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US3884794A - Solvent refined coal process including recycle of coal minerals - Google Patents

Solvent refined coal process including recycle of coal minerals Download PDF

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US3884794A
US3884794A US446973A US44697374A US3884794A US 3884794 A US3884794 A US 3884794A US 446973 A US446973 A US 446973A US 44697374 A US44697374 A US 44697374A US 3884794 A US3884794 A US 3884794A
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Prior art keywords
preheater
coal
solvent
temperature
dissolver
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US446973A
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Willard C Bull
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 US446973A priority Critical patent/US3884794A/en
Priority to CA199,898A priority patent/CA1014501A/en
Priority to ZA00744085A priority patent/ZA744085B/xx
Priority to AU70554/74A priority patent/AU486179B2/en
Priority to DE2431872A priority patent/DE2431872A1/de
Priority to DD179793A priority patent/DD113565A5/xx
Priority to BR6236/74A priority patent/BR7406236A/pt
Priority to SU7402049686A priority patent/SU563920A3/ru
Priority to PL1974173333A priority patent/PL99575B1/pl
Priority to FR7428817A priority patent/FR2263291A1/fr
Priority to JP9628374A priority patent/JPS5714718B2/ja
Priority to GB28159/74A priority patent/GB1499331A/en
Priority to CS75348A priority patent/CS191256B2/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
    • C10G1/006—Combinations of processes provided in groups C10G1/02 - C10G1/08
    • 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
    • C10G1/04—Production of liquid hydrocarbon mixtures from oil-shale, oil-sand, or non-melting solid carbonaceous or similar materials, e.g. wood, coal by extraction
    • C10G1/042—Production of liquid hydrocarbon mixtures from oil-shale, oil-sand, or non-melting solid carbonaceous or similar materials, e.g. wood, coal by extraction by the use of hydrogen-donor solvents
    • 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

