Search Images Maps Play YouTube News Gmail Drive More »
Sign in
Screen reader users: click this link for accessible mode. Accessible mode has the same essential features but works better with your reader.


  1. Advanced Patent Search
Publication numberUS3892654 A
Publication typeGrant
Publication dateJul 1, 1975
Filing dateMar 4, 1974
Priority dateMar 4, 1974
Also published asCA1014094A1, DE2431949A1
Publication numberUS 3892654 A, US 3892654A, US-A-3892654, US3892654 A, US3892654A
InventorsPastor Gerald R, Wright Charles H
Original AssigneeUs Interior
Export CitationBiBTeX, EndNote, RefMan
External Links: USPTO, USPTO Assignment, Espacenet
Dual temperature coal solvation process
US 3892654 A
A solvation process for producing deashed solid and liquid hydrocarbonaceous fuel from coal. Raw coal is slurried with a solvent comprising hydroaromatic compounds in contact with hydrogen in a first zone at a relatively high temperature to dissolve hydrocarbonaceous fuel from coal minerals by transfer of hydrogen from hydroaromatic solvent compounds to hydrocarbonaceous material in the coal. The solvent is then treated with hydrogen in a second zone at a lower temperature to replenish the solvent with hydrogen. Forced cooling of the slurry between zones accomplishes many significant improvements in the process.
Previous page
Next page
Description  (OCR text may contain errors)

United States Patent 1 1 Wright et al.

[ DUAL TEMPERATURE COAL SOLVATION PROCESS [75] Inventors: Charles H. Wright, Overland Park,

Kans; Gerald R. Pastor, Takoma, Wash.

[73] Assignee: The United States of America as represented by the Secretary of the Interior, Washington, DC.

[22] Filed: Mar. 4, 1974 [21] Appl. No.: 446,971

Harris et al. 208/8 Bull et al.

Primary ExaminerDelbert E. Gantz Assistant Examiner-James W. Hellwege 5 7 ABSTRACT A solvation process for producing deashed solid and liquid hydrocarbonaceous fuel from coal. Raw coal is slurried with a solvent comprising hydroaromatic com pounds in contact with hydrogen in a first zone at a relatively high temperature to dissolve hydrocarbonaceous fuel from coal minerals by transfer of hydrogen from hydroaromatic solvent compounds to hydrocarbonaceous material in the coal. The solvent is then treated with hydrogen in a second zone at a lower temperature to replenish the solvent with hydrogen. Forced cooling of the slurry between zones accomplishes many significant improvements in the process.

15 Claims, 7 Drawing Figures FIG.2



DUAL TEMPERATURE COAL SOLVATION PROCESS This invention resulted from work performed under Contract No. l4-0l-000l-496 between The Pittsburgh and Midway Coal Mining Co., a subsidiary of Gulf Oil Corporation, and the Office of Coal Research in the Department of the lnterior entered into pursuant to the Coal Research Act, 30 USC 661 to 668.

This invention relates to a non-catalytic liquid solvent dissolving process for producing reduced or low ash hydrocarbonaceous solid fuel and hydrocarbonaceous distillate liquid fuel, from ash-containing raw coal. Preferred coal feeds contain hydrogen, such as bituminous and sub-bituminous coals, and lignites. The process produces deashed solid fuel (dissolved coal) together with as much coal derived liquid fuel as possi ble, with an increase in liquid fuel product being accompanied by a decrease in solid fuel product. Liquid fuel is the more valuable product but the production of liquid fuel is limited because it is accompanied by pro duction of undesired by-product hydrocarbon gases. Although liquid fuel is of greater economic value than deashed solid fuel, hydrocarbon gases are of smaller economic value than either deashed solid fuel or liquid fuel and have a greater hydrogen to carbon ratio than either solid or liquid fuel so that their production is not only wasteful of other fuel product but is also wasteful of hydrogen.

Hydrocarbon gases are produced primarily by hydrocracking, and since their production is undesired in this process no external catalyst is employed, since catalysts generally impart hydrocracking activity in a coal solvation process.

When raw coal is subjected to solvation at a relatively low temperature, the dissolved product comprises in major proportion a high molecular weight fuel which is solid at room temperature. When the mixture of solvent and dissolved coal is subsequently filtered to remove ash and undissolved coal and the filtrate is then subjected to vacuum distillation, this high boiling solid fuel product is recovered as the vacuum bottoms. This deashed vacuum bottoms product is referred to herein as either vacuum bottoms or deashed solid fuel product. This vacuum bottoms is cooled to room temperature on a conveyor belt and is scraped from the belt as fragmented deashed hydrocarbonaceous solid fuel.

As the temperature of the solvation process is progressively increased, the vacuum bottoms (deashed solid fuel), which is a high molecular weight polymer, is converted to lower molecular weight hydrocarbonaceous liquid fuel which is chemically similar to the process solvent and which has a similar boiling range. The liquid fuel product is in part recycled as process solvent for the subsequent pass and is referred to herein as either liquid fuel product or excess solvent. Production of liquid fuel occurs by depolymerization of solid fuel through various reactions, such as removal therefrom of heteroatoms, including sulfur and oxygen. As a result of the depolymerization reactions, the liquid fuel has a somewhat higher hydrogen to carbon ratio than the solid fuel and therefore exhibits a correspondingly higher heat content upon combustion. It is desirable in the process to convert as much of the vacuum bottoms (solid fuel) product to solvent boiling range (liquid fuel) product, since liquid fuel is economically more valuable than solid fuel. As the temperature of solvation continues to be increased, an increasing proportion of vacuum bottoms fuel is converted to solvent boiling range fuel until a temperature is reached at which conversion of vacuum bottoms to liquid fuel occurs only at the price of excessive and wasteful production of relatively hydrogen-rich by-product hydrocarbon gases due to the onset of excessive thermal hydrocracking. The present process produces 20 or 40 to weight percent of deashed solid fuel on an MAF (moisture and ash free) basis, the remaining product being primarily liquid fuel.

it is the purpose of the present invention to avoid thermal hydrocracking as much as possible and at least to the extent of avoiding excessive production of hydrocarbon gases since production of gases diminishes the yield of desired deashed solid fuel and liquid fuel products. This purpose is accomplished according to the present invention by performing the solvation process in two separate stages, each stage employing a dif ferent temperature. In one embodiment of this invention, less than 6 weight percent of hydrocarbon gases, based on MAF coal feed, is produced. The production limit of hydrocarbon gases establishes the production limit of liquid fuel product and therefore also the production limit of solid fuel product.

A further and very important advantage of the dual temperature method of this invention is that a high temperature stage is made possible whereby product sulfur level can be reduced. Relatively high temperatures are required for sulfur removal whereas temperatures below the required level are not as effective for sulfur removal. The high temperatures required for effective sulfur removal also induce hydrocracking but the hydrocracking reaction is more time dependent and by rapid reduction of the high process temperature reduction of sulfur level is achieved with a minimum of hydrocracking.

