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Fuel Processing Technology, 1 (1977) 3- -20 © Elsevier Scientific Publishing Company, Amsterdam -- Printed in The Netherlands EARLY COAL HYDROGENATION CATALYSIS E.E. DONATH P.O. Box 1068, Christiansted, St. Croix 00820 (U.S.A.) MARIA HOERING 6900 Heidelberg 1 (G.F.R.) (Received September 1976) ABSTRACT Three inputs were necessary to make catalytic hydrogenation of coal possible. One was the ammonia synthesis which, in 1910, introduced high pressure and temperature into the chemical industry. The second was the experimentation by F. Bergius who showed, in 1913, that coal can be liquefied by adding hydrogen at high pressure and temperature. The liquid products were similar to coal tar. They were not of the quality required for gasoline or diesel fuel production. The use of catalysts to refine the coal oil appeared then to be hopeless since coals contained sulfur, a poison for all then known hydrogenation catalysts. The third input was methanol synthesis in 1923. M. Pier found selective, oxidic catalysts that were less sensitive to sulfur than e.g. the metallic catalyst for the ammonia synthesis. In 1924 M. Pier, in the laboratories of the BASF, prepared sulfur resistant coal hydro- genation catalysts: sulfides and oxides of molybdenum, tungsten, and the iron group metals. With these catalysts it became possible to add hydrogen; split carbon--carbon bonds; and eliminate such heteroelements as sulfur, oxygen and nitrogen from coals and oils. Thus fuels were produced that met petroleum fuel specifications. Optimum catalyst action was achieved by subdividing coal hydrogenation into two stages. The coal was converted, with a dispersed catalyst in the "liquid phase", into middle oil. This was then hydrogenated over fixed bed catalyst, in the "vapor phase", to gasoline. On this basis a large~scale demonstration plant for the liquefaction of central German brown coal was erected in 1927. The development of catalysts for these two stages proceeded on different routes. Liquid phase catalysts were discarded after one pass through the reactor. They were cheap, or used in very small amounts. It was found soon that coal of different rank required different catalysts, and that the mineral matter of the coal played an important role. The first commercially used vapor phase catalysts were of the hydrorefining type. Hydro- cracking activity was achieved by using high temperatures. A great step forward was made in 1930 when a special preparation of tungsten disulfide permitted hydrocracking activity at low temperatures. Thus the first essentially dual function catalyst was found. Its hydro- cracking activitity was further increased, and gasolines with a higher octane number were obtained by using it on acidic supports such as materials containing aluminalilica. Such supported catalysts were poisoned by the nitrogen compounds present in coal oils. Therefore a refining step for these oils was needed. The vapor phase was subdivided into the "prehydrogenation" (hydrorefining) and "splitting hydrogenation" (hydrocracking) steps. Further development of catalysts with specific functions for these two steps proceeded rapidly. In addition, separate catalysts were developed for the production of gasolines with a high content of aromatics.

Early coal hydrogenation catalysis

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Page 1: Early coal hydrogenation catalysis

Fuel Processing Technology, 1 (1977) 3- -20 © Elsevier Scientific Publishing Company, Amsterdam -- Printed in The Netherlands

EARLY COAL HYDROGENATION CATALYSIS

E.E. DONATH

P.O. Box 1068, Christiansted, St. Croix 00820 (U.S.A.)

MARIA HOERING

6900 Heidelberg 1 (G.F.R.)

(Received September 1976)

ABSTRACT

Three inputs were necessary to make catalytic hydrogenation of coal possible. One was the ammonia synthesis which, in 1910, introduced high pressure and temperature into the chemical industry. The second was the experimentation by F. Bergius who showed, in 1913, that coal can be liquefied by adding hydrogen at high pressure and temperature. The liquid products were similar to coal tar. They were not of the quality required for gasoline or diesel fuel production. The use of catalysts to refine the coal oil appeared then to be hopeless since coals contained sulfur, a poison for all then known hydrogenation catalysts. The third input was methanol synthesis in 1923. M. Pier found selective, oxidic catalysts that were less sensitive to sulfur than e.g. the metallic catalyst for the ammonia synthesis.

In 1924 M. Pier, in the laboratories of the BASF, prepared sulfur resistant coal hydro- genation catalysts: sulfides and oxides of molybdenum, tungsten, and the iron group metals. With these catalysts it became possible to add hydrogen; split carbon--carbon bonds; and eliminate such heteroelements as sulfur, oxygen and nitrogen from coals and oils. Thus fuels were produced that met petroleum fuel specifications.

Optimum catalyst action was achieved by subdividing coal hydrogenation into two stages. The coal was converted, with a dispersed catalyst in the "liquid phase", into middle oil. This was then hydrogenated over fixed bed catalyst, in the "vapor phase", to gasoline. On this basis a large~ scale demonstration plant for the liquefaction of central German brown coal was erected in 1927.

