AN AMERICAN FISCHER-TROPSCH PLANT

An American

P 1%

Fischer.

Tropsch Plant Associate Editor in oollnboration with

L. L. HIRST AND R. G. DRESSLER U. S . Bsreait

P-sing

of Mines. Louisiana, Mo.

Towera of Distillation Stmctvw

A Sdaff-IndusdrmCollaborative Report.

T .

HE pmduction of liquid fuels from nonpetroleum materials mntmuee to be a popular b p i c of discuseion in Weal circles. Ever eince the Field Intelligence Agency (FIAT) teams came back from Germany in 1945 with reports that the German chemical industry bad succeeded in producing annually more than 4,500,000 tons of liquid fuels from coal, the “informed gueeaers” have been dsmning or promoting the possibility of a synthetic fuels induatry for the United States (2, 11, 34). Even the p m botern fd-metimes With a considerable craeh-into two classes advocating gas syntheais or direct hydrogenation. The G8nnan industry began with a proo~88 developed by Fischer and Tropech (20) as a modification of the synthol p r o w for producing heavy aliphatic alcohols from coal (19). The FiacheFTropech process gasi6ea me.1 to produce a synthesis gas which in then polymeriaed catalytically to form liquid hydrocarbons. F. Bergiua, about 10 yearn before, developed a pro,rocess in which the 4 wan hydmgenated directly in the solid phsae to yield liquid products ( 8 , 9 ) . The hydrogenation procesn was devebped more dowly but proved to yield a much highex grade motor fuel tban the gas synth& prooess, and by the middle of World War I1 the German fuel induntry was rapidly converting to thin type of operation (1). when teebnologkta from American industry began to &ne the results of the German experience they immediately reoogn i d the expedient of nubstituting natural gas for coal for furniebing synthesis w ta the Fiscber-Tmpacb p m c m . Cone+ quently, synthetic fuels -ch in t h industrial ~ labamtonen in

..

America was at first concentrated almost exclusively on the Fischer-Trope& reaction. The first full scale plant in the United States was designed to utilize the natural gas modification of the Fi-Tmpeoh process ($8). This plant was built by Cartbage Hydmwl Co. in Brownsville, Tex., and is now m partial operation. A similar plant was planned by Btanolind Oil and Crae Co. but was never constructed (S7). Recently chemical producers have become interested in the cod hydmgemtion process as a source of certain mal chemicalq moat of which are now obtained from by-product coke ovens. If chemicals were the primary objective, liquid fuels would still be p r o d u d in substantial quantity but would be treated an a byproduct. Carhide and Carbon Chemicals Corp. wan the Grst to build a Semicommercial plant on this basia. Their plant at Institute, W. Va., is just cointo production. The U.8.Bureau of Mines, w h m object wan to 6nd an ultimate alternative to petroleum as a source of liquid fuela (46,46), was not dimctly involved in the development of the natural gas Fincher-Tropnch modilication Bince the supply of natural gas waa subject to Bsaentially the name limitationsan crude oil. The data originally available from the German experienoe did not demonstrate a conclusive economic advantage for either of the two alternative pmoes~eaunder conditions in the United States. The bureau proceeded to explore both processes in theh labaratones (28). Bwic reaearch and pilot plant studies were conducted at the Pittabugh laboratories and later at Bruceton, Pa (7). A coal gasification pilot plant was built at Morgantown, W. Va.

450

I

March lsS2

I N D U S T R I A L A N D E N G I N E E R I N G CHEMISTRY

(S7). Ultimately two demonstration plants were mnstructed a t Louiaians, Mo.+ne for each p r o w . Them plants were denigned to be of the minimum sire that would provide operating wnditinns comparable to those expected in full scale commercial V h t a

-W.TL"

In the middle stegae of the program the hydrogenation pmceas mmed to have a Wt economic advantage over the gas synthesis p r o o a a i f m a l w a a u d a a t b e m w m a ~ . However,subaeqnent step, which s~rmntafor 60 to improvements in the @cation 70% of the total production oost of the gas synthesie operation, have nanowed thin differential and may have already eliminated it (S). The msl hydrogenation demonahtion plant a t Louiniama, Mo., has been in essentially regular operation since 1949 (96). The gas eyntbesis plant hss just wmpleted ita first trial runs. With the performanca of them two plants now available, it may 80011 be m b l e to eliminate mme of the conjectureif not the controversy from the discussions of h e r i m synthetic fuels production.

"ECLSLE

0.8

PROCESS DEVEWPMEWC

I

The synthasia of liquid hydrqwbons from a mixture of carbon monoxide and hydmgen, aa it ia d e d out a t the Bureau of Minea demonstration plant a t Louisiana, Mo.,is the m l t of a Series of modifications and variations of the p r o w used by Germ . ~ industry Except for ita basic principle it bears no reBemblanca to the original FiwberTmpsoh pand will be referred to in this d c l e aa the "gas synthesis proceas." It wnSista of two, relatively independent, steps: @cation of the coal and syntheainof the liquid fuel.

I

I i

G-iliCetioil The original German synthesis plank made their synthesis gaa in water gas seta or other standard gaa producers, sometimes sdding oxygen to the air feed to increaee the capacity of the unita. However, gas from such units is too expensive in this country to produce liquid fuels a t anytbhg approaching a resaollsble price. New tachniquea of g d c s t i o n had to be developed before the gas synthesis p r o m became an economic possibility for producing liquid fuel from mal. %me German experiments had beem conducted on direct gasilkation,but none of the experiments on maw d e d powdered mal Bssiscation with oxygen waa ever used in a production unit. One of these designs wan obtained by the Koppers Co. and used aa a basis for the construction of the first g d e r installed a t Louisiana. Thia unit WBB a horisontal drum

*uc

Figure 2.

V s r t i d Guifi.r

geometry of a gasification chamber; and other details of the gasification. On the basis of this experience the present, sl&g type, vertio4 gasifier WBB designed (Figure 2). Synthesis

.*YTYE)*I

aAs

u ( L E I

Figure 1. liorimntal Gasifier with inlets for oxygeu, &am, and pulverized coal a t ea& end (Figwe 1). Synthmis &sa WM diaabarged from the a e n h of th0 top of tbe drum. Opuation of this unit mvd to reveal the meabaniw of the r& action between mal and 0xyge.n and steam; the && of the

The synthesis step of the process has KOUe through a rather extended evol&on. The o;ipnsl German plants all used 6xed beda of granular or pelleted catalynt~(b, S4. In order to remove the large amounts of heat libersted by the synthesis reaction a large number of w a t e m l e d finned tubes were inserted into the catalyst bed or expensive doublstube, boiler-type con-

verters were used. Theee designs did not prevent hot spots in the bed. and it waa ultimately found desirable to recycle 1 to 5 volumes of tail gas to temper the reaction (40,U). Btill Later it WBB found that by incredng the recycle rllb to about 100 times the freeh feed rate, the heat of reaction wuld be removed from the wnverter aa sensible heat in the gas stream and the internal oooling tubes could be eliminated. The heat waa removadfrom the mcirmlating gas in extend heat exh g - (17, &). Even thin system did not giva complete local temperature control and I G.Farbenindustrie began to experiment with the cir-

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INDUSTRI‘AL A N D E N G I N E E R I N G C H E M I S T R Y

culation of cooling oils through the reactor. They built several pilot plants employing circulating coolant oils, but none was ever developed to oommercial &e (24,27,@).

Val. 44, No. 3

can be obtained by changing the flow rate of the cooling oil. The constant motion of the catalyst particles keeps them free from gum formation and eementing. This is the aystem uaed at

Louisiana Industrial laboratories here and abroad have tended to favor Euidkd catalyst system and the commeroial p s y ~ t h e S i aplant at Brownsvilleusesthis technique. The reactor in 8 fluid catalyst gassynthesis system is 8iIniIar to tbat found h petroleum cracking. The catalyst p “ e s downward through the rising stream of synthesis gas which suspends it in a fluid state. In thia technique the heat of the reaction is absorbed by the catalyst particles themselves which are in turn cooled by heat exchangers ($7, .%). Slurry systems in which the catalyst is in a finely divided form are still under investigation in private laboratories and at the Bureau of Mines. In most of these systems the h e l y divided catalyst is circulated with the cooling oil (16,9f). Bureau of Mines technologists think they show considerable pro&.

7. ,-COAL

HOPPER

-.

The Jiggling Bed

Figure 3. Coal Feeder

The Bureau of Mines, in its postwar program, centered its attention on such oil-mold sy~tems. In the fmt experiments oil was sprayed over the tops of the catalyst trays, but even the high specific heat and heat of vnporkation of the oil did not provide sufficient cooling. They then tried floodingthe catalyst chamber with coolant ciroulated countercurrent to the process flow. This gave adequate cooling, but uniform flow rates were not obtained. All the oil-cooled systems suffered from exceeaive rise of pressure drop with time, caused by the cementing of catalyst particles by wax formation and swelling (29). Finally, a concurrent flowsyskm was developed in which the catalyst, instead of being held in layers by retaining screens, is charged into the vessel in bulk. The rising stream of gas and oil lifts the catalyst and -em the individual particles to move continually with a f i l i n g motion. 8ensitive temperature control

The Louisiana unit is the only large Fiecher-Trope& installation to use the “jiggling” catalyst bed. Its main advantage eeems to he that it produces only about 50% as much methane as the conventional fluid system. T h e reason for this dSerence has never heen sati8factorily explained. Jiggling bed reactors also use leas hydrogen than fluid systems. They operate at o p t h m efficiency with a 1 : l hydrogen-carbon monoxide ratio in their feed. Fluid catalyst reactors must have a hydrogen-carbon monoxide ratio of about 1.6:1,which usually requiree an external source of hydrogen or a shift reactor in the feed line from the gasifier. At lower ratios carbon monoxide is adsorbed in the pores of the fluidued catalyst particles and is reduced to carbon. As this carbon deposit expands it cracks the catalyst particles until they are so small they cannot be separated from the product. Prob ably the %me m o u n t of adsorption and reduction oocurs on the wtalyst in a jiggling bed, but the larger piecea of catslyst have suffieient physical strength to resist the splitting force. Pilot plant studies at the Bruoeton laboratories indicate that the slurry system may have still further advantages (56). A eecond pbaae utilizing this technique was originslly scheduled for theplant a t Louisiana, but it was decided to concentrate on the

Synthesis Gas Compreaors (left); Electrostatic Prdpitatore for Cleaning Gas Prior to Compression (right)

c

March 1962

I N D U S T R I A L A N D E N G I N E E R I N G CHEMISTRY ~

~~

-yyy u s e of operation until i t had been thoroughly explored before introducing another variable. slurry reactors Beem to produce even less methane than the jigglioe bed reactors. Bureau engineers also believe that a slurried catslyat would &er less mechanical attrition, giving it a longer life and eliminatingthe problem of the csrry-o*.er of catalyst Gnss into the product linea.

backed by a filter (BE) which separates the dust from the air. The coal powder +pa into a storage bin where i t is held under a nitrogen blanket until charged into, the gasi6er ($E). The raw cad has a m o h r e content of 10 to 14% by weight. The dryer-pulverieer reduces this to about 5%. Gasi6cntion

GAS SYNTHESIS DEMONSTEATION PLANT, WUISIANA, MO.

