DIRECT COAL LIQUEFACTION PROCESS
A direct coal liquefaction process capable of producing unexpectedly high levels of C5/650° F. product, which process employs a relatively high ratio of solvent plus bottoms product recycle to feed coal.
This is a continuation-in-part of U.S. Ser. No. 14/147,542 filed Jan. 5, 2014.
FIELD OF THE INVENTIONThe present invention relates to a direct coal liquefaction process capable of producing unexpectedly high levels of C5/650° F. product, which process employs a relatively high ratio of solvent plus bottoms product recycle to feed coal.
BACKGROUND OF THE INVENTIONFrom time to time, increases in cost and future shortages of petroleum often leads to increased interest in coal as a fuel source given the vast, easily accessible deposits of coal that exist in several parts of the world. Various processes have been proposed for converting coal to liquid and gaseous fuel products, including transportation fuels, and in some cases, to other products such as lubricants and chemicals. Problems that have hampered the commercial liquefaction of coal include the relatively low thermal efficiency of indirect coal-to-liquids (ICTL) conversion methods, such as Fischer Tropsch (FT) synthesis and methanol-to-liquids (MTL) conversion, as well as high water use and CO2 emissions.
Direct coal liquefaction (DCL) methods have typically involved heating the coal in the presence of a hydrogen donor solvent, and optionally a catalyst, in a hydrogen containing atmosphere to a temperature in the range of about 700° to 950° F. This results in a break-down of the coal structure into free radicals that are quenched to produce liquid products. The catalyst typically contained finely divided iron, molybdenum, or mixtures thereof. One source of the molybdenum catalyst is via in situ formation from a phosphomolybdic acid (PMA) precursor.
Conventional direct coal liquefaction process units typically require passing spent solvent to a hydrotreater for preparing a hydrogen donor solvent that is fed back to the input of the DCL unit to act as a solvent and to increase conversion in the liquefaction process. Unfortunately, solvent hydrotreating requires separation of at least a faction of the solvent, additional equipment, and additional hydrogen-rich treat gas. Addition of an external solvent hydrotreater increases the required investment and decreases thermal efficiency. This is an important reason why low donor solvent-to-coal ratios (typically 1.2 to 1.5) are used in coal liquefaction process units using a solvent hydrotreater.
In addition, hydrotreating reduces the viscosity and lowers the aromatics content of the solvent, which reduces its ability to suspend ash in the slurry and its compatibility with coal. The reduction in ability to suspend ash results in an increased likelihood of solids buildup, deposits, or plugging of high pressure feed pumps, transfer lines, heat exchangers, furnace tubes, and reactors. Hydrotreated solvent also has a higher cracking rate during liquefaction. This can result in the liquefaction process producing insufficient product solvent for recycle for preparing a slurry with the feed coal. For such a situation, an external source of solvent (such as coal tar) is required.
Use of donor solvents also results in higher gas hold-up in the liquefaction reactor, which in turn, requires a larger reactor volume to achieve adequate coal residence time in the reactor. Further, high recycle gas rates are also required because treat gas must be provided for both the solvent hydrotreater and the liquefaction reactor.
In order to reduce gas hold-up and avoid solids build up, ebullated bed reactors are preferred. Because of the high liquid recycle in ebullated bed reactors, the reactors are substantially fully back-mixed, which results in an increase in reactor volume versus plug flow reactors.
As much as 50% of the heat of reaction is moved from the liquefaction reactor to the solvent hydrotreater. Thus, additional fuel must be fired for preheating of the feed to liquefaction and for the solvent hydrotreater. This results in lower thermal efficiency.
A further issue limiting the application of DCL processes is that lower quality coals, such as those having inertinite content higher than about 12 vol. % have been considered unsuitable for use as a DCL feedstock. High inertinite coals are found in many parts of the world, including the United States and China. Many of these coals, such as that in the Ordos basin in China have inertinite contents of more than 25% and have a low ratio of atomic hydrogen to carbon (H/C). Such coals have historically been unacceptably more difficult to liquefy by DCL than higher quality coals that have a substantially low vitrinite content.