Definitions

  • This invention relates to a 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 subbituminous coals, and lignites.
  • the process produces deashed solid fuel (dissolved coal) together with as much coal derived liquid fuel as possible, 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 production of frequently 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 scrapped 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 production limit of hydrocarbon gases establishes the production limit of liquid fuel product and therefore also the production limit of solid fuel product.
  • the dual temperature method Anv important advantage of the dual temperature method 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 temperatures reduction of sulfur level is achieved with a minimum of hydrocracking.
  • the preheater effluent should preferably be quenched at least about 25C. and as much as 50 or C. before entering the dissolver. In some cases a smaller extent of cooling, such as at least 10, 15 or 20C. below the maximum or outlet preheater temperature can be effective.
  • the first reactor stage of the present process is a tubular 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 reactor stage employs a relatively longer residence time in a larger vessel maintained at a substantially uniform temperature throughout.
  • a regulated amount of forced cooling occurs between the stages so that the second stage temperature is lower than the maximum preheater temperature.
  • 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 in 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 are inadequate facilities to handle hydrocarbon gaseous by-product, 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 advantageously 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 stages 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 standard 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/100L).
  • the weight ratio of recycled solvent product to coal in the feed slurry ranges from 0.521 to :1, generally, and
  • the solvent used at process start-up is advantageously derived from coal. Its 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 viscosities being measured at 99C.
  • Relative Viscosity as used herein is defined as the viscosity at 99C., of an incrementof solution, divided by the viscosity of the solvent alone fed to the system measured at 99C., i.e.
  • the Relative Viscosity can be employed an an indication of the residence time for the solution in 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.
  • the slurry would set up into a gel.
  • 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 from again 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 L5 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 gel 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 preheater to 350C. causes the gel to decompose, evidencing completion of hydrogen transfer, producing a deashed solid fuel, liquid fuel and gaseous producta 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 carbon-carbon 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.
  • residence time in the preheater for a feed increment to achieve this maximum temperature is about 0.01 to 0.25 hours.
  • the hydrocarbon gas yield under these conditions is less than about 6 weight percent while excess solvent (liquid fuel) yield is above 10 or 15 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 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 kg/cm generally, and 50 to 200 kg/cm preferably.
  • the solvent hydrogen content tends to adjust to about 6.1 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 of liquid fuel, which has a higher hydrogen content than 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 hydroaromatic 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.
  • the combined effect of time and temperature in the preheater stage is important in the present process.
  • the desired temperature effect in the preheater 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 least 1,000, preferably, so
  • 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.
  • the residence time is extended for a duration which is longer than the preheater residence time.
  • the middle boiling fraction of the vacuum distillation which is liquid at room temperature, is normally at least partially recycled for slurrying with feed coal particles to function as the solvent for the process.
  • the vacuum bottoms which is solid at room temperature, is inferior as a solvent to the normally liquid middle boiling fraction of the vacuum distillation.
  • ash is an advantageous recycle component.
  • the iron in the ash helps to break down extracted heavy coal molecules to provide an extracted fuel of lower average molecular weight. Therefore, when an ash-containing product stream is recycled, the relative yield of middle boiling fraction is enhanced while the relative yield of vacuum bottoms is decreased, as compared to the use of middle boiling fraction without ash-containing recycle material as thesolvent for the process.
  • An additional advantage to the use of an ash-containing recycle stream is that subsequently produced vacuum bottoms has a significantly lower sulfur content.
  • Another advantage obtained by the use of an ashcontaining solvent recycle stream is enhancement of activity for hydrogenation of the solvent to replenish hydrogen lost from the solvent by hydrogen donor activity.
  • a transfer of hydrogen from partially hydrogenated aromatic solvent molecules is necessary.
  • Gaseous hydrogen reacts indirectly with 'the coal and the extracted carbonaceous fuel from the coal by first chemically combining with the aromatic solvent molecules. An increased rate of combination of hydrogen with the aromatic solvent molecules in turn increases the rate at which the gaseous hydrogen reaches the coal.
  • the initial hydrogenation activity of start-of-run solvent obtained from a source outside of the process is not of importance if sufficient time is allowed to permit the process to achieve a steady state by recycle of ash-containing process solvent.
  • the ash content in the feed slurry gradually builds up to an equilibrium level and as the equilibrium level is being achieved the benefits associated with the present invention become apparent.
  • the process of the present invention utilizes a preheater, a dissolver and a vacuum distillation tower in series.
