molecular sieve may further be characterized by pores or channels having a crystallographic free dia meter in the range of from about 4.0 to about 7.1 .ANG., and preferably in the range of 4.0 to 6.5 .ANG.. The crystallographic free dia meters of the channels of molecula r sieves are published in the "Atlas of Zeolite Framework Types", Fifth Revised Edition, 2001, by Ch. Baerlocher, W. M . Meier, and D. H. Olson, Elsevier, pp 10 15.

Hydroisomerization catalysts useful in the present invention optiona lly comprise a catalytically active hydrogenation meta l. The presence of a cata lytically active hydrogenation metal leads to product improvement, especially VI and stability. Typical catalytically active hydrogenation metals include chromium, molybdenum, nickel, vanadium, cobalt, tungsten, zinc, platinum, and palladium . The meta ls platinum and palladium are especially preferred, with platinum most especially preferred. If platinum and/or palladium is used, the total amount of active hydrogenation metal is typica lly in the range of 0.1 to 5 weight percent of the total catalyst, usually from 0.1 to 2 weight percent, and not to exceed 10 weight percent. The refractory oxide support may be selected from those oxide supports, which are conventionally used for catalysts, including silica, alumina, silica-a lumina, magnesia, titania and combinations thereof.

The conditions for hydroisomerization will be tailored to achieve an isomerized liquid intermediate with specific bra nching properties, as described above, and thus will depend on the cha racteristics of feed used. I n general, conditions for

hydroisomerization in the present invention are mild such that the conversion of hydrocarbon materials boiling below about 700°F is maintained above about 50 to about 80 wt % in producing the intermediate isomerates. Mild hydroisomerization conditions are achieved through operating at a lower temperature, genera lly between about 390 and 650F at a LHSV generally between about 0.5 hr ¹ and about 20 hr^"1. The pressure is typica lly from about 15 psig to about 2500 psig, preferably from about 50 psig to about 2000 psig, more preferably from about 100 psig to about 1500 psig. Low pressure provides enhanced isomerization selectivity, which results in more

isomerization and less cracking of the feed, thus producing an increased yield.

Hydrogen is present in the reaction zone during the hydroisomerization process, typica lly in a hydrogen to feed ratio from about 0.5 to 30 MSCF/bbl (thousand

standard cubic feet per barrel), preferably from about 1 to about 10 MSCF/bbl.

Hydrogen may be separated from the product and recycled to the reaction zone.

The process and apparatus of the present invention enjoy a number of advantages. In addition to those described above, such advantages include:

Making maximum value use of the byproduct C3+ that is produced during the DCL liquefaction process and the upgrading of multi-ring aromatics present in coal to single and double ring naphthenes.

Can achieve a fully synthetic jet fuel only yield of up to 60wt% for the coal used in this embodiment.

If BTX or gasoline blend stock by-product is desired, up to an 85% Jet/15% BTX product split is possible.

The yield of jet fuel can be as high as 3 barrels per ton of coal on a MAF basis.

Can achieve a fully synthetic diesel fuel yield of over 67wt% for the coal used in this embodiment which is equivalent to 4.5 barrels per ton of DAF coal.

The invention provides the most efficient route for the production of BTX from coal.

The current commercial approach for producing BTX is via low efficiency, catalytic conversion of methanol.

The thermal efficiency of the overall plant 100 can exceed 60 percent. The jet or diesel fuel produced by the process of the invention has an energy content that is about 5 percent points higher than conventional jet or fuel produced from petroleum and meets all JP-A/JP-8 product specifications and Euro 4 diesel specific gravity and cetane number product specifications.

The resultant DCL jet fuel is thermally stable at a higher temperature and has a lower freeze point then conventional petroleum produced jet fuel, and thus is ideal for use in high speed aircraft where the fuel is used for cooling. The DCL jet meets or exceeds JP-8 specifications, as shown in Example 1.

This combined DCL/F-T process provides several avenues of flexibility to more efficiently tailor jet specifications, thus maximizing the overall jet fuel yield. By this invention a lower density F-T product that does not meet the jet fuel density specification can be blended with a higher density DCL product to meet density specs, as shown in Example 2. Alternatively, a less isomerized, more n-paraffinic F-T product that does not meet the jet fuel freezing point specification can be blended with DCL product having a freezing point that exceeds the applicable specification, as shown in Example 3.

A broad range of DCL/FT blending ratios are possible, depending on overall plant operating and product slate objectives. Additional blending combinations exist via DCL and F-T processing severities and product cut points/boiling ranges. One such method is to vary the degree of hydrocracking or ring-opening to the DCL distillate, as shown in Example 4.

