THE UNIVERSITY OF MICHIGAN INDUSTRY PROGRAM OF THE COLLEGE OF ENGINEERING A TECHNICAL AND ECONOMIC EVALUATION OF A SOLVENT EXTRACTION PROCESS FOR THE EXTRACTION OF URANIUM FROM DOMESTIC ORES M. E. Weech December, 1957 IP-253

ACKNOWLEDGEMENT The Engineering Research Institute of the University of Michigan sponsored this investigation and has approved distribution of the report by the Industry Program of the College of Engineering.

L The University of Michigan * Engineering Research Institute TABLE OF CONTENTS Page LIST OF TABLES iii LIST OF ILLUSTRATIONS iv 1.0 INTRODUCTION 1 2.0 SUMMARY AND CONCLUSIONS 1 3.0 DISCUSSION OF PROCESS 2 3.1 Present Practices in the Field 2 3.2 Proposed Process Basis 5 5.3 Steps in the Proposed Process 7 3.4 Utility Requirements 35 4.0 EQUIPMENT COSTS 36 4.1 Equipment Pricing Methods 36 4.2 Freight Charges 37 4.3 Delivered Equipment Costs 37 5.0 PLANT COSTS 42 6.0 PLANT OPERATING COSTS 44 6.1 Fixed Charges 44 6.2 Variable Charges 44 7.0 PROBABLE PAY-OUT TIMES AND PROFITS 49 8.0 PROCESS AND DESIGN PROBLEMS IN THE PROPOSED PLANT 50 9.0 BIBLIOGRAPHY 52 ii

t, - The University of Michigan - Engineering Research Institute LIST OF TABLES No. Page I Effect of Mesh Size on Direct Acid Leaching of Raw Carnotite Ore 14 II Time Dependence of Direct Acid Leaching of Raw Carnotite Ore 15 III Material and Labor Index 1950 Through 1956 38 IV Freight Rates on Various Classes of Equipment 40 V Summary of Equipment and Freight Costs 41 VI Summary of Pump Costs, Service, and Rating 43 VII Plant Cost Summary 45 VIII Plant Fixed Charges 46 IX Plant Manpower Requirements 47 X Plant Variable Charges 48 XI Plant Costs on a Per-Ton-of-Ore Basis 49 _ L iii

The University of Michigan * Engineering Research Institute LIST OF ILLUSTRATIONS Figure Page 1 Block diagram of salt-roast, acid-leach process. 5 2 Block diagram of ion exchange process for sulfate solutions. 4 3 Particle size distribution from rod and ball mills. 11 4 Ore extractor (leaching). 13 5 Countercurrent leaching of ore calculation results. 17 6 Stage to stage calculations for the uranium in the extractionscrub sections of the solvent extraction apparatus. 19 7 Solvent extractor. 22 8 Solvent stripper. 24 9 Uranium product evaporator and calciner. 26 10 Aqueous waste evaporator design. 29 11 Waste calciner and cyclone separator. 30 12 Recycle gas heater and waste steam generator. 32 13 Absorption tower. 34 Drawing 461-1128-U-1 Chemical flowsheet. 6 461-1128-U-2 Equipment size and arrangement 9 461-1128-U-3 Utility requirements. 10 iv

The University of Michigan Engineering Research Institute 1.0 INTRODUCTION Due to a keen interest on the part of Engineering Research Institute personnel in uranium ore processing, funds were made available through the Director's office for an evaluation of a process comprising the latest technology available. These funds were to be used in a study of the feasibility and economics of such a process, pointing up both the economic and technical advantages to be gained, and so stimulate sufficient industrial interest to instigate research projects in this area. This report summarizes the results of this study. 2.0 SUMMARY AND CONCLUSIONS A process has been presented for the extraction of uranium from domestic ores utilizing a solvent-extraction separation method and containing provisions for recovery of acids that would normally be wasted as reaction products of the undesired components, This process has further advantages in lower grinding costs and more compactness, and shows promise of producing a product of greater purity than conventional processes. The total investment for a mill utilizing this process has been estimated at $4,385,000. Processing costs per ton of ore are estimated at $5.25, which should enable a plant to pay for itself in less than two years at current prices for the U03 product. Pilot plant operation is believed necessary on this process prior to final plant design. The uncertainities believed to exist in the process and in the equipment design are listedo 1

t Th~e University of Michigan Engineering Research Institute 3.0 DISCUSSION OF PROCESS 351 PRESENT PRACTICES IN' THE FIELD For many years the uraniam-containing ores were processed mainly for their vanadium content with the -uranium having secondary valueo Consequently the early flowsheets worked out by the Uo So Bureau of Mines were designed for vanadium recovery with uranium recovery steps added to their process as its recovery became importanto One of the later flowsheets developed by the Bureau of Mines is shown in Fig. 1, which was taken from Refo lo This process is adapted to the processing of a high-line oreo As can be seen, the chemical requirements for such a process are quite high. Appreciable quantities of soda, nitric acid, caustic, and sulfuric acid are required. In addition there are a number of filtration steps in the process which can be very troublesome on many of the ores, A discussion of the operation and difficulties of such a flowsheet is given by Philippone in Refo 2. Newer processes have'been developed. whi.ch eliminate a few of the problems inherent in the Bureau of Mines flowsheet. These processes are based upon ion exchange resins that are selective in their absorption of a uranium sulfate complex. These are anion resins since the complex of sulfate and uranium ion has a negative charge. The processes are called resinin-puAp or resin-in-column depending upon whether the exchange resins are fixed in a column or are contained in wire-screen baskets and lowered into the slurry of acid and finely ground oreo If a resin-in-column process is used, the leach liquor is separated from the solids by filtration, while in the resin-in-pulp flowsheet, a slurry of solids in the leach liquor is circulated through the resin,. In this method the ore is ground finely enough so that there is no plugging cf the resin bed with. coarse ore particleso A brief schematic flowsheet of a resin. extraction process is shown in Figo 20 This flowsheet and other information on ion exchange processes were taken from Refs. 3, 4, and 50 The difficulty with these processes is that the resin bed slowly accumulates "poisons" which cannot be elu.ted from the bed so that eventual replacement of the resin is necessaryo Further, when. the ores being treated contain considerable lime, the acid costs from these processes are high. As an example, assume an ore of the following composition is being treated. U308 0.3 wt f Ca 0 12.0 Nondissolvab les Remainder _1 L 2

FIGURE-I BLOCK DIAGRAM OF SALTROAST, ACID-LEACH PROCESS NaCI ORE SOLUTION (D I S CARD)..... 3 ~NaCO - _ V SOLIDS 2 O' V PRECIPITATION Fe SO4 SOLUTION Na2C03- FeAI PPT'N. j —- Fe,A1 HYDROXIDES (DI SCARD) (DISCARD) H2SC4 H2S%-4MCO REMOVAL * V CONCENTRATES Na OH -- U PRECI PITAT IO N t U CONCENTRATES

FIGURE-2 BLOCK DIAGRAM OF ION EXCHANGE PROCESS FOR SULFATE SOLUTIONS HCI a NoCI NH3 ORE I mmmmM I PR ECIPITATI ON H2S04 I FILTER F 1 -1 DRY CONCENTRATE RESIDUE EFFLUENT 4

The Ulniversity of Michigan T Engineering Research Institute This ore has a reasonably high lime content'but certainly not the highest that would be encountered~ This ore would require approximately 421 lb of 100% H2S04 per ton for leaching, not counting oxidants that would have to be added for rapid dissolution. Commercial 66~ acid (9352 wt %) is assumed to be purchased at a bulk price of lol cents per Ibo Freight in tank car quantities would add approximately 0o2 cents per lb to the base cost for a 300-mile haul, so the final cost of acid for leaching a ton of the ore would, be $5~87~ Total yearly acid costs for a 1000-ton-ore-per-day plant operating 300 days per year would be $1,761,000o This, of course, is the major cost but additional costs would be incurred for oxidants, eluting acids, and caustic or ammonia. The basic difficulties in the present processes can then be summarized as follows: There are large chemical requirements either in leaching acids or in subsequent steps or bothO Filtration steps within the process are cumbersome and slow. Critical steps within the process such as pH adjustment can be troublesome where close tolerances are required. There are several pH adjustment steps in the Bureau of Mines flowsheet and also some in the ion exchange processo Product purity could be considerably improved in the existing processes 352 PR.POSED PROCESS BASIS In arriving at the process proposed here, attempts were made to eliminate or avoid the difficulties outlined aboveo Each of the steps outlined has been backed up with reference data found in the literature, which are known to be operable from past experience, or have been found to be true from laboratory experimentso While the process steps are firm, the optimum selection, of equipment suitable for final plant operation. remains to be done and should be proven with pilot plant information. The basic process, shown in Dwgo 461-1128-U-1, consists of grinding the ore to an average particle size of 10 mesh and leaching the ore fines with nitric acid. The nondissolvable part of the ore is washed with water to remove entrained acid and then discarded. The rich liquor from leaching is combined with the washes and routed to countercurrent solvent-extraction equipmento In this equipment the aqueous solution is contacted with an organic extractant that is immiscible in the water phase and highly selective in extracting uranium from the aqueous phase as a uranyl nitrateo An organic 5

OVERALL FLOWSHEET 0c ORE BOTTOMS OVERHEAD 2000 lb. SOLVENT 703 lb Ca(N03)2 306 lb H20 CoO 12%(2401b.) 10 VOL. / T.rP SCRUB 306 lb H20 AS CALCINE 3941b 02 0 3%( O601b.) ACID WASH 90VOL %KEROSENE H20 CA(NS2 4 H20 1C200'F 1268.5b U3080.3%( 6 1h) 56.5 WT. %o03 946.31bO 20 0.800. 1.75 pO12.850lft3AS.C. |ARETICLE DEN.289/ 569.4 lb HNO3 93 2 9.29gal. 3.84gl 699al 1 P.0551b/ft3.C. RT75L D2b9c1./. /435 lb H20 113.5 gl. g 1.'E 175 Ib./ft? | 1-.3351 t Co I-89' 1WA9 WAT A TAE TEpa.rt.332 g/c Pbulk B838 to/^3 ___ \ ___ _] _______POELEACH PRODUCTCHE I A 700F~F O CRUSHING GRINDING COUNTER-CURREN T 1200 lb H20 I AQUEOUS RAFFINATE4 C OO L GRND TO MEAN OFI I ACID LEACHING 703 lb Ca(N3)2(3M) SOLVENT EXTN 1200 lb H20 I I ATI ~F o MESH (AVE 3 LEACH STAGES 28.3b HNO 5 STAGESEXTN. 703 Ib CaNO3)2 ( __ EVAPORATION Dp(AVE).0625 in. 2 WASH STAGES 6.0 bU308(3.5 g /cc) 3 STAGES SCRUB 28.3 lb HN03 2 EFFECTS BULK DESITY Ip 1.3409/cc P 1733 gal 97 lb /ft3 173gal 173 gal IVOL 2Q.6 t.3 85g TB-P OVERHEAD o\ ------. 6O56 KEROSEN|6569 KEROSENE 94 lb Hp.oIIo, I TA I LI N GS 23 lb HN03 188 Ilb H20 1694 lb NON-DISS. D DISCARD I 1.02 AIR AT 299 lb H20 107 gal 366 ft3 S-C. 0.33 lb HN03 SOLV EXTN. PRODUCT 0 I 366 ft3 S.C. 85 % SOLIDS 10 VOL % TB. P 1.4 lb HN03 ABSORBER OVHD. STRIP 90 VOL % KEROSENE 8 I I PRODUCT | 2.71 b HN03 1 l H20 I I I P -P= 0.801 656g KERO- 54 lb NO 534 Ib HN03 0' 20 o SENE' HO T., P D 1 —-534 lb H20 64 | |31.5 /_ J - 41bH0ABSO'RBER-,9.2, aI --- - (56.5%HNs3) STRIP PRODUCT I IHN03 I I IP 1.3551 --- --, 6.4 gal H20 I ISTRIPPING I MAKE-UP ||, PRODUCT 5.991b U30e 3-5 STAGES I lb I AIR EVAPORALCNE 94.5 CUNT SOL TO STORAGE 9.AT S.C. I ei CALCINE pz 1.095 CURRENT SOLVENT LEACHING I,,______~24 g KEROSENE,97 g TB.P 6.1 lb UO3 [ SOLVENT n -9^nih/f3 IOVOL % TB.P Pbulk-250 lb/ft3 90 VOL % KEROSENE I IP =0.800 19.15 gal I, WASH WATER ~ ~' -----— _ — 1 1.9 gal. WASH WASTEI AGITATED TANK WASH SOLN. RESEARCH INSTITUTE DRAWN BY R. 1. K. SCAL NONE 3.8 gal 1I 1.9 9 a I I UNIVERSITY OF MICHIGAN CHECKED BY, 6Wl. ^' DATE 5/8/5 7 1 IBRECYCLE T O ^^ ^ ^r - 1 — ----- ARBOJCT r'CH E M I CAL FLOWSHEET SOLV.EXTN. — 461-1128 URANIUM ORE PROCESSING ---- _ 0 "|NA IFE |D-NB 41-I2uPLANT _____' CLASSIFICATION DG I......... owUNCLASSI FI.ED O- B- 461 - 1128-U - I