The development of catalysts for these two stages proceeded on different routes. Liquid phase catalysts were discarded after one pass through the reactor. They were cheap, or used in very small amounts. It was found soon that coal of different rank required different catalysts, and that the mineral matter of the coal played an important role.

The first commercially used vapor phase catalysts were of the hydrorefining type. Hydro- cracking activity was achieved by using high temperatures. A great step forward was made in 1930 when a special preparation of tungsten disulfide permitted hydrocracking activity at low temperatures. Thus the first essentially dual function catalyst was found. Its hydro- cracking activitity was further increased, and gasolines with a higher octane number were obtained by using it on acidic supports such as materials containing a lumina l i l i c a .

Such supported catalysts were poisoned by the nitrogen compounds present in coal oils. Therefore a refining step for these oils was needed. The vapor phase was subdivided into the "prehydrogenation" (hydrorefining) and "splitting hydrogenation" (hydrocracking) steps. Further development of catalysts with specific functions for these two steps proceeded rapidly. In addition, separate catalysts were developed for the production of gasolines with a high content of aromatics.

Page 2: Early coal hydrogenation catalysis

The various catalysts developed primarily for the hydrogenation of coal derived oils introduced hydrogen processing into the petroleum refining industry. There they were further modified and improved for the processing of petroleum. These improved catalysts, in turn, will be of help to a future coal liquefaction industry.

INTRODUCTION

We shall attempt to describe here the discovery of sulfur resistant catalysts, and their use in hydrogenating coal to produce liquid fuels. It began fifty years ago at BASF, later I.G. Farbenindustrie, A.G. (IG), in Ludwigshafen/Rh., G.F.R. We had the privilege of participating from almost the beginning, from 1926 until 1945, as coworkers of the pioneer of catalytic coal hydrogenation, Matthias Pier. We shall describe difficulties that were overcome and the reasoning that led to new developments. We hope that some of these insights may be useful to present workers in this field. This report may be timely considering that personal recollections can still enter into it. We shall not describe the important and outstanding parallel efforts necessary for the development of suitable materials of construction, special machinery and equipment, and processes for the production of hydrogen from coal using oxygen blown producers and fluidized bed gasification of brown coal.

The development of catalytic coal hydrogenation began in 1924 with high pressure flow reactors containing 50 cc of catalyst. It continued with larger units and large scale semicommercial plants at Ludwigshafen/Rh. After three years the start-up of a coal hydrogenation demonstration plant, with a design capacity of 100,000 t/y (2500 B/d) of gasoline, in the Leuna works (then part of the IG) became possible. It took four more years to develop catalytic brown coal hydrogenation into a stable, reliable process in 1931/1932. Hydrogenation of brown coal tar had reached that stage already in 1929. The Leuna plant was later expanded to a capacity of 650,000 t/y (17,000 B/d) of gasoline and diesel fuel.

During the approximately ten years after the demonstration plant had established technical and commercial viability of the process, ten hydrogenation plants using brown coals and bituminous coals and tars as feed materials were completed in Germany. In World War II two more hydrogenation plants were started. Together they reached a total capacity of about 4 million tons per year (about 100,000 B/d) of gasoline and other liquid fuels. Several catalytic high pressure hydrogenation plants were erected outside of Germany.

For ammonia synthesis, A. Mittasch had developed iron catalysts that were activated and stabilized by additives. Later 1~. Pier observed that, in the methanol synthesis from CO and H2, iron was harmful. He found alternatives -- selective oxide catalysts, among them zinc oxide activated by chromic acid. Since this catalyst was less sensitive to sulfur poisoning than the metallic ammonia catalyst, he expected that it would be possible to develop catalysts for the hydrogenation of sulfur containing coal tars or coals into motor fuels of high quality. In spite of the then believed dogma that sulfur is a poison to

Page 3: Early coal hydrogenation catalysis

hydrogenation catalysts, he found such sulfur resistant catalysts. Subsequent to this discovery, C. Bosch, President and Chairman of the

Board of the BASF, approved the start of the coal hydrogenation project. There was a number of reasons for this decision. It appeared of interest to a large chemical concern to open up a new field of coal chemistry which promised not only to supply fuels normally of petroleum origin, but also raw materials for captive use. It was expected that the process would convert not only coal, but also tars and petroleum residues into higher value motor fuels with high yield. Thus it would alleviate the lack of petroleum reserves in Germany and counteract the then existing fears of petroleum reserve exha~_~stion. Introduction of a high pressure process that could utilize equip- ment in existing synthetic ammonia plants offered some insurance in case of a decline of the fertilizer market. In addition it is to be remembered that, at the time, Germany suffered both from an extreme shortage of foreign exchange and from high unemployment. A large scale process in the huge liquid fuels field would alleviate both these problems.

The IG initially carried all the development cost of the new process. Very high expenses arose since the technical development lasted longer than anticipated. At the same time the market price of imported gasoline dropped to an extent that was unforeseeable. Production expansion became possible only when the government increased gasoline import duties and, after 1933, began to underwrite sales volume and minimum price. As it happened, in a short time the production cost in several plants dropped below the guarantee price, producing income for the government.