Okada reported in a paper entitled: “Possible Impacts of Coal Properties on the Coal Conversion Technology”, Coal Science, J. A. Pajares and J. M. D. Tascon, 1995 Elsevier Science, that oil yield from autoclave experiments were inversely proportional to the inertinite content of coal for coals of similar rank. Oil yield ranged from about 67 wt % for a zero inertinite content coal to about 40 wt % for a coal containing about 60 vol % inertinite. He concluded that high inertinite coals, such as that found in the Ordos Basin, are not suitable for direct coal liquefaction.
Also, in the paper “Study on Coal Liquefaction Characteristics of Chinese Coals”, Fuel 81 (2002) 1551-1557, Wasaka published results on 53 runs on 27 coals that were made in a 0.1 t/d pilot plant test program. The program specifically focused on direct liquefaction of high inertinite content coals. At constant liquefaction operating conditions, conversion decreased with increasing inertinite content coals.
Further U.S. Pat. No. 7,763,167 to Zhang et al discloses a DCL process that utilized an iron-containing catalyst and an externally hydrotreated donor solvent. The donor solvent was produced in a suspended bed using a forced circulation reactor (ebullated bed). Although they obtained an oil yield, they did not disclose the boiling point of the oil product.
While the art contains various conventional DCL processes to convert various coals to liquids, there remains a need in the art for a DCL process that is able to achieve unexpectedly high yields of C5/650° F. boiling range products, even with relatively high inertinite-containing coals.
SUMMARY OF THE INVENTIONIn accordance with the present invention there is provided a process for the direct liquefaction of coal, which process is conducted in the absence of added carbon monoxide, and is performed in a coal liquefaction process plant comprising: a slurry mixing zone, a preheating zone, a liquefaction zone, a separation zone capable of separating a gaseous product stream from a liquid/solids product stream, an atmospheric fractionation zone, and a vacuum fractionation zone, which process comprises:
a) introducing into said slurry mixing zone:
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- i) coal having an average particle size of about 75 microns to about 600 microns and a moisture content from about 1 to about 4 wt. %;
- ii) non-donor solvent from said vacuum fractionation zone and non-donor bottoms from said atmospheric fractionation zone, wherein the ratio of non-donor solvent plus non-donor bottoms to coal is from about 2.5 to 1 to about 4 to 1:
- iii) a molybdenum-containing catalyst provided at a make-up rate that is equivalent to about 50 wppm to about 2 wt. % molybdenum on a moisture and ash free (MAF) feed coal basis, wherein the resulting slurry is at a temperature from about 200° F. to about 600° F.;
b) conducting said slurry at a pressure from about 1500 psig to about 3000 psig and an effective amount of treat gas containing at least about 80 vol. % hydrogen, to said preheating zone wherein it is heated to a temperature of about 650° F.;
c) conducting said heated slurry to said liquefaction zone wherein it is reacted at a temperature from about 700° F. to about 950° F. thereby producing reaction products;
d) conducting said reaction products to a separation zone wherein a gaseous product stream is separated from a liquid/solids product stream;
e) conducting said liquid/solids stream to said atmospheric fractionation zone, wherein it is fractionated to result in a C1 to C4 gaseous fraction, a C5/650° F. fraction, and a non-donor 650 F+ bottoms fraction;
f) conducting an effective portion of non-donor 650° F.+ bottoms fraction from said atmospheric fractionation zone to said vacuum fractionation zone thereby resulting in a 1000° F.+ fraction and a non-donor 650° F. to 1000° F. solvent fraction from the vacuum fractionator; and leaving a remaining portion of said non-donor 650° F.+ fraction from said atmospheric fractionator;
g) recycling the remaining portion of said non-donor 650° F.+ fraction from said atmospheric fractionation zone to the slurry mixing zone; and
h) recycling at least a portion of said non-donor 650° F. to 1000° F. solvent fraction from said vacuum fractionation zone to said slurry mixing zone.
In a preferred embodiment of the present invention the ratio of non-donor solvent plus non-donor bottoms to coal is about 3:1 to 4:1.