The first reactor stage of the present process is a tu bular preheater having a relatively short residence time in which a slurry of feed coal and solvent in essentially plug flow is progressively increased in temperature as it flows through the tube. The tubular preheater has a length to diameter ratio of at least 100, generally and at least 1,000, preferably. A series of different reactions occur within a flowing stream increment as the temperature of the increment increases from a low inlet temperature to a maximum or exit temperature, at which it remains for only a short time. The second reac tor stage employs a relatively longer residence time in a larger vessel maintained at a substantially uniform temperature throughout. An important feature of this invention is that a regulated amount of forced cooling occurs between the stages so that the second stage temperature is lower than the maximum preheater temperature. Although the preheater stage is operated with plug flow without significant backmixing, full solution mixing with a uniform reactor temperature occurs in the dissolver stage. Data presented below show that the split temperature coal dissolving process of this invention results in high conversion of raw coal to deashed solid fuel and liquid fuel and the proportion of liquid to solid fuel product is enhanced while avoiding excessive production of by-product hydrocarbon gases. It is shown below that these results are better accomplished by employing a split temperature process than by employing a process having a uniform temperature in the two stages, even when the uniform temperature is the same as either temperature ofa split temperature operation.

The coal solvent for the present process comprises liquid hydroaromatic compounds. The coal is slurried with the solvent for charging to the first or preheater stage. In the first stage. hydrogen transfer from the solvent hydroaromatic compounds to coal hydrocarbonaceous material occurs resulting is swelling of the coal and in breaking away of hydrocarbon polymers from coal minerals. The range of maximum temperatures suitable in the first (preheater) stage is generally 400 to 525C. or preferably 425 to 500C. If there is inadequate facilities to handle hydrocarbon gaseous byproducts, the upper temperature limit should be 470C.. or below in order to minimize production of gaseous product. The residence time in the preheater stage is generally 0.01 to 0.25 hours, or preferably 0.01 to 0.15 hours.

In the second (dissolver) stage of the process of this invention. the solvent compounds, which have been depleted of hydrogen and converted to their precursor aromatics by hydrogen donation to the coal in the first stage, are reacted with gaseous hydrogen and reconverted to hydroaromatics for recycle to the first stage. The temperature in the dissolver stage is 350 to 475C. generally. and 400 to 450C, preferably. The residence time in the dissolver stage is 0.1 to 3.0 hours. generally, and 0.15 to 1.0 hours, preferably. The temperature in the dissolver stage is lower than the maximum temperature in the preheater stage. Any suitable forced cooling step can be employed to reduce stream temperature between the preheater and the dissolver. For example, makeup hydrogen can be charged to the process between the preheater and dissolver steps or a heat exchanger can be employed. Also, the residence time in the preheater is lower than the residence time in the dissolver.

The liquid space velocity for the process (volume of slurry per hour per volume of reactor) ranges from 0.2 to 8.0, generally, and 0.5 to 3.0. preferably. The ratio of hydrogen to slurry ranges from 200 to 10,000 stan dard cubic feet per barrel, generally, and 500 to 5,000 standard cubic feet per barrel. preferably, (3.6 to 180, generally. and 9 to 90. preferably, SCM/lOOL). The weight ratio of recycled solvent product to coal in the feed slurry ranges from 0.5:] to :1, generally, and from 1.0:1 to :1. perferably.

The reactions in both stages occur in contact with gaseous hydrogen and in both stages heteroatom sulfur and oxygen are removed from solvated deashed coal polymers. resulting in depolymerization and conversion of dissolved coal polymers to desulfurized and deoxygenated free radicals of reduced molecular weight. These radicals radical have a tendency to repolymerize at the high temperatures reached in the preheater stage. but at the reduced temperature of the dissolver stage of this invention these free radicals tend to be stabilized against repolymerization by accepting hydrogen at the free radical site. The present process can employ carbon monoxide and steam together with or in place of hydrogen since carbon monoxide and steam react to form hydrogen. The steam can be derived from feeding wet coal or can be injected as water. The reaction of hydrogen at the free radical site occurs more readily at the relatively low dissolver temperature than at the higher preheater exit temperature.

The solvent used at process start-up is advantageously derived from coal. lts composition will vary, depending on the properties of the coal from which it is derived. In general. the solvent is a highly aromatic liquid obtained from previous processing of fuel. and generally boils within the range of about C. to 450C. Other generalized characteristics include a density of about 1.1 and a carbon to hydrogen mole ratio in the range from about 1.0 to 0.9 to about 1.0 to 0.3. Any organic solvent for coal can be used as the start-up solvent in the process. A solvent found particularly useful as a start-up solvent is anthracene oil or creosote oil having a boiling range of about 220C. to 400C. However, the start-up solvent is only a temporary process component since during the process dissolved fractions of the raw coal constitute additional solvent, which. when added to start-up solvent, provides a total amount of solvent exceeding the amount of start-up solvent. Thus, the original solvent gradually loses its identity and approaches the constitution of the solvent formed by solution and depolymerization of the coal in the process. Therefore, in each pass of the process after startup, the solvent can be considered to be a portion of the liquid product produced in previous extraction of the raw coal.

The residence time for the dissolving step in the preheater stage is critical in the process of this invention. Although the duration of the solvation process can vary for each particular coal treated, viscosity changes as the slurry flows along the length of the preheater tube provide a parameter to define slurry residence time in the preheater stage. The viscosity of an increment of feed solution flowing through the preheater initially increases with increasing increment time in the preheater, followed by a decrease in viscosity as the solubilizing of the slurry is continued. The viscosity would rise again at the preheater temperature, but preheater residence time is terminated before a second relatively large increase in viscosity is permitted to occur. An advantageous means for establishing proper time for completion of the preheater step is use of the Relative Viscosity" of the solution formed in the preheater, which is the ratio of the viscosity of the solution formed to the viscosity of the solvent, as fed to the process, both vis cosities being measured at 99C. Accordingly, the term Relative Viscosity" as used herein is defined as the viscosity at 99C.. of an increment of solution, divided by the viscosity of the solvent alone fed to the system measured at 99C., i.e.

Relative Viscosity Viscosity of Solution at 99C.}Viscosity of Solvent at 99C.

The Relative Viscosity can be employed as an indication of the residence time for the solution in the pre heater. As the solubilizing of an increment of slurry proceeds during flow through the preheater, the Relative Viscosity of the solution first rises above a value of 20 to a point at which the solution is extremely viscous and in a gel-like condition. ln fact, if low solvent to coal ratios are used. for example. 0.5:1. the slurry would set up into a gel. After reaching the maximum Relative Viscosity. well above the value of 20, the Relative Viscosity" of the increment begins to decrease to a minimum, after which it has a tendency to again rise to higher values. The solubilization proceeds until the decrease in Relative Viscosity (following the initial rise in Relative Viscosity) falls to a value at least below 10, whereupon the preheater residence time is terminated and the solution is cooled and passed to the dissolver stage which is maintained at a lower temperature to prevent the Relative Viscosity fromagain rising above 10. Normally, the decrease in Relative Viscosity will be allowed to proceed to a value less than 5 and preferably to the range of 1.5 to 2. The conditions in the preheater are such that the Relative Viscosity will again increase to a value above 10, absent abrupt termination of preheater exit conditions, such as a forced lowering of temperature.