Following the BASF board clecision in 1924 to develop the hydrogenation process, the then existing "high pressure experiments" group became a separate entity with M. Pier as manager. In the ammonia plant department of C. Krauch, Pier's deputies in the beginning were Karl Winkler, an organic chemist and W. Rumpf, a mechanical engineer. Both deputies had been Pier's coworkers during the development of the methanol synthesis. This "high pressure experiments" group expanded rapidly and, by the end of 1924, consisted of four chemists, two mechanical engineers and forty workmen. In 1927 this group reached its largest size: about seventy graduate chemists and engineers, and about one thousand workmen. At this time, in addition to small scale experi- ments, several large scale development units were in operation mainly to train personnel for the Leuna demonstration plant. As time progressed the staff changed in size and background. Experienced men were sent to manage and staff new plants. On the average the staff consisted of about sixty graduate chemists, physicists, mechanical engineers and a fluctuating number of laboratory assistants and workmen.

There were several subgroups specializing in small and large scale experiments, catalyst preparation and production, analytical and technical services, etc. These groups reported to M. Pier and frequent, often daily discussions with groups or group leaders, were the rule. In these meetings many of the inventions came to light. The list of IG hydrogenation patents shows that M. Pier was

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often inventor or coinventor; however there were many patents by his co- workers.

The program of "high pressure experiments" from the beginning included not only catalytic hydrogenation of coals and tars but also of petroleum and shale oils and their residuums. In addition to motor and aviation fuels, kerosene, and diesel fuel, lubricating oils and paraffin wax were produced or their quality was improved. The new process found great interest in Germany and abroad. A large number of coals, tars and oils of various origins were received in Ludwigshafen, tested, and evaluated in catalytic hydrogenation experiments.

In order to be able to make comparisons, it became necessary to investigate competit ive processes of the coal and petroleum industries such as extraction, pyrolysis and cracking. M. Pier and his comparatively small staff could not have solved the multiplicity of problems in the short time without extensive assistance of specialized departments of the large IG. Further, since 1927, agreements with Exxon and other oil companies, and later German licensees, accelerated the progress. All the available technical data came to M. Pier. He led the exchange of information about all the problems that arose during the development of the new process. This international cooperat ion in ex- changing results was unique at the time. It proved very fruitful to all parties and lasted until World War II.

INITIAL PROCESS DEVELOPMENT

After describing the framework in which the catalytic hydrogenat ion process was developed, we shall turn now to the experimental work and look first at the initial tests with fixed bed, sulfur resistant hydrogenation catalysts. The equipment used is shown schematically in Fig. 1. Feed stocks flow from a pressurized tank together with hydrogen at 200 atm pressure, through a sight- glass, where the feed rate is set by adjusting the number of drops per minute. The pressure shell of the reactor has an I.D. of 2.5 cm and an insert that contained 50 cc of catalyst. The products leave the reactor through a cooling

C[•L4ig h pressure Hydrog_.~ feed tank

I ISight glass L~ feed flow

,Cat a l y s t ~ ' - ~ ~ l ~ I-1

[Xl GOS ~ exit ~ Cooler

I Sight glass÷ L.~J separator ~ Gasol ine

Fig. I. 50 cc catalyst testing reactor.

Page 5: Early coal hydrogenation catalysis

Fig. 2. Battery of catalyst testing reactors.

conduit and enter a sight-glass separator. From there gas and liquid product are separately depressured, measured and analyzed. A photograph of a "ba t t e ry" of such reactors is shown as Fig. 2.

Raw cresylic acid was used as the first model compound. It was an ash-free, easily handled liquid material and, like coal and tar, it contained oxygen plus sulfur and nitrogen containing impurities as well as an aromatic ring. In late 1924 molybdenum sulfide was found to be an active hydrogenating and reducing catalyst. Other catalysts contained oxides and sulfides of tungsten and of elements of the iron and other groups.

A decisive [1] experiment with brown coal tar was made in early 1925 with a molybdenum--zinc oxide catalyst: waterwhite gasoline was obtained in

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a yield of 100% by volume. This was a great step forward since t rea tment of tar without catalyst or by other methods gave unsatisfactory products and only low yields. Economic evaluation of this and similar experiments showed that the throughput was too small and the hydrogen recycle rate too high for commercial practicability. At higher throughputs catalyst activity deteriorated. It was found that the high boiling fractions, especially the asphaltenes, were the cause. Pier and coworkers concluded that it was necessary to conduct the reaction in two stages:

1. Coal and liquid residues were converted in "liquid phase" with suspended catalyst into middle oil boiling below 325 ° C; and

2. ~ iddle oil was converted into motor fuels in "vapor phase" with a fixed bed catalyst. (Later it became possible to treat oils boiling above 32 5°C with fixed bed catalysts.)

