Chapter 5

FLUID BED ROASTING OF Z I N C CONCENTRATE AND PRODUCTION OF SULPHURIC ACID AND PHOSPHATE FERTIWZER

AT CANADIAN ELECTROLYTIC Z I N C , LTD. Val leyf ie ld , Quebec

K. H. Heino, Manager R. T. McAndrew, Production Superintendent

N. E. Ghatas, Technical Superintendent

Canadian E l ec t ro ly t i c Zinc Limited

a n d

B. H. Morrison, Ass i s tan t Di rec tor of Metal lurgical Operations Noranda Mines Limited

Abstract

Two 200 ~ / d a Lurgi Turbulent Layer fluid-bed roa s t e r s have been i n operat ion a t Canadian E l ec t ro ly t i c Zinc Limited s i nce 1966. Zinc concentrate containing 5 2 - M Zn. 9-11s ~e ,31 -3$ S i s t r e a t ed t o produce ca lc ine containing 6 0 - 6 8 ~ n . 1 0 - 1 8 Fe,.l-.3$ S a s sulphide and . 8 - 1 . 3 S a s sulphate. Each roa s t e r t r a i n comprises a waste hea t bo i l e r , two cyclones i n p a r a l l e l and a two-field e l e c t r o s t a t i c pre- c i p i t a t o r . Dis t r ibu t ion of c a l c ine co l lec ted i s 10-15 percent from bed overflow, 55-65 percent from bo i l e r s , 20-25 percent from cyclones and 2-6 percent from prec ip i ta to rs .

A Monsanto designed contact p l an t produces ac id a t a ra ted capa- c i t y of 365 T/da (100% H2S04 ba s i s ) from r o a s t e r gases d i l u t e d t o 7.0-7.5% S02. The process s teps a r e gas cleaning and cooling by weak ac id scrubbing, acid mis t removal i n e l e c t r o s t a t i c p r ec ip i t a t o r s , gas drying i n 93.5-96.55 H2S04, conversion of sulphur d i o d d e t o sulphur trioxi.de i n a f our-pass converter containing vanadium entoxide cata- l y s t , and absorption of t h e sulphur t r i ox ide i n 98.1-9g.3$ HzSOk. Product acid is 9 3 . 3 H2SO4. Conversion e f f ic iency i s 97-98 percent and o v e r a l l sulphur recovery from concentrate i s 37-90 percent.

The adjacent f e r t i l i z e r g l a n t has a capaci ty of 5T/a of P205. The products a r e Y ) percent p osphoric ac id , run-of-pz e t r l p l e super- phosphate, and diammonium phosphate. The p l an t i s of Ugine-Kuhhan design and cons i s t s of an air-cooled mul t ip le tank reac tor , a t i l t i n g pan f i l t e r , and a forced c i r c u l t a t i o n vacuum evaporator. T r i p l e super-

phosphate i s made on a continuous b e l t and p a r t i a l l y cured i n a ro ta ry dryer. Diammonium phosphate i s produced by ammoniating phosphoric ac id i n a s i n g l e s t age reac tor . The s l u r r y of ammonium phosphate i s

granulated i n a blunger and dr ied.

INTRODUCTION

The o r i g i n a l z inc reduct ion p l an t (1 ) b u i l t i n 1963 with a design capaci ty of 180 T/da zinc did not have on-site roas t ing f a c i l i t i e s . The e n t i r e roas t ing f o r t h i s o r i g i n a l p l m t was done by Allied Chemi- c a l Canada, Limited, located about f i v e miles away.

I n l a t e 1966 t h e f i r s t North Americah i n s t a l l a t i o n of Lurgi tu r - bulent l ayer roa s t e r s f o r t r e a t i n g zinc concentrate came on stream a t t he Valleyfield reduct ian p lan t of Canadian E lec t ro ly t i c Zinc Limited (CEZ). The r ~ a s t i n g p l an t was p a r t of a general p lan t ex- pansion which included construct ion of a Monsanto-designed sulphuric acid p l an t , ' a Ugine-Kuhlman designed phosphate f e r t i l i z e r p l an t and expansion of t he leaching and e l e c t r o l y s i s divis ion.

This new roas t e r i n s t a l l a t i o n of CEZ supplies 5040% of t h e re- duction p l a n t ' s present capaci ty of 365 T/d.3 of z inc metal with A l - l i e d Chemical roas t ing the balance.

CEZ i s owned by f i v e Canadian mining companies and managed on - their behalf by Noranda Mines Limited. CEZ a l s o has an i n t e r e s t i n t he adjacent f e r t i l i z e r p l an t which is operated by S t . Lawrence Fer- t i l i z e r s Ltd. (sLF). Val leyfield i s 56 km south-east of Montreal on t h e S t . Lawrence Seaway. This locatiorl provides convenient ac- cess t o world markets, raw mater ia l and labour supply.

Flowsheet

The t h r ee u n i t processes of z inc co:lcentrate f luid-bed roas t ing , sulphuric ac id production and phosphate f e r t i l i z e r production a r e depicted i n Figure 1. Zinc concentrates a r e roasted t o produce cal- c ine s u i t a b l e f o r recovery of z inc by leaching and electrowinning. Sulphuric acid i s produced from t h e roa s t e r gases. Three f e r t i l i z e r products, phosphoric ac id , t r i p l e supsrphosphate and diammonium phos- phate a r e produced from sulphuric ac id , phosphate rock and l i qu id ammonia. Excess sulphuric acid i s shipped t o o ther customers. Table I presents annual production f i gu re s s ince start-up.

ZINC CONCENTRATE ROASTING

Zinc concentrates a r e received i n V d l e y f i e l d by r a i l from the Noranda Mines Limited Geco operat ion i n northern Ontario and from both t h e Mattagami Lake Mines Limited and Orchan Mines Limited opera- t i o n s i n north-western Quebec.

Figure 1 - General Flowsheet

Table I

Production Summary 1966-9

CEZ t r e a t s mostly Geco and M a t t a g 4 concentrates , while All ied Chemical handles mostly Mattagami and Orchan concentrates. Table I1 summarizes tonnages and average analyses f o r concentrate r e ce ip t s and ca l c ine production during 1969.

Mater ia l

Concentrate roas ted - T

Calcine

I

1967 1968 1966

i 16,600

1969

produced - T 1 14,200

I Conversion concentrate t o ca lc ine - % 1 85.5

Sulphuric acid 1 produced (100% b a s i s - T i 14,700

Recovery I

sulphur i n ac id - % I 87.9

I

54% Phosphoric acid produced - T 1,490

T r ip l e super- phosphate produced T : 2,060 24,600 21,500 25,600

1

I I

! I

120,000 114,500 1 135,600 !