  • a slurry of subdivided coal feed and recycle solvent is passed through the preheater.
  • the preheater has a considerably smaller capacity than the dissolver so that the residence time in the preheater is considerably lower than the residence time in the dissolver.
  • No external catalyst is added to either the preheater or dissolver stage, which are the only reactor stages of the present process.
  • the effluent from the dissolver is filtered and then vacuum distilled.
  • gases are removed overhead, fuel product which is liquid at room temperature is removed from an intermediate position in the column, and fuel product which is solid at room temperature is removed from the bottom of the column. It has hitherto been the practice to recycle to the preheater a portion of the liquid fuel product as solvent for the feed coal.
  • an advantageous modification of the process employs recycle of mineral ash as a solvent component.
  • the relative temperature is higher than in the dissolver while the residence time is lower than in the dissolver in order to enhance hydrogen transfer in the preheater from the recycle solvent to the feed coal, without excessive coking and polymerization.
  • the minerals in the ash product such as iron in the form of ferrous sulfide (FeS)
  • FeS ferrous sulfide
  • the dissolver produces hydroaromatics rather than saturated aromatics since hydroaromatics are hydrogen donors while saturated compounds are not hydrogen donors and, therefore, their formation should be avoided.
  • the dissolver employs a relatively lower temperature than the preheater and a relatively longer residence time by utilizing a larger reactor vessel than the dissolver vessel.
  • the present invention utilizes a solution containing fuel product, unreacted coal and removed coal minerals, as a solvent stream for coal feed to the preheater.
  • the presence of coal mineral is shown herein to be highly advantageous.
  • the removed coal minerals may assist the hydrogen transfer reaction in the preheater.
  • the recycle of coal minerals provides the advantage of increased product yield, as compared to the use of distillate liquid fuel product as a solvent without coal minerals.
  • the maximum temperture that can be employed in the preheater depends'upon the hydrogen donor activity of the solvent.
  • the relatively low temperatures employed in the dissolver relative to the preheater favor rehydrogenation of the solvent and favor the production of depolymerized coal-derived products in the dissolver.
  • the maximum minerals concentration that can be tolerated in the recycle vacuum bottoms stream depends on the pumpability of the recycle stream and the filterability of the recycle stream because a portion of the recycle stream must be removed and filtered when the system achieves suitable process equilibrium.
  • Table l and FIG. 1 show that hydroaromatic activity of the solvent can be correlated with the iron content in the feed to the process.
  • Various coals give data which fall on a smooth curve in this respect.
  • FIG. 1 as minerals level increases there is an upward trend in solvent'hydrogen level.
  • the data in Table 1 were taken at 70 kg/cm hydrogen pressure and are illustrated in FIG. 2.
  • Sicne hydroaromatic activity is enhanced by operating the dissolver at relatively low temperatures for relatively longer holding times, it is important that the dissolver vessel be large compared to the preheater vessel.
  • the separate types of reactions occurring in the preheater and the dissolver are further enhanced by employing a higher hydrogen pressure in the dissolver than in the preheater.
  • the first step in accomplishing solution of coal is the hydrogen transfer mechanism in the preheater. If this does not proceed adequately, the solution tends to degenerate by repolymerization and finally by coking. Many lignitic coals have only moderate amounts of iron, so that reactivity with hydrogen in the dissolver may not be satisfactory. If such coals contain high sodium levels, the use of carbon monoxide and steam to produce hydrogen in situ may be more satisfactory than the charging of hydrogen per se, since sodium is a catalyst for the shift reaction of carbon monoxide and steam to produce hydrogen.
  • the rehydrogenation of the solvent in the dissolver may occur by using hydrogen coal-derived ferrous sulfide as a catalyst or by using carbon monoxide and steam with sodium or ferrous sulfide as a catalyst.
  • Table 2 shows the results of tests conducted with a split temperature process illustrating the effect of ashcontaining heavy vacuum bottoms recycle in a process employing a preheater stage followed by a dissolver stage.
  • the dissolver stage effluent is flashed to remove gaseous and liquid material overhead.
  • An ashcontaining bottoms fraction which is solid at room temperature is recycled and employed as a solvent alone or with a lighter fraction.
  • the recycle stream should contain as much liquids as is practical and the amount of solids recycled should be sufficient to obtain a significant advantage.
  • Table 2 further shows that with increasing proportions of ash in the feed slurry, thepercent of MAF conversion tends to increase.
  • the decrease in MAF conversion in Test is due to an excessively high preheat cal/gm.
  • the recycle of ash results in slightly more hydrogen consumption (2.5 to 3 percent), the hydrogen is gainfully utilized because of an increasing yield of higher heat content fuel.
  • the sch recycled, and 10 to 50 percent preferably, is recyheat temperature should be. below 475C., preferably cled. below 470 or 465C. and most preferably no higher TABLE 2 TEST 1 2 3 4 5 Ash Recycle No Yes Yes Yes Pressure H2, kg/cm 70 70 7O 70 70 Preheater Temp., C. 450 450 450 450 475 Dissolver Temp., C. 425 425 425 425 425 l/LI-ISV; Hr.
  • Test 1 of Table 2 was conducted with a solvent that 40 than 460 or 450C, or the retention time should be dedid not include recycled ash. In Tests 2, 3, 4 and 5 of creased if higher temperatures are employed. Table 2, the solvent included recycled ash. In Test 1, Table 2 also shows that the hydrogen content in the the percent of ash in the feed slurry represents the ash liquid product tends to increase with increasing ash from the feed coal, while in Tests 2, 3, 4 and 5 the per- 4 content in the feed. As noted above, an increased hycent ash in the feed slurry is continuously increasing 5 drogen content in the liquid product signifies an indue to a progressively longer recycle duration. Alcreased heat content upon combustion.