In another preferred embodiment of the present invention the ratio of non-donor solvent plus non-donor bottoms to coal is about 3:1 to about 3.5:1.
In yet another preferred embodiment of the present invention the amount of inertinite in said coal is from about 7 to 14 vol. %.
In still another preferred embodiment of the present invention the amount of inertinite in said coal is greater than about 20 vol. %.
It has unexpectedly been found by the inventors hereof that a dramatic and unexpected increase in C5/650° F. liquid product, along with high thermal efficiency, are achieved by the practice of the present invention. This dramatic and unexpected increase can also be achieved with any bituminous or sub-bituminous coal, even using a coal having a high inertinite content. For example, it is well known in the art that the Wyoming coals (Rawhide and Wyodak) typically have an inertinite content from about 8 to 11 vol. %. Coals from other regions of the Powder River Basin in Wyoming can have inertinite contents of up to 20 vol. %. There are also Chinese coals having very high inertinite content. For example, Yulin coal from the province of Shaanxi has a inertinite content of about 30 vol. % and Guojiawan coal has an inertinite content of about 46 vol. %. It is well known in the art that high inertinite content coals are difficult to convert by direct coal liquefaction processes.
One key to the unexpected results obtained by the direct coal liquefaction process of the present invention is the use of a high ratio of non-donor recycle solvent and non-donor recycle bottoms to coal. That is, high ratios of: i) the 650° F. to 1000° F. fraction from the vacuum fractionator zone, sometimes referred to herein as “solvent”, and ii) the 650° F.+ fraction from the atmospheric fractionation zone, sometimes referred to herein as “bottoms”. Typically, the upper range of the 650° F.+ fraction from the atmospheric fractionator will be about 700° F. The ratios are in excess of those conventionally practiced commercial size pilot plants and process plants by direct coal liquefaction processes. Such plants can typically process at least about 75 lbs of coal per day. This dramatic increase in C5/650° F. product yield of the present invention is particularly present in the instant micro-catalytic direct coal liquefaction process. That is wherein a finely dispersed molybdenum or iron catalysts is used. It is preferred that the catalyst of the present invention be comprised of finely dispersed molybdenum, and that a non-donor product recycle stream produced in liquefaction process be used. The ratio of such recycle solvent and bottoms product to coal at the input to the liquefaction zone, on a moisture free weight basis, is at least about 2.5:1, preferably about 3:1, also about 3.5:1, as well as about 4:1.
The term “non-donor” as used herein means that the recycle solvent and streams have not been processed in a hydrotreater to partially hydrogenate multi-ring aromatic compounds in the stream to produce compounds that can donate hydrogen during liquefaction. Surprisingly, increasing the ratio of the recycled non-donor stream to coal in the instant process does not increase the flow rate of the recycled stream and fresh coal to liquefaction for a given rate of product generation. Instead, surprisingly, less coal is required. Although the recycle stream increases relative to coal, the total feed to liquefaction remains substantially the same. The net impact of higher recycle and lower coal rate is a reduction of energy required in the slurry preheating zone and the size of the vacuum fractionator. Hence, investment and energy requirements are reduced for the liquefaction section of the instant liquefaction process.
Referring now to
Most of the remaining moisture in the coal is driven off in the mixing tank due to hot atmospheric fractionator bottoms feeding into said mixing tank. Residual moisture, as well as any entrained volatiles, are condensed out as sour water (not shown). The coal in the slurry leaving mixing tank 203 has about a 0.1 to 1.0% moisture content. The slurry formed by the coal and the recycle streams is pumped from mixing tank 203 wherein the pressure is raised to about 1,500 to 3,000 psig, preferably from about 2000 to 3000 psig, by slurry pumping system 205. The resulting high-pressure slurry may be preheated in a heat exchanger (not shown), mixed with a treat gas comprised of recycled and makeup treat gas containing at least about 80 vol. % hydrogen, then further heated to a temperature of about 600° F. to about 700° F. in slurry preheating furnace 207. The coal slurry and hydrogen mixture is fed to the input of the first reactor of the series connected liquefaction reactors 209, 211 and 213 at between about 600° F. to 700° F. and 1,500 to 3,000 psig. Reactors 209, 211 and 213 are preferably up-flow tubular vessels. The total length of the three reactors is from about 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 700 to 950° F., preferably from about 800° F. to about 900° F. It is preferred that a portion of the hydrogen-containing 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.