105,300 i

99,200 , 116,900

I

I

87.7 86.6 86.2

110,000 102,200 120,000

90.2 88.3 87.7

1 2,800 0 400

Diarnrnoniwn I

phosphate !

produced - T 210 I I i i 1

1

19,700 30,600 1 23,000 I I I I i

EXTRACTIVE METALLURGY OF LEAD AND ZINC

Table I1

1969 Receipt and Production Data

The r o a s t e r p l a n t operat ion includes concentrate unloading, s t o r - age and processing t o produce z inc ca lc ine and a gas s u i t a b l e f o r su lphur ic acid production. Concentrate i s processed i n two pa ra l l e l r oa s t e r t r a i n s each comprising a feed b in , a Lurgi tu rbu len t l ayer r oa s t e r , a forced c i r c u l a t i o n waste hea t b o i l e r , two dus t cyclones and a Lurgi two-field e l e c t r o s t a t i c p r ec ip i t a t o r . The p l a n t was o r i g i n a l l y ra ted f o r 400 T/da concentrate , bu t can handle up t o 435 T/da of feed.

'Receipts and

Production

CEZ r ece ip t s Geco Mattagami Orchan

Total CEZ

Al l ied r ece ip t s Geco Mattagami Orchan

Total All ied

CEZ production All ied production

Tota l production

Leach res idue

Concentrate Handling

Zinc concentrates a r e unloaded from 65 T flat-bottom railway ca r s

Note: Other cons t i tuen ts i n concentrate r e ce ip t s a r e typ ica l ly : .7-1.5% Si02, .4-.6% MgO, .3-.4% CaO, .l-.2% A1203, .1-.3% Mn, .l-.2% Pb, .02--03% Sb, .03-.04% A s , .002-.003% N i , .0003- .003% CO, .007-.01% F, .0005-.01% C 1 , 1.2-2.0 OZ/T Ag

Tons

83,400

Zn

54.0

Cd

.36 48,600 52.4

2,200 52.4 .12 .09

.27

.36

.11

.09

.11

-31 .12

.22

.26

134,200 53.4

1

SO4/S

- - -

- -

- - -

-

1.08 1.67

1.36

1.06 ,

S/S

33.3

Cu

-77

6,200 1 54.9 51,000 1 52.6 67,200 1 52.1

32.5 32.7

33.0

33.3 33.2 33.5

33.4

.32

.32

.32

.93

Analysis %

Fe

10.0 .36 .40

-62

.58

.36

.38

-38

-74 .45

.60

.97

124,400

116,800 106,300 1

223,100

70,000

10.3 10.2

10.1

9.1 10.1 10.8

10.4

11.6 11.8

11.7

37.1

52.5

61.4 59.4

60.5

21.4

by an overhead crane and subsequently e i t h e r t o a 610 mrn conveyor de l iver ing to four 250 T s torage bins and then t o two 80 T roas t e r feed-bins which a r e considered t r l ive" s torage , o r t o a 2,500 T con- cent ra te s torage Say i n t h e unloading shed. The unloading operation requires two-man crews working two s h i f t s per day, f i v e days per week. The concentrate conveying system ra ted capacity is 180 T/hr and is operated from the cen t r a l control roo:n.

The concentrate moisture is typ i ca l ly 4-6 percent. I n t he winter months frozen concentrates a r e thawed i n a 12-car s torage shed f o r 12-48 hours p r i a r t o unloading. The thaw shed temperature is con- t r o l l e d a t 55-60oC with steam heaters . Av~rage steam consumption i n mid-winter i s equivalent t o about 1,440,000 kcal/hr.

Zinc oxide dross contaixing 80-85% Zn and up t o 1% Cl is returned t o t he roas t e r from t h e zinc melting operation a t 9-14 T/da. This material. i s transported by truck t o t he concentrate s torage a rea , blown i n t o a 180 T s torage bin and added t o t h e concentrate being conveyed so the roas t e r feed bins.

Roas t e r s

The L8wgi turbulent layer reac tors a r e t he hear t of the roast ing process. Figure 2 i s a cross-section drawing showing roas t e r de- t a i l s . he bed g ra t e area is 34 m2 and contains 3,300 'cuyeres of 6.35 inm i.d. mounted f l u sh with t h e g r a t e top.

F lu id iza t ion a i r f o r both roas te rs i s furnished by three 187$kW Howden P,zirsons cent r i fuga l bl3wers of which o?le i s a standby spare. Usual a i r flow r a t e s per roas t e r a r e 15,700-17,300 ~m3/hr with up t o 170 mrn Hg discharge pressure aild a r e automatically compens3ted f o r ambient temperature f luctuat ions. The a i r t o concentrate weight , r a t i o used a t CEZ is about 1.43. However, there i s an empirical absolute mhimum of a i r volume required .to maintain adequate bed f l u id i za t ion and acceptable temperatures i n t h e bed i r r e spec t ive of t he amount of feed. A small amount of bed material co l l ec t s i n the wind boxes and i s removed twice per s h i f t .

Roaster feed i s drawn continuously from each feed bin on 3 vari- ab le speed feed b e l t which discharges t o a weigh b e l t and then t o a high speed ns l ingeru b e l t feeder. The b e l t feeder discharges t o t he bed cent re through a s ide po r t about 1.7 m above the grate . Accurate feed cont ro l i s ma.intained from the control roorn.

The be'd and suspension temperatures a r e important f ac to r s i n determining the ca lc ine qual i ty . In general, operating over 950°C w i l l provide lower S/S with acceptable SOq/S. The bed temperatare is affected by t h e qua l i t y and quanti ty of t h e concentrates fed t o t h e roas t e r and bed cooling.

EXTMCTIVE METALLURGY OF LEAD AND ZINC

A-Gas Outlet G - Wind Box Discharge

6 -Oil Burner Nozzle H - Air Inlet

C - Bed Overflow Discharge I - Bed Coils

D - Underflow Discharge J - Slinger Belt

E - Bed Grate K - Charging Port

L - Safety Valve 0 0.5 lm F -Wind Box MJ Scrnl.

Figure 2 - Roaster Cross Section

The temperature of t h e suspension zone is maintained a t about 9700C and i n l i n e with t h e bed temperature. This i s achieved by spraying water over t he concentrate on t h e feed be l t . This is ex- plained through +he f a c t t h a t t he water ac t s as binding medium t o t h e f i n e p a r t i c l e s of t h e concentrates and tends t o p a r t i a l l y agglomerate them, thus confining them f o r a longer period i n t h e bed, where t he sulphur burns of f r a the r than i n t he suspension zone. The amount of water required f o r suspension zone temperature cont ro l i s a l so a funct ion of t h e concentrate p a r t i c l e s ize. The l imi t ing f a c t o r on t h e amo'mts of water sprayed o r in jec ted i s predetermined by optimum operating temperatures .