I The Ulniversity of Michigan * Engineering Research Institute 1 extractant that fulfills these requirements would be n-tributyl phosphate mixed with kerosene or another suitable diluento The uranium-containing organic solvent is scrubbed with a small stream of distilled water in two or more scrubbing stages to remove impurities that may have been extracted with the uranium and is then routed to a second contactor where the uranium is stripped from the organic phase with distilled water. The uranium stripped out in the aqueous phase is quite pure and is concentrated in an evaporator to a thick syrup of UO2(NO3)2.3H20. This tri-hydrate is calcined at approximately 400~F which decomposes the uranyl nitrate to UO3, water, NO2, and 02. The U03 is packaged and shipped and the gases are reabsorbed in water to form nitric acid which can be reused in the processo The organic solvent after being stripped of its uranium is scruibbed with a sodium carbonate solution and then washed with water. This treatment removes any decomposition products of tributyl phosphate that have formed, as well as any traces of metallic ions that were not removed in the stripping operationo After this treatment the organic solvent can be recycled back into the extraction apparatus. The aqueous stream leaving the solvent-extraction apparatus contains all the components in the ore that dissolve in nitric acido Generally these components would be calcium, some aluminum, iron, vanadium, and traces of other elements depending upon the nature of the ore. Nitric acid was cons~umed in the leaching step forming nitrate salts of these elements. A recovery of this acid is mandatory for an economically operating process. Recovery is accomplished by concentrating this solution in evaporators using steam and calcining the salts at 700-12COOF to decompose the nitrates. JO02, 02,and steam are evolved, leaving the metal ion as the oxide. The gases are cooled and absorbed in water in a multi-stage absorption towero The product of the absorption is nitric acid which is recycled back to the leaching step. Metal oxides are discarded, although some use could be found for them in cement or plaster manufacture. A detailed discussion of the process steps follows. 353 STEPS IN THE PROPOSED PROCESS 3.1 Crushing and Grinding.-Crushing and grinding requirements are set both by the high uranium recoveries required in the leaching step and the least possible amount of fines carried over into the solvent-extraction stepo Leaching requirements are best met by fine grinding, while the solvent-extraction step works best if very few fines are carried over from leaching to solvent extractiono I - 7

The University of Michigan T Engineering Research Institute 6 It appears that for the secondary ores such as carnotite that one stage of leaching for a -4 mesh in HC1 would give uranium recoveries' of 96% after leaching for 4 hr at 8o~C, For -20 mesh ore at 80~C, 95% recoveries of uranium were obtained in 1/2 hr with no significant improvement in recoveries being noted for additional leaching times of up to 4 hr. When the ores being treated are pitchblend, the leaching time is apparently quite different 7 times of up to 16 hr being required depending upon acid, oxidant strength, and temperature. However, some essential data regarding mesh size and method of leaching are missing from this reference. Since most ores to be processed will be western ores, an average 10-mesh particles size was selected, recognizing that the information upon which this selection is based is meager and that additional data on leaching rates for all the varieties of ore being processed will be required. The equipment selected in this study is shown in Dwgs. 461-1128U-2 and U-5o Equipment numbers given here refer to these flowsheetso The equipment is composed of a coarse ore screen (1), equipped with electrically driven vibrator, through which the ore is daumped from carryall loaders. A combination scalper, feeder, grizzly unit (3) feeds the ore from the ore bin (2) into the grinding system. This scalper feeds the coarse ore into the jaw crusher (5) while pieces smaller than 3 ino fall through the hopper (4) and onto the conveyor belt (6), The underflow from the jaw crusher falls directly onto the conveyor belto It is assumed that this section of the ore-crushing equipment composed of items (1) through (5) is located in close proximity to the ore storage areas. The belt conveyor with a length of approximately 125 ft conveys the crushed ore to the subsequent grinding equipment The discharge from the belt conveyor falls into bin (7) which guides the ore into the core crusher (8). This crusher reduces the ore size from a maximum of 5 in. to approximately 1/2 in. with the fines falling into the electrically vibrated bin (9). This bin feeds the ore into the rod mill (10), which reduces the particle size from 1/2 in. to a nominal -10 mesh. A bucket elevator (11) conveys the ground ore into the fine ore storage bin (12), where a hold-up of 4 hr is provided to allow some operating flexibility between the crushing grinding equipment and the subsequent extraction steps. All the crushing grinding equipment has been sized for greater than a 1000-ton-per-day capacity, with these sizes being the basis for subsequent pricing. The selection of a jaw crusher rather than a gyratory was made on the basis of a lower first costo This selection is subject to review if operating costs indicate that a gyratory is more desirable, A rod mill was selected for final grinding rather than a ball mill since the spread in mesh size from a ball mill is narrowero This is illustrated in Fig. 35 The data in this figure were furnished by the Allis-Chalmers 8

g i 4 g l

) 1 T tiORb \I' I S t ITlh \rL PUMPS 5555. 0/E, sr _ L " -- -~-l- 4 p.",,,l~.OQ ~\~40 TO Nn ST OU^m |r -- 1U Pm W- / 4 M *STW lTooEM STEAM 11 U 00 _ I-OE XEE-e -- *WD CONW lA *0,000 l4n50 2-ORSE S0IERIE 31 — -OUCTB OREAO COOST PPn SOML _A _ 4-SLY-PSOEUCTCAU:INER OV0NERIEWO ESUR j I f -CWCIS IESPlfOW n N 3STEAM ~1123-05_ S CO_ OESR [ E I w~ | r OT ClS II S-NCETEL SEBIR 41-AGEOUS WAS42 TE,OVIEKHU OONoeSSUTE PUMP-a *LP f O-pi WEATR I-S^IE OF DER FE4 55AOUCESTE PA ff0 PE-ATRCIORERO.. C EOVERESD CODE IS-SASS PUMP 45-2 RECIROSL^"G WAIT! EWDOO NFUS T-: OWET OMVE TAK W-MCLOSE 9 sOLUT -UPS100 ~-HNOS~FEED PUMP 4~3-CYOUCLSTE CAPUOERWOR O H NE 1~ 000.| 9-CM CRUSHER UNDMFLOW 39-STEYM VETOR COND0SM 1 0- W 12-SOE S! SORMA TUENE 4-REI fCUAS2/ Tlf T20 wFE M -TEAM EE lO- _) 1I-SULlE! FE DER SI-IACI ER PROUKV.UCT R0 - AlOE S AORSE S TA' -IsW o PUMP 45-ZM AQUBEUl WASTE EAP WQ H4-C03M.MOE TAW NER 54 PIMP 56 1II-HN0FEED PUMP 41.CYCL IE W.0 *W STU/b MAKE-UP 19- SOLVENT EXTRACTOR R.. 21-_SOLENT FEED PUMP 51- CALCImNR PIOUC T COOLER *C0NDNSER NNO so 2 2-AQOS p PI'CUCINERS PROUCT COOLDER SCONENLE CONOEMIE PUW 23- SOLVEIN STRIPPER, -ASSORBER COLUN 24-01DtLLED M*TER STORAMEIaNK 4-ASONE tATER FEED PUMP 25-DDIITILLED mElR REED W - T PUMP R-SOVEIT S!RSORUMR Sn-OvKonT SSKM-BOILER FEEDOmTER TREIRS O2R -.WS!_A oNIEES 5coN 5-COOUIN TOWER D-PROSUCr APOATOR OVERHE COE O-COUN WAFER PO MP 30-PROC! EVAPORATORS PSOEUCT MALCEIMER 61- SrTEAM CONSESTE RETURN PUMP& WEILL ~COSMATO 0LLE~tIOTAMR ~ - WAIHEO OLVIIT PUMP G- PRODUT PUMP ABOMVIATIMS I$.&TSTANLE M STUL 304.34T.34 KS.-kH(LD STEEL

0 NO. 558 SEMI-LOGARITHMIC-3 CYCLE.0 IN U t. c at -4 ) (O - _FlI4f S X 70 DIVISIONS THE FREDERICK POST CO. CHICAGO. ILL. AMOUNT RETAINED BY SIEVE, % ON w 4 UI 0 4 CD 0 - N.....................................~ 0 v 0 4 + a s I tt ffl I-' H m ut I - z) I ON m -__ - m o, - - H a - PRNTOD IN U.S.A. Iu n(n

The University of Michigan T Engineering Research Institute Manufacturing Companyo It is evident that the ball mill grinding to a nominal -8 mesh particle size would result in 1.6% passing through a 200mesh screen. There would also be about 2.o3 retained on a number 8 screeno The rod mill would give about 0.6% -200 mesh and a smaller proportion of oversize particles it is important that the quantity of solids of -200 mesh be held to as low a value as possibleo The reasons for this will be discussed in the section on leachingo Since it is unimportant in subsequent steps just what the particlesize distribution is, as long as the -200 mesh fraction is minimized and the +8 mesh fraction is not excessive, there is no need for a recirculation system and consequently no closed circuit grinding is requiredo Obviously this simplifies the grinding section of the plant and reduces the over-all plant costs 35352 Leachingo —The ground ore is transported by a bucket elevator (11) into a fine ore storage bin (12). This bin is sized to provide a 4-hr ore hold-up at full plant capacity so that some degree of flexibility will exist in matching the crushing and grinding rates to the plant throughout. The fine ore is fed from this bin by a vibratory feeder into a continuous leaching extractor (14) where the ore flows downward and the leaching liquor flows upward. The basic design of the ore extractor is illustrated in Fig. 4. The ore extractor consists of a tall cylindrical vessel with a conical bottom. Down the center of this vessel is a shaft to which is attached the scraper bladesO As the shaft rotates, the solids are scraped toward the outside of the tray and fall over the edge and into the conical chamber below. Solids pile up at the downcomers from each tray to form a seal to prevent liquid by-passingo The solids are routed downward from tray to tray in this manner until they are discharged out of the bottom of the vessel as a slurryo Leaching acid enters below the fourth conical section from the top and combines with part of the wash water which flows upward after being introduced at the bottom conical sectiono The combined solution flows upward with alternate agitation. and settling at each conical section until the solution, runs out of the overflow at the top of the vesselo Sufficient wash water is introduced at the bottom to furnish adequate slurry liquid for discharge of the spent ore sandso Thus the bottom two conical sections are washing stages and the upper three sections serve as leaching stages. The size selection of this unit will be discussed in the following paragraphs. The problems in ore leaching in this type of apparatus are in matching the dissolution rate the nre t throughout and at the same time keeping the velocity of the liquid through the unit low enough so that fines carried over from the ore are not excessiveo 12