When a slug of hydroaromatic solvent and coal first experience heating in the preheater, the first reaction product is a gei which is formed in the temperature range 200 to 300C. Formation of the gel accounts for the first increase in Relative Viscosity. The gel forms due to bonding of the hydroaromatic compounds of the solvent with the hydrocarbonaceous material in the coal and is evidenced by a swelling of the coal. The bonding is probably a sharing of the solvent hydroaromatic hydrogen atoms between the solvent and the coal as an early stage in transfer of hydrogen from the solvent to the coal. The bonding is so tight that in the gel stage the solvent cannot be removed from the coal by distillation. Further heating of a slug in the pre heater to 350C. causes the gel to decompose, evidenc ing completion of hydrogen transfer, producing a deashed solid fuel, liquid fuel and gaseous products and causing a decrease in Relative Viscosity.

A decrease of Relative Viscosity in the preheater is also caused by depolymerization of solvated coal polymers to produce free radicals therefrom. The depolymerization is caused by removal of sulfur and oxygen heteroatoms from hydrocarbonaceous coal polymers and by rupture of carboncarbon bonds by hydrocracking to convert deashed solid fuel to liquid fuel and gases. The depolymerization is accompanied by the evolution of hydrogen sulfide, water, carbon dioxide, methane, propane, butane, and other hydrocarbons.

At the high temperatures of the preheater outlet zone, repolymerization of free radicals is a reaction which is favored over hydrogenation of free radical sites and accounts for the final tendency towards in crease in Relative Viscosity in the preheater to a value above ll). This second increase in Relative Viscosity is avoided in accordance with the present invention. The elimination of sulfur and oxygen from the solvated (leashed solid fuel is probably caused by stripping out of these materials by thermal rupture of bonds leaving free radical molecular fragments which have a ten dency towards subsequent repolymerization at elevated temperature conditions. The drop in stream temperature by forced cooling following the preheater step in accordance with this invention tends to inhibit polymer formation. The observed low level of sulfur in the liquid fuel product, which for one coal feed is about 0.3 weight percent, as compared to 0.7 weight percent in the vacuum bottoms (solid fuel) product, indicates that sulfur is being stripped out of the solid fuel product leaving low sulfur smaller molecular fragments as free radicals.

We have found that maximum or exit preheater temperatures should be in the range of 400 to 525C. The residence time in the preheater for a feed increment to achieve this maximum temperature is about 0.01 to 0.25 hours. At this combination of temperature and residence time. coke formation is not a problem unless flow is stopped, that is, unless the residence time is increased beyond the stated duration. The hydrocarbon gas yield under these conditions is less than about 6 weight percent while excess solvent (liquid fuel) yield is above l() or l5 weight percent. based on MAF coal feed, while the solid fuel product is above 20 weight percent. High production of gases is to be avoided because such production involves high consumption of hydrogen and because special facilities are required. However, a gaseous yield above 6 weight percent can be tolerated if facilities to handle or store and transport the gas are available.

The relatively low sulfur content in the vacuum bottoms (deashed solid fuel) product of the present process is an indication that the reaction proceeds to a high degree of completion. It is also an indication that the vacuum bottoms product has been chemically released from the ash so that it can be filtered therefrom.

The hydrogen pressure in the present process is 35 to 300 kglcm generally, and 50 to 200 kg/cm, preferably. At about kg/cm hydrogen pressure, the solvent hydrogen content tends to adjust to about 6.] weight percent. If the hydrogen content of the solvent is above this level, transfer of hydroaromatic hydrogen to the dissolved fuel tends to take place. increasing production ofliquid fuel, which has a higher hydrogen content that solid fuel. If the solvent contains less than 6.1 weight percent of hydrogen, the solvent tends to acquire hydrogen from hydrogen gas at a faster rate than the fuel product. Once the solvent is roughly adjusted to a stable hydrogen level, conversion appears to depend on the catalytic effect of FeS. derived from the coal ash. Some deviations from this basic situation are observed in response to temperature and time variables. Higher temperatures tend to lower the by droaromatic content of the system while rapid feed rates may preclude attainment of equilibrium values (not sufficient time). in addition. higher pressures tend to favor more rapid equilibrium and tend to increase the hydroaromatic character of the system.

In the dissolver stage of the present process, aromatic compounds which have surrendered hydrogen in the preheater are reacted with hydrogen to again form hydroaromatic compounds. Hydroaromatic compounds are partially (not completeiyl saturated aromatics. The chemical potential in the dissolver is too low for full saturation of aromatics to be a significant reaction. This is important because while hydroaromatics are ca pabie of hydrogen transfer, saturated aromatics are not. Most of the saturates observed in the dissolver tend to be light products derived from ring opening of iiq uid product or solvent, or derived from aliphatic side chain removal. Solid fuel product aromatic species tend to remiari aromatic or hydroaromatic.

This invention is based upon the use of the effect of time in conjunction with the effect of temperature in the preheater stage. It is based upon the discovery that the desired temperature effect in the prebeater stage is substantially a short time effect while the desired temperature effect in the dissolver requires a relatively longer residence time. The desired low preheater residence times are accomplished by utilizing an elongated tubular reactor having a high length to diameter ratio of at least 100. generally. and at ieast L000, preferably, so that rapidly upon reaching the desired maximum preheater temperature the preheater stream is discharged and the elevated temperature is terminated by forced cooling. Forced cooling can be accomplished by hydrogen quenching or by heat exchange. Thereupon, in the dissolver stage, wherein the temperature is lower, the residence time is extended for a duration which is longer than the preheater residence time.

The data in Table 1, show that there is an adverse effeet in employing excessively high preheater temperatures.

TABLE 1 TEST NUMBER 1 2 3 4 5 H, Pressure. kg/cm" 70 70 70 70 70 Max. Preheater Temp..C. 450 500 450 450 475 Dissolver Temp..C. 450 450 425 425 425 l/LHSV; Hr. 0.52 0.98 1.79 1.89 1.79 GHSV 304 239 342 342 342 Solvent/MAP Coal/H O (wt) 2.50/1/008 2.49/1/006 2.49/1/005 2.49/1/.05 2.49/1/00 Ash in Feed Slurry 5.0 5.0 5.285 7.42 10.65 11: Coal Derived Feed 33.3 33.3 34.8 48.7 69.9

YlELDS ON MAF COAL BASIS-1b CO 0.23 0.42 0.51 0.27 0.28 CO- 1.12 1.20 0.64 0.68 [1,28 H 5 2.32 2.12 2.04 1.62 1.95 Hydrocarbon Gas 5.28 8.89 5.73 5.80 7.1 1 Gas Not Identified 12.87 H 0 360 4.10 3.82 1.22 1.81 Excess Solvent 5.36 15.10 31.98 62.08 49.37 Vacuum Bottoms 68.12 56.81 48.66 30.36 21.06 lnsol. Organic Matter 14.91 13.83 1 1.59 4.99 9.07

TOTAL 100.94 102.47 104.97 107.92 103.20


Recovery. weight 71: 97.94 96.59 95.91 92.63 93.35 MAF Conversion, weight 7r 85.09 86.17 88.41 95.01 90.93