T h e coal chosen for the initial nrrm of&

gaa s y n t h h demon-

Btration plant was a semicohg type obtaltled from Rock Springs, Wyo. A typical analysis on a mointursfree basis gives: Prorirmta. % Free Hydmtilmbon h h m .n

Vola41.0

51.8

6.6--4.8

Car. bon

mti-.te. % Nitro. O w m mo -n .

72.8

1.5

13.2

calorific ValUS.

8ul- B.T.UJ fur

1.1:

Lb.

iz.0~0

Other coals can be u d without modification of the process. The composition of the coal emms to have no effecton the operation of the gasifier. The coal is received by rail a t the demonstration plant, crusbed mine run, and dmpped directly into an unloading pit where it is picked up by a belt conveyor and taken to the primary crusher (1E). The crusber, a grooved mller'type with '/,-inch clearance, may discharge either to the outdoor storage area or to a bucket elevator whicb lifta the coal into indoor storage bins. These bins feed the pulverizer where the coal is dried an well an powdered. The pulverizer in a vertical ball mill. Air heated to about 250-,F. in a ateam preheater (7E)and to 450' F. in a gas fired furnaoe (=E), sneepe tamugh the mill at about 21,000 pounds per hour (4300cubic feet per minute). The -2IM-meah material is carried out of the top ($E) to a cyclone separator

Powdered coal is metered to the gasifier by three parallel %inch diameter screw feeders (Figure 3). Coal from storage gow to a weigh hopper and then into a feed bopper which wrvea the screws. Intermittent blasta of nitrogen delivered at the bottom of the feed bopper and finger agitatorsprevent compacting of the powdered coal. Thee feeders and three feed linea are used to reduce the magnitude of any 5ahhacLs and reduce the seriounneas of possible feeder failures. A stream of oxygen enters the cad line just a t the end of the feeder screwn and picha up the cad to carry it in a fluidized state into the g d u (WE). A special annular noazle carriea the gas m u n d the cad atream to ensure complete fluidization. The oxygen need is generated by a Linde-Frankl plant 4 s cent to the gaei6er and is abont 95% pure. The remainder of the atream is mostly nitrogen and some argon (see Figure 4). In the type operation to he described, the triple feedera supply about la00 poundn of coal (at 5% moisture) per hour to the prwea. Oxygen is fed at 18,000standard cubic feet per bow. The madmum capacity of the gasifier is about 2500 poundn of coal per how. The g& now being used at Louisiana is a vertical oylinder lined with rammed aluminum oxide. The.oxygen-coal mixture enters the d o n chsmber through a triple nozzle about 2 feet from the bottom and tangent to a circle slightly smaller than the

I N D U S T R I A L A N D E N G I N E B R I N G CHEMISTRY

454

i d d e diameter. A quarter of the way around the oircnmferanca

of tbe c y b h a ntrmn~of mperheated steam is intmduced (nee pisure 2B). An the pdveri6ed coal spirala up the inside of the

cos

reaction obsmk and is gaai6ed the noncarbo-s w m p mpts form amolten dagon thewalls. Thisslagws down to the b o w of tbe obsmber and in wntinuous operation would be tapped ofi at portm provided for this purpose, 6 incha abve the obsmber Boor. The longest w n t i n u ~luna made on this equip ment eo far have bean of the order of 48 hours in which time the

CE

H,

co

(3%

2

28

VoL 44, No. 3

Per cent 14.8 37.2 43.6

.... .... 4.0 0.8 0.1 0.1

The made gan leaving the gasi6er is diluted with 20 volume % of cool, cleaned ayntheeis gas. In the event of an upset in the gasifier which would allow wmacted oliygen to p a s out of the reaction %onethis recycle stream would bum in the mixing tea

'

O

101

t

l

I

-

t

I

t

l

I

slag level just barely regched the tapping ports. However, examinationof theinaideof theveeselaftertherunindicatedthat the dag bad NU tluidly and it is anticipated that the tapping operstion w i l l proceed an predieted. The Wyoming coal being 4 ruw 6% e&, 80 to eO% of which wmea out an dag. The remaming ssh and theuaoonverted wal are carried over in the exit gas stream and must be removed by other mearm. Steam for the &is baa@ to Q0Ooto1000' F. bypsar6ng down thmusa a mil of inoresiag diameter in a vertical, psbsrea funuroe. ~ ~ ~ w % 3to permit ~ experimental l e d pmdnction of astream of mal and steam made directlyfrom a ccal-water slurry. W i t h mch a modification oxygen would enter the gani6ar an a separate stream. &me ~xperimenmwork hsa bean done on thin modi6cation and it oontinues to &ow pmmi6e. In prpamt opeaationstesm is introducedat a rate of about 1600 poundaper hour ( 0 . 8 S p m d o f steam per pound of dry coalpro+ ecaed). Thin mtm of feed m l b in a gasi6er temperature of about a5oo'F. and the gadcation of an much 811 97% of the carbon m the mel intro.$uced. About 70,000 cubic featper hour of syntheia gea learn the topof the g d i e r . A trpical ms88 spectmmeie analysis a d d be:

w n d g the oxygen and preventing it fmm forming explosive mixtures with the gas in the woling and purifying units. The combined streams at 1800' F. psss through a ahell-andtube waate heat boiler @.E)which pmducen 2O00 pounds of 275 pound per equare inch ateam per hour or about 1pound of steam

Much 1952

INDUSTRIAL AND ENGINEERING CHEMISTRY

455

for every pound of coal gasi6ed. A cyclone separator (36E) removes moat of the entrained fly & and d u s t A direct contact, o d e - t y p e waaher-mler packed with fourteen banks of wooden hurdles, m b s the gaa with 250 gallona of water per minute t o take out the rest of the solids and reduce the temperature to 100'to 150* F. The gaa leaving the rasher-cooler contains less than 100 grains of ash per 100 cubic feat. It paBaes through an dectmstatic precipitator (88E) which removes any tam that may have carrid thmugh and earn of the remaining aah. An exhauster pump (88E) maintains mro preaaure (+l inch of water) on the gasifier and scmbbing ayatem and is followed by another precipitator which cleans out the last remainii solids. Both precipitatore were originally installed dter the exhauster, but it was found that tars were wllecting in the pump and one of the precipitating units was moved up to p r o w it. The recycle flow to the gasifier o u t let in tdren off the diacbarge of the pump. The dustfree gaa (lea than 0.2 grain per 100 standard cubic feet) is sent either to a aoO,IMO-cubic foot gas holder floating on the line or directly to the suction of the synthesis gas compressore (a@). Them m threa of these c o m p m r s : in normal operation one compramr is NU at full capacity, one at about half capacity, and the other is held as a standby. The operating dincharge pressure of the compresson is 350 pounds per square inch gage. The mmpnasors am steam-driven, reciprocating type, and are somewhat unusual in that they have only two stag- rather than the usual three. Gaa EuriGcation

The ~ynthesiagas must be carefully punlied before it is introduced into the synthesis reactor berause the catalyst ia quite sensitive to poisoning. Sulfur compounds were known to be particularly harmful to the catalyst, and experimental NIX at the

Diethnnolunine Sorubbsr md Reactivator for Gross Sulfur Runoval from Synthesis Gaa

1

the top of the DEA stripping column is cooled with water (S4E) to reflux water and DEA to the tower. The noncoudenaables

am disposed Of at th0 h S k C k

COQlNC UL

OAS

Figure 7.

Miring Tee

Bruceton laboratories showed that carbon dioxide also has a deleterim e6ect. To remove these impurities the gaa p d thmgh m b b e m countercurrent to a &gallon per minute stream of 50% diethanolamine (DEA) in water solution. The DEA from the bottom of the scrubber is let down to atmospheric PI%% m,preheated by lean DEA to 200 F., and stripped with steam. The abmrbeut is then recycled thmugh the rich DEA exchanger (SSE), pumped back up to the pressure of the p r o w gas (IaE), and returned to the top tray of the scrubber. The acid gaa from

The absorbing unit now requires almost no DEA make-up. The original design called for return of the lean absorbent to the fourth tray from the top of the column, but this method of opem thn rcwlted in conaidemble loan of absorbent. The n u d e 8yDtbeSia gaa wntdw about 15% carbon dioxide and 140 grains of hydrogen sulfide per 100 cubic feat. After aerubbmg, the carbon dioxide content in reduced to about 2% and the hydrogen sulfide to 5 to 10 grains per 100 cubic feet. The balance of the inorgmic sulfur is removed by p&g the &as in series through two absorbers c o n t a i n i i wood chips impregnated with imn oxide. Aa a final purification step the gaa is paaeed through one of two absorbers packed with activated carbon to reducetheorganiosulfurcontentfromabout20~sper100cubic feet to 0.05 grain. The tolerance of the synthesis catalyst for total sulfur is believed to be 0.1 grain per 100 cubic feet. The iron oxide absorbent is discarded when it has had the equivalent of 4 to 5 months continuousuee,but the activated carbon ia regenerated by releaaing it from presmre and steaming it at 550' F. for 4 to 6 hours. The two carbon absorbers rue used alternately on a 24-hour cycle although &hour cycles are equally effective. Originally, a fourth purification step utilizing a mixture of iron oxide and soda ash followed the carbon boxes to remove final traces of organic sulfur. Synthesis The conversion of the purified synthesis gas to hydrocarbons, oxygenated organic compounds, carbon dioxide, and water takes place at 450" to 525' F. with the help of a reduced iron catalyst. The operation of the Louisiana reactor at 300 to 350 pounds per square inch falls in what the Ger~nanaconsidered the middle