  • the liquid product is a 801-fural value particularly to the solid deashed coal vent boiling range material which has a somewhat product. i greater hydrogen to carbon weight ratio and an ele- Table 2 shows that an optimum advantage in the revated heat content (above 17,000 BTU/lb. or 9,450 cycleof ash-containing heavy vacuum bottoms is cal./gm.), as compared to the solid fuel (vacuum botachieved by controlling the temperature in the pretoms) product which has a somewhat longer hydrogen heater.
  • Table 2 shows that gaseous hydrocarbon prodto carbon ratio and a somewhat lower heat content uct decreases with increasing ash content in the feed upon combustion (above 16,000 BTU/lb. or 8,800 slurry at a preheater temperature of 450C, but increases at a preheater temperature of 475C.
  • An increase in hydrocarbon gas product is wasteful since it represents loss of desired liquid fuel product.
  • a 6 weight percent hydrocarbon gas yield on an MAF basis is a suitable upper limit for gas production unless gas storage and transporting facilities are available.
  • the liquid product (excess solvent) yield diminishes sharply when the preheater temperature increased the effect of ash recycle upon the hydrogen content in the liquid product (excess solvent) of the process.
  • Table 4 is an extension of the data of Tests 1 thru 5 of Table 2 and shows that as the percent of iron present O 3 45 d 2 H
  • Test 2 of Table 3 shows that employment of a split rectly via hydrogen donor activity between the partially temperature between the preheater and the dissolver hydrogenated aromatic solvent and the coal. advantageously increases liquid product (excess solvent) yield, increases MAF conversion and decreases TABLE 4 sulfur content in the vacuum bottoms product, as com- I Fe in Increase in H pared to Test 1 wherem the preheater and dissolver are Test of Preheatar Above H in operated at the same temperature.
  • Test 3 Table 2 Feed Original Solvent shows that when the preheater temperature and the dis- 1 l 33 0 33 solver temperature are split but the preheater tempera- 2 1:87 0:46 ture is not as high as 500C., the advantages tend to be 3 2g: 8?; enhanced even without solvent recycle.
  • Table 4 shows the results of tests made to illustrate erating the dissolver at a lower temperature than the preheater, thereby favoring the accumulation of high concentrations of hydroaromatic material in the dissolver product for recycle.
  • the preheater is advantageously operated at a higher temperature than the dissolver to more efficiently transfer hydrogen from the hydrogenated aromatic to-the coal feed in the next pass.
  • high temperatures favor depolymerization, and favor sulfur and oxygen removal re actions.
  • a high concentration of transferable hydrogen favors formation of liquid fuel product and tends to prevent coking.
  • the maximum temperature required for the preheater depends on the activity of the hydro gen in the solvent which is available for transfer.
  • moderate temperatures favor hydrogenation of both solvent and depolymerized coal. In general, it is favorable to allow the catalytic ash minerals to build up to the highest concentration which can be managed, with pump and filter characteristics being the limiting factors.
  • FIG. 3 schematically shows the process of the present invention.
  • pulverized coal is charged to the process through line 10 and contacted with recycle hydrogen from line 40 to form a slurry with recycle solvent together with coal ash minerals which is charged through line 14.
  • the slurry with hydrogen passes through line 16 to preheater tube 18 having a high length to diameter ratio of at least 100, generally, and at least 1,000, preferably, 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 preheater is then passed through line 22 where it drops in temperature before reaching dissolver 24 due to the addition of cold makeup hydrogen through line 12.
  • the residence time in dissolver 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 backmixing 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 26 to flash chamber 48.
  • Liquid and gaseous material is removed overhead in flash chamber 48 through line 50 and passed to distillation column 28.
  • Ash-containing solid fuel is removed from the bottom of flash chamber 48 and a portion is recycled to the preheater through line 52 while the remainder is fed to filter 54 through line 56.
  • Ash is removed from filter 54 through line 58 while filtrate is passed through line 60 to distillation column 28.
  • 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 remove 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.
  • a portion of the vacuum bottoms containing removed coal ash minerals is passed through line 52 for recycle through line 14. If the recycle stream in line 52 contains sufficient hydroaromatic solvent liquid, flow through line 44 can be dispensed with. Otherwise, liquid solvent flowing through line 44 is blended with mineral recycle flowing through line 52. Blended streams from both lines 44 and 52 result in a more easily pumpable recycle stream than if the stream in line 52 were recycled in the absence of the stream in line 44.
  • a process for preparing deashed solid and liquid hydrocarbonaceous fuelfrom hydrocarbonaceous 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 between 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 passage through the preheater from a low inlet temperature to a maximum temperature at the preheater outlet, the maximum temperature at the preheater outlet being 400 to 5 25C., the viscosity of an increment of the slurry in passage through the preheater increasing intially to a value at least 20 times the viscosity of the solvent along when each is measured at a temperature of 99C., the viscosity of the slurry when measured a 99C.
  • the residencetime of the slurry in the dissolver being greater than in the preheater, removing the slurry from the dissolver, separating from the slurry an ash-containing stream, a gaseous stream, a fuel product which isliquid at room temperature, and a deashed fuel product which is solid at room temperature, and recycling a portion of the separated ash to said preheater step.