The effluent from the last reactor is separated into a gaseous stream and a liquid/solid stream. The liquid/solid stream is let down in pressure separation and cooling zone 215, also sometimes referred to herein as the separation zone. The gaseous stream is cooled to condense out liquid vapors, such as H2O, naphtha, distillate, and solvent. The remaining gas is then processed to remove H2S 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 at least a fraction of the hydrogen contained therein, which is then recycled to be mixed with the coal slurry. Any remaining processed gas can be purged to prevent buildup of light ends in the recycle loop. Hydrogen recovered therefrom can be used in any downstream hydro-processing upgrading system.
The resulting depressurized liquid/solid stream and the hydrocarbons condensed in separation and cooling stage 215 are passed to atmospheric fractionation zone 219 where they are separated into light ends and liquid fractions. The liquid fraction is separated into a light ends fraction, a C5/650° F. minus liquid fraction, and a 650° F.+ fraction. The light ends are processed to recover hydrogen and C1-C4 hydrocarbons that can be used for fuel gas and other purposes.
An effective amount or portion of the of the 650° F.+ fraction from atmospheric fractionation zone 219 is passed to the vacuum fractionation zone 221 wherein it is separated into a non-donor 650° F. to 1000° F. solvent fraction and a 1000° F.+ fraction. By effective amount or portion we mean that amount required to purge ash from the system. The remaining portion of 650° F.+ stream from the atmospheric fractionation zone and the 650° F. to 1000° F. solvent fraction from the vacuum fractionation zone are recycled to slurry mix tank, also sometimes referred to herein as the slurry mixing zone 203.
At least a fraction of the 1000° F.+ fraction from vacuum fractionator 221 is sent to be gasified by partial oxidation zone 223 to generate hydrogen for use in the present liquefaction process. A portion of the coal from gas sweep mill 201 is preferably fed to partial oxidation zone 223 to produce additional hydrogen. Alternatively, instead of the partial oxidation zone 223, the 1000° F.+ bottoms fraction from vacuum fractionator 221 can be processed in a Circulating Fluid Bed boiler, a cement plant, or sold as a feed for asphalt paving or for electrode manufacture.
Hydrogen for liquefaction and upgrading can also be produced by any suitable technology. For example, Steam Methane Reforming of a stream such as natural gas, shale gas, or coal mine methane, which is well known in the art, can be used. Such a technology is utilized worldwide in refineries and offered by many commercial vendors such as Haldor-Topsoe. Also, catalysts useful in DCL processes of the present invention 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 herein by reference in their entirety.
A preferred embodiment of the liquefaction process of the present invention combines several elements that contribute to maximum premium C5/650° F. fuels product production and maximum thermal efficiency. These include: the recycle of a non-donor streams, preferably including atmospheric fractionator bottoms, to maintain a ratio of the recycle stream to coal at the input to reactors 209, 211, 213 that is at least 2.5:1 on a moisture free weight basis, preferably between 3.0:1 and 3.5:1, as well as 4:1; and the use of a micro-catalyst in the form of finely dispersed molybdenum. Also, use of 650° F.+ recycle, and multiple slurry reactors in series contribute to the benefits of the instant process. It will be understood that the expression (S+B) which is sometimes used herein is meant to mean (Solvent+Bottoms), both of which are non-donor recycle streams.
Use of a micro-catalyst, 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 associated with the use of a donor solvent such as required by prior DCL systems. Firstly, energy is lost during the preparation of the donor solvent. Secondly, 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 present process in the liquefaction reactors during operation with a 650° 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 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 650° F.+ 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 incorporated herein by reference in their entirety.