For the liquid phase hydrogenation of coal a coal--oil paste, as shown al- ready by Bergius, was used to introduce coal into the high pressure reactor. The initial coal hydrogenation tests with brown coal from the seams near Leuna differed from this practice. A tilting reactor shown schematically in Fig. 3 was used for first testing of different coals, of operating conditions and of catalysts and their method of applications to the coal. This reactor permitted rapid heating of the coal by sliding the coal cartridge into the hot zone by tilting the reactor. Accurate selection of the residence time was possible. It gave coal oils that were not diluted by products from the pasting oil and showed differences in the reactivity of different coals. However, the system was limited to small coal samples because of the heat of reaction and to non-swelling or melting coals.

More extensive experiments for the testing of catalysts were made in rotating autoclaves. Different coals, pasting otis, and methods of catalyst

Screen ~ ~ ~ I ''Hyd['OgenotiOn

I ~ Position for reaction

Fig. 3. Tilting coal hydrogenator.

Page 7: Early coal hydrogenation catalysis

Hydrogen H°t iiltil r A

Fig. 4. Liquid phase pilot plant.

addition were investigated. It became apparent later that the catalyst activity found in autoclaves needed confirmation by experiments in cont inuous flow reactors with pasting oil recycle.

For the further testing of coals, catalysts, operating conditions, and solid separation methods, continuous flow reactors of 1--10 liters volume were developed and used routinely. In addition reactors of several hundred liters volume were used at Ludwigshafen. Those larger reactors confirmed data before new coals or practices were introduced into commercial use. A schematic of a small liquid phase unit is shown as Fig. 4. Some of these units were equipped with stirrers, others with recycle pumps to return the hot separator bo t toms to the reactor increasing thereby the linear velocity and to avoid settling out of solids. Feed and hydrogen were heated together. After passing the reactor, they entered the hot separator and were removed as oil and gas in the overhead product , and as residue. The overhead product was cooled and,

Catalyst

I Heavy oil Centrifuge I

Kiln

Hydrogen Coal Recycle gas

of, gos I I

I oo, ydrogeoat,on I "ig t o" ungt l -~

:asting{oil Distillation

I ~ASN, coke Off-gas spent catalyst Fig. 5. Coal hydrogenation flowsheet.

Middle oil

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10

after separation from the hydrogen, distilled into middle oil p roduct and heavy oil. The latter was recycled, the solids from the hot separator bot toms were recovered by various methods.

Mechanical methods such as filtration and centrifuging with various diluents and diluent rates were studied as well as thermal methods such as distillation and pyrolysis. All were investigated under different conditions. Continuous De Laval centrifuges combined with pyrolysis of the centrifuge residue in newly developed equipment became the commercial method. The equipment toler- ated only limited concentrations of asphaltenes and asphaltene--solid ratios. It was a lengthy task to adjust throughput, temperature and catalyst concen- tration in the reactors to keep the oil recovery system workable. Operation on the commercial scale was necessary to give the final adjustments. The schematic flowsheet, Fig. 5, was used with some variations especially in the pyrolysis units in the German coal hydrogenation plants.

CATALYST DEVELOPMENT FOR THE LIQUID PHASE

We shall now describe the catalysts used for the hydrogenation of brown coal tar residue in the liquid phase. Initially finely ground molybdenum catalysts suspended in the residue in a high concentration of about 25% were used. The bo t tom product from the hot separator containing the catalyst was recycled. Only sufficient fresh feed was added to replace the oils distilled off in the hot separator. It was found that at 200 arm the catalyst retained extended activity at temperatures below 450 ° C. This comparatively modest temperature kept the throughput at an undesirably low level. At higher temperature and throughput the catalyst replacement rate became too high. Regeneration of the catalyst, e.g. by oxidation, was a t tempted but was not successful because inorganic material coming from the tar deposited on the catalyst that reacted with it during oxidation. The main cause of catalyst deactivation was the formation of high molecular weight, dehydrogenated material that was selectively absorbed by the catalyst. In the high boiling fractions of the products, among other compounds, the highly condensed aromatic coronene and homologs were found. It was concluded that in spite of the high hydrogen pressure the formation of such material was unavoidable but could be made harmless by the use of small amounts of an absorbent catalyst that was discarded after one passage through the reactor. Satisfactory [2] results were obtained with brown coal char as low cost support. It was produced in the fluidized bed gasifier developed by Fritz Winkler for the product ion of hydrogen. Two percent molybdenum was impregnated from a solution of ammonium molybdate after the calcium compounds in the char were neutralized by dilute sulfuric acid. Catalyst amounts of one percent or more gave at 480°C sufficient hydrocracking activity and reduction of asphaltenes.