Q'ON'9MC LIQUID IN PLAN OF SETTUNG TRAY SOLIDS SEAL BETWEEN STAGES EXTRACTED SOLIDS IN WATER SLURRY BUILT BY THE GAS MACHINERY CO. OF CLEVELAND,OHIO DSIGND Y M.E APPROVED BY ENGINEERING RESEARCH INSTITUTE DRWN R.I.. CA I - UNIVERSITY OF MICHIGAN CHECKED.B M. E.W. ATE 5/3 / 57 ANN ARBOR MICHIGAN TITLE PRo 6 46-1128 ORE EXTRACTOR (LEACHING) a I....... U N C A. S SIF I I I A """,- "" UNCLASSIFIED FIGURE-4 I SSUE A I i I

I - The University of Michigan * Engineering Research Institute Saine and Brown give data on the leaching rate of a number of western United States ores. Their data on ore from the Uravan district in Colorado are summarized in Table Io With the exception of experiment B-16 TABLE I Effect of Mesh Size on Direct Acid Leaching of Raw Carnotite Ore (80~C Leaches) 1st Leach 2nd Leach Tails Expo Ore % % U V % U' V No. Mesh of of of of of % of % of 0 of ______ _ lTotal Total Total Tal Wt Total Wt Total B-15 -4 96 42 0o5 7 Oo004 4.0 ) 485 51 B-16 -8 93 36 L0 6 0.015 6~o0 o 525 58 L-365 -20 97 42 05 11 00o 2 5 0o445 47 B-17 -60 97 32 0 5 13 Oo 004 2 5 0o520 55 3B-18 -80 97 36 1o5 12 o C002 1o5 0o500 52 B-19 -100 98 37 o4 10 o o002 o5 o 495 53 Average analysis of head samples: L3(-,B: C0l5% V205 l o 7% Ore Sample: 100 g of raw Uravan ore Acid Leac'h: 200 ml of 5% HC1 agitated 4 hr at 80~C in these data, it apparently makes some difference whether the ore is -4 or -100 mesh, The second leach contributes to over-all uranium recovery but not to any great extento If sample B-16, by comparison with the rest of the runs, is indicative of the experimental errors in the data, then the conclusion might be reached that mesh size is unimportant in leaching this particular oreo However, if B-16 were discarded as being nonrepresentative, then grinding to -100 mesh would represent a 2.5% gain in the uranium recovered from the tails. it should be noted that these data are based on batch leaching of the ore and not on a continuous countercurrent system. While batch data are indicative of gross recoveries that could be achieved, there should be -i

The University of Michigan Engineering Research Institute some fundamental advantages in a countercurrent system which would result in greater uranium recoverieso There remains the question of how much leaching time is required per stage. Saine and Brown also investigated this and their results are given in Table IIo In this table, runs for equal acid concentrations are compared for leaching times varying from 1/2 to 4 hro It can be seen that the quantity of uranium leached from the ore does not vary significantly for times within this period. This is evident particularly in runs S-375, L-364, and L-365. The data also show that the acid used for leaching makes little TABLE II Time Dependence of Direct Acid Leaching of Raw Carnotite Ore (80~C Leach) _ L-550 L- 351 L -555 L-363 s-572 S -375 L-364 L-365 L -567 1% 1% 5% 5% 20%o 1% 2% 5% 5% 5%0 20Go HaSO4 H2S04 H2S04 H2SO4 H2sC HC1 HC1 HC1 HC 1 HC 1.H~fC 1 1 4 1 4 4 4 2 1/2 1 4 4 90 90 95 96 96 90 93 95 95 97 96 7 3 12 6 30 0.7 39 0.4 51 0.3 -_ 5 27 2 22 0.7 22 1 5 42 0.4 -- 03 8.009 12 006 ~. oo6 10 J005 9.005. oo6 9. oo6 --. 006 8 006 7. 005 9.005 11.005.oo005 7 4 4 4 4 5 5 4 4 3 5 0.790. 560 0.550 o0.510 0o400 O. 56,, 565 0.665 o0445 85 76 65 51 40 65 71 69 47 Ore Sample 100 gm of Uravan ore ground to -20 me sh. Leach Description: 300 ml of leach solution per 100-gm ore at 800~C 15

The University of Michigan ~ Engineering Research Institute difference in the recoveries attained. For this reason no differences are expected in reaction times when HN03 is used rather than HC1, or H2S04. On the basis of these data three countercurrent leaching stages were selected having a solids hold-up of 10-15 min per stage. It may be necessary to increase this time when more data are available on a wider variety of ores. It is interesting to note that Saine and Brown found carnotite ore very easy to grind down to -80 mesh, whereas extensive ball-mill grinding was necessary to reduce the size to -100 mesho This supports the authorcs opinion that the process proposed here would result in significantly lower grinding costs than conventional processes. The leaching conditions were selected on the basis of the available data. It was necessary, of course, to alter the acid concentration in the leach liquor to higher concentrations to permit dissolution of all the lime in the particular type ore selected. This change should result in greater rather than lower recoveries. The conditions selected were used in a stageto-stage leaching calculation to determine probable composition of the liquid in each stage and to determine the number of wash stages required to remove essentially all the acid from the sands. The results of this calculation are portrayed in Fig. 5, In this calculation only the lime was considered as a dissolvable component since only its performance was completely predictable at this time, Most of the lime is dissolved in stage 3, where fresh acid contacts the solids coming from stage 2. It is assumed that this will be true of the uranium in the ore alsoo It should be noted that there is still unreacted ITi03 present in the leach product, so leaching will also occur in stages 1 and 2. This is shown also in the decrease in lime concentration in the solids stream from stages 1 and 2. This calculation has also been valuable in predicting the number of washing stages requiredo It is evident that two such stages, 4 and 5, will reduce HNI03 losses to a negligible quantityo The internal area of the countercurrent leaching equipment is adjusted so that the upward flow of the leach liquor does not carry appreciable quantity of fines with it. This liquid velocity was set at a nominal value of 5 x 10-3 ft/sec. Even with this low velocity, solid particles below -200 mesh would be carried out of the leaching equipment and into *the solventextraction apparatus. This is the chief reason for selecting grinding equipment that will give a minimum of fines in the -200 mesh region~ Referring to Fig. 3, it is evident that up to 006 wt % of the ore could be expected to be carried over when a rod mill is used as the final grinder. This may be a pessimistic figure since some filtering action would occur as the liquor 16

FIGURE-5 COUNTERCURRENT LEACHING OF ORE CALCULATION RESULTS LEAH PROUCT ACID FEED 1200 lb i0 HN03 569.4 Ib/ton ore 28.3 lb HN03 HO02 435 Ib/ton ore -703 lb Co(N03)2 38~HN03 565 wt.% 6 lb U538 p 1355 g/cm3 3M Ca(N03) WASH WATER 946.3 Ib/ton ore TAILINGS 1694 lb non-diss. 0.33 lb HN03 299 Ib HfO INTER STAGE CONCENTRATIONS 240 lb CoO 6 lb U3% (OS 40 lb H20 1694 Ib non-diss. p = 2.8 g/cm3. - 1694 235 106 2.5 lb non-di.s lb CoO lb H20 lb HN03 (S - 1694 lb non-diss. 93.5 Ib H20 68 lb HN03 BULKp S 97 Ib/ft3 10 MESH VL 1264 lb HO0 42.6 lb HN03 GL 740.8 lb H20 6.5 lb HN03 1694 216 104 35 lb non-d"s. Ib CoO lb H20 lb HN03 1694 lb 935 lb 1.1 lb non-diss H20 HN03 740.8 lb H20 0.8 lb HN03 1256 lb H20 86.91 b HN03

The University of Michigan * Engineering Research Institute I'L passes through the ore fineso Certainly, however, any plant design should be based upon some anticipated solids carryover. Several other types of equipment could be considered for the leaching application discussed above. It is the customary practice in the present plants to use large agitated vessels for leachingo These vessels are charged with the ground ore and the appropriate solution pumped in, and the ore and solution agitated for as long as 8 hro The ore is allowed to settle and the supernate routed t'o ion-exchange extraction apparatuso The ore sands are then washed, resettled and decanted, and the sands pumped outo Present plants do not generally use countercurrent leaching of more than two stages with just one stage being the rule rather than the exception. The disadvantages of this system are the very large tanks required and the comparatively low throughput compared to the continuous type of equipment for equivalent capital investment. Another type of equipment that could be used is a percolation tower where ore is fed into the top of a tall column and the leach liquor percolated up through the fines. The conventional Dorr-Oliver leaching equipment could also be used. This type of equipment is not normally enclosed, so covers would have to be placed over it to prevent the escape of N02 fumes. Comparing the advantages and disadvantages of all the possible types appears to indicate that the type portrayed here has a higher throughput and is comparable in first cost to other typeso Maintenance on the unit selected should not be significantly more than on the other types with the possible exception of the simple tank-agitator variety. 3o3.3Solvent Extraction and Strippingo-The rich leach liquor overflows from the extractor (14) to the solvent extractor (19). Thus the solvent-extraction aqueous feed rate is actually controlled by the controllers on the acid and wash streams to the extractor (ioeo, leaching apparatus). The solvent-extraction equipment contains five extraction stages and three scrub stageso The function of the extraction section is to extract uranium from the aqueous phase into the organic phaseo Extraction efficiency of uranium is measured by a parameter called a distribution coefficiento This coefficient is the ratio of uranium concentration in the organic phase to the uranium concentration in the aqueous phase, using consistent unitso The value of this coefficient varies with the nature of the organic extractant, the nitrate concentration in the aqueous phase, and the uranium concentration in both phaseso The effects of other components in the system are relatively minor compared to those mentioned aboveo Distribution ratios are given for a wide variety of conditions by Moore8 and Granquist and Merrill9 The data presented in Refs. 8 and 9 permit the calculation of extraction and stripping stage requirements in this processo This calculation is shown in Fig. 6 and will be discussed in later paragraphs. I 18

FIGURE-6 STAGE TO STAGE CALCULATIONS FOR THE URANIUM IN THE EXTRACTION-SCRUB SECTIONS OF THE SOLVENT EXTRACTION APPARATUS t' -- =: I-: 71-1- +-, I - -:t -'? i ti77:7 w.i4 — " z -- 0., LI) FH -... 7 - II I I i E34fl412.I4- I-, I -; +-;z: —- I: - - — I —-— I — - F --— 4 —---- r- I — -- 47 j4 it 1 — RI-vr tI:-:-:l yz I y I I -I i "' fwTrwI| I L -1<-4 II111.1.. I:~~~~~~.J 1 —:A 1. I APPR(K 4 SCR~~~~~~~~~~~~i --;; I- -.; 1.: I.t: 9 V3-.1; 3~7-:: -- 4 1i-/:c 4- { — I --- -- 7 —-- I.: —'- i...: 1-44! 4-~-4A t I - -- - I' - -II c ot 5 c N Ol in.i = is' r3 < ---—' —-— I-~'- I i-I. I I: --- id'I I Ii - M f;-i- i-~l- jI- i ~-1 — I i -tFI —.00 1 n I n I IA -I I _ _ __VW -F-, ~-,;_ _...._ _ _ _ _ I; ~i~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~,::.l~~i _.. 1__1_: 1-' _I_$-_: —T- F__E~~I', _ _ _ _ _ _ _ _ _ _ _ _ _ _ - i 2~~~~~~~~~~~~~~~~~~~~~~~~~~~0: L l:.. 1i I__~~~~~~ ~ _ _ _ _ _ _ _ t__ 2 i'- -j- _____- ____ =I j_ — H.I.l K K I r T~r~tK-T i 1 _ _ _ _ _ _ _ j A A I~~~~~- 4 73 7 6 8 1 4 5 6 7 9 12 8 4.VV I.VI g. U X, ij- AQUEOUS PHASE Ku,