COMPOSlTlON OF LIQUID AND VACUUM BOTTOM FUEL PRODUCT Carbon. Weight 89.68 89.40 89.72 90.65 Hydrogen, weight 9: 5.94 5.93 6.20 6.54 Nitrogen, weight 70 0.979 1.06 1.15 1.31 Sulfur, weight 0.46 v 0.410 0.420 0.438 Oxygen. weight 4.13 5.00 2.51 1062 VACUUM BOTTOMS FUEL PRODUCT COMPOSITION Carbon, weight 87.32 89.03 88.54 88.71 91.12 Hydrogen. weight 70 5.11 5.12 4.74 5.35 5.10 Nitrogen, weight 70 1.91 2.02 2.22 2.10 2.22 Sulfur. weight 7c 0.944 0.719 0.676 0.606 0.488 Oxygen, weight 70 4.58 3.04 3.619 3.156 1.00 Ash. weight 0.133 0.067 0.205 0.078 0.075

As shown in Table 1, an increase in maximum pre heater temperature from 450C. to 475C. or 500C results in an increased yield of hydrocarbon gases to a level above 6 weight percent based on MAF coal feed. A 6 weight percent hydrocarbon gas yield on a MAP basis is a suitable upper limit for gas production unless gas handling facilities are available. Not only are hydrocarbon gases considerably lower economic value than liquid and deashed solid fuel product. but they contain a considerably higher ratio of hydrogen to carbon than either liquid or deashed solid fuel product. Therefore. excessive production of hydrocarbon gases not on'i'. signifies a depressed yield of liquid and solid fuel prod uct but also constitutes an unnecessary consumption of valuable process hydrogen due to hydrocracking of higher molecular weight fuel to produce the by-product hydrocarbon gases. Table 1 shows that preheater temperatures of 475C. and 500C. both result in a hydro ture is below 475C. (for example, below 470C). Table 1 shows that the use of a preheater temperature of 450C. and a dissolver temperature 0f425C., rather than a temperature of 450C. in both stages, results in an increase in the ratio of excess solvent (liquid fuel product) yield to vacuum bottoms (solid fuel product) yield, an increase in conversion of MAP coal, and a reduction in the sulfur content of the excess solvent plus vacuum bottoms product and of the vacuum bottoms product itself. These data clearly show the considerable advantages realized by employing a higher temperature "1 the preheater exit than in the dissolver, as compared employing a uniform temperature at the preheater exit and in the dissolver, even though the split temperature results in an overall lower average temperature in the process.

The data of Table 1 show that when employing a maximum preheater temperature of 475C. and a lower dissolver temperature, the ratio of solvent to vacuum bottoms yield, the percent MAF conversion, the total liquid and vacuum bottoms sulfur content and the vacuum bottoms sulfur content all improved when compared to a test employing a uniform maximum preheater temperature and dissolver temperature, but the improvement is achieved at the expense of excessive hydrocarbon gas yield, which ranged above 6 weight percent. Based upon the data of Table l, the preheater temperature preferably should not be higher than 460 or 470C. The preheater effluent should preferably be quenched at least about 25C. and as much as 50 or 100C. before entering the dissolver. In some cases a smaller extent of cooling, such as to at least l5 or C. below the maximum or outlet preheater temperature can be effective.

Test 4 of Table 1 shows an especially advantageous split temperature test of the present invention because the liquid product yield (excess solvent) is greater than the vacuum bottoms yield (solid deashed fuel product) ing. considerable backmixing occurs in the dissolver which contributes to a uniform temperature throughout the dissolver.

In the dissolver, the reactions occurring require a temperature lower than the maximum preheater temperature. Rehydrogenation of the aromatics in the solvent to replenish hydrogen lost from the solvent by hydrogen donation reactions in the preheater requires a longer residence time than is required in the preheater. but proceeds at a temperature lower than the preheater temperature. After the solvent is hydrogenated in the dissolver to reconvert aromatics to hydroaromatics. it is in condition to be recycled to the next preheater pass for hydrogen donation reactions. A coincident reaction which occurs in the dissolver in addition to formation of hydroaromatics is the removal of additional sulfur from the extracted coal. The relatively higher preheater temperatures are more effective for sulfur removal than the lower dissolver temperatures. However. some of the sulfur cannot be removed at the low resi- TABLE 2 TEST NUMBER l 2 3 4 5 H2 Pressure. kg/cm 7U 70 70 70 70 Max. Preheater Temp..C. 450 450 450 450 LHSV 28.36 28.36 28.35 l5.23 7.74 GHSV 3035 2964 2953 30l2 3083 l/LHSV; Hr. 0.035 0.035 0.035 0.066 0.12)

YIELDS ON MAF COAL BASIS% CO 0.03 0.03 0.07 0.00 (1.25 CD. 0.35 0.35 0.35 (1.45 0.51 H. .S 0.94 (1.97 1.73 1.49 2.66 Hydrocarbon Gas 0.3] 0.31 0.2! 0.65 0.76 H. 0.66 l .42 L1 1 0.9! 4). l 2 Excess Solvent -38.47 59.30 -39! 7.84 l [.80 Vacuum Bottoms 65.57 l36.2l I053! 75.78 69.93 Insoluble Organic Matter 71.06 23.20 31.34 13.21 15.23 TOTAL 100.45 100.35 r0055 |00.3Q 1mm DATA Recovery. weight 94.14 95.46 96.80 90.68 951 4 MAF Conversion. weight 28.94 76.80 (18.66 86.79 84.77

VACUUM BOTTOMS FUEL PRODUCT PROPERTlES Carbon. weight L72 84.00 Hydrogen. weight 71 5. 1 Nitrogen. weight L93 Sulfur. weight L38 Oxygen, weight .5 Ash. weight 72: 0.39

and the hydrocarbon gas yield is relatively low. dence time of the preheater. but requires an extended Higher temperatures can be advantageously emresidence time. Therefore, additional sulfur in the coal ployed in the preheater than in the dissolver only in product is removed during the extended residence time conjuction with a lower residence time in the preheater utilized in the dissolver. A third and highly important than in the dissolver. Most of the residence time in the reaction occurring in the dissolver is the addition of hypreheateris employed in heating the coal-solvent slurry drogen to free radicals formed in both the preheater mixture to the maximum preheater temperature, which and the dissolver to arrest polymerization of molecular is the exit preheater temperature. Preheater reactions fragments to high molecular weight material. are rapid and tend to occur at the required maximum Table 2 shows the results of preheater tests which temperature with a low residence time. On the other were all conducted at 450C. Certain of the tests were hand. the reactions which occur in the dissolver are conducted at a very low preheater residence time of slower reactions. Therefore, the dissolver not only op- 0035 hours and other tests were performed at someerates at a lower temperature than the maximum preheater temperature but also at a longer residence time. Although the preheater substantially avoids backmixwhat longer preheater residence times.