Further, the 650° F.+ fraction from atmospheric fractionator 219 and the 650° F. to 1000° F.− fraction from vacuum fractionator 221, as the non-donor streams being recycled to the slurry mixing tank 203 provides preheat for the coal and solvent in slurry mix tank 203. This preferably raises the temperature in the mixing 203 tank to about 200° F. to 600° F., preferably 300° F. to 600° F., more preferably about 300° 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 micro-catalyst is entrained in the 600 to 700° F.+ fraction recycled from the atmospheric fractionator 219, so that recycling a larger portion of such fraction increases the catalyst concentration in the liquefaction reactors, thereby decreasing the requirement for the addition of fresh catalyst precursor and increasing the conversion efficiency of the process.
Use of the non-donor 600° F. to 700° F.+ stream, preferably 630° F. to 670° F.+, and more preferably a 650° F.+, process derived recycle solvent in the DCL process of the present invention reduces cracking, relative to use of a donor solvent, and thus produces a 650° F.− product having a greater fraction of diesel and less light gases and naphtha. The 650° F.− product can be selectively upgraded by conventional processing to finished products in fixed bed upgrading reactors.
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 maintained to maximize selectivity to liquids. Thus, the use of three serially connected reactors are preferred. 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 is disclosed in U.S. Pat. No. 5,198,099, the disclosure of which is hereby incorporated herein by reference in its entirety.
The diesel product, after upgrading, will have a Cetane number of between 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. A higher Cetane Number is required for Euro 4 diesel, thus a Fischer-Tropsch facility producing a 70-75 Cetane Number diesel blend stock may be added to the plant operating in accordance with the present invention. The gasoline produced by upgrading the relevant portion of product of the process of the present invention will meet all current gasoline specifications, or can be upgraded to a Research Octane of 106 if desired. This will permit the blending of the low octane naphtha into the gasoline pool while maintaining adequate octane for the blended fuel. Also, the upgrading process can also be operated to maximize the production of jet fuel or gasoline. The jet fuel produced will meet all Military JP-8 specifications.
Continuous, integrated pilot plants are utilized to define commercial coal conversion and product yields. These facilities typically include distillation towers for recycle of solvent, bottoms, and catalyst contained in the bottoms. It generally takes seven or eight days for catalyst concentration, coal conversion, composition of streams, and yields to equilibrate and provide real world data for commercial coal liquefaction plants. Small batch reactors, such as tubing bombs, mini-bombs, and autoclaves, that are typically used for examples in most coal liquefaction patent applications, merely provide basic information to researchers. They are not capable of providing reliable data for use in designing a commercial direct coal liquefaction plant. Further, such small reactors are not designed to be continuous reactors, but represent a batch experiment with catalyst, stream composition, and relative rates that are not equilibrated because of their designed short run times. In addition, because of their small size, extraction with solvents such as cyclohexane or THF (tetrahydrofuran) are utilized for determining conversion rather than distillation (fractionation) which is a very important step in a commercial plant to identify a range of products. Typically, only total conversion is capable of being determined with use of such small once-through reactors. Thus, the present invention applies only to those direct coal liquefaction plants containing both an atmospheric fractionator and a vacuum fractionator. It is preferred that the size of the liquefaction plant be at least about 75 lb per day coal feed, and more preferably at least about 240 lb per day of coal. It will be understood that the terms atmospheric distillation, atmospheric fractionator, atmospheric distillation tower, and atmospheric pipe still can be used interchangeably herein. Also, the terms vacuum distillation, vacuum fractionator, vacuum distillation tower, and vacuum pipe still can be used interchangeably herein as well.
The following examples provide coal conversion and yield data for either a 75 lb/day or 240 lb/day pilot plant. Both units have been designed or modified to operate as a continuous, integrated pilot plant that utilizes both atmospheric and vacuum distillation for separation of products and recycle streams.