While t h e amount of i n j ec t ion waker t o control t h e bed tempera- t u r e may be ad ju i ted according t o t h e amount of concentrate fed, t h e cooling r a t e pe r bed c o i l i s constant as it i s t i e d i n with the for- ced c i r cu l a t i on system of t h e respect ive waste heat b o i l e r and, thus, has a cons-cant amount of water flowing through a t a l l times because of the c r i t i c a l water ve loc i t i e s Ynat have t o be maintained i n order t o prevent c o i l f a i l u r e . Depending on t h e forecas t throughput of t h e roas t e r s , t he rider of bed c o i l s is determined (3s it involves major in te r rupt ion t o t he operation t o i n s t a l l 3r dismantle a bed co i l . CEZ roas t e r s operate normally with f i v e cooling c o i l s per roaster .

The equ i l i b rhm of t h e calcin: discharging -Erom the roas te rs i s mainly determined by the r a t e of agglomeration i n t h e bed. This agglomeration depends on the p a r t i c l e s i ze , chemical analysis , mois- t u r e content of the feed and a i r flow through the bed.

When shut t ing down a roas t e r , a i r f l u id i za t ion is continued f o r 15-20 minutes t o completely oxidize res idua l sulphide a f t e r the feed i s stopped and thereby ensure an un f r i t t ed bed f o r t he next start-up. The roas t e r can be ign i ted d i r e c t l y w i t h concentrate, i f t h e bed temperature is above 7r)0°c. Between 580 and 700°c t he roas t e r should be s t a r t e d w i t h elemental sulphur. Preheating with o i l burners i s requirsd below 5800C.

'The majori ty of t he ca lc ine is car r ied i n t o t h e waste hea t boi l- e r s with t h e roas t e r gases . Dis t r ibu t ion of ca lc ine col lected is 10-15 percent from bed overflow, 55-55 percent from bo i l e r s , 20-25 percent from cyclones and 2-5 percent from prec ip i ta tors . Although an underflow o u t l e t was o r ig ina l ly provided, t h e need f o r t h i s ha:; not been required and normal 3verflow control seems t o be suZf i c i en t . The depth of t h e nomal f lu id ized bed i s about 1009 mm. The bed pressure d i f f e r e n t i a l i s typ i ca l ly 935-1215 nun Hg.

Advantage was taken of a radioact ive t r a c e r t e s t a t t he 3rchan concentrator(2) t o attempt t o determine bed residence time and other re levant time f ac to r s i n t he roas te rs . Bed residence times were measured using about 60 T of radioact ive zinc concentrate. A t a concentrate feed r a t e of 10 T/hr t h e ce s t r e s u l t s indicated an aver- age residence time f o r t h e roas t e r bed overflow material of about

five hours and an average bed residence time for the dust carry-over of about one hour(3). Complete renewal of the bed occurred in about 20 hours. Further tests are required to confirm these results.

Table 111 summarizes typical sulphur analyses and particle size data for roaster feed and the calcine fractions.

Table 111

Roaster Products

:Precipitator dust j .1 15.5 1 1 2 1 9 8

(Total calcine : -2 j 1.1 1 11 j 39 ) 61

-

Material

Concentrate

The calcine product should be low in sulphide sulphur to maximize metal leach extractions of zinc and cadmium and low in sulphate sul- phur to minimize sulphate accumulation in the leaching and electro- lyzing solutions. Incorporation of a water-cooled settling chamber as an inlet section of each waste heat boiler is a feature which contributes to producing calcine containing only .8-1.2% SOq/S.

Screen size distribution %

Rapid separation of hot calcine from the roaster gases as quickly as possible is desirable, because the sulphate sulphur increases the further from the roaster the product is collected. CEZ was the first to use the water-cooled settling chamber concept which has gained wide acceptance in subsequent zinc concentrate roaster installation. Lower gas temperatures, finer particle size and longer contact times favour the sulphation reaction. Air infiltration anywhere in the roaster system is akso a major contributing cause to sulphation. Provision to collect and recycle high sulphate precipitator dust through the roaster was made in the original plant but no significant decrease in overall sulphate was achieved and the practice was dis- continued.

Bed Overflow

W.H. boiler dust

!Cyclone dust

Analyses %

Attemptsto control total sulphate in the plant system by the use of spent electrolyte for roaster temperature controlwereunsuccess- ful because this led to increased sulphate content in the calcine and a tendency for the bed to agglomorate excessively. Similarly,

S/S

32.5

-.075 mm

63

+.15nun

11

SOq/S

-

.3

.2

-1

+.075 nun

3 7

2 2

7 7

97

.2 28

.8 1 3 2.5 -

I

78

23

3

the use of weak acid from t h e acid p l an t humidifying tower f o r tem- perature control had t o be discontinued because of excessive dust build-up i n t he bo i l e r s and gas cleaning systems.

Waste Heat Boilers

Each roas t e r has a separate waste hea t recovery system rated a t 13,500 kg/hr of steam production. The steam pressure is reduced from 39.1 atm t o 8.8 atm f o r CEZ and SLF plant use. A waste hea t bo i l e r cons is t s of a water-cooled s e t t l i n g chamber and four bo i l e r tube bundles suspended i n t h e gas stream a f t e r t he s e t t l i n g chamber a s shown i n F i p r e 3. The s e t t l i n g chamber water cooled tubes a r e s e t i n t he re f rac tory walls and roof. There a r e 120 tubes located on 45 mrn centers.

Water a t 250'~ from the steam drurn r ec i r cu l a t e s i n p a r a l l e l th- rough each tube bundle, s e t t l i n g chamber cooling tube and roas t e r bed c o i l a t 186,750 kg/hr. Make-up water i s preheated i n a de-aera- t o r with process steam. Flow r a t e s a r e f ixed by o r i f i c e p l a t e s in each tube bundle. Well water i s deionized i n s trong base type ion exchange un i t s before being used f o r bo i l e r feed water make-up.

A a l ique f ea tu re or' forced c i r cu l a t i on type bo i l e r s i s t h a t t he steam drum is located ex terna l ly from the b o i l e r proper instead of being p a r t of t h e t o t a l heating surface as i n na tura l c i r cu l a t i on bol le rs . Forced c i r cu l a t i on bo i l e r s , although i n i t i a l l y more ex- pensive than na tura l c i r cu l a t i on bo i l e r s , a r e more e f f i c i e n t because of b e t t e r use of surface area and l e s s water i n t h e c i r c u i t , s a f e r i n case of tube leaks and can be integrated with bed c o i l s f o r roas- t e r temperature control. Gas temperatures decrease about 6104300C through t h e bo i l e r uni ts . The steam t o concentrate r a t i o averages 1.05-1.15.

Boilers must be kept clean of dus t build-up t o maintain maximum roas t ing r a t e s and e f f i c i e n t waste hea t recovery. Each tube bundle is equipped ,with a mechanized rapping device f o r removing loose dus t pa r t i c l e s . The rappers operate automatically on a sequential three- minute cycle. I n addit ion, lancing of t he b i l e r s is done manually each s h i f t with 6 atm a i r and requires 3-4 manhours per boi ler .