As shown in Table 2. at the lowest residence time tests there is a net loss ofliquid solvent (due to binding 1 l of the solvent in a gel). Also the percent conversion of MAF coal is low in the low residence time tests. However, at the higher residence times indicated in Table 2 there is a net production of solvent and the percent Table 4 shows tests performed at a relatively mild preheater temperature of 450C. As shown in Table 4, even at the moderate preheater temperature of 450C, lengthy preheater residence times result in hydrocarconversion of MAP coal is considerably higher. There- 5 bon gas yields above 6 weight percent. Table 4 shows fore, it is apparent that while excessive preheater resithat as preheater residence times increase from about dence time is detrimental in that repolymerization will 0,5 to about 1.3 hours at a constant preheater temperaoccur {as evidenced y a Second in ase in Relative ttir of 450C. the yield of vacuum bottoms (solid fuel) Viscosity to a level above 10), a deficiency in residence product gradually decreases while the yield of solvent time in the preheater at a deslred preheater temto (liquid fuel) product gradually increases. together with perataure is also undersirable in that the process cana disadvantageous increase i h d b gas i 1d not sustain its solvent needs and in that there is inade Al h h not shown b h data, h h d Content quflte Conversion of Coal- Solvem be pp from of solvent produced at a low residence time is low as an external source only at process start-up. and once Compared to Solvent d d at a hi h residence the process achieves equilibrium It Will satisfy its own i5 i at u givgn {temperature so h Solvent producd Solvent requirements and not depflnd p an external at a low residence time tends to be of low quality of hy- Solvem pp ydrogen donation in the preheater stage. Table 4 shows The data in Table 3 illustrate the interchangeability that a 450C. preheater temperature coupled with a of time and temperature in preheater operation. 2 low residence time results in a low yield of hydrocarbon TABLE 3 TEST NUMBER 1 2 3 4 5 6 H Pressure, kg/cm 7o 70 70 70 70 70 Max. Preheater Temp..C. 475 475 475 500 500 450 LHSV 27.57 13.65 7.8] 28.36 7.68 7.74 GHSV 3102 3126 3095 2988 3066 10x3 HLHSV; Hr. 0.036 0.073 0.128 0.035 0.130 0.129

YlELDS ON MAF COAL BASlS-% Co 0.14 0.36 0.25 0.17 0.13 0.25 (0 0.46 0.58 0.63 0.66 0.77 0.51 H I57 259 2.0l l.77 2.94 2.66 Hydrocarbon Gas Product 0.62 1.29 3.27 1.85 5.11 0.76 -2.43 0.29 0.63 150 0.64 0.12

Excess Solvent (Liquid Ftiei Product) 1 1.13 13.09 17.11 13.43 25.45 1 1.80 Vacuum Bottoms (Solid Fuel Product) 101.97 69.57 63.15 71.17 52.56 69.93 insoluble Organic Matter 9.09 [3.38 l4.2l l2.83 I3.l7 l5.23

TOTAL 100.29 101 15 101.26 100.38 100.77 101.02


Recovery. weight 99.85 98.30 96.75 98.38 96.79 95.94 MAP Conversion. weight 90.91 86.62 85.79 87.17 86.83 84.77

VACUUM BOTTOMS (SOLID FUEL) PROPERTIES Carbon, weight 84.62 84.87 86.67 88.30 84.00 Hydrogen. weight 5.50 5.42 5.33 4.96 5.7l I Nitrogen. weight 7! 1.90 2.04 1.78 2.10 1.93 Sulfur, weight 9r 1.34 1.29 1.21 0.80 1.33 Oxygen. weight 7r 6.23 5.42 4.90 4.19 6.59 Ash. weight "71 0.35 0.96 0.1 1 0.15 0.39

Table 3 shows a test performed at 475C. employing gases. These data Illustrating the effect of preheater the very low preheater residence time of 0.036 hour in residence time indicated that with increasing preheater which there is a net loss of solvent in the process. The residence time at a preheater temperature of 450C., loss is probably caused by the solvent being bonded in there is a continuous conversion of vacuum bottoms a gel with the coal from which the solvent has insuffi- 60 (solid fuel) product to solvent (liquid fuel) product, accient time to become disengaged. and from which the companied by a continuous conversion of product to solvent cannot be separated by distillation. However, hydrocarbon gases.

Table 3 shows that at 475C. there in a net production The data of Table 3 show that at high maximum preof solvent in the process when the preheater residence heater tempertures of 475 and 500C. the effect of restime is increased. Table 3 further shows that if the temidence time tends to be greater than at lower preheater perature is increased to 500C. the residence time can be reduced again while obtaining a high production of solvent in the process.

temperatures. Table 3 shows that at 475C. a very low residence time of 0036 hours resulted in highly incomplete conversion and a loss of solvent in the process. As

13 the residence time increased at 475C, conversion increased as indicated by a net production of solvent in the process together with a decrease in vacuum bottoms (solid fuel) yield. Table 3 shows that at 500C.

solid fuel of all the tests of Table 3 (84.6 percent). and this yield is greater than that shown in Tests 3 and be cause of the relatively low yield of hydrocarbon gases in Test 4.

with a 0.036 hour residence time, the increase in tem- 5 Comparing Tables 3 and 4, it is seen that a maximum perature tends to compensate for the low residence preheater temperature of 450C. is sufficiently low that time so that there is a net production of solvent in the lengthy preheater residence times of more than 0.5 process, and the yield is increased at 500C. (accompahour are required to achieve significantly high yields of nied by a decrease in vacuum bottoms yield) by inexcess solvent (liquid fuel product). At higher precreasing the residence time to 0.130 hour. At a resil0 heater temperatures of 475 and 500C, higher excess dence time of 0.130 hour a greater solvent yield and a solvent yields are achieved at lower residence times. lower vacuum bottoms yield is achieved at 500C. than M t importantly, at the elevated preheater temperaat 475C. except that hydrocarbon gas yield is higher tures of Table 3, excess solvent yield is much more sen in the 500C test- F ompa ti p p Table 3 sitive to slight changes in residence time than at 450C. shows a test conducted at 450C. and about the same At 450C, large increases in residence time are reresidence time wherein the Solvent yield s ower and quired to achieve significant increases in solvent yield. the acuum ttOm yield i5 higher ha in m6 50(PCv (3f all the tests in Tables 3 and 4, the combination in test. Furthermore. the 500C. test performed at 0.130 T ble 3 of a preheater temperature of 500C. and a hour Produccd a Vacuum bottoms PmduCt having y preheater time ofO. l30 hour achieved the highest sol- 0.8 percent sulfur which is the lowest sulfur level vacm i ld d h l t va uu b tt m yi ld. Th uum bottoms PTOClUCt Of all the tests Of Table HOW- daia how that that solvent was reduced by one- CV81, the test at the ilOUI' residence time a{ a lower residence time at a preheater tempgra- ShOWS the amount of Vacuum bmmms Produced ture of 500C. and was also reduced by about one-third mlflished faVOT 0f not y liquid Product but 3150 at about the same residence time but at a lower pregaseous product. As the vacuum bottoms level diminheater temperature of 475.. while it was reduced by ishes, it is seen that conversion to gases becomes mmore h ha]f at the Same residence time and a cr asingly favored. and this therefore lends ultlstill lower preheater temperature of 450C. It is seen y limit the extend 0f Conversion of Vacuum that there is a considerable interdependence between toms to liquid product. HO ev r, f facllllles are avallpreheater temperature and preheater residence time. able to Colle t a d purify gas u P the hlgll An important distinction between 500C. preheater level of hydrocarbon gas production can be advantaoperation at which the highest solvent yields are geously utilized as a commercial fuel. achieved as compared with 450C. temperature pre TABLE 4 TEST NUMBER l 2 3 4 5 H2 Pressure. kg/cm" 70 70 70 70 70 Max. Preheater Temp..C 450 450 450 450 450 LHSV 1.96 1. 1 1.34 1.09 0.7 GHSV 208 225 31 228 235 l/LHSv; Hr. 0.510 0.524 0.746 0.917 1.351