456

INDUSTRIAL AND ENGINEERING CHEMISTRY

Vol. 44, No. 3

bined with about an equal volume of recycle gas. The Materials combined stream is then mixed Dimensions, of with 600 to 1800 gallons of a Length X I.D. Designation Lining or Packing Construction Remarks coolant oil in a mixing tee 25,ft. X 4 f t . 1 1 / 2 4-inch Panellay brick- Carbon steel Steam preheater Coil lcngth, 664 ft.; outin. lined; single coil^ 5 put 5,202,000 B.t.u./ (Figure 7). The rate of coolturns 3/4 inch, 23 t h n s hr. 2 incher, 30 turns 3 ant feed is adjusted to liniit inches the temperature rise in the Vertical gasifier 2 0 f t . X 4 f t . 4 i n . Rammed alumina Mild steel Upper section, G ft. X 2 f t . 8 inches converter to about 25" F 61 it. X 6 Et. 6 in. Fourteen banks of Carbon steel Washer cooler Four sprays distribute wooden hurdles (34 wash water 240 gal./ About 50% of the gas dissolves trays total) nun. in the oil. Diethanolamine scrub- 4 5 f t . X 2 f t . Bin. 35 ft. 1 1 1 2 inches of oe- A S T M 285, 90,000 stand. cu. ft./hr. ber ramie Raschig rings Grade C The original coolant was a F.B.Q. Diethanolamine reac- 3 5 f t . X 3 ft. Sixteen trays Mild steel ...... 500" to 800" E'. cut of low tivator sulfur crude oil obtained from Iron oxide absorbers 12 f t . 6 in. X 4 ft. 10 f t . (115 c u . ft.) of dry ASThl ,4212 90,000 stand. cu. f t . / h r . iron oxide a n d wood GradeBfirea t 430 lb./qq. inch (2) the East White Lake field in shavings; 25% FelOs box steel gage by wt. Louisiana. Eventually a Activated carbon ab- 12 ft. 6 in. X 4 f t . 8 ft. of activated charcoal A S T M A212 140,000 stand. cu. ft./hr. similar sulfur-free cut from the (92 cu. ft.) sorbers ( 2 ) GradeB-firea t 450 Ib.'sq. inch box steel gage product of the synthesis plant Reactor 30 f t . 10 in. X 3 1 ft. of graded steel balls: Carbon steel 160,000 stand. cu. ft./hr.; ft. 15 f t . Fe catalyst will be used for this purpose. 900 gal./min. coolant (14,000 Ib., 106 c ~ i . The gas-oil mixture enters ft > ..., 2ft.51/2in. X l f t . ...... Separator Molybdenum Swarthout horizontal the bottom of the reactor and steel separator, 160,000 stand. cu. ft./hr. passes up through the catalyst 57 ft. X 3 ft. Thirty tra 7s sixteen 4 1 / ~ Carbon steel Primary fractionator , . . . . bed. Its upward velocity lifts inch b u t h i e caps/tray Six t r a y s ; seven 41/*-inch Carbon steel Heavy distillate strip- 16 ft. X 2 f t . .. . the catalyst particles and holds cadtray per 1 6 f t . X 2 ft. Six trays; seven il/n-inch Carbon steel Diesel oil stripper them in semisuspension, result....., caps/tray ing in a 20 to 30% increase in ...... 10 f t . X 3 it. 6 in. Gasoline sevarator Carbon steel , . . . . (740 gal.) the height of the bed. The Monoethanolamine 45 f t , X 2 ft. 6 in. 35 f t . ,1 1 1 4 inches of ce- ASTM A-285, ,..... scrubber ramic Raschig rings individual particles (8 to 18 Grade C (7375 Ib., 171.5 cu. ft.) Firchox steel mesh) have limited mobility Monoethanolamine re- 54 f t . X 3 ft. 6 in. Twenty-three bubble cap ASTM A-285, ...... activator trays, twenty-nine 4Grade C, in the expanded bed; this reinch caps/tray Flange steel Gasoline absorber 58 ft. 6 in. X l f t . 36 ft. of a/einch gal- Carbon steel 2 ft. G inches between :ecsults in continuous effective 6 in. vanised spirals, 3"long tions contact with the gas flow and Diesel oil absorber 17 it. X 1 ft. G in. 12 ft. of a/r-inch galva- Carbon steel niaed spirals. 3" long prevents cementing of the 30 it. of :/a-inch galv- A S T N A-286, Stabilizer 4ft. x 2ft. anized spirals 3" long particles with wax or tars. Grade C F.B.Q. The cooling oil removes apPolymerization reactor 7 f t . X 2 ft. 6 in. 400 Ih. of '/~-~/r-inch ASTM A-285, , . . . . . crushed firebrick; 2000 Grade C proximately 7000 B.t.u. for lb. (34.4 cu. ft.) UOP F.B.Q. KO.1 polynier catalyst every pound of hydrocarbon Bauxite reformers ( 2 ) l o f t . X 3 f t . 6 inches of Insulcrete lin- Carbon steel ...... product produced. It overing: 4-inches of '/2inch alumina halls: 4flows from the top of the inches of a/s-inch alumina balls; 4 inches of reactor through a conical baffle '/&-inch alumina balls. to minimize carry-over of 4 ft. 9 inches of 4-16 mesh bauxite catalyst particles. Volatile Debutaniaer 63 f t .6 i n . X 1 ft. 21 f t . 6 inches of a/,-inch Carbon stecl Feed a t the center constituents are separated in 6 in. galvanized spirals, 3" long a knockout drum. The coolRe-run column 37ft.8in. X 1 ft. 25 Et. of a/,-inch galva- Carbon steel .~.... 4 in. nized spirals, 3" long ant then enters the circulation Pebble superheater Upper, 11 ft. X I.'pper, 41/?incheseach of Carbon steel 2900 lh./hr.; inlet, 225O 4 f t . 6 inches t o high alumina and firepumps, which force it througb F.; outlet 2500" I:. 8 it. 5 inches; brick; center, 9 inches a waste heat boiler which procenter, 7 ft. X of high alumina, and 3 ft.; lower, 8 41/2 inches of firebrick; duces up to 6500 pounds of f t . X 4 ft. 3 lower, 9 inches of fireinches brick; 9000 Ib. of 3 / ~ 275-pound per square inch gage inch aluminape bbles steam per hour. oirculatedat 6150lb.(hr. Horizontal gasifier 12 ft. 6 in. X 6 f t . Silica brick: firehrlck; in- Mild steelshell 38 f t . high 12 inches in diIt was originally planned to 6 in. sulating brick ameter; explosion riser bleed off a portion of the ciron t o p Top, 7 Et. l'/z in. F i r e h r i c k - l i n e d ; w a t e r - X i l d steel Gas generator and~ pass X 6 f t . 10 in.; jackpted ~ ~ ~ t ~ ~ & ~ ~ ~ cculating e ~ g coolant, s ~ c p ~ ~ ait ~ bottom. 5 ft. f o r gas production through a vacuum distillation 91/zin. X 7 ft. Catalyst reducer 20 it. 7 in. X 2 f t . ...... Carbon steel column to remove the waxes and other heavy hydrocarbons formed in the synthesis r e a c t i o n . These waxes do range of operating pressures. However, the Bureau of Mines not pass out of the react'or with the volat,ile overhead product equipment has been designed for operation up to 600 pounds presand will build up in the cooIant. However, experimental insure and will eventually be investigated at thi, pressure (Figure vestigations have revealed that the build-up of the wax constituents jn the coolant will proceed only to an equilibrium 5). About 7 tons of catalyst are placed in the reactor on top of a concentration of about 80%. At this concentration the cornbed of iron spheres graduated from s /inch ~ to 3 inches in diameter position of the coolant will become essentially stable, and no (see Figure 6). These spheres act both to support the catalyst additional waxes will be formed, presumably because of mass above the level of the gas inlet to prevent the formation of a action effect. Operation with t.his high wax concentration in the pocket where gums could form and to distribute the gas evenly coolant does necessitate that the coolant circulating line he kept through the area of the reactor. The volume of the catalyst is either heat'ed or drained since the equilibrium composition coolapproxiniately 116 cubic feet. The gas from the gasifier is comant has a melting point above 200' F. TABLE I. VESSELDIMENSIOFS

,

March 1952

INDUSTRIAL AND ENGINEERING CHEMISTRY

457

Distillation Columns ( l e f t ) for Refining Synthetic Crude into Finished Products; Synthesis Structure ( r i g h t ) Shows Heat Exchange Equipment; Dark Vessel at Right is FischerTropsch Reactor