I n case of power f a i l u r e , a d i e se l 210 kW generator automatically provides power f o r p l an t l igh t ing and the waste heat bo i l e r feed- water and c i r cu l a t i ng pumps.

Gas Cleaning System

Exhaust gases from each waste hea t bo i l e r divide between two p a r a l l e l 8.2 m x 2.5 m diameter dust cyclones. Blockage of t he cyclones was an i n i t i a l problem which was corrected by removal of

- A- Gas lnlet

B - Settllng Chamber

C- Water Wall

D - Steam Drum

E - Tube Bundles

F - Mechanosal Shaker Location

G- Gas Outlet

H - Hopper Partitsons Cast Iron

I - Discharge Hoppers

J - Drag Conveyor

Refractory

- A i r Lancing Doors

0 0 5 Irn -- I-.,.

I - . -

Figure 3 - Waste Heat Boiler Cross Section

I A- Gas Inlet

B - Dstr~but~on Plats ' C - Rapper Assembly

D - Inlet Field I

E -Outlet Field

F -Gas Outlet 1 G - Dlrchargs Hopper

H - Orag Conveyor

I

Figure 4 - Electrostat ic Precipitator Cross Section

t he vortex breakers. Some continuing problems i n dust build-up i n t he ducts between cyclone o u t l e t s and c o t t r e l l i n l e t s occur periodi- c a l l y and clean-out is required severa l times per year.

Gases from the two cyclones combine before enter ing a Lurgi two- f i e l d e l e c t r o s t a t i c p rec ip i t a to r shown i n Figure 4. Each f i e l d sec- t i o n contains 16 frames of co l lec t ing p l a t e electrodes and 15 frames of discharge wire electrodes. Each discharge frame supports 24 wires. The frames a r e v e r t i c a l l y suspended a l t e rna t e ly i n t he d i rec t ion of gas flow. A perforated d i s t r i bu t ion p l a t e a t t he i n l e t of each pre- c i p i t a t o r i s i n s t a l l e d i n an e f f o r t t o maiatain uniform flow through the uni t . Excessive dus t build-up a t t he d i s t r i bu t ion p l a t e has been a problem. Each s e t of frames has a separate automatic sequenced mechanical rapping system. Gases e x i t t he p rec ip i t a to r s a t 320-340°c and contain about 500-1000 mg/~m3 of dus t , although the u n i t s - a r e designed f o r about 180 mg/~m3.

The most troublesome operating problem i n t he roas t e r p l an t is dust build-up i n t h e e l e c t r o s t a t i c prec ip i ta tors . Each p rec ip i t a to r must be cleaned every 8-12 weeks because n d u l a r build-up on t h e discharge wire electrodes i n t he i n l e t sec t ions causes shor t c i r cu i t - ing between wires and p la tes . During p rec ip i t a to r cleaning, gases from both roas t e r s a r e directed through the second p rec ip i t a to r , but dust losses increase even a t reduced roas t ing ra tes . The dust build- up on t h e wires i s removed manually with fork-ended rods. Total cleaning time is 30-36 hou-s using six-man crews with only t h ree men ab l e t o work i n a p rec ip i t a to r a t one time. Repeated cleanings have decreased the wire tension n d thereby decreased the effect iveness of t he rapping system. A i r i n f i l t r a t i o n i s a contr ibut ing cause of t he build-up and scrupulous care must be taken t o prevent leakage pa r t i cu l a r ly around the p rec ip i t a to r i n l e t dampers.

Single phase 550 V power i s fed t o t h e p rec ip i t a to r s through two Westinghouse ac-dc power un i t s , each combining a 45 kVA step-up transformer with full-wave s i l i c o n diode r e c t i f i e r rated f o r an average dc output of 700 mA a t 45 kV. The i n l e t f i e l d sect ions of both p rec ip i t a to r s a r e connected t o one r e c t i f i e r while both t he out- l e t f i e l d sec t ions a r e connected t o t he o ther r e c t i f i e r . This i s a d e f i n i t e disadvantage as each p rec ip i t a to r sec t ion should have sepa- r a t e power packs f o r maintaining maximum operation. Typical opera- t i n g e l e c t r i c a l outputs a r e 100 mA dc a t 300-400 V ac and 10 A ac. Appropriate control c i r c u i t s allow f o r e i t h e r automatic o r manual operation.

Since so much d i f f i c u l t y has been experienced with the dust build- up on the wires and the subsequent high dust losses it has been de- cided t o make changes t o t he ex is t ing gas cleaning equipment. Sepa- r a t e power packs w i l l be provided f o r each p rec ip i t a to r sec t ion so t h a t maximum col lec t ing e f f ic iency can be maintained even though one sectio.7 may be faul ty. Design and ma-cerials of t he rapping system a r e t o be improved upon a s these proved t o be inadequate. The cross-

over duct , a s shown i n Figure 5, w i l l be eliminated t o prevent any a i r i n f i l t r a t i o n poss ib i l i ty . The cyclone e f f ic iency w i l l be increased from about 80-85 t o 95 percent i n order t o reduce t h e dust loading t o t h e prec ip i ta tors . The e n t i r e discharge wires and frames w i l l be re- placed i n t he summer of 1970 because frequent cleaning has loosened t h e wires and thereby reduced t he rapping effect iveness . Design does not permit individual replacement of wires. Minimizing t he dust build-up on t he wires w i l l decrease t he required cleaning frequency and thereby maintain adequate wire tension f o r e f f ec t i ve rapping.

I f t he above measures f a i l , i t may become necessary t o add more p r ec ip i t a to r capaci ty with a t h i r d un i t o r more sec t ions t o t he pre- sen t un i t s .

1 A-waste "eat Boilers E - cross-Over ~ u c t

- Cyclones - 5 6 - F - Start-up Fan ,

D -Prec~p~tators

I

I

0 05 lm MJ Scale

Figure 5 - Plan View of Roaster Gas Ducting Arrangement

Calcine Handling

Calcine co l lec t ion and t r a ~ l s f e r i n t he roas te r p lan t i s by a ser- i e s of 10 enclosed Buehler drag-type chain conveyors. E i ther of two main co l lec t ing conveyors t r ans fe r s t he various ca lc ine products t o a 1525 x 915 mm diameter b a l l m i l l . There a r e water cooling jackets on t h e bed overflow, bo i l e r discharge and main co l lec t ing conveyors. Usual ca lc ine temperature a t t he b a l l m i l l i s about 200°C.

The b a l l mi l l discharge is conveyed t o an 80 T surge bin before being pneumatically pumped 183 m t o three of f i v e 775 T storage silo:; a t the leaching p lan t , using two Fuller-Kinyon dust pumps.

