US3892654A - Dual temperature coal solvation process - Google Patents

Dual temperature coal solvation process Download PDF

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US3892654A
US3892654A US446971A US44697174A US3892654A US 3892654 A US3892654 A US 3892654A US 446971 A US446971 A US 446971A US 44697174 A US44697174 A US 44697174A US 3892654 A US3892654 A US 3892654A
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preheater
temperature
solvent
slurry
hydrogen
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Charles H Wright
Gerald R Pastor
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US Department of the Interior
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US Department of the Interior
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Priority to US446971A priority Critical patent/US3892654A/en
Priority to CA199,910A priority patent/CA1014094A/en
Priority to ZA00744086A priority patent/ZA744086B/xx
Priority to AU70557/74A priority patent/AU487833B2/en
Priority to DE2431949A priority patent/DE2431949A1/de
Priority to DD179788A priority patent/DD113564A5/xx
Priority to FR7425613A priority patent/FR2263295B1/fr
Priority to BR6114/74A priority patent/BR7406114A/pt
Priority to SU742053471A priority patent/SU910125A3/ru
Priority to PL1974173334A priority patent/PL99576B1/pl
Priority to JP9628174A priority patent/JPS5714716B2/ja
Priority to GB28160/74A priority patent/GB1499332A/en
Priority to CS75347A priority patent/CS191255B2/cs
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    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G1/00—Production of liquid hydrocarbon mixtures from oil-shale, oil-sand, or non-melting solid carbonaceous or similar materials, e.g. wood, coal
    • C10G1/06—Production of liquid hydrocarbon mixtures from oil-shale, oil-sand, or non-melting solid carbonaceous or similar materials, e.g. wood, coal by destructive hydrogenation
    • C10G1/065—Production of liquid hydrocarbon mixtures from oil-shale, oil-sand, or non-melting solid carbonaceous or similar materials, e.g. wood, coal by destructive hydrogenation in the presence of a solvent
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G1/00—Production of liquid hydrocarbon mixtures from oil-shale, oil-sand, or non-melting solid carbonaceous or similar materials, e.g. wood, coal
    • C10G1/006—Combinations of processes provided in groups C10G1/02 - C10G1/08

Definitions

  • 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.
  • 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.
  • the dissolved product comprises in major proportion a high molecular weight fuel which is solid at room temperature.
  • 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.
  • 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.
  • the present invention 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.
  • 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.
  • 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.
  • 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.
  • makeup hydrogen can be charged to the process between the preheater and dissolver steps or a heat exchanger can be employed.
  • the residence time in the preheater is lower than the residence time in the dissolver.
  • the liquid space velocity for the process 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 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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.
  • 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 Max.
  • Table 3 shows a test performed at 475C. employing 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.
  • 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.
  • 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.
  • 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!
  • Table 5 shows that adequate residence time Table 6 shows the results of tests conducted with must lapse to provide a net production of solvent.
  • 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.
  • l able 6 illustrates the interchangeability of preperature of the coaLsolvent slurry starts to rise in iii: i'iuitltt temperature and preheater residence time.
  • 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.
  • FIG. 4 shows the fraction of organic sulfur removed from the vacuum bottoms (deashed solid fuel) product versus residence time at various temperatures.
  • 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. 450 450 500 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
  • 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.
  • 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.
  • 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.
  • the bottoms product 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.
  • 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.
  • 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.
  • 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

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Priority Applications (13)