The Ulniversity of Michigan T Engineering Research Institute Unfortunately, while the extraction of uranium by this system is very efficient, it is not highly selectiveo Some metal ions present in the rich leach liquor will also extract. Metal ions of vanadium, iron, titanium, and also boron extract to some degreeo These contaminant ions would result in an impure uranium product if they were not removed from the processo Removal is accomplished by the scrub section in the solvent-extraction apparatuso Since the distribution coefficient of these ions is not as high as for the uranyl ion, removal can be accomplished in the multi-stage counterflow apparatus by scrubbing the uranium-containing organic phase with a small stream of water. After scrubbing, the uranium in the extractant (organic phase) is in a pure state and becomes a very suitable product. A calculation of stage requirements in this apparatus is shown in Fig. 6. The operating lines in this figure are arrived at by taking material balances around the bottom (extraction section) and top (scrub section) of the solvent-extraction apparatus. A uranium distribution ratio organic to aqueous of 30 was assumed in the extraction section.. This value appears justified from the data in Refs. 8 and 9. In the scrub section, where no Ca(N3s)2 is present, equilibrium is not constant but varies with uranium concentration. These data were taken directly from the references aboveo The stages are stepped off on the diagram as dotted lines. The u.=mber of stages required in the scrub section is indeterminate since a "pinch" exists in this sectiono Three scrub stages is well into the pinch and this is believed adequateo In the extraction section, the stages can be easily determined since no pinches existo The stages required in this section are counted down from the point where the scrub equilibrium and operating lines intersect or I = 54.5g U/lo Six stages are adequate to reduce losses to.it is interesting to note the amount of uranium reflux that occurs in. the scrub section and the magnitude of the turanium concentration change between the scrub and extraction changesO The maximum uranium concentration in the solvent-extraction apparatus occurs at point B, which is just above the feed pointo At this point the uranium concentration is 116 g/1 in the aqueous phase and 54,5 g/l in the organic phaseo At the feed point the uranium concentration is about 6 g/l, which is where the extraction section begins. Thus the reflux in the scrub section varies from 116 to 6 g/l. It is this refluxing action that is important in removing the impurities that are extracted from the uranium. Product and concentrations are represented by point A on the diagram where Yp = 31.5 g/l and is the product stream concentration leaving the apparatus. At the same point but in the aqueous phase, the uranium concentration is 79 g/l. The scrub and extraction section operating lines should intersect at the feed stream concentration if the extraction diagram is drawn correctly. This intersection is shown as point C and the aqueous phase concentration at this point is 3.5 g/l which checks the flowsheet I 20

L The University of Michigan * Engineering Research Institute feed concentration. It is also necessary for the first stage in the extraction section to include this concentration. Clearly stage 1 does this, as shown by the dotted line. Uranium losses from the extraction section are represented by the magnitude of XW. In Figo 6, the flowsheet valve of.00355 g/l is showno This concentration would occur somewhere between stages 5 and 6 in the apparatus. Five stages are shown on the flowsheet and in the drawing on the extractorO It would be advantageous in a final design to increase this to 6 stages based upon Fig. 6. The stripping apparatus stage requirements are much simpler to compute since this apparatus is not complicated by a compound function such as scrubbing and stripping. No computation of stages is included here because of its simplicity. Calculations by the same methods as in Fig. 6 indicate that 3 stripping stages are adequate. It is advisable to include extra stages so 3 to 5 stages are shown on the flowsheet. The selection of equipment for extraction and stripping has been left open up until now. Among the possible choices are packed columns, spray columns, or pulsed columns. Other types that could be considered are the Schiebel contactors, and the Fenske contactoro Various types of mixer settlers are also available as well as the high-speed centrifugal type of contactorso From the standpoint of reliability, simplicity, and ruggedness, the choice for this application appeared to be between columns of various types and mixer settlersO Most columns operate at maximum rates of 1000 gal/hr ft2 sum of both phases which for these flow rates would set the column diameter at 352 ft. Since there is no way to predict theoretically the height of a stage in a column, experimental data must be available before a height can be seto The largest column known by the author to have been operated on this system was 2.0 ft in diameter. Extrapolation from 2.0 ft to 3~2 ft was considered unwarranted without pilot plant information. Consequently, a simple mixer-settler system was chosen which requires a minimum of extrapolations and pilot plant information. Tests were made on the equilibrating times, settling times, and maximum agitation for this system, and from these data an average time of agitation of 3350 sec and a 5-min settling time were selected as a design basis. A tentative design for the solvent extractor is shown in Fig. 7. The body of this unit contains conical separators which serve as the settling sections between stages. The bottom of these cones is fitted with valves so that solids being accumulated from leaching (ore extraction) carryover can be flushed from the unit. On the outside of this unit are mounted the interstage mixers, which are small cylindrical vessels with a flange-mounted propeller-type agitator mounted on the top. The aqueous feed (leach product) enters the mixer on the left and is mixed by the agitator with organic. solvent being pumped up from the settling section below where the two phases separate. The aqueous phase flows over the weir down into the next mixer, where it is mixed with organic solvent from the third settling section where the sequence is repeated. _ 1 21

IRIDtERICK POST CO. 179H 500 9-54 SCRUB MIXI LEACH PROD. PHASE PUMP 4~.J v AQUEOUS "HA SOLVENT- L AQUEOUS WASTE DIM ON UN OTRW I MUT H7LD TO A TOERNC - CTON, DECIMA 005" NGULAR AUL DIMENSIONS UNLESS OTHERWISE SPECIFIED MUST BE HELD TO A TOLERANCE - FRACTIONAL ~ ^4." DECIMAL ~.005." ANGULAR ~ U* O > _f~ ~ wn ~~~ IDESIGNED BY M.EW. APPROVED BY ENGINEERING RESEARCH INSTITUTE DRAWN BY R.E.K. SCALE /4' DRAWN BY R. I.K. SCAL 4 UNIVERSITY OF MICHIGAN CHECKED BY ME.W DATE /27 ANN ARBOR MICHIGAN TITE.. TITL" PROJET 461-1128 SOLVENT EXTRACTOR IUE D CLASIICATION UNCLASSIFIED FIGURE-7 ISSUE DATE 22

The University of Michigan Engineering Research Institute Thus, the aqueous flow proceeds downward and the organic or extractant flow is upwards. Two stages are shown as an illustration for the scrub section, and since the aqueous flow is much smaller here, the over-all size of the equipment is reduced. It is to be noted also that the organic phase is pumped between stages while the aqueous flow is downward by gravity. The solvent extractor is item 19 on the flowsheets D-461-1128-U-2 and U-3. The organic stream leaving the solvent extractor contains uranium at approximately 31 g/l. It is necessary to transfer this uranium to an aqueous phase before subsequent treatment can take place. This transfer is accomplished in a solvent stripper (23). This piece of equipment is very similar in construction and operation to the solvent extractor except that it is smaller in size and contains fewer stages, as shown in Fig. 8. Since the N03 ion concentration is low in this apparatus, the uranium is preferentially extracted into the aqueous phase. The actual equilibrium conditions for this situation are the same as for the scrub section of the solvent extractor except operation is in the more dilute regions. Referring to Fig. 6 for the scrub-section equilibrium line, it is evident that, for uranium concentrations below 30 g/l (aqueous phase), extraction would be preferentially from the organic to the aqueous phase. In the solvent stripper the 19.2 gal/tonr"of organic phase is contacted in a 3-to-5 stage countercurrent operation with 6.4 gal/ton of distilled water. The uranium-rich aqueous phase flow from the solvent stripper to the product evaporator (26) and the uranium-free organic phase flows to the solvent washing apparatus (27) and (28). Stage requirements in the stripper were calculated in the same manner as for the solvent extractor, but here the calculations are much simpler since the stripper does not serve a dual purpose. Agitation and settling times within the stages are tentatively the same as for the extractor. 3.3.4 Solvent Treating. —Since tributyl phosphate is an ester of n-butyl alcohol and orthophosphoric acid, it reacts in a manner typical of these compounds, It can decompose by hydrolysis giving mono- and di-butyl phosphates as well as free phosphate ion and alcohol. These decomposition products are undesirable since they interfere with subsequent extraction and stripping of uranium, and so must be reduced down to acceptable levels. In addition, the kerosene diluent has a tendency to nitrate at the double bonds. These nitrate compounds appear to have detrimental effects during the extraction-stripping operations, and so must also be removed. Removal of these undesirable components is accomplished by mixing the organic phase with a 5-10% solution of sodium hydroxide, settling and decanting the phases, and then repeating the operation using a water wash. These washes are done in the solvent scrubber (27), using the caustic, and the solvent washer (28), using the water. The residence time in the scrubber _ 23

z o MIXER-SETTLER - PHASE PUMP - S TRIP PRODUCT U SOLUTION ALL DIMKNSIONS UNLESS OTHKRWIS SPKCIFIKD MUST *E HKLD TO A TOLERANCI - FRACTIONAL:k %4," DECIMAL.005," ANGULAR -~ * ENGINEERING RESEARCH INSTITUTE DEIWND Y M.E.W. APPRO /4= I'y UNIVERSITY OF MICHIGAN CHNCKD BY M. EW. DATE 5/2 9/)57 ANN ARBOR MICHIGAN TITL PROJECT | "461-1128 SOLVENT STRIPPER ISU DATu C F^" UNCLASSIFIED FIGUFE-8 24

t The University of Michigan * Engineering Research Institute I 1 and washer were tentatively set at 5-min agitation and 30-min settling. Flow ratios of 10 organic to 1 wash solution were set by past experience. Caustic and water washes are combined and discarded with the organic extractant routed to a solvent make-up storage tank (34). It is pertinent at this time to discuss solvent losses in the solvent extraction section of this process. Cost of replacing lost kerosene will not be significant, but since tributyl phosphate costs 50 cents/lb ($4,10/gal less freight) in tank car lots at the factory, significant losses of this chemical could be a serious economic factoro Solvent losses can occur by three different mechanisms: 1) solubility of solvent in aqueous solutions contacting it, 2) hydrolysis or decomposition, and 3) entrainment losses in one of the waste aqueous streams. It has been found by experiment that solubility losses are substantially greater than hydrolysis losses. Fromanorganic phase containing 10 vol % TBP in kerosene and equilibrated with an aqueous phase, the TBP will dissolve to the extent of 0o14 g/l in the aqueous phase. Taking the sum of all aqueous streams contacting the organic phase (183 gal/ton ore), the TBP losses are seen to be about 97 g TBP/ton ore, or about 1/4 lbo This is small enough to be almost negligible; however, solubility losses can be greatly exceeded by losses due to maloperation or poor design of the equipment. At the present time, these losses are not known, but it is definite that TBP losses can approach about 1/4 lb/ton ore under proper conditions. Kerosene losses will amount to only 1-1/2 lb/ton ore as a minimum. 3.3.5 Product Evaporation and Calcining.-The product stream from the solvent stripper contains approximately 94.5 g U/l which must be concentrated and de-nitrated into an oxide to become an acceptable product. This is done in the product evaporator (26), and in the product calciner (32). Sketches of these items are shown in Fig. 9. The product evaporator receives 2430 lb/hr of solution from the product stripper, and concentrates this solution until its composition is roughly that of U02(N03)2 *6H20 with a boiling point at atmospheric pressure of 244~Fo Using 60-psig steam, the At across the heating surface will be 64~F. Heat load in this unit is 2,220,000 BTU/hr and the area required 140 ft2. A thermosyphon type evaporator was selected since circulation is fairly rapid, and excellent control and uniformity is achieved without external pumpso It is characteristic of this solution that little or no scaling occurs on the heat transfer surfaces. The overhead vapor from the product evaporator is condensed in the product-evaporator overhead condenser (29). This unit is cf standard tube shell, single-pass in tube construction. L I - 25