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  • Chemical & Material Sciences (AREA)
  • Engineering & Computer Science (AREA)
  • Oil, Petroleum & Natural Gas (AREA)
  • Life Sciences & Earth Sciences (AREA)
  • Wood Science & Technology (AREA)
  • Chemical Kinetics & Catalysis (AREA)
  • General Chemical & Material Sciences (AREA)
  • Organic Chemistry (AREA)
  • Production Of Liquid Hydrocarbon Mixture For Refining Petroleum (AREA)
  • Solid Fuels And Fuel-Associated Substances (AREA)
US446973A 1974-03-04 1974-03-04 Solvent refined coal process including recycle of coal minerals Expired - Lifetime US3884794A (en)

Priority Applications (13)

Application Number Priority Date Filing Date Title
US446973A US3884794A (en) 1974-03-04 1974-03-04 Solvent refined coal process including recycle of coal minerals
CA199,898A CA1014501A (en) 1974-03-04 1974-05-13 Solvent refined coal process including recycle of coal minerals
ZA00744085A ZA744085B (en) 1974-03-04 1974-06-25 Solvent refined coal process including recycle of coal minerals
AU70554/74A AU486179B2 (en) 1974-03-04 1974-06-27 Solvent refined coal process including recycle of coal minerals
DE2431872A DE2431872A1 (de) 1974-03-04 1974-07-03 Verfahren zur herstellung von aschearmen festen und fluessigen kohlenwasserstoff-brennstoffen
DD179793A DD113565A5 (pt) 1974-03-04 1974-07-08
BR6236/74A BR7406236A (pt) 1974-03-04 1974-07-30 Processo para a preparacao de combustivel hidrocarbonaceo solido e liquido sem cinzas
SU7402049686A SU563920A3 (ru) 1974-03-04 1974-08-02 Способ производства обеззоленного твердого и жидкого топлива из угл
PL1974173333A PL99575B1 (pl) 1974-03-04 1974-08-08 Sposob wytwarzania pozbawionego popiolu paliwa stalego i cieklego z wegla
FR7428817A FR2263291A1 (pt) 1974-03-04 1974-08-22
JP9628374A JPS5714718B2 (pt) 1974-03-04 1974-08-23
GB28159/74A GB1499331A (en) 1974-03-04 1974-12-25 Continuous solubilization processes for preparing de-ashed solid and liquid hydrocarbon fuels from coal that contains ash
CS75348A CS191256B2 (en) 1974-03-04 1975-01-17 Process for preparing fly-ash handled solid and liquid hydrocarbon fuel from coal

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JP (1) JPS5714718B2 (pt)
BR (1) BR7406236A (pt)
CA (1) CA1014501A (pt)
CS (1) CS191256B2 (pt)
DD (1) DD113565A5 (pt)
DE (1) DE2431872A1 (pt)
FR (1) FR2263291A1 (pt)
GB (1) GB1499331A (pt)
PL (1) PL99575B1 (pt)
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ZA (1) ZA744085B (pt)

Cited By (23)