CO 0.21 0.18 0.47 0.25 0.37 CO. 1.26 1.11 1.45 1.02 1 37 H5 1.74 2.58 1.50 2.42 2.86 Hydrocarbon Gas Product 4.8) 3.36 6.3! 5.34 8.83 H 0 6.00 7.08 1.81 3.75 2.30 Excess Solvent (Distillate Fuel Product) 6.00 11.73 14.74 16.98 18.97 Vacuum Bottoms (Solid Fuel Product) 69.78 65.02 64.30 61.32 58.83 Insoluble Organic Matter 1 1.10 10.62 10.60 10.12 9.64

TOTAL 100.98 102.18 101.24 101.20 103.17


MAF Conversion. weight *7 88.90 89.38 89.40 89.88 90.36

VACUUM BOTTOMS PROPERTIES Carbon. weight a 87.94 86.85 87.57 87.57 88.57 Hydrogen. weight it 5.07 5.47 5.44 5.37 5.24 Nitrogen, weight "71 2.02 2.07 L96 2.0l 2.01 Sulfur, weight 6 1 1.04 1.01 0.85 0.80 0.70 Oxygen. weight '71 3 82 4.40 4.05 4.00 3.36 Ash. weight 7r 0.1 1 (H4 0.13 0.25 0.12 H/C Atomic Ratio 0.346 0.379 0.373 0.368 0.355

Table 3 indicates that at temperatures as high as 475C. or 500C, gas production can exert a limitation on total fuel product (liquid fuel plus solid fuel). Test 4 ofTable 3 realized the highest yield of both liquid and heater operation, is that the low temperature operation is not conducive to as rapid polymerization of free radicals produced in the solvation operation. Tables 3 and 4 show that solvent-insoluble organic matter. which tends to be produced by free radical polymerization in the process and which decreases desired fuel product, tends to be higher in the 500C. tests than in the 450C. tests, even though the preheater residence times are during onset of hydrogen transfer. This binding is so tight that. the solvent involved in the gel cannot be distilled from the get at this stage of the reaction. As the temperature of the increment continues to increase very long in the 450C. tests. Furthermore, very careful 5 along the length of the preheater to a level of about control of residence time in the preheater at 500C. ophy r ge transfer o the Sol e t t the Coal eration is required if plugging of the tubular preheater proceeds further to an extent that the gel is broken and due to coke formation is to be avoided. which is a less high viscosity hydrocarbon polymer is dissolved out of sever problem in 450C. operation. the coal and enters into solution with the solvent, caus- Tabk 4 hows data to illustrate the effect upgn prod l0 ll'lg the Relative Viscosity of the solvent SOlUtlOl'l C0"- uct sulfur level when increasing residence time in the mining this p y to decilme to a Value below preheater at a constant preheater temperature of With continued flow to a higher preheater temperature 450C, region, heteroatom sulfur and oxygen are removed AS Shown In Table he p n MAF (JOHVBYSIOT! IS from the dissolved polymer causing the polymer to c:- substantially maximized at all the residence times l polymerize and form free radicals, resulting in a further tested. Ho r. pro lf r yield n g ous y drop in the viscosity of the solution. Ultimately, the decrease with increasing residence imes Relative Viscosity drops to a value below generally, Table 5 shows the results obtained when varyzng the or to a value b l 5, or even b l 2, Th f di. outlet or maximum preheater temp r ure WIthOI-It cals of depolymerized coal will tend to repolymerized varying total preheater residence time. under the elevated temperature conditions near the TABLE 5 TEST NUMBER 2 3 4 5 H2 Pressure, lag/cm" 7O 70 70 70 70 Max. Preheater Temp.,"C. 200 300 350 400 450 LHsv 28.09 28.09 27.96 28.36 28.36 GHSV 2978 2978 978 2987 1035 WWW; Hr. 0035 0.035 0.036 0.035 0.035

YlELDS ON MAP COAL BASIS-5% CO 0.00 000 0.00 0.24 003 C0 0.07 0.14 0.11 0.28 035 H s 0.04 0.07 0.07 0.66 0.94 Hydrocarbon Gas Product 0.00 000 0.00 0.03 0.31 H .0 0.95 0.25 0.35 0.147 0.66 Excess Solvent (Liquid Fuel Product) 153.a1 172.31t -l34145 129.56 -3a47 Vacuum Bottoms (Solid Fuel Product) 147.91 193.61 167.60 017.51% 65.57 Insoluble Organic Matter 106.101 78.52 6703 40.21 71.06 TOTAL 100.00 100.21 100.21 100.31 100.45


Recovery. weight 95.2l H858 27.02 9219 9414 MAF Conversion, weight I? 61.84 2 l .48 32.97 59.79 28.94

As shown in Table 5, at a constant residence time of preheater exit unless the holding time in the preheater 0.035 hour, solvent is consumed due to gel formation is terminated and the solution temperature is forced to at low preheater temperatures. The solvent loss tends drop. At this time, the preheater stream is withdrawn to diminish with elevation of preheater temperatures, from the preheater, quenched or otherwise cooled and but even at higher preheater temperatures the employpassed to a lower temperature dissolver stage before merit of extremely low residence times does not permit the free radicals can repolymerize to an extent that the complete breaking of the gel and release of the solvent. Relative Viscosity increases again to a value above l0. The data of Table 5 show that adequate residence time Table 6 shows the results of tests conducted with must lapse to provide a net production of solvent. m xim m preheater temperatures f 450C, d whereby the process can be self sustaining in solvent. 500C. and with variable preheater residence times. Minimum preheater residence times must be adequate As shown in Table 6. at the preheater temperature of to at least achieve a net production of solvent. 450C. and the low residence time of 0.035 hour, there The data in Table 5 illustrate the earlier explanation is a net loss of solvent. Table 6 shows that the preheater of rise in Relative Viscosity as the feed slurry begins its is capable of a net production of solvent either by transit through the preheater to plug flow. The increase lengthening the residence time at a preheater temperain Relative Viscosity in a stream increment to a valu: ture of 450C. or by increasing the final preheater temabove 20 is due to formation ofa gel between the pui' c'ftillrC to 500C, without increasing the residence verized or ground coal feed and the solvent as the tort 3. l able 6 illustrates the interchangeability of preperature of the coaLsolvent slurry starts to rise in iii: i'iuitltt temperature and preheater residence time.

preheater. This gel is caused by the onset of hydrogen donation from the hydro-aromatic solvent to the coal and the gel forms by binding of the solvent and coal FIGS. through 6 illustrate the effects of varying certain parameters in the process of this invention. FIG. 7 presents a schematic diagram of the present process.