The following examples describe jet and diesel fuel compositions that can be produced by the present invention. The jet fuel compositions of Examples 1 through 4 can be contrasted with the 25-45% n-paraffins, 15-30% i-paraffins, 30-50% cycloparaffins, and <5% aromatics disclosed as the optimum composition for high thermal stability jet fuel by Liu GuoZhu et al., Sci China Ser B-Chem, Feb 2008, vol. 51, no. 2, 138-144. In the current invention, the DCL/F-T ratios are >1 with a component balance of <50% paraffins and >50% cycloparaffins , whereas

GuoZhu's jet fuel is comprised of <50% cycloparaffins. Example 1

Example 2

Example 3

Example 4

Neat DCL diesel can meet US #2 low sulfur diesel specifications, as shown in Example 5 below. However, blending with F-T diesel is required to meet more stringent density, Cetane Number, and Cetane Index specifications, such as for Euro 4 diesel. For a coal-only DCL/F-T balanced plant, Euro 4 diesel specifications can be met using a range of 7-to-2 DCL to 1 F-T blend, depending on the degree of F-T isomerization, the DCL cut points, and the degree of DCL distillate hydrocracking. Two examples of blends that meet Euro 4 specifications are shown in Examples 6 and 7. I n Example 6, an 80/20 DCL/F-T blend produced with a partially isomerized F- T blendstock is possible, while in Example 7 an 85/15 DCL/F-T blend with no isomerization is also possible.

The CI calculation formula and the resulting restrictions on the DCL/F-T blending ratios in Examples 6 and 7 are questionable for DCL liquids given their relatively high densities. The reason for having lower CI for higher density petroleum-based diesel is because of increasing aromatics, which have poor ignition quality. By contrast, the higher density of DCL-based fuel is because of increasing cycloparaffins, which provide improved ignition quality. Therefore because increased density decreases CI for a given diesel boiling curve, using CI as an indicator of ignition quality is not meaningful for highly cycloparaffinic DCL stock or DCL/F-T blends. A CI correlation, such as the one developed by Syncrude (2005) for oils sands bitumen-derived diesel fuels, is likely necessary for the highly cycloparaffinc DCL diesels. Syncrude's correlation adds aniline point (i.e. aromatics indicator) to the existing petroleum 4-variable CI equation and is reported to significantly improve the CI fit to CN. An appropriate CI for DCL diesel needs to add a term involving cycloparaffins. Additional input would also be needed to cover a DCL/F-T blend.

If the currently formulated Cetane Index were no longer a critical spec, use of a cetane improver to increase the Cetane Number of the DCL blend stock would allow additional blending freedom. Cetane Index can be met at higher diesel initial cut points, but if no longer critical, initial diesel cut points as low as 180F with a small amount of cetane improver will meet the remaining Euro 4 specs. Additional DCL diesel property adjustments, such as boiling range and aromatics content, can be made to further adjust the blending ratio. For example, increasing the DCL Cetane Number to 48 allows a 12 to 1 DCL to F-T blend to achieve a 51 CN at a density of 0.839.

Example 5

GCD, F @ wt%Off D-2887/TBP

572 / 673

90 589 min/max

95 612 680max

Copper Strip Corrosion Rating D-130 la No. 3 max

Flash Point, F D-93 169 125min

Water & Sediment, vol% D-1796/2709 0.005 0.05max

Kinematic Viscosity, cSt @ 40C D-445 3.1 1.9 / 4.1 min/max

Aromaticity, vol% D1319 2 35max

Lubricity HFRR @ 60C, micron

wear D6079/D7688 460 WSD 520max

Ramsbottom Carbon, wt% on

10% btms D-86/D-524 0.04 0.35max

Example 6

Example 7

While much of the above diesel examples and discussion relates to the production of the diesel fuels meeting the stringent Euro 4 diesel requirements, it should be noted that US #2 low sulfur diesel specifications are met at all of the above discussed conditions.

As further illustrations, Graph 1 below shows a ~350F initial diesel cut point would be required to meet the 46 CI specification for Euro 4 diesel with a diesel yield of 63wt% on MAF coal. Again, if the CI specification were to be waived or revised, a 250F initial cut point meets the 51 CN specification with an increased diesel yield of 67wt% on DAF coal. Although diesel yield decreases as CN increases, this illustrates a range of conditions to suit plant operations are possible by varying initial diesel cut point. As shown on the graph, DCL distillate hydrocracking can also be utilized to produce Euro 4 diesel with heavier back-end DCL raw distillates.