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US4040957A (en) * 1976-02-20 1977-08-09 The Lummus Company Separation of insoluble material from coal liquefaction product by use of a diluent
US4083769A (en) * 1976-11-30 1978-04-11 Gulf Research & Development Company Catalytic process for liquefying coal
US4176041A (en) * 1977-02-24 1979-11-27 Kobe Steel, Ltd. Method for reforming low grade coals
WO1980000156A1 (en) * 1978-07-03 1980-02-07 Gulf Research Development Co Combined coal liquefaction-gasification process
US4210518A (en) * 1977-01-24 1980-07-01 Exxon Research & Engineering Co. Hydrogen-donor coal liquefaction process
US4210517A (en) * 1977-10-31 1980-07-01 Mitsui Mining Co. Ltd. Preparation of carbonaceous products
US4226698A (en) * 1978-08-04 1980-10-07 Schroeder Wilburn C Ash removal and synthesis gas generation from heavy oils produced by coal hydrogenation
US4244812A (en) * 1978-12-28 1981-01-13 Kerr-Mcgee Corporation System for producing a powdery composition comprising coal products in a coal deashing process
WO1981002305A1 (en) * 1980-02-05 1981-08-20 Gulf Research Development Co Solvent refining of coal using octahydrophenanthrene-enriched solvent and coal minerals recycle
WO1981002304A1 (en) * 1980-02-05 1981-08-20 Gulf Research Development Co Coal liquefaction process employing octahydrophenanthreneenriched solvent
US4312746A (en) * 1980-02-05 1982-01-26 Gulf Research & Development Company Catalytic production of octahydrophenanthrene-enriched solvent
WO1982000830A1 (en) * 1980-09-09 1982-03-18 Pittsburgh Midway Coal Mining Controlled short residence time coal liquefaction process
WO1982000831A1 (en) * 1980-09-09 1982-03-18 Pittsburgh Midway Coal Mining Short residence time coal liquefaction process including catalytic hydrogenation
EP0051345A2 (en) * 1980-11-03 1982-05-12 Exxon Research And Engineering Company Donor solvent coal liquefaction with bottoms recycle at elevated pressure
WO1983000874A1 (en) * 1981-09-03 1983-03-17 Pittsburgh Midway Coal Mining Improved coal liquefaction process
US4558651A (en) * 1983-10-19 1985-12-17 International Coal Refining Company Fired heater for coal liquefaction process
US4617105A (en) * 1985-09-26 1986-10-14 Air Products And Chemicals, Inc. Coal liquefaction process using pretreatment with a binary solvent mixture
US5256278A (en) * 1992-02-27 1993-10-26 Energy And Environmental Research Center Foundation (Eerc Foundation) Direct coal liquefaction process
US20070295590A1 (en) * 2006-03-31 2007-12-27 Weinberg Jerry L Methods and systems for enhancing solid fuel properties
US20090038213A1 (en) * 2003-12-12 2009-02-12 Weinberg Jerry L Pre-burning, dry process methodology and systems for enhancing metallurgical solid fuel properties
US20090119981A1 (en) * 2006-03-31 2009-05-14 Drozd J Michael Methods and systems for briquetting solid fuel
US20090272028A1 (en) * 2006-03-31 2009-11-05 Drozd J Michael Methods and systems for processing solid fuel
US20170342326A1 (en) * 2014-12-05 2017-11-30 Posco Method and apparatus for manufacturing cokes additive

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AU506174B2 (en) * 1976-05-28 1979-12-13 Kobe Steel Limited Coal liquefaction
DE2654635B2 (de) * 1976-12-02 1979-07-12 Ludwig Dr. 6703 Limburgerhof Raichle Verfahren zur kontinuierlichen Herstellung von Kohlenwasserstoffölen aus Kohle durch spaltende Druckhydrierung
RU2458975C2 (ru) * 2006-03-31 2012-08-20 Коултэк, Инк. Способы и устройства, повышающие качество твердого топлива

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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
US3617465A (en) * 1969-11-20 1971-11-02 Hydrocarbon Research Inc Coal hydrogenation
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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Publication number Priority date Publication date Assignee Title
US3341447A (en) * 1965-01-18 1967-09-12 Willard C Bull Solvation process for carbonaceous fuels
US3617465A (en) * 1969-11-20 1971-11-02 Hydrocarbon Research Inc Coal hydrogenation
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 (35)