Excessive wear on t h e calcine pump screws has been a major main- tenance problem. The b a l l m i l l was designed only t o break r e l a t i v e l y s o f t ca lc ine agglomerations present mainly i n t he bo i l e r calcine product and therefore does not have s u f f i c i e n t capacity t o adequately grind t h e coarse bed over-flow material. To help a l l ev i a t e pump wear and produce a f i n e r ca lc ine f o r leaching, the load on the ex is t ing b a l l mi l l w i l l be reduced by i n s t a l l i n g an a i r c l a s s i f i e r i n c l ~ s e d c i r c u i t with the m i l l . Cyclsne and p rec ip i t a to r dust w i l l be col- lected i n one of the main co l lec t ing co:lveyors and fed d i r e c t l y t o the c l a s s i f i e r with the b a l l m i l l discharge. Bed overflow material and the plus 1.5 mrn f r ac t i on from the bo i l e r discharge w i l l go t o t he b a l l m i l l and a l l o ther products t o t h e c l a s s i f i e r . The system is scheduled f o r i n s t a l l a t i o n i n 1970. A calcine product of 98 per- cent minus .15 mm and 90 percent minus .075 mm is expected.

General

An operator i n a cen t r a l coiltrol room monitors and regulates the complete roas t e r p lan t operation. Indicat ing, recording o r control instruments a r e provided f o r temperature, pressures, flow r a t e s , sulphur dioxide ana lys is , and p rec ip i t a to r e l e c t r i c a l data. The acid p lan t operation is a l so monitored from t h e cont ro l room.

A superintendent ,assis ted by a general foreman, supervises both t he roas t e r and acid p lan ts as a s ing l e department. A foreman, two operators and two helpers comprise a normal s h i f t crew. Concmtrate unloading normally rec~ui res two men f o r ten s h i f t s per week. Two department mechanics handle regillar maintenance requirements. Major r epa i r s and modi f ica~ions requi re t he Plant Engineering Department's assis tance. P lan t t e s t i n g , instrumentation maintenance and labora- to ry serv ices a r e provided by the Metallurgical Department.

Table IV suimnarizes the important roas t e r p lan t s t a t i s t i c s f o r a typical. operating month.

Table IV

Roaster Plant Operating Statistics - November 1969

Moisture in concentrate feed

Average water injection

I I I 1 Quantity

St~am production

Steam to concentrate ratio

Operating time

Concentrate receipts

Concentrate to roasters

Average temperature: I

I

hr 1 1,421 % t 98.7

T 1 13,248

T 1 13,596

1 bed OC 950

Dross to roasters I T i 303

Calcine production I I

T 1 11,969 I

Calcine S/S 1 C I /O I .30 I

Calcine S04/S % ' .98

Moisture in concentrate receipts O/ 10 I 4.1

1 boiler inlet OC ' 965

( precipitator outlet OC I 330

I oc ball mill outlet 201

(~verage pressure : I

bed differential

roaster draft

boiler differential

1 cyclone differential I i nun H g 1 3.10 I

) precipitator outlet draft I m m H g 1 8.33

Power consumption I kwh Operating labour I hr Maintenance labour 1 ilr

714,000 3,676 1,234

SULPHURIC ACID PRODUCTION

Sulphuric acid is produced from t h e roas t e r plant exhaust gases i n a Monsanto-designed contact plant . The general pr inc ip les and p rac t i ce of producing sulphuric acid from sulphide concentrate roas- t e r gases were presented by Monsanto a t t h e 1967 AIME Annual Meet- ing(4) . The CEZ operation i s a t yp i ca l example of the process. The o r ig ina l ra ted plant capacity was 365 T/da (100% H2S04 bas is ) but t he p lan t can produce a t 19-15 percent overload.

Gas Pur i f ica t ion

The roas t e r p lan t e x i t gas is typ i ca l ly 320-340°C and contains 10-11% S02, 5 6 % 02 and 9-11% H20. Gas flows from the two dust pre- c i p i t a t o r s through a common duct t o t h e humidifying tower where re- c i rcu la ted weak ac id scrubs t h e gas t o remove the dust p a r t i c l e s , sulphur t r i ox ide and o ther gaseous impurities. The temperature of the gases leaving t h e humidifying tower is 65-800C. Weak acid i s a l so pumped through Karbate tube and s h e l l heat exchangers i n paral- l e l to a packed spray cooling tower where the gases from t h e humidi- fying tower a r e fu r the r cooled t o 38-45W.

Weak acid volume increases because of moisture condensatiorl and is discarded regular ly. Fresh water may a l s o be added t o t he weak acid pump tank t o cont ro l t he acid and impurity concentrations. Typical weak acid analyses a r e 3 6 % H2S04, 2-8 g/dm3 C1, .3-.5 g/dm3 F, and 15-40 g/dm3 Zn. The weak ac id discard is neutral ized with l h e s t o n e and pumped t o t h e SLF gypsum residue pond. Figure 6 i s the weak acid c i r cu l a t i on flowsheet. The l i f e of t h e antimonial lead spray nozzles i n t h e humidifying to:.jer i s l imited t o 6-12 months by the abrasion of suspended so l id s in t h e weak acid. Carpenter 20 a l l oy t e s t nozzles had shor te r l i fe t imes because of excessive corro- sion. The p a r t i a l l y pur i f ied gases from t h e cooling tower pass th- rough two p a r a l l e l Koppers e l e c t r o s t a t i c p rec ip i t a to r s t o remove acid mist and rem2ining t races of dus t and fume. Each p rec ip i t a to r chamber cons is t s of a group of 122 v e r t i c a l 254 mm i.d. lead tube co l l ec t i ng electrodes each concentric with an i ron core lead wire discharge electrode. The mist laden gas en ters a t the bottom and t h e mist co l l ec t s on the tubes and dra ins t o a d r i p tank. The d r ip acid is pumped t o t h e weak acid pump tank. Power f o r t h e prec ip i ta - t o r s i s provided with two Westinghouse ac-dc power uni t s each com- bining a 30 kVA step-up transformer with a full-wave s i l i c o n diode r e c t i f i e r . Typical operating e l e c t r i c a l outputs a r e 100 mA dc a t 400-450 V ac and 30-35 A ac. The p rec ip i t a to r s operate a t over 99 percent eff iciency and o u t l e t mist loadings a r e l e s s than 70 mg/Nm3. Daily ha l f hour washing periods remove dust and fume residue from t h e e1,sctrodes and t h e washings a r e diverted t o a s e t t l i n g pond.

The gas is dried i n a drying tower where moisture is removed t o 40-50 mg/Nm3 a s t h e gas passes upwards through the tower counter-

c u r r e n t t o t h e drying acid. Acid concentrat ion i s usual ly maintained a t 95% H2SO4 b u t can vary between 93.5 and 96.5 percent. The ac ld temperature is maintained a t a minimum of 35%. The acid i s c i r c u l a - t e d over t h e towers a t 270 m3/hr from a 60 m3 pump tank.