FIG. 1 shows the relationship between percent conversion of MAP coal and maximum preheater temperature at a space time of 0.035 hour. FIG. 1 shows that very high yields are obtained at temperatures of at least 450C. at a constant low residence time.

FIG. 2 shows percent conversion of MAP coal as a function of residence time in the preheater. FIG. 2 is based on data taken at 450C. and shows that substantially maximum conversion (above 80 or 85 percent) is achieved very quickly in the preheater and that continuance of preheater holding time for a considerably greater duration has a very small effect on total conversion. Therefore, at at 450C. preheater temperature, after about 0.05 or O.l hour the preheater time is substantially removed as a process factor in regard to conversion.

[ residence time.

FIG. 4 shows the fraction of organic sulfur removed from the vacuum bottoms (deashed solid fuel) product versus residence time at various temperatures. As shown in FIG. 4, a high level of sulfur removal is least 5 dependent upon residence time at elevated temperatures while residence time becomes increasingly impor- TABLE 6 TEST NUMBER 1 2 3 4 H- Pressure. kg/cm" 70 70 70 70 Max. Preheater Temp..C. 450 450 450 500 LHSV 28.36 28.35 15.23 28.36 GHSV 2964 2953 3012 2988 l/LHSV; Hr. 0.035 0.035 0.066 0.035

YIELDS ON MAF COAL BASIS% CO 0.03 0.07 0.06 0. l7 C0 0.35 0.35 0.45 0.66 H S 0.97 L73 L49 L77 Hydrocarbon Gas Product 0.3l 0.2I 065 L85 2 l .42 l.ll 0.9l l .50

Excess Solvent (Liquid Fuel Product) -59.30 39.6I 7.84 23.43 Vacuum Bottoms (Solid Fuel Product) 13621 I053! 75.78 7!.l7 Insoluble Organic Matter 23.20 3! .34 l3.2l 12.83

TOTAL 100.35 l0().55 100.39 l0().38


Recovery. weight 7? 95.46 96.80 90.68 i8 38 MAF Conversion, weight 70 76.80 68.66 86.79 87.17

VACUUM BOTTOMS (SOLID FUEL PROPERTIES Carbon. weight A 86.67 Hydrogen. weight 5.33 Nitrogen. weight 1.78 Sulfur. weight 7 l.2i Oxygen, weight 4; 2 Ash. weight 7! (Ll i H/C Atomic Ratio 369 FIG. 3 shows the sulfur content in the deashed coal product as a function of total preheater and dissolver residence time at various maximum preheater temperatures. FIG. 3 shows that residence time exerts a greater effect on sulfur level in the vacuum bottoms product at high temperatures that at low temperatures. FIG. 3 shows that if significant sulfur is to be removed without utilizing relatively high temperatures. a prolonged residence time must accompany low temperature operation. Therefore, the present process employs a dissolver at a relatively low temperature and a relatively long residence time to accomplish a degree of sulfur removal beyond that which is possible in the preheat alone. which operates at a higher temperature at which long residence times are prohibitive due to the onset of hydrocracking. In this manner, a dual tempe ature process produces a product having a lower sulfur level than the sulfur ievcl that is obtained by operating both tant to a high level of sulfur removal at lower tempera tures. FIG. 4 again illustrates the basis for employing a relative y low temperature dissolver coupled with an extended residence time in cooperation with a relatively high temperature preheater coupled with a relatively short holding time.

FIG. 5 illustrates the relationship of hydrocarbon gas yield to preheater outlet temperature of 0.035 hours and shows that hydrocracking to gases increases rapidly as the temperature is increased above 400C. and especially above 450C. The present invention permits the achievement of high conversion without excessive hydrocracning by utilizing a high temperature only for a short duration (preheater stage) followed by a rela tively low temperature for a longer residence time (dissolver stage). In this manner, a high conversion is achieved without a high yield of hydrocarbon gases. Production of hydrocarbon gases constitutes a waste of product and a needless consumption of hydrogen unless gas handling facilities are available.

FIG. 6 illustrates the effect of temperature and rcsidence time on hydrogen consumption and shows that at low residence times hydrogen consumption in not affected by temperature but at higher residence times (above 0.4 or 0.5 hours) hydrogen consumption is affected by temperature. Either low residence times or low temperatures favor low hydrogen consumptions. Therefore, in the present process. a low residence time is employed in the relatively high temperature preheater while a relatively long residence time is em ployed in the relatively low temperature dissolverv FIG. 7 shows schematically the process of the present invention. As shown in FIG. 7, pulverized coal is charged to the process through line 10, contacted with recycle hydrogen from line 40 and forms a slurry with recycle solvent which is charged through line 14. The slurry passes through line 16 to preheater tube 18 having a high length to diameter ratio which is greater than 100, generally, and, preferably, greater than 1,000 to permit plug flow. Preheater tube 18 is disposed in a furnace 20 so that in the preheater the temperature of a plug of feed slurry increases from a low inlet value to a maximum temperature at the preheater outlet.

The high temperature effluent slurry from the pre heater is then passed through line 22 where it is cooled before reaching dissolver 24 by the addition of cold makeup hydrogen through line 12. Other methods for cooling can include water injection, a heat exchange or any other suitable means. The residence time in dis solver 24 is substantially longer than the residence time in preheater 18 by virtue of the fact that the length to diameter ratio is considerably lower in dissolver 24 than in preheater 18. causing backrnixing and loss of plug flow. The slurry in dissolver 24 is at substantially a uniform temperature whereas the slurry in preheater 18 increases in temperature from the inlet to the exit end thereof.

The slurry leaving dissolver 24 passes through line 48 to flash chamber 50 from which lighter overhead stream passes through line 64 to vacuum distillation column 28 while ash-containing heavy fuel is removed as flash residue through line 52 and passed to filter 58. Ash is removed from the flash residue through line 60 while the deashed residue is passed to vacuum distillation column 28 through line 62.

Gases, including hydrogen for recycle. are removed overhead from distillation column 28 through line 30 and are either withdrawn from the process through line 32 or passed through line 34 to scrubber 36 to removed impurities through line 38 and prepare a purified hydrogen stream for recycle to the next pass through line 40.

A distillate liquid product of the process is removed from a mid-region of distillation column 28 through line 42 and recovered as liquid product of the process. Since the process produces sufficient liquid to be withdrawn as liquid fuel product plus sufficient liquid to be recycled as solvent for the next pass, a portion of the liquid product is passed through line 44 for recycle to line 14 to be employed to dissolve pulverized coal in the next pass.

Vacuum bottoms is removed from distillation column 28 through line 46 and passed to moving conveyor belt 54. On conveyor belt 54 the bottoms product is cooled to room temperature, at which temperature it solidifies. Deashed solid fuel containing as low as ash content as is practical is removed from conveyor belt 54 by a suitalbe belt scrapper means, as indicated at 56. As shown in FIG. 7, no material is removed from the process between the preheater and the dissolver and all material entering the preheater passes through both the preheater and dissolver before any product separation occurs.