The products of the reaction together with the unreacted gas leave the top of the reactor, pass through the exchanger which heats the incoming synthesis gas, are condensed in a water-cooled, shell-and-tube condenser, and are deposited in an overflow type separator vessel (Table 11). In the separator a water layer settles out which contains almost all the oxygenated compounds formed in the synthesis. At present this water layer is drawn off the bottom of the vessel and put into a storage tank. It is available to any laboratory for experimental purposes. The bureau has no plans for intensive study of this product and hopes t o eventually supply it to a commercial plant for recovery. The noncondensables from the product stream are drawn off the top of the separator and boosted 20 to 50 pounds per square inch to full reactor feed pressure by three steam-driven, reciprocating compressors (%E). This gas goes to a monoethanolamine, countercurrent absorber essentially identical t o that used to scrub the synthesis gas. It uses a 30% solution of monoethanolamine which takes up more than 90% of the carbon dioxide in the stream, reducing the concentration to 2%; 85% of the carbon dioxide-free gas from the top of the absorber is returned to the reactor feed. The remainder is sent to the distillation section of the plant for the recovery and processing of the Ca and C4hydrocarbons. The liquid synthesis product, decanted from the separator, is reduced to atmospheric pressure in a letdown drum and proceeds to the distillation area a t about 100" F. Primary fractionation in a 40-tray, bubble-cap column yields four fractions: gasoline, Diesel oil, heavy distillate suitable for coolant oil or cracking stock, and bottoms comprising mostly paraffin-type waxes which may also be subsequently cracked. The bottoms are at present sent directly to storage. Further examination of their composition should reveal whether it will be most profitable to refine them for direct sale or crack them for recycle to the reactor. The heavy distillate, withdrawn a t the ninth tray, is steam stripped of its light ends in an auxiliary stripping column, cooled, and sent to storage. It also may be cracked and recycled or used as is for fuel oil or coolant oil. The Diesel oil cut, taken a t the nineteenth tray, is also stripped in an auxiliary column. About 75y0 of the Diesel product goes through the Diesel oil absorber to assist in recovery of light ends and then returns to the fractionator as a reflux. The other 25%

is cooled and passed through a 10% caustic solution spray to neutralize and remove acid constituents. The fraction is then suitable for direct consumption a's Diesel fuel. It was originally thought that it might be necessary to hydrogenate the Diesel oil cut to stabilize it and raise its cetane number to make it a suitable fuel. The quality of the product obtained from the process so far has made this step unnecessary, The present product does have a slight sharp odor which could undoubtedly be removed by hydrogenation, but the same result can probably be achieved much less expensively by an activated carbon treatment. The overhead from the primary fractionator contains the primary product of the synthesis, gasoline, and water and light ends. These products are treated by conventional gasoline refinery practices, including polymerization and catalytic reforming, to produce a satisfactory motor fuel (Figure 8). A normal product distribution would be: Gasoline (to 400' F.) Diesel oil (400-600° F.) Craoking stock and coolant oil (600-842' Wax (above 842' F.)

F.)

Volume Yo 62.8 12.0 11.1 14.1 100 0

-

Alternate Gasifiers

In addition t o the vertical gasifier already described, the gas synthesis demonstration plant has two other gas producers which may be used t o provide feed to the synthesis reactor-the original horizontal gasifier and a standard Kerpely generator. The Koppers unit. This is a horizontal steel cylinder lined with ceramic brick and is B1/2 feet in diameter and 9 feet long on the inside. Coal-oxygen fluid and superheated steam were fed into the centers of both ends of the cylinder. The same coal feeding equipment was used that is now installed on the vertical gasifier. The steam nozzle is annular and surrounds the triple coal-oxygen nozzles. The made gas discharge is in the center of the top of the cylinder. The gasifier was designed to be nonslagging and although some molten residue was encountered most of the nonreacting constituents of the coal either passed out with the product gas as fly ash or collected in three ash legs provided a t the bottom of the cylinder. Steam for the horizontal gasifier was superheated in a gas-fired

458

INDUSTRIAL AND ENGINEERING CHEMISTRY

Vol. 44, No. 3

I

. .

INDUSTRIAL AND ENGINEERING CHEMISTRY

460

double column as in the Linde procees, but in place of the conventional tubular heat exchangers, corrugated aluminum strips packed in towers act as heat resavoim (50). By a reversing. flow cycle the stripe are alternately cooled by the product nitrogen or oxygen and heated by the incoming air to be fractionated; by an unusual flow system 7% of the air intake is compressed to 3ooo pounds per square inch gage, 73% is compressed to 65 pounds per square inch gage, and the remaining 20%is compressed to only 12 pounds per square inch gage. All air streams are purified and cooled before entering the fractionation columna. The two high preaeure streams enter the lower column, and the low pressure stream is introduce3 into the upper column (86). Oxygen of better than 98% punty is produced routinely COKE I" by this plant; 99.5 and 95% pure nitrogen may also he obbtained. A t p r e e ent the nitrogen is used as a purge gas throughout the plant Or exhausted ta the atmosphere. Catalyst Production The German

CommercialFiScher-

urn JMFT

z::FM

Tropsch plants all used cobalt c a b lysts of type looco: 5ThOI: :MgO: 200 K i e s e l g u h r (58). However, when the Bureau of M i n e s undertook their study of the l i q u i d fuel from Coal proce98es their DUTDWe . . WE4 to explore the posaibilities of extensive industrial utilization. A cssual survey of the known cobalt sources revealed that there is not enough of this metal available at any price to supply catalyst to an industry comparable in &e to the petroleum industry. Consequently, the bureau turned its atteution to the iron camlysts which they had h t uaed in small scale trials at the Pittsburgh laboratories in 1928 and which were later used experimentallyin Germany. These catalysts were virtually identical to those commonlyusedforammoniasynthesia,anditwasdiscoveredthatcommercially available ammonia catalysts would serve equally well for this purpose. However, few manufacturers produce thistype catalyst for general eale and in order to ensure a supply and to have control over the composition and characteristics of their material the bureau undertook to manufacture their own (Figure 10). This type catalyst is manufactured by fusing iron oxides, usually magnetite, with promoters and then crushing the pig and reducing with hydrogen. Activity experiments revealed that a mmt effectivecatalyst could he made by using ordinary mill scale as a source of iron and magnesium oxide and potassium carbonate as alkaline promoters. The optimum mix has 93.7% mill ecale, 4.6% magnesium oxide, and 1.7% potassium carbonate. The ingredients are mixed in an ordinary concrete mixer and poured into 8 half-cyhh'ical, steel trough about 10 feet long and 2 feet in diameter. T w o water-cooled electrodes are placed in the mix abwt 9 feet apart, and four small air lances are inserted into the charge spaced evenly between the electrodes. Current is supplied by a battery of three .Wkv.a. trnnsfomers. Pieoes of iron wire are laid just under the surfaceof the charge to c a n y the cur-

Vol. 44, No. 3

rent until fusion begine. The voltage may be varied between 35 and 120 volts. The current is initially introduced a t 110 volts and held a t that voltage until the total current reaches 300 amperes; this usually requires 20 to 25 minutes. The voltage is then adjusted to maintain this maximum amperage for another 60 to 80 minutes. When the fused zone in the furuace a p proachea 8 inches in diameter the current is shut off and the pig allowed to cool. The cool pig, weighing about 600 pounds, is lifted out of the furnace and broken with mauls until the pieces are 4 to 6 i n o h or less in size. After the initial mushing the mated sized to 8 to 18 megh in two smooth-roll crushers in series. Oversize material is returned to the mill. Undersize goes back to the melting furnace. The raw catalyst has the composition: Total Fe Fe++

E?? SiO,

Per Cent 69.2 39.5 4.6 0.6 0.9

This material is reduced with hydrogen at 800' to 1oOO' C. in a vertical, steel vessel which can be opemted either batchwise or continuously. The single unit has a capacity of 2ooo pounds per day. Hydrogen is obtained by the catalytic cracking of natural gss in the presence of steam and shifting the carbon monoxide to carbon dioxide in a shift converter. Carbon dioxide can then be partially eliminated by scrubbing with sodium hydroxide to give a reducing gas with the composition: 93.7% H,, 0.5% COS 0.8% CO, and 0.6% CH,. The feed rate is controllea to give 1500 standard cubic feet per minute of hydrogen. The bydrogen is preheated to lMM' F. in a g F 6 r e d furnace and tlows upward. The heat for the endothermic reaction is supplied completely by the sensible heat of the hydrogen. Catalyst is fed into the top of the reducer and removed from the bottom through a star wheel valve and immediately immersed in coolant oil to prevent contact with the atmosphere. 8esled in steel d r u m the catalyst in oil will retain ita full activity for at least a year. It is charged to the reactor while still amended in the coolant oil. The final activated by a &day period of stepwise h e k n g in the presence of coolant oil. After the catalyst charge is in the m o t o r and the vessel is seaied, the pressure is raised to 300 pounds per square inch gage by the introduction of synthesis gas, and circulation of the cooling oil is begun. The circulating oil is heated to 400" to 425" F. over a period of about 12 hours. About 25% of the carbon monoxide and hydrogen in the feed stream should be reacting by this time; if not, the temperature is raised further until 25% reaction is obtained. This temperature is beld for 20 hours and then the temperature is increased to the point a t which 40% of the reactive gases bave reacted. This temperature is usually in the range 435' to 460' F. and is maintained for an additional 20 hours. For the next 20 hours the temperature is

TABLE111. Honrsonrmr. GASIFEXPWFORMWCE DATA Steam temp. F.

Gasification temp. Feed rates

&rod

F.

lb hr

Owpan ( d p & d , atand. om ft./hr.

Steam Ib./hr. Orypeb (100%) ta dry -I ratio. stand. ou.ft. fib. Steam to dry oosl ratio, Ib./lb. Starrm deaomposed, % Product. Total stand. on ft./hr. E, atnnd. CU. ft./hr. Lb. dry aoa1/1000 stnnd. E". ft. of E, CO Stand. OU. ft. oxypen/1000 atand. CY. ft. of

+ %,

+ E*+ co Lb.ofsteam/lOOOatand.Eu.ft.ofH, + CO % of tots1 oarbon .asifled

1700

2400

zag; 21,601 1806

9.0

0.8 1 72.50 68 24 48. 849 30.4 8a.4

460

INDUSTRIAL AND ENGINEERING CHEMISTRY

rent until fusion begine. The voltage may be varied between 35 and 120 volts. The current is initially introduced at 110 volts and held at that voltage until the total current reaches 800 amperes; this ueually requires 20 to 25 minuks. The voltage is then adjusted to maintain this maximum amperage for another 60 to 80 minutes When the fused eone in the furnace a p proaches 8 inchen in diameter the current is shut off and the pig allowed to cool. The 0001 pig, weighing about 600 pounds, is lifted out of the furnace and broken with mauls until the pieces are 4 to 6 i n c h or leas in size. After the initial crushing the materid is shed to 8 to 18 mesh in two amooth-roll crushem in aeries. Oversize material ia returned to the mill. Undersize goeg back to the melting furnace. The raw catalyst has the composition:

double column as in the Linde process, but in place of the conventional tubular heat exchangers, corrugated a l d u m drips packed in towers act ~dheat reservoirs ($0). By a reversing flow cycle the Strips are alternately cooled by the pmduct nitrogen or oxygen and heated by the incoming air to be fractionated; by an unusual flow system 7% of the air intake is compressed to 3000 pounds per square inch gage, 73% is w m p r d to 65 pounds per square inch gage, and the remaining 20% is compreeaed to only 12 pounds per quare inch gage. All air streams are purified and cooled before entering the fractionation columns. The two high presrnve atreama enter the lower column, and the low pressure stream is intmduced into the upper column (86). oxygen of better than 98% purity is pmduced routinely by this plant; 99.5 and 95% pure nitmgen may ala0 he obtained. A t p r e b ent the nitrogen is used 88 a purge gaa tJlnmghout the plant or exhausted to the atmo%phem.