Similarly brown coal hydrogenation with high concentrations of finely ground suspended molybdenum catalyst gave excellent results. However, no

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11

TABLE 1

Brown coal hydrogenation 1924 to 1932

Result Initial 1932

Liquefaction (%) 65 98 Middle oil and naphtha in oil product (%) 70 100 Residue feed to centrifuges t/t oil product 2.5 0.7 Asphaltenes in oil product (%) 10 0

solution to the problem of catalyst separation and recovery was apparent at that time. Tests with then available fixed bed catalysts gave unsatisfactory catalyst life at acceptable rates of conversion to middle oil. Therefore the use of low cost catalysts or small amounts of catalyst on a once through basis became necessary again.

In small scale experiments at Ludwigshafen impregnation of brown coal with 0.02% molybdenum as ammonium molybdate gave satisfactory results. Considerable further improvement was achieved by neutralizing 50% of the calcium humates of the coal with dilute sulfuric acid. Catalyst and acid were sprayed on the coal before drying. In large scale operation higher concen- trations of the catalyst were needed to assure satisfactory operation of the centrifuge and long operating periods of the pyrolysis kilns. The progress up to 1932 is illustrated in Table 1. It was achieved by correlation of small scale results with those of the commercial plant plus a careful analysis of the in- fluence of changes of the operating conditions in all parts of the system. Other requirements included careful heating of the coal paste in the presence of hydrogen, suitable residence time and linear velocity in reactor and heating zones, and improved temperature distribution in the reactor. The middle oil produced was a satisfactory feedstock for vapor phase hydrogenat ion and the pasting oil quality was stable. The process had sufficient elasticity to accommodate other but similar brown coals.

Because of cost and limited availability of molybdenum it became necessary to find other liquid phase catalysts. Low cost iron ores and especially Bayermasse were found to have sufficient catalytic activity when used in amounts of 2% or more by weight of the coal. With these catalysts, neutralization of the calcium humates was not required; in fact it was harmful. Occasionally an increase of the asphaltene purge by direct kilning of part of the residue wi thout centrifuging was used to reduce the asphaltene content and improve the stability of the pasting oil. Similarly in brown coal tar residue hydrogenation in liquid phase the molybdenum in the char supported catalyst was gradually replaced by iron compounds.

For the brown coal from central Germany -- with a high hydrogen, bi tumen and sulfur content -- a hydrogenat ion reactor inlet pressure of 230 atm was sufficient. For the brown coal from the Rhine district -- with low hydrogen and sulfur content and high oxygen content -- an operating pressure of 650

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12

atm was selected for the liquid phase part of the commercial plant. It maintained a satisfactory asphaltene level and stability in the pasting oil. In addition it was necessary to supplement the activity of the iron catalyst by adding elemental sulfur.

Experiments with bituminous coals [3] made along with those with brown coals showed that sufficient reduction of asphaltenes required higher catalyst activity than was supplied by 0.02% molybdenum as in brown coal. Observation [4] at the ICI laboratories in Billingham, England had shown that a chlorine containing coal gave much better conversion than expected for its rank when using 0.06% tin oxalate as catalyst. This combination of tin and hydrochloric acid was found to be an active catalyst with all tested bituminous coals. It was adopted in England for the first commercial bituminous coal hydrogenation plant in Billingham. It was used in the first German bituminous coal hydrogenation plant at Scholven, Ruhr, at an operating pressure of 300 atm in liquid and vapor phase. Hydrogen chloride is very corrosive, especially in the cooling zone in the presence of liquid water or amines. Therefore the products leaving the reactor were neutralized with a soda ash--oil slurry, usually injected into a second hot separator.

In order to avoid corrosion problems, and to have greater freedom in the selection of catalysts, a nominal 700 atm operating pressure was selected for the liquid phase units of the later German plants. A comparison of the results [ 5 ] obtained at the two pressure levels and with different catalysts is shown in Table 2. The data indicate the improvement attained by the higher pressure: Throughput and coal liquefaction were increased and off-gas yield and asphaltene level decreased. Today pressures in the 200--300 atm range would be more economical since hydrocarbon gases are more valuable products than formerly and methods of utilizing the asphaltic residue have been improved.

Further improvement of the iron ore catalyst was achieved by replacing it partially by ferrous sulfate impregnated on the coal. For coals containing

TABLE 2

Hydrogenation of Ruhr coal

Pressure (atm)

300 700 700

Catalyst Tin oxalate (% of coal) HCI Iron oxide

Coal liquefaction (% maf) Oil product to 325 ° C (kg/l'h) reaction space (lb/cf" h) Off-gas % of liquefaction products Asphaltenes in heavy oil from hot separator

0.06 -- 0.06 0.75 -- 0.75 -- 3 --

93 96 97 0.18 0.27 0.32

(11.2) (16.S) (20.0) 25 21 18 18 11 9

Page 11: Early coal hydrogenation catalysis

13

1.0

0 .9

._o O.B

u

T 0 .7

E o

4~ j J • 4 9 0

489 / ST: /

• 5 0 2

0.6

400

•465 /

/ /

~&o 66o ~So Pressure, ATM

hi° .6 coal

~ ke oven r pitch

800

Fig. 6. Coal hydrogenation pitches and pressure. Figures indicate operating temperatures (°C).

chlorine, somet imes sodium sulfide was added to avoid corros ion . Di f fe ren t coals requi red d i f fe ren t amoun t s and ratios o f these catalysts to op t imize o u t p u t and rel iabi l i ty o f the high and low pressure processing units.