0 z P EVAPORATOR I W 1 51;C ~' 12" DIA. CALCINER ALL DIMENSIONS UNLESS OTHERWISE SPECIFIED MUST BE HELD TO A TOLERANCE - FRACTIONAL ~ I4," DECIMAL ~.005," ANGULAR f %o ENGINEERINGRESEARCH INSTITUTE DESIGNED BY M.E.W. APPROVED BY ENGINEERING RESEARCH NSTITUTE DRAWN BY R.. K SCALE I/"' UNIVERSITY OF MICHIGAN CHECKED BY M. E W DATE 5/24/57 ANN ARBOR MICHIGAN TITL - - PROJECT 4? URANIUM PRODUCT _4_61-1128 EVAPORATOR & CALCINER ISSUE | DATE CLASIFICAT' UNCLASS IFI ED IGURE 9 26

The University of Michigan * Engineering Research Institute 1 Entrainment of concentrated liquor in the product-evaporator overhead can be held to 105 Ib liquor/lb overhead by the design of the de-entraining section. Losses of uranium due to entrainment are thus negligible. Concentrated U02(N03)2.6H20 solution flows from the product evaporator to the product calciner where the uranyl nitrate is converted to the desired U03. This conversion occurs at about 400~F and requires 1650 BTU/ lb feed heat input. About 695,000 BTU/hr is required to dewater the feed with the remainder used to decompose the nitrate. Since much of the heat transfer is to a granular solid, and caking or lump formation is undesirable, a steam-heated jacketed screw-type calciner was tentatively selected. Other calciners, such as the vertical tray type common in ore roasting, a fluidized bed unit, a rotary type kiln, and possibly other types could be considered for this service. It would be desirable to reabsorb the N02 and 02 formed from the uranyl nitrate decomposition into water for acid recovery, and for this reason heating with combustion gases appears impractical due to the large dilution and subsequent absorption difficulties of the N02. Dowtherm, or molten salt, and direct fired heating outside a jacket would be worthy of consideration. In this unit, steam at 250 psia was selected for heating since this is the generation pressure from the boilers. Both the screw and the jacket are steam-heated. The screw will make approximately 6.2 revolutions/hr to deliver the 254 lb/hr of U03. The U03 product falls directly from the calciner into a loading hopper equipped with an intermediate gas lock and is charged directly from this hopper into a scale-mounted drum for final weighing, sampling, and packagingo The evaporated water, N02, and 02 evolved from the calcining is condensed in the product calciner overhead condenser (33)- Item 33 is deliberately overdesigned from the heat-transfer standpoint so that liquid and gas hold-up times will be long. This gives the N02 and 02 long enough to be reabsorbed into the condensed water giving HNO3 About 80% of the NO2 should be absorbed in the condenser. The remaining gas is discharged out of the stack since it is small in quantity, and it is not considered economical to compress this gas so that it can be totally absorbed in the absorption tower later in the process.'The condensate from the product-evaporator overhead condenser (29) and the product calciner overhead condenser (33) is collected in a tank (30) and is pumped back into the system or can be discarded. 3.3.6 Waste Evaporation and Calcining —The aqueous waste stream from the solvent extractor (19) contains all the components of the ore that are soluble in nitric acid. These would be largely the lime and to a lesser degree the iron, vanadium, and aluminumo If this stream were to be discarded, large losses of nitric acid would result, particularly when the ore being treated has a high lime content. These large losses would occur 27 J

The University of Michigan T Engineering Research Institute whether the leaching acid is sulfuric, hydrochloric, or nitrico The major advantage of this process lies in its ability to recover the acid in the waste stream, and is the chief justification for using nitric acid for leaching because this acid can be easily recovered by the proposed methodso The aqueous waste stream leaves the bottom of the solvent extractor (19) and is pumped by pump (22) to the second waste evaporator overhead to aqueous waste heat exchanger (37). Here part of the second evaporator vapor is condensed, and exchanges its heat with the waste stream heating it from 80~F to 144~F. About 17,090,000 BTU/hr of heat are saved by this interchange. The waste stream then goes to the aqueous waste feed preheater (42) where it is heated to the boiling point and some 5% of the stream is evaporated. The purpose of item 42 is to: 1) volatize off any Cl-that may be contained in the stream and thereby decrease corrosion of the subsequent evaporation heat transfer surfaces; 2) volatize off any organic material that may be dissolved or entrained in the waste stream, and 3) precipitate any S04 present as CaS04 that has a negative temperature solubility coefficient on the theory that item 42 can be more easily cleaned or spared if necessary than the subsequent evaporators. Overhead from the aqueous feed preheater (42) is condensed in the aqueous waste feed preheater overhead condenser (43). This condensate is then discarded. The aqueous waste stream then flows to the waste evaporators at its boiling point, 221~Fo The waste stream evaporators are two multi-effect units (items 44 and 45), The first of these is steam-heated with the overhead from this unit supplying the heat for the second. The necessary temperature At's are obtained by a barometric leg (40), pulling through the second waste evaporator overhead condenser (38). Inert gases and the startup air are removed by a steam ejector and condenser unit (39). The heatload split between the two evaporators is shown in 461-1128-U-30 Condensate from both evaporator effects is pumped back into the treated water system of the processo Both evaporators (44 and 45) are similar in design and differ only in the heat-transfer surface contained. The tentative design of these evaporators is shown in Fig. 10o Here again, and for the same reasons given for the product evaporator (26) thermosyphon type of evaporators are used. Entrainment from these evaporators is not as serious as from the product evaporator, so more leeway can be taken in the design of the vapor section. It may be possible to eliminate the mesh de-entraining section. as shown in Fig. 10 in the final design. The bottoms from the second aqueous waste evaporator (45) are pumped to the spray nozzle in the calciner (47). The calciner details are shown in Fig. 11. This calciner operates at 150 psig so that its size is reasonable and is composed of a large chamber into which the concentrated ---- - - 28

z > WIRE MESH DE-ENTRAINER ALL DIMENSIONS UNLESS OTHERWISE SPECIFIED MUST BE HELD TO A TOLERANCE - FRACTIONAL ~ %j4" DECIMAL ~.00S." ANGULAR ENGINEERING RESEARCH INSTITUE DLESIGNED BY M.E.W APPROVED BY ENGINEERING RESEARCH INSTITUTE DRAWN BY R. I.K. SCALE 1/4" - I UNIVERSITY OF MICHIGAN CHECKED BY M E.W DATE 5/2/57 ANN ARBOR MICHIGAN PROJECT= AQUEOUS WASTE 461-1128 EVAPORATOR DESIGN ISSUE AT CLA81IFICATION UNCLASSIFIED FIGURE-10 Icr~r I ATEUNCLASSIFIED 29

_____ _1T_ -E DESIGNED BY M.E.W. APPROVED BY ENGINEERING RESEARCH INSTITUTE DRAWN BY R.. K. SCALE 1/4 I UNIVERSITY OF MICHIGAN CHECKED BY M E.W. DATE 5/31 /7 ANN ARBOR MICHIGAN TITLE -- -- PROJECT 4611128 WASTE CALCINER & _ 46"1-1128 CYCLONE SEPARATOR ISSU DATE S UNCLASSIFED FIGURE-II DATE 30

The University of Michigan * Engineering Research Institute aqueous waste salts are sprayed. This spray nozzle breaks the solution up into droplets in the 50-to-20-micron range. As the droplets fall, they are heated and dried by a hot gas stream flowing upward through the calciner temperature. As the particles are dried, they are carried upward by the hot gas stream and pass through a cyclone separator where the solids and gases separate~ The solids fall to the conical bottom of the cyclone and are discharged through a solids-collection lock out of the bottom of the unit and are hauled away. The gases are recirculated through a recirculating gas blower (49) and through heating surfaces contained in the recirculating gas heater and steam generator (50). Gas temperatures into the calciner are 1200~F and leave the cyclone at 700~Fo The hydrated salt Ca(N03)2)o4B20 is decomposed at temperatures of 700~F into steam, NV2, and 02o Residence time in the calciner is approximately 1.2 sec, which should be adequate for the decomposition to take place. It is these gases that are recirculated through the calciner and thus serve as a heat-transfer medium. A calcining system such as this is discussed in Refo 10o The NO2 content of the resultant gas is high and consequently makes an ideal feed to an absorption tower for the formation of nitric acido Pressure in the calciner unit is controlled by a pressureregulating valve set at 150 psig; thus the gas discharge from the system is automatic. It is believed that a spray-dryer system as discussed is mandatory for this calciner system. Other types of calciners became prohibi>-. tively large in size or result in gases so dilute in N02 that difficulties would be encountered in the subsequent absorption stepo The recirculating gas blower (49) poses a design problem, although not an insurmountable one, because it operates at a 700"F temperature and >i 150-psig inlet pressure~ This blower would have to be constructed of austenitic stainless steel to prevent corrosion and this type of material shiould also easily meet the temperature conditions. The seal problem would, however, require considerable attention. This blower is rated at 10,U00 cfm at 150 psig against a 1o25-psi pressure differential. About 75 hp is required to drive it. A centrifugal type of blower would be necessary to meet this service. The recirculated gas from the calciner is heated in special heattransfer surfaces placed in the recirculating gas heater and steam generator (50). This unit is illustrated in Fig. 12. The gas heating surface is place!d adjacent to the steam envelope so that the 1200~F outlet temperature can be achieved, This heating surface is formed by rolling and welding the tubes into pipe headers at each endO The whole unit is supported at each end and with two central hanging supports on the tubeso Tubes are approximately 16 ft long and have a 1-ino OD, The total surface of this unit is 2420 ft2 31

iii 0 z > /TO SACKOS FEED WATER HEATE STEAM -I - 250 psig! 400 F AMGENERATO I OO~RFCYC4E GA HiT (CLE GA - I I _ OQoF^" 7^ I FEED WATER 70-80~ F RECI 12 *-RECYCLE GAS 150 psig, 700~ F r') CN I OUTER SHEATH IS ORDINARY BRICK WITH METAL SUPPORT INNER ROW FIRE BRICK, -t' ~NATURAL GAS BLOWER d ALL DIMENSIONS UNLESS OTHERWISE SPlCIFIED MUST BE HELD TO A TOLERANCE - FRACTIONAL 1,." DECIMAL.00S," ANGULAR ~ 5o ID.ION.. BY M.E.W. APPROVED BY Y. ENGINEERING RESEARCH INSTITUTE DAN Y K. I CROV DRAWN BY R I. K. SCALE V/i'' ^ NAUNIVERSITY OF MICHIGAN CHKCKKD BY M, E.W. DAT 5/27/55 ANN ARBOR MICHIGAN PROJECT ---— L RECYCLE GAS HEATER E 461-1128 WASTE STEAM GENERATOR ISUE DATE CASSIFIATION UNCLASSIFIED FIGURE - 12 32

The University of Michigan * Engineering Research Institute The combustion gases after passing through the gas heating surface are still at about 1500~F. This heat from 1500~F down to 400OF is used to generate steam at 250 psia for use in the process and in building heating. As can be seen in Dwg 461-1128-U-3, there is sufficient heat in the combustion gases after supplying the calciner requirements to supply 10,387 lb/hr of steam more than is required by the process or in heating loado This suggests that some pumps and blowers could be steam-driven; however, no provisions have been made to do so in this study. In Fig. 12, the fuel used is natural gas and this is used at 3000-scfm rate supplied at 25 psig. Oil or some other convenient fuel could be considered. Steam, NO2, and 02 must now be converted into nitric acid as it leaves the calciner (47). As the pressure tends to increase in this unit above 150 psig, a pressure-regulating valve bleeds off the gas which then flows to thebcalciner product cooler and condenser. In this unit the steam is condensed and subcooled to 100~Fo Two heat exchangers are used here in series to accomplish the condensing subcooling function. Simultaneously with condensing and subcooling, considerable absorption of the NO2 and 02 occurs to form nitric acid in these exchangerso The condensate from the condenser-subcooler is pumped to the second tray from the bottom of the absorber column (55), where the liquid phase acid concentrations are equal. The unabsorbed gas phase goes in below the bottom of the absorber column and passes upward through the bubble trayso Distilled water is pumped by a pump (54) to the top of the absorber column as the absorbing stream. Nitric acid at 56.5% by weight is pumped from the bottom of the column to the nitric acid storage tank (17). Figure 13 shows the absorption tower (53). This tentative tower design was arrived at from equilibrium data in Refo 11 and from the reaction kinetics given in Ref. 12 and pilot plant operating data from Ref. 13. Considerable heat is evolved from the N02-water absorption reaction and increasing temperature is a disadvantage in further absorption. For this reason every other tray in the absorber column is a cooling tray. Details of these trays are shown in Fig. 13. Twenty bubble trays and 19 cooling trays are thus used. The column dimensions then become 3 ft 6 in. in diameter and 64 ft high, which includes bottom surge capacity. This is a very small column for such a large acid-production capacity; however, the feed is very concentrated in NO2, and the column is operated under 150psig pressure which aids absorption, and the downcoming liquor is cooled after each bubble tray. With this provision, it should be readily possible to meet the desired rates. Compressed air is blown into the bottom of the column which aids in absorption since excess oxygen drives the reaction in the desired directiono The air compressor is item 56, and it supplies 10 cfm at150 psig and 100~Fo 53o.7 Storage and Make-Up Tankage and Auxiliary Equipmento-The process equipment as a whole has been discussed with the exception of the tankage and necessary auxiliary equipmento The tanks provided are the 33