Vol. 44, No. 3

COW I"

Per Cent

k

;K*E

I

0 0

SiO,

This material is reduced with hydrogen at 800' to 1ooO' C. in a vertical, steel veaeel which can be op.rated either batchwise or continuously The single unit haa a capacity of 2OOO pounds per day. Hydrogen is obtained by the catalytic cracking of natural gaa in the presence of ateam and ahifting the carbon monoxide to carbon dioxide in a shift converter. Carbon dioxide can then be partially eliminated by scrubbing with sodium hydroxide to give a reducing gas with the wmpnsition: 93.7% I&, 0.6% C a , 0.8% CO, and 0 6% CH,. The feed rata is controlled to give 1500 standard cubic feet per minute of hydrogen. The hydrogen is preheated to 1ooO' F. in a gagfired furnace and flows upward. The heat for the endothermic reaction is supplied completely by the sensible heat of the hydrogen. Catalyst is fed into the tap of the reducer and removed from the bottom through a star wheel valve and immediately immersed in coolant oil to prevent contact with the atmosphere. Sealed in steel druma the catalyst in oil will retain its full activity for at least a year. It ia charged to the reactor while still suspended in the coalant oil. The h a 1 catalyat ia 85 to 95% reduced. In the reactor the catalyst in preaent practice is still further activated by a M a y period of stepwiae heating in the presence of coolant oil. After the catalyst charge is in the react07 and the vessel is sealed, the preasure is raised to 300 pounds per quare inch gage by the introduction of synthesis gan, and circulation of the cooling oil is begun. The circulating oil is heated to 400" to 425" F. over a period of about 12 hours. About 25% of the carbon monoxide and hydrogen in the feed stream should be reacting by this time: if not, the temperature is raised further until 2.5% reaction ia obtained. This temperature is held for 20 hours and then the temperature is increased to the point a t which 40% of the reactive gases have reacted. This temperature is usually in the range 4 3 5 O to 460' F. and is maintained for an additional 20 hours. For the next 20 hours the temperature is

Catalyot M u m tion The German COUlmercialFiacberTropech plants all used cobalt cab lysts of type looco: 5 T h 0 , : :MgO: 200 K i e s e l g u h r (58). However, when the Bureau of Minea undertook their study of the l i q u i d fuel from Figure 9. Kcrpcly Gas Pmduar coal processes their purpase was to explore thepossibilitienofexkmive indvtrial utilization. A casual survey of the known cobalt murcea revealed that there is not enough of this metal available a t any price to supply catalyst to an induatry comparable in size to the petroleum industry. Conmuentlv. the bureau turned ita attention to the iron camlysts which the; had frmt aaed in small scale trials a t the Pittsburgh laboratorien in 19!B and which were later used experimentally in Germany. Theee catalysts were virtually identical to t h m commonlyusedforammoniasynth&,andit wasdiscoveredthat commercially available ammonia catalysts would serve equally well for this purpase. However, few manufacturers produce thia type catalyst for general sale and in order to ensure a aupply and to have mntrol over the wmpnsition and charscteristica of their material the bureau undertook to manufacture their own (Figure 10). This type catalyst is manufactured by fusing iron oxides, usually magnetite, with promoters and then crushing the pig and reducing with hydrogen. Activity experiments revealed that a mmt eEeetive catalyst could be made by using ordinary mill scale as a m r c e o f iron and magnesium oxi& and potaseium carbonate as alkaline promotem The optimum mix haa 98.7% mill male, 4.6y0 magnesium oxide, and 1.7% potassium carbonate. The ingredients are mixed in an ordinary concrete mixer and poured into a half-cylindrical, ateel trough about 10 feet long and 2 feet in diameter. Two water-cooled electroden are plsoed in the mix about 9 feet apart, and fcur small air lanoes are inserted into the c m spaced e d y between the electrodes. Current is sup plied hy a batof thrw Nbkv.8. traneformers. Pieces of iron wire nra hid just under tbe avface ofthe Charge to carry the cur-

69 2 39 5 46 06

Total Fe

TABLE 111. HORIZONTAL GASIFIERPERFORMANCE DATA steam tamp. F. W w t i o n tamo. Feedrate

*

.

imo

F.

2480

F a i d Ib hr Oxygen ( b 6 d p & ) , stand. (I".ft./hr. 8 t a m Ib./hr. Oxy& (100%) to dry car1 ratio. stand OU. ft. Pb. S t a m to dry -1 nrtia. Ib.pb. 8-m deoom-d. % produst. Total atand. ou ft./k. E, atand. OU. ft./hr. Lb.dry d / 1 0 0 0 e k d . CY. It. ofE, CO Stand. ou. ft. oxyg,en/1000 stand. cu. ft. of

+ %.

+

&+co Lb. of Stam/1000 atand. en, ft. of ffi + CO

% of total wrhon .asiBed

2297

m.mo 1SW 9.0

0.8 16

72,500 6e.a~)

as. s u9

30.4 88.4

INDUSTRIAL'AND ENGINEERING CHEMISTRY

46a

MAGNESIA

I

conveniently thought of as being in two parts-the furnishing of purified s y n t h i s gas and the synthesis of that gas into hydrocarbons. A t Louisiana, advantage has been taken of this fact to develop the two phases of the process independently. As won as ju&ed, the making of synthesis gas from coal will be integrated with the synthesis step into a combined Nn. Of the two primary reactions in the gasifier, the one between the steam and the coal is endothermio and that between the coal and oxygen is exothermic. At 2200' F., for exnmple:

C

+ H*O = CO + H?

c + ZH,O = co, + 2n, c + l/% = co

c+o,=co,

B.t.u./Lb. Coal Converted -4,842

-3,755

(1)

+4,125 +14,180

Feed proportions are adjusted so that the heat generated by Reactiom 3 and 4 is sufficient to promote Reactions 1 and 2. As the rate of oxygen feed is increaeed the ratio of carbon dioxide to carbon monoxide in the product is increased along with the amount of both compounds that are formed. Since carbon dioxide is useleas or even harmful to the ultimate synthesis reaction, this is not d&ble. However, incrpased oxygen feed also increases the percentage of coal gasilied and maintains the ternture of the reaction chamber at a point where the reaction between carbon and steam is pmmoted. There are other ways of mnhtairiihg the reaction temperature which are less expensive than the introduction of additional oxygen. Preheating the feed componenta is the most practical of these. In the original experiments at h u i b the steam WBB preheated to 88 high 88 2400" F.in order t o eave oxygen. However, heiting to such high tempernturn requires special equip

Vol. 44, No. 3

m t and thin proved to be more exptmsivethan additidoxygen. hp~toperationsthebesteam is preheated to about looOo F., which can he done in conve.ntiod ateam equipmeut, and the rate of oxygrm feed has h e n raised. &me addittonal oxy&8n can aW1 he saved by preheating the coalsnd~aawell~theBtesm,aodmdilications to provide such pre6eating are now being studied. The higher*the mtio of ateem to coal in the feed the more Btesm m&a to produce h y d m to be used in the synthesis reactton. Honeper, ea this ratio increases, the percmtagv of ateem that reacta decreases. At a poUna-for-poUOa ratio about 20% of the atesm normslly reaatk Since the hydwgem-forming reaction is endothermic, the temP e I a in ~ the rsactor will drop astbeamount of Bteam introduced is i n d . some heat is also lost through the heating of umeacbd stesm. Consequently, the rate of introduction of stesm is u d l y kept relatively low to conmve heat in the gaeifier. Btesm-coal ration 88 low ea 0.31 have been used with high degrees of superheat, The deeree of contact hetween the coal particles and the steam hae an importgnt effect on the amount of steam that is d e c o m p d in the gasifier. Thus, the rate of this reaction is related m v d y to the &e of the coal particles. This effect is accentuated at very m a l l particle &ea. Here agiu an ewnomic balance is reaohed be-

tween the value of the hydrogen p r o d u d and the cost of pulverizing the coal. Although originally intended to use larger &ed coal, the h u b ana plant now grinds to 70% through 325 meah, nhich 888m8 to be about the optimum &e. To data about 15 ahohtime NOS have been made on the vertical gasifier. In carrying out these NUS usually about a !&hour natural gas heatup is allowed, followed by a 6 1 ~c4o a l - ~ ~ g eheatup u to bring the gspifier to equilibrium temperature, and finally a data collecting period of 4 to 12 hours normally. After such a NU the femt.01 can be coaled and opened for internal inspection. The program is of a development nature rather than one of demonstration, because there are many new featurea that have been incorporated in this unit. The methcd of feeding with an intimate m*ture of oxygen and coal tangentially in a vertical gasi6er of this size baa only meager background. The e m t methods of mixing in the steam is ale0 in the exploratory stage. Thetypeofliningforthihisuseisnew. In this development there are eeveral objectives: 1. To study the thermodynamics, kinetics and mechanics of gasification in general performed at atmospheric pressure. &me of the information obtained is expected to have application to preseure gaeification. The plant data studies are accompanied by thco&ical studies on the thermcdynamics and kinetics of coal gasiscation (6). 2. T o waluate the chemical and physical factors involved in the operation of tbe vertical gaaiSer at hand. This differs from the &at objective in that the results and observations may be applicable only to the specific equipment being used. Certain physical limitations that are d-ed into this particular gasifier may in some caeea limit the general application of the result~. Tbe experiments to date have been carried on with Rock Springs, Wyo., coal, e0 ae to be comparable torasults from previous work. Other coals are to be teeted, eapecially e0me of those with lower melting ash, 80 88 to facilitate slag tapping. 3. T o determine the usefulness of this gasifier, with only minor mcdifications, 88 a supplier of synthesis gas to the FiscberTropsch demonstration reactor.