Data presented in the following examples was obtained at liquefaction temperatures of 780° F. to 880° F. and operating pressures of 2,000 to 2,500 psig. Catalysts used include (1) iron, (2) iron and molybdenum, or (3) molybdenum alone. Summaries of the data obtained in the following examples are presented in Table 1 and in
Wasaka et al, discloses in “Study on Coal Liquefaction Characteristics of Chinese Coals; Fuel 81 (2002); 1551-1557, the direct coal liquefaction of 27 Chinese coals having inertinite contents between 0.8 vol. % and 46.1 vol. %. Solvent was recycled at a ratio solvent to coal of 1.5 for each run with each coal. All runs were performed in a 240 lb/day pilot plant in China.
The pilot plant was comprised of a slurry preparation and feed section, a liquefaction section, and a product separation and withdrawal section. A slurry of catalyst, donor solvent and coal was prepared in the slurry preparation and feed section and was blended with hydrogen from a hydrogen feed section, then sent to a liquefaction section. It was retained in the liquefaction section for about 1 hour at a temperature of about 878° F. (470° C.) and a pressure of about 2466 pia.
As an example, Guojiawan was one of the coals tested by Wasaka et al. This coal contained 46.1 vol % inertinite. The donor solvent was hydrotreated to produce a hydrogen donor solvent prior to preparation of the slurry with the feed coal. A catalyst make-up rate of up to 3 wt. % Fe2O3 was used as a catalyst in addition to the donor solvent. Both the iron catalyst and the donor solvent are known routes for increasing coal conversion, but both are unacceptable with the present process because the solvent used in the present invention is a non-donor solvent and the catalyst is molybdenum.
The liquefaction products were analyzed and it was found that total conversion was about 70 wt. % and it was estimated that the C5/650 F reaction product would be about 37.5 wt. %, which are both shown in Table 1 and
Two coal liquefaction experiments were performed on a high inertinite content (about 29.9 wt. %), Yulin coal from Inner Mongolia. The recycle solvent+bottoms to coal ratio used was 1.2 for Example 2 and 1.5 for Example 3.
Both examples were performed in a 240 lb/day continuous coal liquefaction pilot plant in Beijing, China. The pilot plant was modified to have substantially the same configuration as the 75 lb/day continuous coal liquefaction pilot plant hereof for examples 4 and 5 hereof. For example, the 75 lb/day pilot plant contained a slurry preparation and feed section, a liquefaction section, and a product separation and withdrawal section which were also included in the 240 lb/day plant. For this example, three liquefaction reactors were operated in series at temperatures of 788° F., 842° F., and 842° F. respectively. Pressure for each reactor was 2,500 psig with 120 minutes nominal run time, and a catalyst make-up of 300 wppm molybdenum for Example 4 and 1000 wppm molybdenum for Example 5.
The liquefaction products were analyzed and it was found that total coal conversion was only 57.2 wt. % for Example 2, and 57.8 wt. % for Example 3. C5/650° F. yields was 41 wt % and 41.2 wt. % for Examples 2 and 3, respectively.
The low total coal conversion resulting from these two examples is consistent with the higher inertinite content of the Yulin coal. It is conventional wisdom in the art that it is difficult to convert a substantial fraction of high inertinite coal in DCL (Direct Coal Liquefaction) processes.
Example 4 and 5 (Comparative)Two coal liquefaction experiments were performed at a recycle (solvent+bottoms)/coal ratio of 1.65.
Both experiments were performed with Rawhide coal, which had a inertinite content of about 9.5 vol. %, in the same 75 lb/day continuous coal liquefaction pilot plant that included fractionation of the resulting liquid product and recycle of solvent and bottoms. Reactor conditions were: 801° F. and 842° F. in two stages of liquefaction; each at a pressure of 2500 psig; 50 minutes nominal run time, and a make-up rate of mixed catalyst containing 100 wppm molybdenum and about 1 wt % Fe2O3.
The resulting liquefaction products were analyzed and it was found that total coal conversion was 89.3 wt. % for Example 4 and 88.6 wt % for Example 5. C5/650° F. yield was 46.1 for Example 4 and 44.2 wt. % for Example 5.
Examples 6 to 11 (Comparative)Six experiments were performed using the same Rawhide coal used in Examples 4 and 5 hereof but at recycle (solvent+ bottoms)/C of 1.7, 1.71, 1.74, 1.74, 1.83 and 1.91, respectively.