Application Number Priority Date Filing Date Title
US446971A US3892654A (en) 1974-03-04 1974-03-04 Dual temperature coal solvation process
CA199,910A CA1014094A (en) 1974-03-04 1974-05-13 Dual temperature coal solvation process
ZA00744086A ZA744086B (en) 1974-03-04 1974-06-25 Dual temperature coal solvation process
AU70557/74A AU487833B2 (en) 1974-03-04 1974-06-27 dual TEMPERATURE COAL SOLVATION PROCESS
DE2431949A DE2431949A1 (de) 1974-03-04 1974-07-03 Verfahren zur herstellung von aschearmen festen und fluessigen kohlenwasserstoff-brennstoffen
DD179788A DD113564A5 (es) 1974-03-04 1974-07-08
FR7425613A FR2263295B1 (es) 1974-03-04 1974-07-24
BR6114/74A BR7406114A (pt) 1974-03-04 1974-07-25 Processo para a preparacao de combustivel hidrocarbonaceo
SU742053471A SU910125A3 (ru) 1974-03-04 1974-08-02 Способ получени из угл обеззоленного твердого и жидкого топлива
PL1974173334A PL99576B1 (pl) 1974-03-04 1974-08-08 Sposob wytwarzania pozbawionego popiolu paliwa stalego i cieklego z wegla
JP9628174A JPS5714716B2 (es) 1974-03-04 1974-08-23
GB28160/74A GB1499332A (en) 1974-03-04 1974-12-25 Two-temperature coal solubilization process
CS75347A CS191255B2 (en) 1974-03-04 1975-01-17 Process for preparing fly-ash handled solid and liquid hydrocarbon fuel from coal

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Cited By (16)

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US4057484A (en) * 1975-12-15 1977-11-08 John Michael Malek Process for hydroliquefying coal or like carbonaceous solid materials
US4111786A (en) * 1975-04-16 1978-09-05 Mitsui Coke Co., Ltd. Process for liquefying coal
US4201655A (en) * 1976-12-17 1980-05-06 Continental Oil Company Process for making metallurgical coke
US4243488A (en) * 1975-05-21 1981-01-06 Mitsui Coke Co., Ltd. Coke compositions and process for manufacturing same
US4303527A (en) * 1979-03-09 1981-12-01 Linde Aktiengesellschaft Surge control in the biological purification of wastewater
US4314898A (en) * 1979-05-01 1982-02-09 Kobe Steel, Ltd. Process for reforming coal
US4330389A (en) * 1976-12-27 1982-05-18 Chevron Research Company Coal liquefaction process
US4330391A (en) * 1976-12-27 1982-05-18 Chevron Research Company Coal liquefaction process
US4391699A (en) * 1976-12-27 1983-07-05 Chevron Research Company Coal liquefaction process
US4396488A (en) * 1981-10-08 1983-08-02 Electric Power Research Institute, Inc. Process for coal liquefaction employing a superior coal liquefaction process solvent
US4421630A (en) * 1981-10-05 1983-12-20 International Coal Refining Company Process for coal liquefaction in staged dissolvers
US4534847A (en) * 1984-01-16 1985-08-13 International Coal Refining Company Process for producing low-sulfur boiler fuel by hydrotreatment of solvent deashed SRC
US4639310A (en) * 1984-08-04 1987-01-27 Veba Oel Entwicklungs-Gesellschaft Process for the production of reformer feed and heating oil or diesel oil from coal by liquid-phase hydrogenation and subsequent gas-phase hydrogenation
DE3527129A1 (de) * 1985-07-29 1987-01-29 Inst Vysokikh Temperatur Akade Verfahren zur herstellung fluessiger produkte aus kohle
AU2010263737B2 (en) * 2009-06-22 2013-01-31 Kabushiki Kaisha Kobe Seiko Sho Method for producing carbon materials
CN104039937A (zh) * 2011-12-28 2014-09-10 株式会社神户制钢所 无灰煤的制造方法

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AU506174B2 (en) * 1976-05-28 1979-12-13 Kobe Steel Limited Coal liquefaction
JPS55139489A (en) * 1979-04-16 1980-10-31 Mitsui Cokes Kogyo Kk Liquefaction of coal or the like
GB2053955B (en) * 1979-07-17 1983-01-26 Coal Industry Patents Ltd Coal extraction
WO2011136695A1 (ru) * 2010-04-27 2011-11-03 Stepanenko Yury Mikhailovich Установка для получения композиционного топлива на основе промышленных и органических отходов
JP2013136691A (ja) * 2011-12-28 2013-07-11 Kobe Steel Ltd 無灰炭の製造方法
JP2013136692A (ja) * 2011-12-28 2013-07-11 Kobe Steel Ltd 無灰炭の製造方法