I 1 > COOUNG 44 WATER CONN. 1 ",150 lb 38 REdD. 3'6" — ] I (BUBBLE TRAY-20 REQD. I00I 00 o0) |I - I z -3"~PRESSED STEEL T~I~~~ SELF SUPPOR TING CAPS - 29/TRAY COND. COOLING TRAY —19 REQb. AIR + N 02. s IAT U ALLL DIESO2 3 OL~ NCAKE 4 HIGH ALL DIMENSIONS UNLESS OTHERWISE SPECIFIED MUST BE HELD TO A TOLERANCE - FRACTIONAL ~ I-4," DECIMAL ~.001," ANGULAR - jo DESIGNEDBY M.E.W APPROVED BY ENGINEERING RESEARCH INSTITUTE DRAWN BY R 1. K. SCALE V/4= I UNIVERSITY OF MICHIGAN CHECKED BY M.E.W. DATE 5/29/57 ANN ARBOR MICHIGAN TITLE 4 1PROJECT 6 2 ABSORPTION TOWER 461- 1128UNCLASSIFIED,UE | DAT" CLASSIFICATION UNCLASS IFIED F I G U R E - 13 _ _~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~~I i - 34

The University of Michigan * Engineering Research Institute treated water storage tank (16), which holds the clarified water for use in washing stages of the extractor and for general plant use where nonpotable water is required; the nitric acid storage tank (17), which has been discussed; the solvent storage tank (20), which holds up approximately three process hold-up volumes of the TBP-kerosene mixture; the distilled water storage tank, which collects steam drips from the condensers for use in the solvent stripper and solvent washing-scrubbing operations; the solvent makeup tank (34); and solvent wash solution make-up and storage tanks (36 and 35), which serve to make the solvent wash caustic solution to the proper concentration and volume requirements. These tanks are covered further in the cost and specification sheets in a later section. Also required will be feed-water treatment facilities for the steam generatoro This unit is a standard package unit that is in no way unique, so its details are not discussed. Costs for the unit have been allowed for, assuming it to be part of the recirculating gas heater and steam generator. A water-cooling tower (59) is shown on flowsheet 461-1128-U-3. It was deemed advisable to include this in the plant costs since -uranium mills are normally located in arid regions where water is scarce. This unit is a forced-draft package unit that can be furnished by several suppliers. It cools 5294 gpm from 120~ to 80~F, and will require 120 gpm of make-up water during normal operation. A cooling-water pump that supplies 3294 gpm cooling water at 60 psig for all plant cooling-water services is also shown. Provisions for furnishing clarified nonpotable water and potable water are not specifically shown. The specific design varies widely depending upon local water conditions. Costs were allowed in the plant-utilities estimate for average water conditions such as pumping from a well or from a stream. Admittedly, these costs could be in error for any specific situation, but are believed correct for most probable situations. Other facilities required by the plant would be vehicles, ore loaders, fork-lift loading trucks, and loaders and track cars for calcine disposalo Costs have been allowed for these items, and since use is so specific, no detailed description of their function is requiredo 354 UTILITY REQUIREMENTS Utility requirements have been discussed in the process discussion; however, the total requirements were not presented. These are given in the paragraphs belowo 3.4,1 Fuel Requirements. —Plant vehicles are assumed to be dieselengine-driven and will consume approximately 10,000 gal/yr of diesel oil. This assumes 24 hr/day operation for 300 days per year. 35

t The University of Michigan ~ Engineering Research Institute I Natural gas is used for process fuel at the rate of 3000 scfm delivered at 25 psig. 3.4.2 Electrical Power.-The electrical power load of 120-440 volt a-c power, including building and grounds lighting as well as all process and plant loads, will be 825 kw. 3.4.3 Process Water Requirements.-Process water requirements will amount to approximately 200 gpm. This water is clarified, but not biologically purified. This water will supply the boiler feed-water purifiers. Potable water at 30 gal/day per person will be approximately 1.5 gpm. 4.0 EQUIPMENT COSTS Equipment costs mean very little without an accompanying equipment specification. However, inclusion of a complete specification would make this report very bulky and since most of the major equipment is illustrated in figures together with the throughput and heat-transfer rates, it is believed that the costs given can be related to the equipment without additional specificationo There is another factor that has not been discussed: the materials of construction. Since nitrates or nitric acid is handled all through the process except when raw ore is handled, it is easier to point out those items that are not constructed of stainless steel rather than those that are. The items up to item 16 are of mild or low alloy steel construction. This includes all the ore handling, crushing, and grinding equipment. Item 16, a treated water storage tank, is of galvanized sheet iron construction. It is of course understood that all supports, platforms, walkways, and structures are mild steel. Where heat exchangers are shown with the process fluid in tubes, the shell and baffles may be of mild steel, and where warranted the heads and tube sheets are stainless-steel clad. Steam and feed water heat-transfer surfaces in item 50, the recirculating gas heater and steam generator, are of low alloy steel, All steam, water, and air lines are mild steel. All items not specifically mentioned above can be assumed to be constructed of either 304, 347, or 316 stainless steel. 4o1 EQUIPMENT PRICING METHODS The best method of pricing equipment is, of course, by direct quotation from an equipment supplier. Quotations were obtained from suppliers on the ore crushing and grinding equipment, as well as the belt conveyor and the vibratory feeders. However, quotations on all items would be very time-consuming and probably unnecessary. Prices on the other equipment were I 36

r — The University of Michigan * Engineering Research Institute obtained using the data in Refso 14, 15, 16 and 17. Data collected by the author were also used where such data were believed to be more reliable or more recent than data given in the references. Since the cost information used may be several years old, it is necessary to bring it up to date. This was done by multiplying the cost data given by the ratio of an averaged labor and materials index for 1956 and the year in which the data were prepared. Since the labor and material indexes are not the same, a mean value of the two must be used. Where mild or low alloy steel was used in the equipment, a 50%-labor, 50%-material cost for the item was assumed and the labor-material index averaged accordingly. Where stainless steel equipment is priced, a 40%-labor-60%-material factor was used. These indexes up to May, 1950, are given in Refo 15. Table III gives the indexes up to December, 1956. These data are from the U. S. Department of Labor, Bureau of Labor Statistics, and from the wholesale price index for metals and metal products. 4.2 FREIGHT CHARGES The plant pricing methods use as a base the delivered equipment costs which include a freight charge. Freight charges vary with route, nature of shipment, and quantity of shipment and of course the type of shipper. For the purpose in mind it would be impossible to arrive at accurate freight rates without knowing the plant location and the location of the equipment supplier. Freight costs for an average condition are as accurate as can be obtained at this time. If it is assumed that the plant is located in the Salt Lake City, Utah, area, then the equipment would probably be shipped from one of these areas: Chicago, San Francisco, or Los Angeles. With this assumption, freight rates were supplied by the Union Pacific Railroad for various classes of equipment. These data are shown in Table IVo Additional freight cost information is available in Ref. 18. In using this information in this reference, it is necessary to make adjustments for increases occurring since the data were tabulated. 4,3 DELIVERED EQUIPMENT COSTS Using the methods described above, equipment and freight costs were determined. These costs are given in Table V. This table does not include pump costs. Pump service and ratings are listed separately in Table VIo The price given includes pump, motor and mounts, and freight. The delivered equipment costs, the sum of totals of Tables V and VI, is then $826,546. 37

The University of Michigan * Engineering Research Institute TABLE III Material and Labor Index 1950 Through 1956 Year Material Labor Year Material Labor Month Index Index Month Index Index 1950* Jan, Feb. Mar Apr. May June July Aug. Sept. Oct. Nov. Dec. 1952 Jano Febo Maro Apro May June July Aug. Sept. Octo Nov Dec. 1043. 104.9 o04.8 105.2 106.6 108.8 109.2 113.3 ll6.1 117.7 121.9 122.4 122.6 122.6 122.5 121.8 121.8 121.9 124.1 124.6 124.1 123.9 124.0 148.6 148.3 148.6 149.8 150.9 152.4 153.3 153.9 156.1 157.7 158.8 161.5 172.5 172.9 174.4 174.2 174.6 174.7 173.3 176.8 181.o 181.9 183.0 184.3 1951 Jan. Feb. Mar. Apr. May June July Aug. Sept. Oct. Nov. Dec. 1953 Jan. Feb. Mar. Apr. May June July Augo Sept. Oct. Nov. Dec. 124.0 123.7 123.2 123.3 123.2 122.7 122.3 122.2 122.1 122.4 122.5 122.5 124.0 124.6 125.5 125.7 125.7 126.9 129.3 129.4 12804 127.9 127.9 127.5 162.7 163.7 165.1 165.9 166.4 168.4 168.2 168.3 170.3 170.3 170.5 184.8 184.8 185 185 186 187 188 188 189 189 189 190 1955 Jan. Febo Mar Apr. May June July Aug. Sept Oct. Nov. Dec. 127.2 126.2 126.3 126.8 127.1 127.1 128.0 128.6 129.1 129.7 129.9 129.8 191 190 190 190 191 191 191 191 192 193 194 195 Jan. Feb. Mar. Apr. May June July Aug. Sept. Oct. Nov. Dec. 38 130.1 131.5 131.9 132.9 132.5 132.6 136.7 139.5 141.9 142.3 142.9 144.9 196 196 197 197 199 199 202 201 204 206 206 206

The University of Michigan Engineering Research Institute TABLE III (Concluded) Year Material Labor Year Material Labor Month Index Index Month Index Index 1956 1956 (concluded) Jan. 144 9 205 July 144 9 207 Febo 145o1 205 Augo 150 2 210 Maro 146 5 206 Sept. 151 9 213 Apr. 147.7 208 Oct. 152.2 215 May 146.8 208 NovO 152,1 216 June 145.8 209 Dec, 152.4 218 *The wholesale price index was revised in 1952 to a 1947-49 = 100 basis. These figures are based on that revision. 39