* Cited by examiner, † Cited by third party
Publication number Priority date Publication date Assignee Title
US4040957A (en) * 1976-02-20 1977-08-09 The Lummus Company Separation of insoluble material from coal liquefaction product by use of a diluent
US4083769A (en) * 1976-11-30 1978-04-11 Gulf Research & Development Company Catalytic process for liquefying coal
US4210518A (en) * 1977-01-24 1980-07-01 Exxon Research & Engineering Co. Hydrogen-donor coal liquefaction process
US4176041A (en) * 1977-02-24 1979-11-27 Kobe Steel, Ltd. Method for reforming low grade coals
US4210517A (en) * 1977-10-31 1980-07-01 Mitsui Mining Co. Ltd. Preparation of carbonaceous products
WO1980000156A1 (en) * 1978-07-03 1980-02-07 Gulf Research Development Co Combined coal liquefaction-gasification process
US4203823A (en) * 1978-07-03 1980-05-20 Gulf Research & Development Company Combined coal liquefaction-gasification process
US4226698A (en) * 1978-08-04 1980-10-07 Schroeder Wilburn C Ash removal and synthesis gas generation from heavy oils produced by coal hydrogenation
US4244812A (en) * 1978-12-28 1981-01-13 Kerr-Mcgee Corporation System for producing a powdery composition comprising coal products in a coal deashing process
US4347117A (en) * 1979-12-20 1982-08-31 Exxon Research & Engineering Co. Donor solvent coal liquefaction with bottoms recycle at elevated pressure
WO1981002305A1 (en) * 1980-02-05 1981-08-20 Gulf Research Development Co Solvent refining of coal using octahydrophenanthrene-enriched solvent and coal minerals recycle
WO1981002304A1 (en) * 1980-02-05 1981-08-20 Gulf Research Development Co Coal liquefaction process employing octahydrophenanthreneenriched solvent
US4312746A (en) * 1980-02-05 1982-01-26 Gulf Research & Development Company Catalytic production of octahydrophenanthrene-enriched solvent
US4323447A (en) * 1980-02-05 1982-04-06 Gulf Research & Development Company Coal Liquefaction process employing octahydrophenanthrene-enriched solvent
US4322284A (en) * 1980-02-05 1982-03-30 Gulf Research & Development Company Solvent refining of coal using octahydrophenanthrene-enriched solvent and coal minerals recycle
US4328088A (en) * 1980-09-09 1982-05-04 The Pittsburg & Midway Coal Mining Co. Controlled short residence time coal liquefaction process
US4330388A (en) * 1980-09-09 1982-05-18 The Pittsburg & Midway Coal Mining Co. Short residence time coal liquefaction process including catalytic hydrogenation
WO1982000830A1 (en) * 1980-09-09 1982-03-18 Pittsburgh Midway Coal Mining Controlled short residence time coal liquefaction process
WO1982000831A1 (en) * 1980-09-09 1982-03-18 Pittsburgh Midway Coal Mining Short residence time coal liquefaction process including catalytic hydrogenation
EP0051345A2 (en) * 1980-11-03 1982-05-12 Exxon Research And Engineering Company Donor solvent coal liquefaction with bottoms recycle at elevated pressure
EP0051345A3 (en) * 1980-11-03 1982-09-22 Exxon Research And Engineering Company Donor solvent coal liquefaction with bottoms recycle at elevated pressure
EP0073866B1 (en) * 1981-09-03 1988-06-01 The Pittsburg & Midway Coal Mining Company Improved coal liquefaction process
WO1983000874A1 (en) * 1981-09-03 1983-03-17 Pittsburgh Midway Coal Mining Improved coal liquefaction process
US4377464A (en) * 1981-09-03 1983-03-22 The Pittsburg & Midway Coal Mining Co. Coal liquefaction process
US4558651A (en) * 1983-10-19 1985-12-17 International Coal Refining Company Fired heater for coal liquefaction process
US4617105A (en) * 1985-09-26 1986-10-14 Air Products And Chemicals, Inc. Coal liquefaction process using pretreatment with a binary solvent mixture
US5256278A (en) * 1992-02-27 1993-10-26 Energy And Environmental Research Center Foundation (Eerc Foundation) Direct coal liquefaction process
US20090038213A1 (en) * 2003-12-12 2009-02-12 Weinberg Jerry L Pre-burning, dry process methodology and systems for enhancing metallurgical solid fuel properties
US8579998B2 (en) 2003-12-12 2013-11-12 Coaltek, Inc. Pre-burning, dry process methodology and systems for enhancing metallurgical solid fuel properties
US20070295590A1 (en) * 2006-03-31 2007-12-27 Weinberg Jerry L Methods and systems for enhancing solid fuel properties
US20090119981A1 (en) * 2006-03-31 2009-05-14 Drozd J Michael Methods and systems for briquetting solid fuel
US20090272028A1 (en) * 2006-03-31 2009-11-05 Drozd J Michael Methods and systems for processing solid fuel
US8585786B2 (en) 2006-03-31 2013-11-19 Coaltek, Inc. Methods and systems for briquetting solid fuel
US8585788B2 (en) 2006-03-31 2013-11-19 Coaltek, Inc. Methods and systems for processing solid fuel
US20170342326A1 (en) * 2014-12-05 2017-11-30 Posco Method and apparatus for manufacturing cokes additive

Also Published As

Publication number Publication date
FR2263291A1 (pt) 1975-10-03
SU563920A3 (ru) 1977-06-30
JPS50119808A (pt) 1975-09-19
JPS5714718B2 (pt) 1982-03-26
GB1499331A (en) 1978-02-01
ZA744085B (en) 1975-06-25
CS191256B2 (en) 1979-06-29
AU7055474A (en) 1976-01-08
DD113565A5 (pt) 1975-06-12
DE2431872A1 (de) 1975-09-11
CA1014501A (en) 1977-07-26
BR7406236A (pt) 1976-03-23
PL99575B1 (pl) 1978-07-31

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