Di lu t ion a i r is added t o the gas stream p r i o r t o t h e drying tower t o a d j u s t t h e sulphur dioxide concentrat ion t o an optimum 7.0-7.2 per- cen t f o r subsequent oxidat ion. The E l l i o t t blower is i n s t a l l e d a f t e r t h e drying tower and has a r a t e d capac i ty of 72,000 m3/hr a t 37.4 mm Hg i n l e t s u c t i o n and 159 mm Hg o u t l e t presswe. A 735 kW General E l e c t r i c 2,300 V induct ion direct-connected motor dr2:zz t h e f a n a t 3,580 rpm.

D r ~ p Tank

Weak Acid,

umertone Pump Tank

Flush Tank

Figure 6 - Weak Acid Ci rcu la t ion

Conversion and Absorption

The clean dry gas-from t h e drying tower passes t o a four-stage converter containing vanadium pentoxide c a t a l y s t t o acce lera te t he oxidation of sulphur dioxide t o sulphur t r ioxide. The overa l l con- version e f f ic iency averages 97.0-97.5 percent. The hea t generated by the exothermic reac t ion in each s tage or pass i s used t o preheat t h e sulphur dioxide gas passing t o t he converter through a s e r i e s of t h r ee v e r t i c a l tube and s h e l l heat exchangers. The temperature r i s e i s a measure of t h e conversion done i n each converter stage. The r e l a t i v e amount of t h e overa l l reac t ion achieved i n each successi.ve pass i s normally about 65, 25, 8 and.2 percent respectively.

I n l e t gas temperatures t o the f i r s t s tage must be 4150C minimum t o ensure i n i t i a t i o n of t he oxidat ion reaction. Outlet gas from t h e f i r s t pass i s typ i ca l ly 580600oC and i s cooled i n t h e "hot" heat exchanger t o 4.20-4500C before enter ing t h e second converter stage. Temperatures over 600°c may damage the ca t a ly s t . Outlet gas from the second pass i s typ i ca l ly 500-5200C and i s cooled i n t he " inter- mediatew heat exchanger to 415-4250C before enter ing the t h i rd con- ve r t e r stage. Outlet gas from t h e third-pass i s typica l ly 435-4500C and i s cooled i n an atmospheric cooling duct t o 415-420°C before enter ing the four th stage. The temperature r i s e during t h e four th pass r a r e ly exceeds 50C. Outlet gas from t h e four th s t a g e - i s cooled i n t he llcoldll hea t exchanger t o 240-250T before passing t o t he cab- sorbing tower. The sulphur t r i ox ide gases from the converter pass through the tube s ide of t he t h r ee hea t exchangers and a r e cooled oil the s h e l l s i de by sulphur dioxide gases from the drying tower. Each hea t exchanger i s provided with two by-pass ducts with control valves f o r cont ro l l ing t h e converter bed temperatures.

A preheater i s a l so provided t o heat the converter c a t a l y s t t o t he 4150C minimum ign i t i on temperature a f t e r an extended shutdown. The preheater cons is t s of a br ick o i l - f i red combustion chamber and a tube and s h e l l heat exchanger. The hot cornbustiorl gases pass t i - rough the tubes which hea t dry a i r from the acid p lan t main blower. Depending on t h e ambient temperature, shutdowns of l e s s than 4-6 hours do not usual ly necess i ta te preheating.

Sulphur t r i ox ide contained i n t h e converter e x i t gases a f t e r cool- ing i n t he cold hea t exchanger i s absorbed by approximately 98 per- cent acid flowing counter-currently through the packed absorbing tower. The exhaust gases from the absorbing tower contain .15-.25% SO2 and pass tlhrough a recen-ily i n s t a l l e d demister t o t h e 64 m x 1.4 m i.d. mild s t e e l stack.

A s tack plume was a chronic problem p r io r t o t he demister i n s t a l - l a t i o n mainly because of acid mist formation i n t he absorbing tower. The plume was minimized but not eliminated by coiltrolling t h e ahsor- bing acid a t 98.2-98.3% H2S04 and 85-900C and by maximizing the acid circula. t ion r a t e . These condikions produced an acid m i s t of 130-350

mg/~rn3 with 10 percent minus 3 microns. The P lan t Engineering De- partment designed and constructed a demister using York demister pads which reduced t h e ac id mist t o 30-50 mg/~m3 with 40 percent minus 3 microns and eliminated t h e plume. Figure 7 shows some of the demister d e t a i l s . Two -15 m x 2.4 rn 0.d. wire mesh demister pads of Carpenter 20 a l l o y a r e mounted i n an acid r e s i s t a n t b r i c k l ined s h e l l mounted on top of ths absorbing tower. Care was taken t o i n s u l a t e a l l i n t e r n a l m e t a l l i c components t o avoid galvanic cor- rosion. A by-pass duct i s provided so t h a t any required maintenance can be c a r r i e d ou t on t h e demister without incur r ing excessive down time.

A - G a s Outlet

B - Low Density Pad

C- ~ i ~ h Density Pad

D - Support Frame

E - S t e e l Shell

F - Acid Proof Brick

G- Insulating Sleeve

H - Drain Pipe

I - Gas Inlet

o 2 5 _ _ Jm

Scale - -

Figure 7 - Demister Cross Sect ion

Acid Ci rcu la t ion and Storage

The concentration of absorbing acid i s corltrolled by d i r ~ t i o n with water and a cross-flow of 93 percent drying acid pumped from the drying ac id p u q tank. The concentrat ion of drying acid i s con- t r o l l e d by f o r t i f i c a t i o n with 98 percent absorbing acid pumped from the ,&sorbing acid pump-tank. Drying acid is pumped continuously through a cross-flow s t r ipp ing tower t o the product acid pump-tank. Additional water may be added d i r e c t l y t o tkLe product acid pump-tank t o maintain t he product concentrat ion a t Ikhe desired 93.2 percent. Product ac id i s pum~ed t o four 2,790 T s torage tanks. A l l t r a n s f e r s from the t h r ee pump-tanks a r e continuous and automatically control led a t such flow r a t e s as t o maintain constant pre-set l eve l s i n t he tanks. Product acid and r ec i r cu l a t i ng ac id t o each of t he absorbing and drying towers .are pumped through separa te c a s t i r o n pipe sec t ions of a water-spray cooli.?g tower f o r temperature control . Figure 8 is the concentrated acid flowsheet.

Dilut ion a i r required t o cont ro l t he sul2hur dioxide coilcentra- t i o n i? the gas fed t o t h e converter i s added p r i o r t o t he drying tower through the cross-flow s t r i pp ing tower. Dissolved sulphur dioxide i s t h ~ ~ s removed f ro~n the product acld and t he ove ra l l p l an t acid y i e ld i s increased. Product acid contains l e s s than .008% S02.

Ex11 Gases To Atmosphere

f To Converter

4 rnbfi Cooler

I . . ' I

Gas From Mist

I Drying Tower 4 :

.