The University of Michigan * Engineering Research Institute - I I TABLE IV Freight Rates on Various Classes of Equipment Statement of rates in cents per 100 lb to Salt Lake City from Chicago, Los Angeles, and San Francisco on commodities indicated below, M/W for From LCL CL 40-ft Car Heat Exchangers: (Item 28212) UoF.Co Noo 3 Chicago Los Angeles San Francisco Chicago Los Angeles San Francisco 540 585 385 54o 385 385 286 204 204 272 204 204 24,000 24,000 24,000 24,000 24,000 24,000 Grinding or Crushing Equipment: (Item 29662, UoFoCo No. 3) Fabricated Tanks: 1/4 inch or thinner but not thinner than 16 gauge Fabricated Tanks: Thicker than 1/4 inch (Item 41705 uoF.C. No. 3) Chicago 635 (381 (254 (272 454 (182 14,000 24,000 14,000 24,000 I Los Angeles) San Francisco) Chicago 540 Los Angeles) San Francisco) 385 *(1l5 *(218 40,000 20,000 Distilling Apparatus, NOIBN: Iron or Steel (Item 28525 UoFOC, No. 3) Chicago Los Angeles San Francisco 635 454 454 518 227 227 24,000 24,000 24,000 Bulk Chemicals: in tank cars (Item 1090-E) PoSoFoBo Tariff 260-E Los Angeles San Francisco 57 57 * Applies on tanks, plate or sheet, No. 2 gauge or thicker (Item 8970-E, PSFB Tariff 260-E)o Rates include all ex parte increases

The University of Michigan Engineering Research Institute TABLE V Summary of Equipment and Freight Costs It em Initial Delivered No. Item Cost, $ Wt,lb Freight Cost, $ 1 Ore Screen 150 500 22 70 173 2 Ore Hopper 677 971 37o00 714 3 Grizzly-Scalper-Feeder 3619 7100 38o0oo 3999 4 Grrizzly UJnderflow Hopper 15 300 1400 164 5 Jaw Crusher 32760 38000 1030o00 33790 6 Belt Conveyor 12000 15000 810o00 12810 7.Cone Crusher Bin 233 800c 3600 269 8 Cone Crusher 33280 59000 1600o00 34880 9 Cone Crusher UTnderflow Bin 650 1050 40.00 690 10 Rod Crusher 93100 150000 - 93100 11 Bucket Elevator 3500 6000 320.00 3820 12 Fine Ore Storage Bin 4930 20000 548.00 5478 13 Fine Ore Feeder 1794 3500 189o00 1933 14 Extractor 48900 29760 607.00 49507 16 Treated Water Storage Tank 3100 8500 390.00 3490 17 Hl3O Storage Tank 13300 8200 380o00 13680 19 Solvent Extractor 31100 14297 292.00 31392 20 Solvent Storage Tank 1850 1700 77.00 1927 23 Solvent Stripper 9000 3425 155.00 9155 24 Dist. Water Storage Tank 5900 7500 340.00 6240 26 Product Evaporator, 2230 1930 74o00 2304 27 Solvent Scrubber 650 241 9o00 659 28 Solvent Washer 650 241 9o00 659 29 Product Evap. Ovhd. Condenser. 2230 1500 58.0c 2288 30 Prod. Evap. and Prod. Calciner Condo Collection Tank 300 60 300 35053 32 Product Calciner 4300 1500 58o00 4358 33 Prod. Calciner Ovhd. Condenser 2730 1200 50)00 2780 34 Solvent Make-Up Tank 925 850 39500 964 35 SolvO Wash Solution Stor. Tank 1200 900 41o00 1241 56 Solv. Wash Solution Make-Up 950 850 39-00 989 37 Second Waste Evapo Ovhdo to Aqueous Waste Heat Exchanger 7000 6290 240.00 7240 38 2nd Waste Evapo Ovhdo Condenser 31600 14000 540o00 32140 39 Steam Ejector and Condenser 1785 500 27o00 1812 40 Barometric Leg 354 250 10o00 364 =1

t The University of Michigan * Engineering Research Institute l TABLE V (Concluded) Item Initial Delivered No. Item Cost, $ Wt,lb Freight Cost, $ 42 Aqo Waste Feed Pre-heater 43 Aqo Waste Feed Pre-heater Overhead Condenser 44 1st Aqueous Waste Evap. 45 2nd Aqueous Waste Evap. 47 Calciner 48 Cyclone Separator ) -- 49 Recirculating Gas Blower) 50 Recirculating Gas Heater and Steam Generator 51 CalcoProdoCooler and Condo 53 Absorber Column 56 Air Compressor 58 Boiler Feedwater Treaters 59 Cooling Tower 5920 3500 135.00 6055 6200 21700 20300 2000 8500 8000 77.00 330,00 310.00 6277 22030 20610 64578 47000 960 o0 65538 137300* 28400 30900 2500 24110 15000 12700 15900 4500 30000 31200 490o00 720 00 12350 612 00 636.00 137300 28890 31620 2673 24722 15636 1 MISCELLANEOUS Ore Trucks (Hydraulic lift) Prod. Packaging Facilities Fork Lift Truck Calcine Dump Cars and Track Plant Vehicles Power Substation $20/kw 6500 2500 1500 21200 11000 16000 1000 1500 20000 8000 54o00 81o00 770 00 308.00 6500 2554 1581 21970 11000 16308 TOTAL $ 786,576 *Installed Cost 5.0 PLANT COSTS Methods of estimating plant costs from delivered equipment costs as well as on other bases are given in Refs. 14, 15, and 16. However, the method that appears best and has the greatest amount of supporting data is Lang's Refo 16, Part. I. Lang has taken cost data from fourteen plants of various types and broken these costs down into factors times the delivered 42

TABLE VI Summary of Pump Costs, Service, and Rating Item TteRtnanFucinMtrasadyp Delivered Item Title Rating and Function Materials and Type Delivered No. Cost, $ 15 Sand Pump 18 Nitric Acid Feed Pump 21 Solvent Pump 22 Aqueous Waste Pump Pumps 70% solids by vol: 75 gpm, 40-psig head, 2.5 hp 66 gpm of 56.5% HNO3, p = 1.335, 1 hp, 25psig head Pumps 13.2 gpm, solv. p = 0.84 25-psig head, 1/2 hp Solv. of Ca(N03)2 and HN03, 120 gpm, 60-ft head, p = 1.338, 2 hp Cast iron body, stainless steel shaft, open centrifugal, 1750 rpm 304 s.s. casing and impeller and shaft, 1750 rpm 304 s.s. bowl, impeller and shaft, 1750 rpm 304 s.s. bowl, impeller and shaft, 3600 rpm $ 1,420 (spared) 1,462 (spared) 882 (spared) 2,001 (spared) 25 Distilled Water Feed Pump Pumps combined scrub, strip and solv. wash., 8.5 gpm H20,60-ft head, 1/2 hp M.S. 859 31 Product Ovhd. Evap. Pump 41 Aqueous Waste Ovhd. Condensate Pump 46 Calciner Feed Pump 52 Calciner Prod. Cooler and Cond. Condensate Pump 54 Absorber Water Feed Pump 55 Absorber Product Pump 57 Boiler Feed Water Pump 60 Cooling Water Pumps 61 Steam Condensate Return Pump and Well Pumps 4.5 gpm of 94.59 w/l as nitrates. p = 1.095, 1/3 hp H2C and HN03 solv., p = 1.02, 35 gpm, 25-psig head, 1/2 hp 47.9 gpm at 235~F, p = 1.75, 233-ft head, 8 hp 65 gpm of 50% HNO3, 40-ft hd, p = 1.31, 3/4 hp 15.2 gpm H0O, 100-ft head, 1/2 hp 66 gpm of 56.5% HNO3, p = 1.335, 25-psig hd, 1 hp 97 gpm feed water to 250 psig, 12 hp Fluid cooling, HE0-3294 gpm,60-ft hd 65 gpm I O and Dil HN03, 80-ft hd, p = 1.02, 2.5 hp Same as No. 21 Same as No. 31 304 s.s. bowl, impeller and shaft, 1750 rpm Same as No. 31 s.s. const., 3600 rpm s.s. const., 1750 rpm Bronze impeller s.s. shaft, 3600 rpm 304 s.s. casing, impeller and shaft, 1750 rpm Mild steel or bronze s.s. shaft, 3600 rpm M.S. const., 1750 rpm s.s. const., 1750 rpm Same as No. 21 Same as No. 21 442 2,359 442 (spared) 1,367 (spared) 1,446 (spared) 1,245 (spared) 562 62 Washed Solv. Pump 442 6-) Product Pump 442 $20,021 TOTAL 43

The University of Michigan T Engineering Research Institute plant costs or a factor times the total plant investment. The values given seem to be consistent with data collected by the authoro Consequently Lang's method of estimating plant costs were largely adhered to. Unfortunately, the factors to be applied for any item of plant cost can have a rather wide range. A question also arises about just what category a uranium-processing plant fits into. Cost data are given on plants handling only solids, solids and fluids, and only fluids. In these areas the element of judgment enters into the selection of just what should and should not applyo In this estimate the solids (ore crushing and grinding) section was treated as a solids-handling plant, and the remainder of the plant as a fluids-handling plant. Appropriate factors from Lang were then used for each section to arrive at a plant cost.:hI Table VII, the various costs that comprise a total plant investment and just what each cost includes, are listed. Plant costs will be $3,814,000, total cost, $4,385,000, and physical cost for tax purposes, $2,311,320, using the basis discussed above. 6.0 PLANT OPERATING COSTS 6.1 FIXED CHARGES Plant fixed charges are shown in Table VIII. The items in this table are believed self-explanatory. The percentage factors used in arriving at costs are based upon data in Refo 15. 6o2 VARIABLE CHARGES A major factor in variable costs will be salaries. Manpower requirements are shown in Table IX. The costs per shift for operating personnel from this table is $77,430. All variable charges are shown in Table X, which is believed to be self-explanatory. The variable charges at 1000 tons of ore per day, 300 operating days per year, will be $950,964. Operating costs for other plant throughput rates can be easily arrived at by using the information presented in Tables VIII and Xo

r The University of Michigan * Engineering Research Institute I TABLE VII Plant Cost Summary Delivered Equipment Cost(l) Installation Costs(2) Piping(3 ) Electrical Eqpto and Installation(4) Process Control Instruments(5) Service Facilities 6) (7) Buildings(7) Grounds Improvement(8) Land(9) Total Physical Cost Engineering and Construction(l0) Contingency (11) Size Factor 12) Total. Plant Cost $ 806,597 246,903 523,320 129,900 63,200 253,900 196,900 70,600 20,000 $2,311,320 694,000 462,000 347,000 $3,814,320 Working Capital(13) 571,000 I TOTAL INVESTMENT $4,385,320 (1) Sum of totals from Tables V and VI. (2) The Rod Crusher (10) and the recirculating gas heater and steam generator (50) were quoted on an installed basis, so these items were excluded from installation costs. Installation costs were then taken as 30% of delivered equipment costs less the two items mentioned. Installation includes foundations and supports, ladders, platforms, walkways, and thermal insulation. (3) Based on 7o4% of installed equipment costs for solids section, and 62% of same for fluids sectiono Cost includes all internal process, and auxiliary lines, but not outside lines. (4) Based on 13o53 and 12% of installed equipment costs for solids and fluids sections, respectively. Cost includes lines from substation to plant (but not substation itself), switchgear, internal conduit, motor starters, internal lines, lighting, and switcheso (5) Includes all process control instruments, panels, accessories, spare parts, calibration, and service equipment~ Based on 6% of installed equipment costso 45

The University of Michigan * Engineering Research Institute 1 TABLE VII (Concluded) (6) Based on 11.1% and 27.8% of installed equipment costs for solids and fluids sections, respectively. Includes plant external piping such as steam and water supply lines, sewers, process drains, cooling water lines and fire protection apparatus. (7) Includes process, laboratory, and office buildings, and office furniture. Based upon 23.5% and 17.53% of installed equipment costs for solids and fluids sections, respectively. (8) Includes grading and leveling, fences, roads, sidewalks, parking areas, and landscaping. Based upon 15% and 4.3% of installed equipment costs for solids and fluids sections, respectively. (9) Assumes a total of 20 acres required at $100/acre. This applies to undeveloped arid land in an area where a mill Of this type would probably be built (10) Taken as 30% of total physical cost. Includes engineering fees, field and home office expense, contractor fees, and all other construction costs. (11) Assumed as 20% of total physical cost. (12) Assume as 15% of total physical cost. This factor allows for process changes and capacity adjustments after engineering has been completed. (13) Working capital at 15% of total plant costs allowed. This covers normal inventory and other operating burdens encountered during operation. TABLE VIII Plant Fixed Charges Item Cost, $/yr AMORTIZATION-Average life of 15 years assumed = 211320 $154,000 15 INTEREST-On plant investment at 6% =.06 x $4,385,000 263,000 TAXES-State and local at 2% physical cost =.02 x 2,311,320 46,200 INSURANCE-At 1% total physical cost =.01 x 2,311,320 23,100 SALARIES-Of nonoperating personnel (see Table IX) 110,210 LABOR OVERHEAD —At 15% of salaries =.15 x 110,210 16,500 OFFICE OVERHEAD-At lO1o of salaries =.10 x 110,210 11,000 TOTAL FIXED CHARGES $624,010 I I