The f o r ma t ion o f very r e f r ac to ry , low h y d r o g e n c o n t e n t mater ia l r emoved as oil loss in the r e q d u e t r e a t m e n t units is i l lustrated by da ta ob t a ined in a p i lo t p lant wi th a 6 l i ter r eac to r o f the Koppers C o m p a n y , Inc. in Pi t t sburgh [6] shown in Fig. 6. The 700 a tm pi tch has a h igher h y d r o g e n c o n t e n t t han the coal. The H/C ra t io decreases with decreasing opera t ing pressure and at 300 a tm is s ignif icantly be low tha t o f the coal itself.

The German coal h y d r o g e n a t i o n plants were designed to conver t coal in the l iquid phase in one step in to middle oil suitable fo r vapor phase hydro- genat ion. The possibi l i ty o f p roduc ing heavy oils [7] was invest igated repea ted ly The invest igat ion on a large scale occur red in 1941 wi th a conve r t e r o f 1.6

TABLE 3

Bituminous coal hydrogenation at 700 atm

Products: Middle oil (%) 100 65 54 Heavy distillate oil (%) 0 35, 46

Temperature (°C/°F) 476/890 478/892 480/896

Middle oil in pasting oil (%) 0 15 27 Coal liquefaction (% maf) 96 96 96 Oil product kg/l" h (lb/cf' h)

Middle oil 0.26 (16.2) 0.26 (16.2) 0.22 (13.7) Heavy distillate oil 0 0.14 (8.7) 0.19 (11.8) Total distillate oil 0.26 (16.2) 0.40 (24.9) 0.41 (25.5)

Off-gas of total oil + gas (%) 23 20 18.5 Off-gas of middle oil + gas (%) 23 27.8 29.7 Asphaltenes in heavy oil pasting oil (%) 14.7 18 19.3

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cbm (57 cf) volume. Bituminous coal was used at 700 atm pressure with the above mentioned iron catalysts. Increased throughput to obtain additional heavy oil led to a highly viscous unmanageable pasting oil. Therefore middle oil was used as part of the pasting oil to reduce its viscosity. Some results of these experiments [8] are shown in Table 3. Distillate heavy oil was produced satisfactorily. The evaluation of the economics indicated that an advantage would result if the heavy oil were used directly as navy fuel oil heavier than seawater. Experiments showed that hydrogenation of the distillate heavy oil with fixed bed catalysts was possible and that savings in reaction space and hydrogen consumption would result. However these possibilities were not explored in depth at the time since gasoline product ion had priority and would have suffered temporarily if the necessary changes had been made.

C A T A L Y S T D E V E L O P M E N T F O R THE V A P O R PHASE

Let us now describe the development of vapor phase hydrogenation catalysts [9] . Physical strength and stability of the catalyst shapes and their size were always a matter of great concern and intense study. As a rule the only compositions tested were those that had a high melting point such that the operating temperature of the catalyst did no t exceed 2/3 of the melting point on the absolute temperature scale. Investigation of many elements, especially noble metals, was not emphasized because of limited availability in these years.

After many tests catalyst 3510, a MoO3--ZnO--MgO combination, was adopted for all round use in 1927. It contained equimolecular amounts of the three oxides. Its activity depended greatly on the particle size, structure and purity of the components and the method of manufacturing. The Mo-component was prepared from molybdic acid anhydride. Today 3510 would be called a hydrorefining catalyst. Hydrocracking activity was achieved by using it at temperatures around 450 ° C. Surprisingly use in commercial brown coal hydrogenation showed initially lower activity than expected from smaller scale experiments. This discrepancy was caused by the large amount of ammonia returned to the hydrogenation systems by the more complete gas recycle.

In the used but still active catalysts, the originally present oxides of molybdenum or tungsten were always sulfided. It was found that molybdenum and especially tungsten oxide became very active catalysts when treated with hydrogen sulfide and hydrogen under pressure. The hydrocracking reactions occurred at a lower temperature than with 3510. The use of the successful activation method was impractical on a large scale because of corrosion of the available equipment.

In 1930 it was found that thermal decomposit ion of oxygen-free ammonium sulfotungstate to tungsten disulfide gave a catalyst of outstanding activity, catalyst 5058 [10,11] . The conversion of middle oil to gasoline proceeded at about three times the rate with 3510 and at 400 instead of 450°C. This was essentially what later was called a dual function catalyst [ 12 ].