These experiments were also performed in the same 75 lb/day continuous coal liquefaction pilot plant used for Examples 4 and 5 hereof. Liquefaction temperatures ranged from about 807° F. to about 836° F. in the first stage and 842° F. in the second stage. For each example, the pressure was 2500 psig. Nominal residence time ranged from 41 to 49 minutes. The liquefaction catalysts make-up rate was 100 wppm moly and 0.25 to 1.03 wt. % FeO3.
The resulting liquefaction products were analyzed and it was found that total coal conversion for the six Examples was found to be 85.6, 87.0, 86.9, 85.2, 85.7, and 88.6 wt % respectively. The C5/650° F. yield for the six Examples was found to be 44.0, 43.5, 41.9, 42.4, 44.4, and 48.3 wt. % respectively.
Example 12 to 14 (Comparative)Three experiments were performed using Wyodak coal having an inertinite content of about 10 vol. %. All three examples used a recycle (solvent+ bottoms)/Coal ratio of 2.0.
These experiments were performed in the 75 lb/day pilot plant as previous mentioned in prior examples hereof. The liquefactions conditions for each experiment were: average liquefaction temperature of 797° F.; pressure of 1969 psig; a nominal residence time of about 144 minutes; and a molybdenum make-up rate 100 wppm.
The resulting liquefaction products were analyzed and it was found that the total coal conversion for example 12 to 14 was 77.5, 81.4, and 75.9 wt. % respectively. The C5/650° F. yield was found to be 47.1, 44.2, and 45.1 wt. %, respectively.
Examples 15 and 16 (Comparative)Two experiments were performed using the Wyodak coal of Examples 12 to 14 hereof at a (solvent+ bottoms)/Coal ratio of 2.0 but at higher liquefaction temperatures.
These experiments were performed in the 75 lb/day pilot plant discussed in prior examples hereof. The average liquefaction temperature for each example was 841° F.; at a pressure of 1994 psig for example 15 and 2412 psig for Example 16; a nominal residence time of 48 minutes for Example 15 and 70 minutes for Example 16. Both examples used a molybdenum make-up rate of 100 wppm on MAF coal.
The resulting liquefaction products were analyzed and it was found that the total coal conversion for Example 15 was 78.2 wt. % and 76.1 wt. % for Example 16. C5/650° F. yield was found to be 47.8 wt. % to 43.5 wt. % for Examples 15 and 16 respectively.
Example 17 (Example of this Invention)A coal liquefaction experiment was performed using the same Yulin coal used in Examples 2 and 3 hereof but a recycle (solvent+ bottoms)/Coal ratio of 3.
The same 240 lb/day pilot plant and conditions were utilized as in Examples 2 and 3 hereof except for the higher (Solvent+Bottoms)/Coal ratio of at least 3. The molybdenum catalyst make-up rate was 300 wppm.
The resulting liquefaction products were analyzed and it was found that total coal conversion was unexpectedly found to be 86.5 wt. % versus 57 wt. % as in Examples 2 and 3 hereof. It was also unexpectedly found that C5/650° F. yield increased to 58.3 wt. %. If 650/700° F. is included, the total liquid yield would be 66.4 wt. %.
All coal conversion and C5/650° F. yield data from the 17 Examples is summarized in Table 1 and in
Coal conversion for the low inertinite Rawhide and Wyodak coals was between the high 70's and high 80's depending upon operating conditions and catalyst concentration and were independent of recycle (solvent+ bottoms)/Coal ratio between zero and 2.0. For the high inertinite Yulin coal, total coal conversion was below 60 wt. % at a (Solvent+Bottoms)/Coal ratio of 1.2 to 1.5. This low coal conversion has been reported by other researchers for high inertinite content coals. Increasing the (Solvent+Bottoms)/Coal ratio to 3.0 for Yulin coal unexpectedly increased conversion to substantially the same level as that for low inertinite coals.