The University of Michigan * Engineering Research Institute TABLE IX Plant Manpower Requirements (Operating 3 shifts/day, 300 days/year) OPERATING PERSONNEL No. of Pay Rate People Location Job Description Pa Rate People __$h T 2 Ore Loading 1 Jaw Crusher and Belt Conveyor 1 Cone Crusher and Rod Mill 1 Extractor, Solvent Ext., Stripper, Prod. Evaporator, Prod. Calciner 1 Product Loader 1 Product Handling and Storage 1 Solv. Wash. Solution Make-Up, Tank Farm 1 Aqueous Waste Preheat and Evaporation Tank Farm 1 Calciner and Abs. Clm. 1 Recirc. Gas Heater and Steam Generator 1 Gate 1 Chem. Lab. Truck Operators Machine Operator Machine Operator Chemical Operator Chem. Op. Helper Chem. Op. Helper Chem. Op. Helper Chem. Operator Chem. Operator $3.70 2.28 2.28 2.48 1.98 1.98 1.98 2.48 2.48 2.50 1.75 2.75 2.75 3.50 1.85 oo00 $ 7,800 4,810 4,810 5,240 4,180 4,180 4,180 5,240 5,240 5,270 3,690 5,800 5,800 7,390 3,800 8 440 $ 77,430 Boiler Operator Guard Chemist 1 1 1 1 Calcine Disposal Foreman Shift Supervisor Truck Operator Prod. Supervisor TOTAL OPERATING PERSONNEL/SHIFT _ NONOPERATING PERSONNEL - Maintenance 1 Plant Engineer 1 Millwright 1 Welder 2 Pipefitters at $2.40 2 Maintenance Men at $2.15 2 Janitors at $1.60 Plant Manager Accounting (1) Secretary (1) Warehouse and Shipping (1) NONOPERATING PERSONNEL - Ore Buying 1 Weighman 2 Checkers and Samplers at $2.10 2 Lab Assistants at $2.00 1 Bookkeeper 1 Secretary TOTAL NONOPERATING PERSONNEL 4.25 3.00 2.25 4.80 4.30 3.20 5.65 2.50 1.75 2.25 8,960 6,330. 4,750 10,100 9,060 6,750 $ 45,950 12,000 5,280 3,690 4,750 $ 25,720 4,430 8,870 8,450 4,750 3 700 $ 0,100oo $101,770 2.10 4.20 4.oo 2.25 1.75 Salary scales from Oil and Gas Journal, June 9, 1949. Ratioed up by the labor index to December, 1956. 47

The University of Michigan T Engineering Research Institute TABLE X Plant Variable Charges (Table assumes 3 shifts/day operating 300 days/yr on a 1000-ton/day rate) Item Cost, $ Salaries: Operating Personnel - 3 x $77,430 Payroll Overhead - at 15% of payroll General Plant Overhead - at 40% of operating salaries Repairs and Maintenance - at 5% of total plant physical cost-.05 x 2,311,320 $232,290 34,800 92,800 116,000 Chemicals: HNO3* TBP** Assume 1% loss/ton ore 9.94 lb/ton x 300,000 tons ore x $.022/lb tech. grade + freight at $.57/100 lb Losses 270 g/ton ore or $.312/ton ore $.512 x 300,000 + freight on 178,600 lb/yr at $.57/100 lb 82,600 94,600 Kerosene Losses 2100 g/ton ore or $.104/ton at $.15/gal $.10 x 300,000 + freight on 1,137,000 lb/yr at $.57/100 lb Na2CO3 0.8 lb Na2CO3/ton ore at $1.50/100 lb + freight 240,000 lb at $.57/100 lb Fuel and Power: Fuel Oil - 10,000 gal at $.12/gal Nat. Gas - 3000 scfm at $.15/1000 scf Power 825 kw at $.007/kw hr Product Shipping Costs: 1,830,000 lb U03/yr Freight at $.90/100 lb Drums at $..50/100 lb Water Costs: Process Water - 200 gpm at $.14/1000 gal Potable Water - 30 gal/person/day at $.50/1000 gal TOTAL VARIABLE COSTS 37,680 6,210 1,200 180,000 41,600 _ 1 16,500 2,280 12,080 324 $950,964/yr *Price quote from DuPont Explosives Division for a 56% technical **Price from Commercial Solvents circular P.S. - No. 38A, January cost $.525/lb in drums and carload lots. grade HN03. 1, 1957, TBP 48

The University of Michigan T Engineering Research Institute 1 7.0 PROBABLE PAY-OUT TIMES AND PROFITS For comparison purposes it is convenient to have the charges based on a cost-per-ton basis. This is shown in Table XI below. These costs were taken from the data in Tables VIII and X. The total costs for processing a ton of ore is then $5.25. TABLE XI Plant Costs on a Per-Ton-of-Ore Basis (Table assumes 1000-ton/day plant rate operating 300 days/year) Cost, $/Ton Ore Fixed Charges Amortization.513 Interest.875 Taxes (State and local).154 Insurance.077 Salaries (nonoperating personnel).366 Labor Overhead.055 Office Overhead. 037 $2.0o8 Variable Charges Salaries (operating personnel).774 Payroll Overhead.116 General Plant Overhead.310 Repairs and Maintenance.387 Chemicals.736 Fuel and Power. 744 Product Shipping.063 Water Cost.041 $3.17 TOTAL FIXED PLUS VARIABLE COSTS $5.25 If costs paid to the miner for the ore are based upon the Atomic Energy Commission schedule, the base cost for U308 will be approximately $3.50 per lb. If the bonuses covering higher grades of ore and quantity plus haulage are allowed, the actual cost could approach $3.75 per lb U308. On the basis of a 0.3% ore content assumed for this report, the cost of ore to the processor would be $22.50 per ton for the $3.75 per lb base. The total cost to the processor would be $27.75 per ton of ore. I 49

1 The University of Michigan * Engineering Research Institute I 1 The present price for U308 from the A.EoCo is approximately $10 per lb for a product of 75% minimum U308. The differential per ton is then $32.25 per ton of ore. The gross income before Federal income taxes wouldbe $9,680,000 per year, assuming 1000 tons per day and 300 operating days per year. With income taxes assumed at 52% of profits, the net to the processor would be $4,650,000 per year. This represents a "pay-out time" of one year. Other standards could be applied to evaluate the economic feasibility of such a plant. However, by any other standards the process proposed should appear equally good. The major reason for the high profit margin lies in the low chemical costs inherent in the proposed process as compared to processes currently used. It is possible for a comparatively small investment to add equipment to this process so that a product of much greater purity could be obtained. This would involve a second solvent-extraction-stripping operation similar to the cycle shown on this process but physically smaller. By increasing the purity to meet more rigid specifications, the processor could qualify for a price of $12.50 per lb of U308 rather than $10.00. This addition would be well worth serious consideration if other plant process and design problems are satisfactorily met. No value has been placed on the CaO from the aqueous waste processing It has been assumed that its value would offset the disposal costs. Similarly, no provisions have been made for vanadium processing since many ores with high lime content contain little or no vanadium. The ore composition assumed is one with high lime and no vanadium. 8.0 PROCESS AND DESIGN PROBLEMS IN THE PROPOSED PLANT It should not be construed that the proposed plant is ready to enter a design phase. Rather this study is to serve as a guide in directing a research effort toward a goal that has considerable economic promise. While the steps portrayed in this process are known to be technically sound, the selection of equipment and materials can still be the factors that make or break a plant. A suggested program to be followed for a successful final plant design is as follows: 1) A bench-scale test of each of the proposed steps in the process should be made to confirm or deny the assumptions made hereo 2) Upon concluding bench-scale tests, a pilot plant of a ton-perday size should be designed and operated. Every ore that is to be processed should be tested in this pilot plant since wide variations in characteristics can occur in our western ores. 50

The University of Michigan * Engineering Research Institute The major questions that pilot plant operation should investigate are: 1) required leaching times as a function of particle size, acid strength, temperature, and ore characteristics; 2) uranium distribution ratios and solvent-extraction-stripping stage requirements; 3) optimum selection of uranium product calcining equipment; 4) effects of CaS04 present in the ore upon the evaporator and heat-transfer surfaces in the aqueous waste treating section of the process; 5) corrosion effects of the C- ion upon heat-transfer surfaces; 6) organic solvent losses in the process; 7) HN03 losses in the process; 8) optimum equipment for the waste calcining operation; and 9) evaluation of the assumptions made in this report. It is believed that the questions listed above could be satisfactorily answered in a year's pilot plant operation. 51

The University of Michigan * Engineering Research Institute 9.0 BIBLIOGRAPHY 1o McQuiston, F. W,, Jr., Mining Congress J., 28 (1950). 2. Philippone, Ro Lo, "Operation of the Monticello Mill," Chem. Eng. Prog., 51, 6, p. 261 (1955). 3. Grinstead, R. R., "Recent Developments in the Processing of Uranium Ores and Their Significance in the Extractive Metallurgy of Metals," AECU-3071, Sept. 12, 1955. 4. Shiff, H. H., Hollis, E. T., and Sauer, G. W., "Recovery of Uranium from Vitro Leach Liquors by Ion Exchange. Part I - The Effect of Molybdenum on Uranium Adsorption and Subsequent Cyclic Column Testing of Leach Liquor,"ACCO-355, March 10, 1954. 5o Fisher, S., and McGarvey, Fo, "Recovery of Uranium by Exchange Resins," RMO-2518, March 17, 1953. 6. Saine, Vo L., and Brown, Ko B., "Studies on Recovery Processes for Western Uranium-Bearing Ores," AECD-3241, Oct. 14, 1949. 7. Woody, Ro J., and George, D. R., "Acid Leaching of Uranium Ores," Paper given at Nuclear Engineering and Science Congress at Cleveland, Ohio, Dec. 12-16, 1955, Preprint 329. 8. Moore, R. L., "The Extraction of Uranium in the Tributyl Phosphate Metal Recovery Process," HW-15230, Sept. 1, 1949. 9. Gravquist, D. P., and Merrill, E. T., "Uranium Phase Equilibria in the TBP Process," HW-17747, March 1, 1951. 10. Kapp, N. M., and Weinrich, W. W., "Recovery of Free and Combined Nitric Acid from Metal Nitrate Liquors," U. S. Patent, 2,757,072, July 31, 1956, 11. Burdick and Freed, J. Am. Chem. Soc., 43, 518 (1921). 12. Taylor ITnd. Eng, Chem., 19, 1250 (1927). 13. Taylor, Chilton, and Handforth, Ind. Eng. Chem., 23, 860 (1931)o 14. Chilton, C. H., Chemical Engineering. p. 97, June, 1949. 15. Zimmerman, 0. T., and Larine, I.,' Chemical Engineering Costs, Industrial Research Service, Doror, New Hampshire. 16o "Data and Methods for Cost Estimating," Parts 1, 2, and 3, collection of articles from Chemical Engineering, McGraw-Hill Publishing Co., New York. 17. Nelson, W. L., "Cost-imating," collection of articles from The Oil and Gas Journal, 1946, 18. Carload Waybill Statistics, 1955, Statement SS-6, Interstate Commerce Commission, Bureau of Transport Economics and Statistics. 52