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15

TABLE 4

Hydrogenation with tungsten disulfide catalyst

Feed Product Temperature Reaction conditions (°C)

Pressure Rate Conversion (atm) (kg/1 'h) (%)

Diisobutene Isooctane 216 250 2.0 99 Naphthalene Decaline 336 200 0.9 90 Brown coal Prehydroge- Middle oil hated M.O. 380 200 1.0

Phenols 18% 99.5 N-Bases 4% 99.5 Sulfur 2.5% 99.0

n-Butane Isobutane 400 200 0.5 35 Paraffins Gasoline 260--320 ° C 180 ° C EP 408 220 1.0 90

From catalyst 3510 to 5058 more than 1500 catalysts were tested in high pressure flow reactors. Many of these tests were life tests using recycle of the unconver ted oil. Each of these experiments gave indications of advantages or drawbacks and gave leads for changes in composi t ion and method o f preparation. About five times more catalysts were prepared in the laboratory than were actually tested at high pressure. There were two main reasons for this. One was that other tests may have shown some of the composit ions as unpromising. The other was unsatisfactory physical properties of the sample.

Full scale use of 5058 in 1931 fulfilled all expectations. Some results and reaction condit ions are shown in Table 4. Saturation of double bonds occurs at low temperature. Higher temperatures are needed for the removal of the heteroelements present in coal l iquefaction middle oils. Finally isomerization and hydrocracking reactions require temperatures around 400 ° C. Hydro- cracking is usually preceded by isomerization as can be concluded from the high isobutane con ten t of about 50% of the C4-fraction. This is higher than the equilibrium concentrat ion. In t roduct ion of 5058, instead of 3510, virtually tripled the capacity of the vapor phase reactors.

The catalyst was entirely satisfactory as regards throughput , pur i ty of products and yield. However, its high hydrogenating activity led to gasolines of low aromatic con ten t and octane number. In 1934 Terrana [12] , a mont- moril lonite earth treated with 10% hydrof luor ic acid and containing 10% 5058, was found to be a catalyst that gave gasolines with about a 5 points higher octane number than pure 5058. It became catalyst 6434. It was found in the tenth year of catalyst development. It gave the same gasoline yield at the same temperature as 5058. The isobutane con ten t of the C4-fraction was 75%. There was one impor tan t difference between 5058 and 6434. It was known that nitrogen compounds reduced the activity of 5058. The Terrana

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16

supported catalyst was so sensitive to organic nitrogen compounds that it required practically nitrogen free feed stocks to develop its full activity.

Therefore the vapor phase stage was divided into two steps when using coal derived middle oils:

1. "Prehydrogenat ion" or in present terminology a hydrorefining step; and 2. "Splitting hydrogenation or benzination" a hydrocracking step. At first 5058 catalyst was used for the prehydrogenation. Its drawback was

that it produced some gasoline of low octane number in addition to refined middle oil. At the time, as stated above, we had already recognized that "acidic" supports increased hydrocracking activity and were very sensitive to nitrogen compounds. It was found, too, that "basic" supports enhance the hydrorefining activity. This led to the development of activated alumina based prehydrogenation eatalyst~ in 1937. Finally in 1940f1941 catalyst 8376 was developed as prehydrogenation catalyst. It was adopted for commercial use. It contained 25% 5058 and 3% NiS on activated alumina.

Motor gasolines obtained with these catalysts had the properties shown in Table 5. The octane number improvement achieved with the new catalyst was significant. Equally satisfactory prehydrogenation catalysts were obtained with molybdenum instead of tungsten. However tungsten was more readily available than molybdenum at that time. It was also known that replacing nickel by cobalt in the prehydrogenation catalyst gave products with lower hydrogen content. The refining activity, though still high, was borderline for the reduction of the nitrogen content of coal middle oil for further treat- ment with the 6434 catalyst.

At that time satisfactory motor gasolines as well as leaded aviation gasolines were obtained with the 8376/6434 combination. To obtain gasolines with a higher octane number from coal a bet ter retention of the aromatic character of the coal middle oil was at tempted. Operation with newly developed catalysts at 500 instead of 400 ° C produced highly aromatic gasolines with an additional formation of off-gas of 10%. Two commercial catalysts were used at 300 and 600 atm. The 300 atm catalyst from Ludwigshafen had activated carbon as support for small amounts of Cr- and V-oxides. The 600 atm catalyst was developed at the Ruhroel GmbH and had Terrana as support for Zn, Cr, Mo,

T A B L E 5

Gasol ines from b i t u m i n o u s coal h y d r o g e n a t i o n at 300 a t m

Cata lys t Hy d ro re f i n ing WS 2 WS 2 Hy d ro c rack ing WS; Ter rana- -WS 2

Gasol ine Spec. gray. 0 .735 0 .745 o API 61 58.4 ON-RM 67 75 ON-MM 66.5 74

WS 2 - - N i S - - a l u m i n a Ter rana- -WS 2

0 .770 52.3 78 75

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HF and S. It was used for the conversion of middle oil obtained from coke oven tar. With both catalysts, gasolines with an aromatic content of about 50% were obtained. Blends with 10% isooctane or alkylate and lead had an octane number in the 100 range. The off-gas losses with these "aromatization" catalysts were more than twice those observed with catalysts operating at 400 ° C. Therefore, attempts were made to develop catalysts operating at 400 ° C that combined retention of aromatics with enhanced ability to produce branched compounds at this lower operating temperature. It was found at Ludwigshafen that several specially prepared alumina--silica supports had suitable properties [14]. They gave substantial improvement of the octane number over 8376/6434 but were developed too late to be used commercially for the hydrogenation of coal liquefaction oils.