In addition to increasing coal conversion with high inertinite content coals, operating at high recycle (solvent+ bottoms)/Coal ratio of 3:1 also unexpectedly increased the direct coal liquefaction yield of C5/650° F. in the product. This lower boiling product can be readily converted to gasoline, diesel, and jet fuel using fixed bed upgrading reactors which are common to today's refineries. This evidences that it is possible to even convert a high inertinite content coal to an unexpectedly high yield of C5/650° F. product.
Claims
1. A process for the direct liquefaction of coal, which process is conducted in the absence of added carbon monoxide, and is performed in a coal liquefaction process plant comprising: a slurry mixing zone, a preheating zone, a liquefaction zone, a separation zone capable of separating a gaseous product stream from a liquid/solids product stream, an atmospheric fractionation zone, and a vacuum fractionation zone, which process comprises:
- a) introducing into said slurry mixing zone:
- i) coal having an average particle size of about 75 microns to about 600 microns and a moisture content from about 1 to about 4 wt. %;
- ii) effective amount of recycle non-donor solvent from said vacuum fractionation zone and non-donor bottoms from said atmospheric fractionation zone, wherein the ratio of non-donor solvent plus non-donor bottoms to coal is from about 2.5 to 1 to about 4 to 1:
- iii) a molybdenum-containing microcatalyst provided at a make-up rate that is equivalent to about 50 wppm to about 2 wt. % molybdenum on a moisture and ash free (MAF) feed coal basis;
- b) conducting the resulting slurry from step a) above at a pressure from about 1,500 psig to about 3000 psig and an effective amount of a treat gas containing at least 80 vol. % hydrogen, to said preheating zone wherein it is heated from a temperature from about 200° F. to 600° F. to a temperature of about 650° F.;
- c) conducting said heated slurry to said liquefaction zone wherein it is reacted at a temperature from about 700° F. to about 950° F. thereby producing reaction products;
- d) conducting said reaction products to a separation zone wherein a gaseous product stream is separated from a liquid/solids product stream;
- e) conducting said liquid/solids stream to said atmospheric fractionation zone, wherein it is fractionated to result in a C1 to C4 gaseous fraction, a C5/650° F. fraction, and a 650 F+ bottoms fraction;
- f) conducting an effective portion of non-donor 650° F.+ bottoms fraction from said atmospheric fractionation zone to said vacuum fractionation zone thereby resulting in a 1000° F.+ fraction and a non-donor 650° F. to 1000° F. solvent fraction from the vacuum fractionator; and leaving a remaining portion of said non-donor 650° F.+ fraction from said atmospheric fractionator;
- g) recycling the remaining portion of said non-donor 650° F.+ fraction from said atmospheric fractionation zone to the slurry mixing zone; and
- h) recycling at least a portion of said non-donor 650° F. to 1000° F. solvent fraction from said vacuum fractionation zone to said slurry mixing zone.
2. The process of claim 1 wherein the ratio of non-donor solvent plus non-donor bottoms to coal is about 3:1 to 4:1.
3. The process of claim 2 wherein the ratio of non-donor solvent plus non-donor bottoms to coal is about 3:1 to about 3.5:1.
4. The process of claim 1 wherein the amount of inertinite in said coal is from about 7 to 14 vol. %.
5. The process of claim 1 wherein the amount of inertinite in said coal is greater than about 20 vol. %.
6. The process of claim 1 wherein said treat gas contains at least about 70 vol. % hydrogen.
7. The process of claim 6 wherein said treat gas contains at least about 80 vol. % hydrogen.
8. The process of claim 1 wherein the pressure is from about 2000 psig to about 3000 psig.
9. The process of claim 1 wherein the liquefaction temperature if from about 800° F. to about 900° F.
10. The process of claim 7 wherein the pressure is from about 2000 psig to about 3000 psig and the liquefaction temperature is from about 800° F. to about 900° F.
Type: Application
Filed: Jul 27, 2017
Publication Date: Nov 9, 2017
Inventors: Richard F. Bauman (Billingham, WA), Peter S. Maa (Sugarland, TX)
Application Number: 15/662,113