93% A c ~ d Pump Tank

a Cross Flow 7 1

I I Cooler J===kJ

Product Acid Pump Tank

Cooler I I

98% Acid Pump Tank

To S L F .lT= : * Storage Tank

Figure 8 - Concentrated Acid Ci rcu la t ion

About 65 percent of t he acid production i s piped da i l y t o SLF f o r f e r t i l i z e r production. The remaining acid production I s shipped t o okher customers by r a i l c a r o r tank-truck. A 250 m l i n e a l so allows d i r e c t loading of tanker vessels a t the Por t of Val leyfield.

General

Process tanks and towers 3re general ly constructed of mild s t e e l welded p l a t e and li-led with ac id- res i s tan t brick. A lead membrane i s a l so used i~ t h e humidifying, cooling, drying and cross flow s t r i p - ping towers, but not i n t he absorbing tower. A l l towers, except t he humidifying tower, a r e packed with U.S. Stoneware I n t a l ~ x saddles. The mild s t e e l product s torage tanks a r e not l ined. In t he weak acid sec t ion of Lke process f i be rg l a s s reinforced polyester (FRP) is used f o r t he pump tank, gas duct work between the humidifying and drying towers, and f o r process piping over 50 nun diameter. Polyvinyl ch- l o r i de (PVC) is used f o r smaller diameter piping. Impervious graph- i t e arbat ate) i s used f o r t h e weak acid heat exchangers and circula- thg pumps. The m i s t p r ec ip i t a to r s a r e fabr ica ted of lead while t h e dra in pu-lops a r e polypropylene. Concentrated acid c i r cu l a t e s through c a s t i r on piping. The hot gas heat exchangers and ducts a r e general- l y of n i l d s t e e l but port ions of t h e system have been metallized with a -25 mm coating of aluminum-silicon al loy. F i r e br ick l in ing is used i n t h e f i r s t two converter passes and a l so f o r t h e preheater combu:;tion chamber.

The vanadium ,zatalyst beds become packed through prolonged use and must be per iod ica l ly screened t o prevent excessive pressure drop through t h e converter. Monthly pressure surveys a r e taken as a guide t o t he general condition of t he p lan t and pa r t i cu l a r l y f o r t he con- ve r t e r passes. Such a survey i s included i n Table V. I n t h i s exam- p l e , t h e pressure drop across t he fourth pass i s 2-3 times grea te r than f o r t h e o ther passes and a screening is therefore indicated a t t he next annual shutdown. The fourth pass has not been screened s ince t h e p l m t commenced operation. Experience ind ica tes t h a t screening i s required ,annually f o r t he f i r s t pass , every th ree years f o r t he second and t h i r d passes and every fo-LC years f o r t he four th pass.

The acid p lan t operation i s monitored with the roa s t e r p l an t from the cen t r a l control room by one operator from the combined four-man crew. A second operator makes cont ro l checks and adjustments i n t he f i e l d . He a l s o handles water treatment f o r the waste heat bo i l e r s i n the roa s t e r plant . Instruments ind ica te and record temperatures, tank l eve l s , acid concentrations, d i l u t i on water flow r a t e s , sulphur dioxide concentrations, p r ec ip i t a to r f i e l d s t rengths and cooling water pH. Pump tank t r ans f e r s a r e automatically control led t o main- t a i n pre-set tank leve ls .

FLUID BED ROASTING, ACID, AND FERTILIZER AT VALLEY FIELD

Table V

Typical Gas Pressure and Temperature Survey - October 1969

(Production rate - 390 T/da, 100% H2S04 basis)

Table VI summarizes the acid plant operating statistics for a typical month.

Humidifying tower

Cooling tower

Mist precipitators

Drying tower

Main blower

Cold exchanger - shell side

Intermediate exchanger

shell side

Hot exchanger - shell side

Converter first pass

Hot exchanger - tube side

Converter second pass

Intermediate exchanger

tube side

Converter third pass

Converter fourth pass

Cold exchanger - tube side

Absorbing tower

Demister

In

-12.0

-12.7

-19.1

-21.9

-27.1

115.0

114.6

114.1

95.4

85.1

75.7

66.4

59.8

39.3

15.9

6.5

3.4

In

317

63

39

- -

-

-

- 443

- 449

- 426

423

- 221

-

Pressure,

Out

-12.7

-19.1

-21.5

-26.4

115.0

114.6

114.1

95.4

85.1

75.7

66.4

59.8

51.4

15.9

7.5

3.4

0.4

Temperature

Out

63

39

38

- -

-

-

- 590

- 507

- 445

427

239

- -

mh Hg Dif - ference

0.7

6.4

2.4

4.5

142.1

0.4

0.6

18.7

10.3

9.3

9.3

4.7

8.4

23.4

8.4

3.2

3.0

OC

Dif- ference

254

24

1

- -

-

-

- 147

- 58

- 19

4

- - -

Table V I

Acid Plant Operating S t a t i s t i c s - November 1969

Operating time I Concentrate t o roas t e r s

Concentrate S

Acid production (100% H2S04)

S conversion t o H2SO4

Weak acid discarded (100% H2S04)

Converter i n l e t SO2

E x i t gas SO2

SO2 Conversion t o SO3

Power consumption - plant

Power consumption - blower

Operating labour

Maintenance labour

Unit

h r

%

T

%

T

%

T

%

%

%

kwh

kwh

h r

h r

Quantity

PHOSPHATE FERTILIZER PRODUCTION

When the roaster-acid plant i n s t a l l a t i on a t CEZ was being consi- dered,one of the major concerns was the disposal of the by-product sulphuric acid. The loca l market could not absorb 126,000 tons of new acid and the value of sulphuric acid is such tha t it could not be shipped economically by r a i l more than about 800 ?a. Although low-cost water t ransport through the St. Lawrence Seaway i s avail- able f o r s i x months of the year,the construction of tanks f o r 60- 70,000 tons of winter s torage a t the acid plant and a t the plant of a potent ia l buyer would be costly.

High-analysis phosphate f e r t i l i z e r s such as diammonium phosphate (1846-6) and t r i p l e superphosphate (0-46-0) consume about 2.8 tons of sulphuric acid per ton of P205 converted t o phosphoric acid. Thus, subs tant ia l amounts of sulphuric acid can be consumed. However, t h e cos t of t he raw materials , phosphate rock and ammonia,represent i n excess of 60 percent of the t o t a l manufacturing cost.

There appeared t o e x i s t a s u f f i c i e n t market f o r ammonium phos- phate and phosphate f e r t i l i z e r within an economic radius of Valley- f i e l d t o j u s t i f y t h e i r production. However, t h e present sadly de- pressed pr ices f o r f e r t i l i z e r products i n North America make the venture much l e s s a t t r a c t i v e now than it seemed i n 1965.

Although the p lan t i s located adjacent t o a dock on t he Seaway, the de l ivery of rock by r a i l r a the r than vessel i s l e s s cos t l y under cur ren t condit ions as it avoids t he l a rge c a p i t a l expenditure requi- red t o s t o r e 60,000 T of rock during t he winter when the Seaway i s closed and t h e financing cos t s of a l a rge rock inventory.