:

'

. ,

I N D U S T R I A L ' A N D E N G I N E E R I N G CHEMISTRY

r-

thia proved to be more expenaivethan additionaloxygen. Inpresemtoperstionstheateam is preheatd to about 1000° F., whiah om be d m in c o n v e n t i d ateam equipmmt, and the rataofaxygenfeedhasbeenlaised. Someaddit i d oxy&au om still be e a d by preheetiug the cadand axygen as well as the stesm, and modscations to provide such preheating are now behg studied. The higher-the ratio of ateam to c d in the feed the mom ateam raacts to produce hydrogen to be used in the synthesis reectton. Hoaeoeo; as this ratio increases, the percentag~of ateam that reacta decreases. At a pomd-fmpmnd ratio about zO% of the n o d y reaatk Sinus ihe hydmgen-foresction is endothmnia, ihe tempersture in the reactor will dmp asthesldount merit and

'

OfateamintmdUoediBind

I

m 10. Catalyet Production Flow Diagram

conveniently thought of 88 being in two parte-the furnishing of purified BJmtbeais gas and the synthesis of that gaa into hydrocarbona At Louisiana, advantage baa been taken of this fact to develop the two pbases of the process independently. As won 88 jnstilied, the making of s y n t h d gas from cod d l be integrsted with the s y n t h d Step into a combined run. Of the two primary reactions in the &er, the one between the steam and the coal is endothermic and that between the coal and oxygen is exothermic. At a200° F., for example:

++ HIO = CO + Hs 2HrO = COS + 2Hz c + l/ao, = co c+o,=co,

C C

B.t.u./Lb. Coal converted -4,842 ~

~~~

~~

~

+4,125 +14,180

8

Feed proportions are adjusted e0 that the heat generated by Reactions 3 and 4 is Sufficientto promote Reactions 1and 2. AB the rate of oxygen feed ia increased the ratio of carbon dioxide to carbon monoxide in the product is increased along with the amount of both compounds that are formed. Rmce carbon dioxide is useless or even harmful to the ultimate eynthesis reaction, thia is not desirable. However, increased oxygen feed also inc-es the percentage of coal &ed and maintains the temperature of the reaction ohamber a t a point where the reaction between carbon and steam is promoted. There are other ways of maintaining the reaction temperature which are less expensive than the introduction of additional oxygen. F'reheatiug the feed components is the most practicsl of theae. In the original experiments a t Louisiana the steam was preheetd to as high as uoOoF. in order to eeve oxygen. However, h&ting to such high temperaturea requires special equip-

somehestis

alsolcntthrounhtheheatinrtofunreactedstesm. couaquently,the rata of i n k c t i o n of stesm is usually kept d a t i v e l y low to conaena heat in the g d i e r . Steam-cad ration as low ea 0.31 have been need with high degrea of supheat. The degree of contact between the cad particlee and the steam has an important eEect on the amount of stesm that is d e c o m p d in the gasifier. Thus, the rate of this reaction is related inversely to the &e of the coal particles. This efiect ia accentuated a t very d l particle &ea. Hem again an economic balance is reached between the value of the hydrogen produced and the cost of pulveridng the coal. Although originally intended to use larger sised cad,the Louisiana plant now grinds to 70% h u g h 325 mesh, which

F

Vol. 44, No. 3

=

'

~ e e m ato

be about the optimum &e. N B II have been made on the d u d gaeifier. In uurying out these run8 usually about a 20hour natural gaa heatup is dlowed, followed by a hal cad-axygen heatup to bring the M er to equilibrium temperature, and h a l l y a data coll& period of 4 to 12 hours normally. After such a mn the kactor can be -led and opened for internal inspection. The program is of a development nature rather than one of demonstration, because there are many new features that have been incorporated in this unit. The method of feeding with an intimate mixture of oxygen and coal tangentially in a vertical gssifier of this size has only meager backgmund. The exact methods of mixing in the steam is ale0 in the exploratory stage. The type of l i i for thin n8e is new. In this development them sre w v e d objectives: 1. To study the thermcdynamics, kinetics and mechanics of Basification in general performed a t atmospheric pressure. &me of the information obtained is expected to have application to presaura &cation. The p h t data studies are accompanied by theoretical studies on the thermodynamics and kinetics of coal gaaifioation(8). 2. To evaluate the chemical and physical factors involved in the operation of the vertical gaeifier a t hand. This d8em from the first objective in that the results and observations may be applicable only to the d c equipment being need. Certain physical limitations that are designed into this particular gasifier may in wme c 8 8 e ~limit tbe general application of the resulk The experiments to date have been csrried on with Rock Springs, Wyo.,coal, wasto becomparabletorenultafrompmvious work. Other coals sre to be teeted, +ly nome of thoee with lower melting ash, 80 as to facilitate slag tapping. 3. To determine the ueafdness of this gaa'kr, with only minor modifications, as a supplier of s~mthesiagas to the &herTropscb demonstration reador.

To data about 15 Bhorttime

.

March 1952

INOUSTRIAL AND ENGINEERING CHEMISTRY

The prelnninary indications are that better operations can be obtained from This unit than with the horizontal gasifier previously used, perhaps to the extent of 5 to 10% improvement. This may be attributed to the higher turbulence and better steam mix. Also, there is a difference in the two in that the vertical gasifier is predominantly slagging. In all of the gasification work the objective is not necessarily to obtain maximum gasification, although if too much coal is left ungasified it normally will be entrained in the made gas stream and become a handling and disposal problem. An approximate measure of good gasification is expressed by the minimum coal and oxygen requirements for making a unit quantity of carbon monoxide plus hydrogen. The provisions for the purification of the synthesis gas have been satisfactory. The alkalized iron oxide boxes originally installed to remove the last traces of organic sulfur from the gas stream exhibited a tendency t o local overheating and one of them actually developed a crack in its wall from such a condition. Fortunately, the activated carbon boxes installed for the same purpose proved so effective that the final purification step could be eliminated without impairing the quality of the synthesis gas. The scrubber in the purification section was intended to use triethanolamine as an absorbent. This material was chosen because it was slightly selective and would remove all the hydrogen sulfide and only a minor part (about 35%) of the carbon dioxide. However, further studies showed that carbon dioxide 11 as not desirable in the gas stream. It reduced the capacity of the activated carbon for absorbing organic sulfur compounds and appeared to impair the activity of the reactor catalyst. Consequently, diethanolamine was substituted for this service. This absorbent has a higher capacity for absorbing sulfides than the triamine and absorbs nearly all the carbon dioxide. I n the synthesis section of the plant monoethanolamine is used t o absorb carbon dioxide. The monoamine has greater capacity for carbon dioxide absorption than the di- and tri-materials, but its effectiveness is impaired by the presence of sulfur. No sulfur is present in the product streams. The electrostatic precipitators have proved satisfactory when run on the products of either of the two gasifiers. However, the Rerpely generator operates a t much lower temperatures than the

463

gasifiers and distills over considerable tars with its gas product. These tars condense in the precipitators causing increased pressure drop which often requires cutting them out for cleaning. However, since the Kerpely unit is only a stopgap producer, no steps are contemplated t o correct this situation. The new charge of catalyst recently put in the reactors has performed satisfactorily although it had been stored in coolant oil for 6 months. After the 25-day run it showed no detectable decrease in activity. Conversion efficiencies of more than 80% have been achieved consistently in the Louisiana reactor. With a recycle ratio of 1: 1, this means that 40% of the total gas through the reactor actually reacts. The recycle gas has a higher hydrogen-carbon monoxide ratio (about 4:3) than the fresh feed gas (about 4:5) and may accumulate excess methane (4 to 5%). Because of this difference the rate of recycle can be used t o adjust the composition of the feed to the reactor within certain limits. Pilot plant runs a t Bruceton a t hydrogen-carbon monoxide volume ratios of from 0.7:l to 1.4:l indicate that about 1:l ratio gives the best performance in the jiggling bed type reactor. This ratio can be maintained with the recycle gas. Higher hydro-

TABLEIV.

REACTOR PERFORMANCE

(Half capacity) 61,900 Fresh feed gas flow stand. cu. ft./hr. 0 76 H d C O fresh feed ;as, vol. 1.03 Recycle ratio, vol. Maximum temp., OF. 524 15 Temp. gain in reactor, OF. 329 Maximum pressure, Ib./sq. inch gage Pressure drop throu h reactor, lb./sq. inch 30 Space velocity, vol.,%ol./hr. 53 1 500 Coolant oil circulation rate, gal./min. Conversion, % 85 9 0 71 Ratio of Hn to CO reacted COZin fresh feed gas % 2 0 0 1 C O in ~ recycle gas, d 1 2.2 COz in gas from reactor, % CI Cz yield, grams/cu. meter (CO H P )in fresh 24.6 feed CS yield, grams/cu. meter (CO Hn) in fresh 154.0 feeda (1 Celculated.from theoretical relationships, 3'% conv. X 2O8-(C1 yield) = Ca+ yield.

+ +

+ +

+

C

INDUSTRIAL AND ENGINEERING CHEMISTRY

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gen ratios would require the addition of hydrogen from an external source, probably reformed natural gas, or the passing of the synthesis gas through a shift reactor to decompose more steam (Table IV). When the Louisiana plant was originally designed it was thought that some such step nould be necessary. Fortunately, further studies revealed that additional hydrogen would not be needed with the result that the cost of the synthesis gas was reduced by several cents per thousand cubic feet.

TABLE V. UTILITYREQUIREMEKTS (Design) Steam, Lb./Hour 275 50 lb./sq. lb./sq. inch inch gage gage 1550 3500 2600 16,100 2620 12,500 1840 1900 150

Kater, gal./ hour 2450 57 > 100 10,800 34,000

Power (440 V.) Max. Load, Hp. 425.2 70 196 17.8 66 531 286

Laboro, Men/ Day 17 17 17 17

Gasification Purification Synthesis Refining .. ... Catalyst preparatia 17 7500 .. Oxygen plant 11 950 109,000 4izo General service .. ,.. .. 1500 300 Line losses __ __ 1592 96 14,660 210,850 36 670 Totals From waste h e a t boilers 12,500 Recycled 193,860. Net 24,170 17,000 Steam reacted 1,550 8,850 a Includes all supervision a n d operating labor b u t no maintenance utility personnel.

__

~

-

01’

The major difficulty in the reactor section of the plant has been caused by the physical deterioration of some catalyst and its resultant carryover into the process lines. This carryover was so severe during the initial run that the filters in the cooling oil recirculating pumps, which were intended to remove catalyst fines from the coolant, had to be removed because they were overloaded and became plugged. Some heat exchanger piping recently opened lvas found to be one quarter filled with deposited catalyst fines. It is believed that this condition was severely aggravated by the condition of the initial batch of catalyst placed in the reactor. This catalyst was from the first batches produced and was not completely reduced. It was used because the coolant oil was known to contain some sulfur compounds which would poison any catalyst that was used. Because of this poisoning and the deteriorated condition of the catalyst, the initial reactor run gave a gas conversion efficiency of only about 30%. In an attempt to increase this efficiency the operating temperature of the reactor was raised, rrhich is known to accelerate catalyst deterioration. Later runs made with properly reduced catalyst and sulfur-free cooling oil achieved conversion efficiencies up to SO-SS%, and indicated that even better results may be possible. In the second run, of 25 days’ duration, it was found that the catalyst fines could be allowed to circulate through the coolant oil pumps with the oil without immediate damage to the pumps. At the time of shutdown of the plant, the pump used was still operable but shoxred some wear. This indicated that with proper seals and use of flushing oil it is probable that a pump of this design can be maintained in service for an extended time. Possibilities of Pressure Gasification

Bureau of Mines engineers have believed for some time that the economics of coal gasification could be improved by increasing pressures to 300 to 600 pounds per square inch gage. The present atmospheric pressure investigations have been continued at Louisiana because, for the purposes of studying the reaction phenomena, the lower pressure is much easier and cheaper to work with and some of the observations may be validly extrapolated to high pressure conditions. Hoirever, the pilot plant

Vol. 44, No. 3 * laboratories a t Morgantown are now working primarily on pressure gasifiers and recommend that pressurized processes be u-ed in any subsequent commercial installation. Pressure gasification has the advantage of using much smaller equipment to attain the same gas capacity. One experimental unit having one fourth the capacity of the Louisiana gasifier is about the size of a home hot water heater. Gasification under pressure also saves compression costs. The products of the gasifier must eventually be compressed to the reactor pressure of about 400 pounds per square inch gage. In atmospheric gasification the entire volume of synthesis gas must be compressed. If the gasifier is under pressure only the volume of the oxygen fed to the reaction need be compressed. The introduction of pressure gasification will undoubtedly result in a change of the feed system of the gasifier. A reliable system now in use is a fluidized feeder developed by the Bureau of Mines a t its Morgantown, W. Va., station. This feeder supplies pulverizcd coal at a steady rate to a pilot-scale gttsifier under prcssure. Consistjng essentially of a fluidized bed of coal and a coal delivery tube entering near the bottom of the hed, the feeder produces a uniform ratio cf coal to oxygen. If the pressure drop causing flow and the ratio of coal to gas are constant, the rate of coal delivery to the gasifier is constant and this steady stream can be maintained by suitable instrumentation. Successful tests have been made on feeding 90~-through-200-mcsh coal, lOO%-through-lO-mesh coal, sand, and pulverized limestone. The provisions for removing slag from the present vertical gasifier are rather primitive but are expected to serve for the present. Ultimately the bureau may install an arrangement similar to that now used in some large power plant boilers in which the molten coal ash drops through a small opening in the bottom of the combustion chamber into a pool of circulating water which quenches the slag in small granules which can be removrd continuously. This technique is now used at the Morgantown pilot plant. XIATERIALS OF CONSTRUCTION