If it is desired to maximize the production of BTX product, Graph 2 shows the trade-off between diesel yield and BTX. DCL BTX can be increased by DCL distillate hydrocracking, but plant BTX yield is dependent on feed requirement to the F-T plant and on meeting diesel specifications.

Graph 1

Impact of Cut Point on

Graph 2

Co-Production of Diesel and BTX

2 4 6

BTX Yield, wt% on DAF Coal

As illustrated in Example 8, if the boiling point specifications were relaxed somewhat, a single product diesel/jet blend can be produced similar to that shown in Example 4, which is based on the JP-8 FBP with a resulting lower boiling, high CN, high energy content diesel.

Example 8

Products

F-T/DCL Mid Pt,F FzPt, C CN Density

F-T 15 425 -5 89 0.748

MCL 85 421 -70 45 0.851

Claims

What is claimed is:

1) A method for producing liquids from feed coal, comprising the steps of:

a) supplying feed coal to a direct coal liquefaction (DCL) reactor and operating such reactor in a catalytic process for converting at least 50% of the supplied feed coal on a moisture and ash free (MAF) basis to coal liquids;

b) upgrading coal liquids produced in step a to produce C3/180-400°F and distillate blend stock streams;

c) converting C3/180-400°F produced in step b to syngas;

d) converting syngas produced in step c by a Fischer Tropsch (F-T) process to

predominately paraffinic liquids;

e) upgrading the liquids produced in step d to produce a distillate blend stock stream and lighter products; and

f) blending distillate blend stock streams produced in steps b and e to produce a liquid fuel or fuel blend stock.

2) The method of Claim 1 wherein the stream being converted in step c is a C3/180-350°F

stream.

3) The method of Claim 2 further including the step of recycling the lighter products produced in step e to be converted into syngas and used as a feed to the F-T process.

4) The method of Claim 2 wherein the ratio of distillate blend stock produced by step b that is blended in step f with the distillate blend stock produced by step e is between 2 and 12 to 1 and the blended distillate fuel produced in step (f) is a diesel fuel or diesel fuel blend stock having a Cetane Number of greater than 47. 5) The method of Claim 2 wherein the ratio of distillate blend stock produced by step b that is blended in step f with the distillate blend stock produced by step e is between 4 and 15 to 1 and the blended distillate fuel produced in step (f) is a jet fuel or jet fuel blend stock having a freeze point less than or equal to -47° C.

6) The method of Claim 4 wherein the ratio of distillate blend stock produced by step b that is blended in step f with the distillate blend stock produced by step e is between 3 and 10 to 1 and the blended diesel fuel or diesel fuel blend stock produced in step f has a Cetane Number of at least 50, and a specific gravity of 0.83 to 0.85.

7) The method of Claim 6 wherein the ratio of distillate blend stock produced by step b that is blended in step f with the distillate blend stock produced by step e is between 3 and 7 to 1 and the blended diesel fuel or diesel fuel blend stock produced in step f has a Cetane Number of at least 51, and a cetane index of at least 46.

8) The method of Claim 5 wherein the ratio of distillate blend stock produced by step b with the distillate blend stock produced by step e is between 4 and 10 to 1 and the blended jet fuel or jet fuel blend stock has a specific gravity of 0.775 to 0.840.

9) The method of Claim 2 wherein the inertinite content of the coal is greater than 10%.

10) The method of Claim 2 wherein the DCL reactor contains a molybdenum microcatalyst.

11) The method of Claim 2 wherein the upgrading step b also produces a 160/250°F to

160/300°F feed stock for the production of aromatics for chemicals production. 12) Apparatus for producing jet or diesel fuels or blend stocks comprising:

a) a direct coal liquefaction (DCL) reactor for converting a coal to liquid hydrocarbons; b) and upgrader for upgrading the liquid hydrocarbons produced by the DCL reactor; c) a separator for separating the output of the upgrader into C3/180-350°F, Naphtha and DCL jet or diesel blend stock streams;

d) a Fischer Tropsch system for converting C3/180-400°F product from the separator to syngas and for converting produced syngas to predominately paraffinic jet or diesel blend stock.

13) The Apparatus of Claim 12 wherein the Fischer Tropsch system of element d is for

converting C3/180-350°F product from the separator to syngas and for converting produced syngas to predominately paraffinic jet or diesel blend stock.