Catalysts 5058 and 8376 were used not only for the prehydrogenation of middle oils but also for the refining of higher boiling feed stocks. One such application was the direct hydrogenation of brown coal tar. In 1939 commercial production of gasoline, diesel fuel, lubricating oils and paraffin wax was begun by this route. Removal of ash from the tar before use, and hydrogenation with stepwise increased temperature along the catalyst bed, were necessary. Thereby the most unstable, unsaturated compounds of the tar were stabilized before ,~hey reached higher temperatures necessary for heteroelement removal. Thus condensation reactions leading to the formation of catalyst deactivating deposits were avoided.

Catalysts for the hydrogenation plants were produced in Ludwigshafen by the same group that prepared new experimental catalysts for the development program. Usually only four chemists and engineers were in charge of this work. Physical properties of catalysts by adsorption measurements, X-ray analysis, etc. were made in other laboratories at Ludwigshafen. Daily inputs from the catalyst testing group were vital in finding better catalysts. Similarly cooperation with large scale experimental groups, and even the commercial plants, helped in improving properties of existing catalysts. If the catalyst supply from Ludwigshafen to a hydrogenation plant did not appear attractive or economical, which was especially the case for plants outside of Germany, catalyst manufacturing units were installed there.

CONCLUSION

Continuing development of catalysts together with progress in process and plant engineering were responsible to a large exteat for the rapid acceptance of the process. Fig. 7 shows the row of eighteen high pressure stalls erected in Leuna in 1927 for a design capacity of 100,000 t/y of gasoline. The same stalls were later able to accommodate a production of 650,000 t/y.

It was the development of catalyst 5058 that made it possible to hydrogenate bituminous coal successfully to gasoline. Fig. 8 shows the row of stalls for the hydrogenation of bituminous coal built 1939 at Gelsenberg/Ruhr.

Because of the low cost of crude oil in the 1950s and 1960s, the high pressure hydrogenation of coal did not spread. However the catalysts developed

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Fig. 7. High-pressure stalls (Leuna).

for this process have found worldwide use in the petroleum refining industry. They are at the root of the great variety of improved hydrorefining and hydro cracking catalysts available today. We hope that these improved catalysts in turn will benefit the future coal hydrogenation industry.

ACKNOWLEDGEMENTS

The authors acknowledge helpful suggestions by Drs. Larry Anderson and Alex Oblad and are grateful for the assistance received from Bituminous Coal Research, Inc., Monroeville, Pa., U.S.A.

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Fig. 8. High-pressure stalls (Gelsenberg).

REFERENCES

1 Krauch, C. and Pier, M., 1931. Angew. Chem., 44: 953. 2 Pier, M., 1934. Proc. 1st World Petr. Congr., London, Vol. F, pp. 290--294. 3 Pier, M., 1935. Chem. Fabrik, 8: 45. 4 Gordon, K., 1935. J. Inst. Fuel, 9: 68. 5 Hoering, M. and Donath, E.E., 1958. In: W. Foerst (Editor), Ullmanns Encyklopaedie der

Technische Chemie, Vol. 10, Urban & Schwarzenberg, Munich, Berlin, 3rd ed. p. 495. 6 Donath, E.E., 1961. Fourth Internat. Conf. on Coal Science, Le Touquet, France. 7 Kroenig, W., 1950. Katalytische Druckhydrierung, Springer, Berlin, pp. 69--75.

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8 Hoering. M. and Donath, E.E., 1958. In: W. Foerst (Editor), Ullmanns Encyklopaedie der Technische Chemie, Vol. 10, Urban & Schwarzenberg, Munich, Berlin, 3rd ed., p. 518

9 Donath, E.E., 1956. Adv. Catal., 8: 239. 10 Pier, M., 1949. Z. Elektrochem. Phys. Chem., 53: 291. 11 Reitz, O., 1949. Chem. Ing. Technik, 21/22: 413. 12 Heinemann, H., Mills, G.A., Shalit, H. and Briggs, W.S., 1954. Brennst. Chem., 35: 366. 13 Pier, M., 1938. Z. Angew. Chem., 51: 606. 14 Hoering, M. and Donath, E.E., 1958. In: W. Foerst (Editor), UUmanns Encyklopaedie

der Technische Chemie, Vol. 10, Urban & Sehwarzenberg, Munich, Berlin, 3rd ed., p. 506.