We claim:

1. A process for preparing deashed solid and liquid hydrocarbonaceous fuel from hydrocarbonactous feed coal containing ash comprising contacting the feed coal with hydrogen and a solvent for the hydrocarbonaceous material in the coal to form a coal-solvent slurry in contact with hydrogen, passing the slurry and hydrogen through a preheater for a residence time be tween 0.01 and 0.25 hours, said preheater having a length to diameter ratio of at least to inhibit backmixing so that an increment of said slurry gradually increases in temperature in passge through the preheater from a low inlet temperature to a maximum temperature at the preheater outlet, the maximum temperature at the preheat outlet being 400 to 525C, the viscosity of an increment of the slurry in passage through the preheater increasing initially to a value at least 20 times the viscosity of the solvent alone when each is measured at a temperature of 99C., the viscosity of the slurry when measured at 99C. subsequently dropping to a vaiue lower than 10 times the viscosity of the solvent alone when each is measured at 99"Cv in continued passage through the preheater, the viscosity of said slurry finally tending to increase to a value greater than 10 times that of the solvent alone when each is measured at 99C. at the exit temperature of said preheater but the slurry and hydrogen being removed from said preheater after the relative viscosity drops to a value below l0 but before the relative viscosity finally increases to a value of IO. forcing reduction of the temperature of the slurry at least i0C. to a temperature at which the viscosity of the slurry does not increase to a value above l0 times that of the solvent alone when each is measured at 99F, passing the cooled slurry to a dissolver maintained at a temperature between 350 and 475C. and which is below the temperature at the outlet of the preheater, the residence time of the slurry in the dissolver being greater than in the preheater, removing the slurry from the dissolver and separating the slurry into gaseous product, a fraction which is liquid at room temperature and a deashed fraction which is solid at room temperature, recycling hydrogen contained in said gaseous fraction to the preheater, and recycling at least a portion of said liquid fraction as solvent for said preheater step.

2. The process of claim 1 wherein the maximum temperature in the preheater is 425 to 500C.

3. The process of claim wherein the temperature is the dissolver is 400 to 450C.

4. The process of claim 1 wherein the forced reduction in temperature between the preheater and dis solver is at least 50C.

5. The process of claim 1 wherein the preheater length to diameter ratio is at least 1,000.

6. the process of claim 1 wherin the residence time in the preheater is 0.0] to 0.15 hours.

7. The process of claim l wherein the dissolver residence time is 0.] to 3.0 hours.

8. The process of claim 1 wherein carbon monoxide and steam are used together with or in place of hydrogen.

9. The process of claim 1 wherein the yield of deashed solid fuel is 20 to 80 weight percent based on moisture and ash free coal feed.

10. The process ofclaim 1 wherein the yield of liquid hydrocarbon fuel is above weight percent based on moisture and ash free coal feed.

11. The process of claim 1 wherein the viscosity of the slurry in the preheater falls to a value at least as low as 5 times the viscosity of the solvent alone when each is measured at 99C.

12. The process of claim 1 wherein the viscosity of the slurry in the preheater falls to a value at least as low as 2 times the viscosity of the solvent alone when each is measured at 99C.

13. The process of claim 1 wherein said forced temperature reduction is accomplished by injecting makeup hydrogen into the slurry stream.

14. The process of claim 1 wherein said forced temperature reduction is accomplished by cooling the slurry stream in a heat exchanger.

15. The process of claim 1 wherein the gaseeous product comprises less than 6 weight percent of hydrocarbon gases based on moisture and ash free coal feed l k UNITED STATES PATENT OFFICE CERTIFICATE OF CORRECTION PATENT NO. 3 892 654 DATED l July 1, 1975 |NVENTOR(5) Charles H. Wright and Gerald R. Pastor It is certified that error appears in the above-identified patent and that said Letters Patent are hereby corrected as shown beiow:

Col. 13, line 8, after "yield" insert -of solvent--.

Signed and Scaled this third Day of February 1976 [SEAL] Attesr:

RUTH C. MASON C. MARSHALL DANN Atr sll'ng ffic Commissioner of Parents and Trademarks

Patent Citations
Cited PatentFiling datePublication dateApplicantTitle
US3341447 *Jan 18, 1965Sep 12, 1967Bull Willard CSolvation process for carbonaceous fuels
US3594304 *Apr 13, 1970Jul 20, 1971Sun Oil CoThermal liquefaction of coal
US3645885 *May 4, 1970Feb 29, 1972Exxon Research Engineering CoUpflow coal liquefaction
US3808119 *Oct 12, 1972Apr 30, 1974InteriorProcess for refining carbonaceous fuels
Referenced by
Citing PatentFiling datePublication dateApplicantTitle
US4057484 *Dec 15, 1975Nov 8, 1977John Michael MalekProcess for hydroliquefying coal or like carbonaceous solid materials
US4111786 *Mar 26, 1976Sep 5, 1978Mitsui Coke Co., Ltd.Dispersing in oil, depolymerizing, distillation
US4201655 *Dec 18, 1978May 6, 1980Continental Oil CompanyBlending bottoms with char, coke, and coal
US4243488 *Jul 17, 1978Jan 6, 1981Mitsui Coke Co., Ltd.High strength, metallurgical
US4303527 *Mar 10, 1980Dec 1, 1981Linde AktiengesellschaftMeasuring organic impurities in storage tank to allow proper flow
US4314898 *Apr 29, 1980Feb 9, 1982Kobe Steel, Ltd.Process for reforming coal
US4330389 *Dec 27, 1976May 18, 1982Chevron Research CompanyCoal liquefaction process
US4330391 *Feb 25, 1980May 18, 1982Chevron Research CompanyCoal liquefaction process
US4391699 *Feb 3, 1982Jul 5, 1983Chevron Research CompanyCoal liquefaction process
US4396488 *Oct 8, 1981Aug 2, 1983Electric Power Research Institute, Inc.Aromatic hydrocarbon
US4421630 *Oct 5, 1981Dec 20, 1983International Coal Refining CompanyDesulfurization at higher temperature; hydrogenation at lower temperature
US4534847 *Jan 16, 1984Aug 13, 1985International Coal Refining CompanyProcess for producing low-sulfur boiler fuel by hydrotreatment of solvent deashed SRC
US4639310 *Aug 2, 1985Jan 27, 1987Veba Oel Entwicklungs-GesellschaftProcess for the production of reformer feed and heating oil or diesel oil from coal by liquid-phase hydrogenation and subsequent gas-phase hydrogenation
DE3527129A1 *Jul 29, 1985Jan 29, 1987Inst Vysokikh Temperatur AkadeProcess for making liquid products from coal
U.S. Classification208/408, 208/416, 208/419, 208/414, 208/423, 208/409, 208/431
International ClassificationC10L9/04, C10G1/06, C10L9/00, C10G1/00, C01G1/00
Cooperative ClassificationC10G1/006, C10G1/065
European ClassificationC10G1/00D, C10G1/06B