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US3341447A (en) * 1965-01-18 1967-09-12 Willard C Bull Solvation process for carbonaceous fuels
US3594304A (en) * 1970-04-13 1971-07-20 Sun Oil Co Thermal liquefaction of coal
US3645885A (en) * 1970-05-04 1972-02-29 Exxon Research Engineering Co Upflow coal liquefaction
US3808119A (en) * 1972-10-12 1974-04-30 Pittsburgh Midway Coal Mining Process for refining carbonaceous fuels

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US3341447A (en) * 1965-01-18 1967-09-12 Willard C Bull Solvation process for carbonaceous fuels
US3594304A (en) * 1970-04-13 1971-07-20 Sun Oil Co Thermal liquefaction of coal
US3645885A (en) * 1970-05-04 1972-02-29 Exxon Research Engineering Co Upflow coal liquefaction
US3808119A (en) * 1972-10-12 1974-04-30 Pittsburgh Midway Coal Mining Process for refining carbonaceous fuels

Cited By (19)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US4111786A (en) * 1975-04-16 1978-09-05 Mitsui Coke Co., Ltd. Process for liquefying coal
US4243488A (en) * 1975-05-21 1981-01-06 Mitsui Coke Co., Ltd. Coke compositions and process for manufacturing same
US4057484A (en) * 1975-12-15 1977-11-08 John Michael Malek Process for hydroliquefying coal or like carbonaceous solid materials
US4201655A (en) * 1976-12-17 1980-05-06 Continental Oil Company Process for making metallurgical coke
US4330391A (en) * 1976-12-27 1982-05-18 Chevron Research Company Coal liquefaction process
US4391699A (en) * 1976-12-27 1983-07-05 Chevron Research Company Coal liquefaction process
US4330389A (en) * 1976-12-27 1982-05-18 Chevron Research Company Coal liquefaction process
US4303527A (en) * 1979-03-09 1981-12-01 Linde Aktiengesellschaft Surge control in the biological purification of wastewater
US4314898A (en) * 1979-05-01 1982-02-09 Kobe Steel, Ltd. Process for reforming coal
US4421630A (en) * 1981-10-05 1983-12-20 International Coal Refining Company Process for coal liquefaction in staged dissolvers
US4396488A (en) * 1981-10-08 1983-08-02 Electric Power Research Institute, Inc. Process for coal liquefaction employing a superior coal liquefaction process solvent
US4534847A (en) * 1984-01-16 1985-08-13 International Coal Refining Company Process for producing low-sulfur boiler fuel by hydrotreatment of solvent deashed SRC
US4639310A (en) * 1984-08-04 1987-01-27 Veba Oel Entwicklungs-Gesellschaft Process for the production of reformer feed and heating oil or diesel oil from coal by liquid-phase hydrogenation and subsequent gas-phase hydrogenation
DE3527129A1 (de) * 1985-07-29 1987-01-29 Inst Vysokikh Temperatur Akade Verfahren zur herstellung fluessiger produkte aus kohle
FR2585717A1 (fr) * 1985-07-29 1987-02-06 Inst Vysokikh Temperatur Akade Procede d'obtention de produits liquides a partir de charbon et produits obtenus par ledit procede
AU2010263737B2 (en) * 2009-06-22 2013-01-31 Kabushiki Kaisha Kobe Seiko Sho Method for producing carbon materials
CN104039937A (zh) * 2011-12-28 2014-09-10 株式会社神户制钢所 无灰煤的制造方法
CN104039937B (zh) * 2011-12-28 2016-02-10 株式会社神户制钢所 无灰煤的制造方法
US9382493B2 (en) 2011-12-28 2016-07-05 Kobe Steel, Ltd. Ash-free coal production method

Also Published As

Publication number Publication date
FR2263295B1 (es) 1979-10-19
SU910125A3 (ru) 1982-02-28
CS191255B2 (en) 1979-06-29
GB1499332A (en) 1978-02-01
ZA744086B (en) 1975-06-25
FR2263295A1 (es) 1975-10-03
BR7406114A (pt) 1976-03-09
DD113564A5 (es) 1975-06-12
DE2431949A1 (de) 1975-09-11
JPS5714716B2 (es) 1982-03-26
PL99576B1 (pl) 1978-07-31
AU7055774A (en) 1976-01-08
CA1014094A (en) 1977-07-19
JPS50119806A (es) 1975-09-19

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