The gas synthesis process as practiced a t Louisiana picsents no particular structural problems. The pressures encountered are moderate. The hottest part of the process is the coal gasifier 15 here temperatures up to 3000” F. must be anticipated, but this is still within the range commonly served by commercial insulations and lining materials. Neither corrosion nor erosion problems are unusual, although some pumps are fitted with stainless steel trimmings to reduce the ear caused by ash and bolid particles in certain process streams. The original horizontal gasifier was constructed of mild s l e d and lined with I l l 2 inches of insulating firebrick, 41/2 inches of common firebrick. and 9 inches of silica brick. The inner lining of silica brick resists corrosion by the acid coal ash which would rapidly dissolve an alkali brick. The acid brick, however, has the disadvantage of being highly sensitive to thermal shock. Consequently, it was necessary to heat and cool the vessel very slowly. Heating to operating temperature from complete shutdown required 3 weeks and a like period was required for complete cooling. Even with this gentle handling there was some deterioration of the surface, presumably caused by frequent start-up and shutdown. The geometry of the new vertical gasifier permitted the application of a recently introduced lining technique. A wooden shell was built inside the vessel and a “ram mix” composed principally of tabular alumina was poured into the space between the shell and the outside of the reactor. About 2 or 3 feet of the mix was put into the annulus and then men with viooden tampers rammcd it down as much as possible. This procedure was repeated until the entire annulus was filled. The wooden shell and its internal wooden cross-bracing were then burned out by introducing a natural gas flame into the reactor. The gasifier was then filled with

March 1952

INDUSTRIAL AND ENGINEERING CHEMISTRY

465

Linde-Frankl Oxygen Plant High pree%urecompressor (foreground) and low tempera ture fractionating column (background)

aluminum oxide pebbles to give even heat, distribution and fired with natural gas for 2 days at 3100" F. to dry the liner and set the binders. Aluminum oxide fuses a t 3700" F. This lining was furnished and installed under a cooperative agreement between the Bureau of Mines and the Aluminum Co. of America. Some deterioration of the lining has been noted in the area directly across from the oxygen-coal burner and restricted areas adjacent. At other points in the gasifier, where temperatures have normally not exceeded 2800" F. for prolonged periods, the lining is in excellent condition. On one occasion when slag accumulations on the burner nozzles diverted the coal and oxygen flame onto the vessel wall a spiraling groove was burned into the lining. However, the operating engineers estimate that such direct impingement of the flame must have produced temperatures up to 5000' F. The groove was subsequently filled with more of the ram mix and set with a gas flame. This repair has performed satisfactorily. Most of the equipment in the gas-synthesis plant is fabricated from carbon steel, or where physical strength requires it, low alloy steel. The most notable exceptions are the receivers of the vapor product from the primary fractionation column. The feed heater exchanger has tubes of Type 304 stainless steel. The condenser which follows the heat exchanger has tubes made of admiralty metal (70% Cu, 1% Sn, 29% Zn) and naval bronze tube sheets. The coal feed pipes are also made of Type 304 stainless to minimiee erosion, and, more important, to provide greater high temperature strength in the event of a blowback. I n spite of this precaution, a blowback did actually burn through one of the coal feed lines. The Linde-Frank1 oxygen plant is fabricated almost completely of pure copper because of this metal's superior shock resistance a t low temperature. Current American practice is to make such installations of aluminum, but there seems to be little difference in performance between the two metals.

ECONOMICS

Since the several processes for the manufacture of synthetic liquid fuels each affords considerable freedom in choice of method, raw materials and other conditions, it is not surprising that there are differences of opinion concerning their relative economics, The Bureau of Mines made comprehensive studies of the cost of producing synthetic fuels from oil shale and by coal hydrogenation in commercial size plants. The studies on coal hydrogenation have been released in the form of a report on a pro osed coal hydro enation plant of size to produce 30,000 barreg per day of profuct (43). The National Petroleum Council has also made an economic study of the coal hydrogenation process which produced figures considerably different from those submitted by the Government. At least part of this differential can be attributed to basic disagreements on appropriate accounting practices, the extent t o which community facilities must be charged to the fuel plant, and the ultimate value of the nonfuel roducts of the process. Still another survey, this one by an injependent consultant, is now under way. No one has yet published a detailed economic study of the coal gasification process at its present state of development. As a result, discussions of relative economics of the roduction of liquid fuel from coal by hydrogenation or gas syntgesis, from shale, or directly from petroleum are still based largely on conjecture. Until more operating data from actual plants are available, and until some general agreement is reached on the roper techniques for calculating the cost of the process, such jiscussions will continue t o include much opinion. The completion of the gas synthesis plant a t Louisiana is a big step toward supplying some of the €actual data which will make it possible to determine the economic feasibility of an American synthetic fuels industry. BIBLIOGRAPHY

(1) Alden, R. C., Petroleum Enor., 18, 148-58 (January 1947).

(2) Atwell, H.V., and Schroeder, W. C., U. S. Dept. of Commerce, OTS Rept. PB 367 or 412 (May 15, 1945). (3) Batchelder, H.R., Dressler, R. G., et al., U. S. Bureau of Mines Rept. I m e s t . 4775 (March 1951). (4)Batchelder, H.R.,Dressler, R. G., et al., "Operation of Kerpely Producer with Oxygen Enriched Blast," reprint of paper presented to American Gas Assoc., Atlantic City, N. J. (October 1950).

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(5) Batchelder, H. R., and Ingols, H. A . , U. S. Bur. Mines, Rept.

Invest. 4781 (March 1951).

(6) Batchelder, H. R., and Sternberg, J. C., IND.ENG.CHnx, 42, 877 (May 1950). (7) Benson, H. E., Crowell, J. & et I al., ., Proceedings Third World Petroleum Congress, Sect. IV, Subsect. I, Preprint 4 (Nov. 15, 1950). (8) Bergius, F., J. Gasbeleucht., 54, 784-9 (1912). (9) Bergius, F., Billwiller, J., Ger. Patent 301,231 (Aug. 1, 1913); [Brit. -4ppl. 18,232 (Aug. 1, 1914)l. (10) Blatchford, J. W.,“Water Gas Production with Tonnage Oxygen,” presented a t Prod. and Chem. Conf., Americah Gas Assoc., New York (May 1950). (11) Colten, E., NatE. PetroEeum News, 38, R-425-34 (1946). (12) Crowell, J. H., Benson, H. E., et al., IND. ENG.CHEM.,42, 237684 (1950). (13) Dressler, R. G., and Bircher, J. R., U . S.Patent 2,574,469 (Xov. 13. 1951). (14) Duftschmid, F., Linchkh, E., and TVinkler, F., U. S. Patents 2,159,077 (May 23, 1939); 2,207,581 (July 9, 1940);2,287,092 (June 23, 1942); 2,318,602 (May 11, 1943). (15) FIAT films,Library of Congress, Washington, D. C., Keel Yo. K-31, Frames 1108-15. (16) Ibid., Reel No. X-115, Frames 1388-1447, 1593-1672. (17) Faragher, W., and Foucher, J., “Syntheses from CO + Hz at I. G. Farbenindustrie,” 1,part. C (1947) (prepared at Ludwigshafen, translation, FIAT final report No. 1267, PB 97368). (18) Faragher, W., Horne,.W. A,, et al., G. S.Dept. of Commerce, OTS Rept. PB 1366 (1945). (19) Fischer, F., Tropsch, H., Brennstoff-Chem., 4, 276-85 (1923); 5, 201-8 (1924); 5, 217-27 (1924). (20) Ibid., 7, 97-104 (1926). (21) Hall, C. C., Craxford, S. R., et al., H.M. Stationary Office, London, BIOS Final Rept. 1712. Hall, C. C., and Haensel, V., U. S. Dept. of Commerce, OTS Rept. PB 415 (1945); Rept. 30, File 27-69 (January 1946). Hall, C. C., and Jones, J. P.. Ibid., PB 294 (1945) [Gas Age, 97, No. 3 (1946)l. Hall, C. C., and Powell, 9.R., U. S.Dept. of Commerce, OTS Rept. PB 286 (1945). Kastens, M. L., IND. ESG. CHEM.,40, No. 6, 1OA (1948). Kastens, M. L., Hirst, L. L., and Chaffee, C. C., IND.ENG CHEM.,41, 870-85 (1949). Keith, P. C., Chem. Eng. News, 25, 1044 (1947). Keith, P. C., OiE Gas J . , 45, No. 6, 102 (June 15, 1946). Latta, J. E., and Walker, S.W.,Chem. E n g . Piogiess, 44, 173 (1948). Lund, G., and Dodge, B., IND. ENG.CHEM.,40,1019-32 (1948). Mitchell, R. F., Can. Chem. Process I n d , 30, 34 (August 1946). Neumann. R.. and Schroeder. W.C.. U. S. DeDt. of Commeice. OTS R&t. PB 1279 (1945). Pier, M.. and hlichael, W., U. S. Patent 2,167,004 (July 25, 1939). Reichl, E. H., U. S. Dept. of Commerce, OTS Rept. 22841 (1945). Russell, R. P., Chem. Eng. News, 25, 86-7 (1947). Schlesineer. M. D.. Crowell. J. H.. et al.. IXD.ENG.CHEX. 43, 1474-9 (1951). Sebastian, J. J. S.,Edeburn, P. W., et al., U. S. Bureau of Mines, Rept. Inuest. 4742 (January 1951). Storch, H. H., Golumbic, N., and Anderson, R. E., ”The FischerTropsch and Related Syntheses,” New York, John Wiley &Sons, Inc., 1951. Storch, H. H., Powell, A. R., and Atwell, H. V., U. S. Dept. of Commerce, OTS Rept. PB 2051 (1945). Technical Oil Mission films. Library of Congress, Washington, D. C., Reel 37, Bag 3451, Item 21. Ibid., Reel 47, Bag 3446, Item 96. Ibid., Keel 134, Item I-a, No. 8, Item I-b, No. 11, 12, 13. Ibid., Reel 12, Bag 3043, Item 7, Frame 150. U. S. Bureau of Mines, Intrabureau Report, “Cost Estimate for Coal Hydrogenation” (October 25, 1951), revised Jan. 11, 1952. U. S. Congress, 78th Session, Public Law 290 (April 1944). U, S. Congress, 80th Session, Public Lam 443 (hlarch 15, 1948). Wright, C. C., Barclay, K. RI., and Mitchell, R. F., IXD.ENG. CHEM.,40, 592 (1948).

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Processing Equipment (IE) American Pulverizer Co., St. Louis, Mo., ring mill crusher

WC-18, 15 ton/hour to - 3 / a inch, 20-hp. motor drive, 700 r.p.m. (2E) Babcock & Wilcox, Inc., 85 Liberty St., New York, KO.138, Type B pulverizer, 75-hp motor drive, 1170 r.p.in.

Vol. 44, No. 3

(3E) Ibid., No. 48 centrifugal type exhauster, 21,200 lb. of air/hour

a t 40 inches water pressure differential, BO-hp. motor drive a t 1765 r.p.m. (4E) Ibid., special Stirling. water-cooled, hopper-bottomed. wastc heat boiler, 1772 sq. f t . of heating surface. (5E) Brown Fintube Co., Elyria, Ohio, No. lh.-1322D. (6E) Ibid., No. 1AA-2420. (7E) Buffalo Forge Co., Pittsburgh, Pa., Xo. 72, Type F Acrofii~ coil heaters, 21,200 lb. of air/hour to 271’ E’. with 100 lb./sq. inch steam. (8E) Dracco Corp., Cleveland, Ohio, five-compartment AAA Dracco bag filter, low pressure type, 6-foot cyclone precollector included, 3 tons of dust/hour. (9E) Foster-Wheeler Corp., Pittsburgh, Pa., Ic’Y-420-1087-D, tencoil vertical heater, 190,000 B.t.u./hour output, SA T22 tubes(2-2.5 Cr, 1.0 Mo). (10E) Ibid., S-70 Dowtherm vaporizer, 2,700,000 B.t.u./hour output. (11E) Ibid., natural gas-fired process furnace, 240,000 B.t.u./hour a t 100 lb./sq. inch and 915O I?., 2-inch tubes, SA T22 steel (2-2.5 Cr, 1.0 RIo). (12E) Gould Pumps, Inc., Seneca Falls, N. Y., Figure 1713, 3 fcot X 6 inch iron triplex positive displacement, 31 gal./min. at 475 lb./sq. inch gage. (13E) Ibid., Figure 3360, 2-inch, eight-stage turhine-driven ccntrifugal, 75 hp., 3550 r.p.m., 85 gal./min. at 680 lb./sq. inch gage, all iron. (14E) Griscom-Russell Co., Ic’ew York, N. Y., S o . 17-8-168, Typo SSE,shell-and-tube heat exchanger. (15E) Ibid., No. ASS5, Twin-0-Fin heat exchanger. (16E) Ibid., No. 2555-248, Twin-G-Fin heat exchanger. (17E) Ibid., No. CS2S-248, Twin-G-Fin heat exchanger. Ibid., No. AA85-248, Twin-G-Fin heat exchanger. Ibid., KO.19-6-192, Type FST, shell-and-tube heat exchanger. Ibid., No. C112A-248, Twin-G-Fin heat exchanger. Ibid., S o . CSS-248, Twin-G-Fin heat exchanger. Ingersoll-Rand Co., Pittsburgh, Pa., Cameron No. 4, Class SFL single-stage centrifugal pumps, 600 gal./min. a t 600” F., 50 hp. motor drive, 12% Cr hardened stainless steel internal parts. Ibid., Class 63-PL-16, horizontal pipeline-type cooler, 90,000 stand. cu. ft./hour with 24 gal./min. water a t 90’ F. Ibid., Type XPV-3, horizontal, two-stage, size 14 X 30X 23l/r X 10 X 24, reciprocating compressors, 250 lb. steam drive, capacity 900,000 stand. cu. ft./hour a t 450 lb./sq. inch gage. Ibid., Class ES-2, 6 X 2 5 / 8 X 7, two-stage horizontal centrifugal pumps, 25 hp., 1750 r.p.m., squirrel-cage induction motor drive. Research Corp., Bound Brook, N. J., SK-3929-L, 5152h precipitators, 1630 cu. ft./min. a t looo F., met flushing type, 35,000 kv. max. Spencer Turbine Co., Hartford, Conn ., multistage centrifugal, blower, 416 cu. ft./min., 25-hp. motor drive, 3500 r.p.111. Type 316 stainless steel internal construction. Ibid., 2800 cu. ft./min. booster compressor, centrifugal t w c , 40-hp. motor drive, 3500 r.p.m., Type 316 stainless steel internal construction. Struthers-Wells Corp., Warren, Pa., Type GY11-6H heat exchanger. Ibid., Type U42-6H, shell-and-tube heat exchanger. Ibid., 11-inch 0.d. X 10-foot kettle type, No. IJ-12-6H, shclland-U-tube heat exchanger. Ibid., Type T16-6H, shell-and-tube heat exchanger. Ibid., Type Tll-6H, shell-and-tube heat exchanger. Ibid., Type 28H-6H. shell-and-tube heat exchanger. Surface Combustion Co., Toledo, Ohio, special G.C. indirect duct type, gas-fired heaters, 21,200 Ib. of air/hour from 250n to 450’ F. Western Precipitation Co., Los Angeles, Calif., Type 9VG-12, Size 9-3 Mdticlone dust collectors. Worthington Pump and Machinery Co., Pittsburgh, P a . , Type HS-2, two-stage, steam-driven, reciprocating comprcssora, 9 X 11/5 X 11, 268 cu. ft./min. a t 320 lb./sq. inch discharge pressure. Ibid., Type HS, 9 X l l / 5 X 11, two-stage, steam-drivon 250 lb./sq. inch gage centrifugal, 268 cu. ft./min. at 320 lb./sq. inch gage. (39E) Ibid., Type HS one-stage, steam-driven, centrifugal romprrssora 9 X 10 X 11, 300,000 stand. cu. ft./hour. R E C E I V E for D review September IS. 1951.

Accepted February 1, 1032.