Raw Materials

Phosphate Rock: A t present , t he p lan t i s using calcined North Caro- l i n a rock a s a source of phosphate. The spec i f ica t ions f o r t h i s mater ial a r e shown i n Table V I I . The bulk dens i ty i s 1600-1760 kg/m3. This rock d i f f e r s from Florida rock i n t h a t it contains f r e e lime and l i t t l e organic matter as a r e s u l t of i t s ca lc ina t ion ; consequent- l y , subs t an t i a l l y more hea t i s evolved when the rock i s attacked with sulphuric acid but no foam i s produced i n t h e reactors . The i r o n and alumina content i s lower than most North American rocks of t he same grade and t h i s , with t he absence of organic matter, allows t h e production of a good qua l i t y of 54% P2O5 acid.

I n common with most calcined rock a small amount of hydrogen sul- phide i s evolved when it i s reacted with sulphuric acid. This H2S has not been a nuisance i n t he p lan t a rea but i t i s believed t o have accelerated t he corrosion of s t a i n l e s s s t e e l equipment i n t h e phos- phoric acid production sect ion.

Table V I I ( a )

Calcined Phosphate Rock Typical Screen Size Dis t r ibu t ion

I Microns I Percent I

EXTRACTIVE METALLURGY OF LEAD AND ZINC

Table VII (b)

Calcined Phosphate Rock Typical Chemical Analysis

Component Percent Dry Basis

Bone Phosphate of Lime (BPL)

Phosphate (P205)

'~cid Insoluble

Calcium (CaO)

Iron (Fe293)

Aluminum (~1203)

Total Carbon as C

Carbonate Carbon as C

organic Carbon as C

I Fluoride (F)

Piagnesium (Mg0 )

\Chloride (Cl 1 I I Potassium (~20)

Sodium (Na20 )

Total Sulphur ( S )

Sulphate Sulphur (S)

Acid Evolved Sulphur (5)

Hydrogen Sulfide ( S)

Bulk Moisture as shipped (H20)

The rock i s received i n 10-car l o t s of covered hopper c a r s weigh- i n g 90 T each and unloaded by g r av i t y through a t r a ck hopper and b e l t conveyor system t o a 12,500 T capaci ty s i l o .

Ammonia: Ammonia i s received a s anhydrous l i q u i d i n 72 T c a r s and a small vapour compressor i s used t o t r a n s f e r t h e l i q u i d by displace- ment t o a 200 T pressurized s to rage sphere.

Sulphuric Acid: Sulphuric ac id i s de l ivered by p ipe l ine a s 93 per- c en t product from t h e su lphur ic ac id p lan t about 300 m away. Two, 330 T- tanks a r e ava i l ab l e - a t t h e f e r t i l i z e r p lan t f o r t h e s torage of acid.

Phosphoric Acid Production

The manufacture of phosphoric ac id a s shown i n Figure 9 involves t h e d ige s t i on of ground phosphate rock with sulphuric acid. The p r i nc ipa l r e ac t i on i s represented by t h e following equation:

L 20% Ron Acld Recycle 1

F igure 9 - Phosphoric Acid Production

Rock Grindinq: To obtain maximum ext rac t ion of P2O5 and b e t t e r con- t r o l of t h e r eac t i on , t he rock i s ground zo a f ineness of about 70 percent minus 74 microns by an air-swept, 3.05 m diameter by 2.14 m cy l i nd r i ca l length Hardinge b a l l m i l l driven by a 373 kW motor. The a i r i s c i r cu l a t ed through the m i l l and c h s s i f i e r by a 112 kW fan. Ground rock i s elevated t o a small s to rage b in ahead of t h e a t t a ck sec t ion by a bucket e levator .

Attack Section: 'The rock d iges t ion equipment a t SLE i s a mult iple tank system.. That i s , t he r e a r e four tanks i n s e r i e s which hold the s l u r r y i n t he d i g e s t i x c i r c u i t about s i x hours, each tank has a ca2aci ty of 71 m3. The f i r s t two r eac to r s a r e equipped with dual 450 pitched turb ines 1.66 m i n diameter r o t a t i ng a t 68 rpm. The No. 3 reac tor has a s i ng l e 450 pitched turb ine of t h e same diam<?ter ro- t a t i n g a t t h e same speed. I n t h e No. 4 r eac to r which a l so serves a s a s l u r r y holding tank ahead of the f i l t e r , t h e reac t ion goes t o com- p l e t i on allowing t h e dihydrate gypsum c r y s t a l s (CaS04.2H20) t o s t a - b l l i z e . This tank i s ag i ta ted by a la rge paddl'? mixer 4.3 m i n dia- meter turning a t 25 rpm. The a g i t a t o r s i n t h e f i r s t t h r ee r eac to r s a r e dr iven by 55 kW motors and absorb about 95 percent of t h e f u l l load power. The ag'tator i n t he fo .x th r eac to r i s driven by a 45 kW motor and absorbs only 55 percent o f f u l l load power. A l l t h e agi- t a t o r s a r c dr iven through f l u i d couplings.

The o r i g i n a l ag i t a to r s w<ere constructed of Type 316 s t a i n l e s s s t e e l and showed littl.3 corrosion a f t e r about 18 months of p lan t operation on Moroccan rock which was t he i n i t i a l source of phosphate. However, when t h e change t o North Carolina rock was made, these ag i t a to r s were severely attacked within t h r ee o r f 0 . x months. It i s believed t h a t t he small amount of hydrogen sulphide generated by t he North Carolina rock created a reducing condit ion i n t h e s l u r r y so t h a t Type 316 s t a i n l e s s s t e e l was no longer passive. The wetted pa r t s of t he ag i t a to r s have been replaced by Jessop 700 a l l oy which has a much b e t t e r r e s i s t ance i n these conditions. Wnile awaiting de l ivery of the Jessop 700 ag i t a to r s , rubber covered mild s t e e l bla- des wers used; these showed much b e t t e r r e s i s t ance than might be expected.

'The f i r s t th ree reac tors a r e air-cooled t o l i m i t the s l u r r y tem- pera ture t o a maximum of 7g°C. A teinperature above 80°C induces t h e formation of a semi-hydrate ( c ~ s o ~ . % ~ ~ o ) c r y s t a l which i s d i f - f i c u l t t o f i l t e r . To reinove t h e hea t of reac t ion , about 15,000 m3/hr of a i r a t 56 mm Hg pressure i s in jec ted through nozzles located about 457 mm above t h e surface of t he s lur ry . A f i be rg l a s s reinforced p l a s t i c exhaust fan r a t ed a t 2 7,000 m3/hr withdraws t h e sa tura ted a i r and fumes From the reac tor through a rubber-lined s t e e l , Model 535 Aerornix gas scrubber.

Ground phosphate rock and &out 90 m3/hr of s l u r r y from the No. 3 reac tor a r e premixed i n a s l u r r y mix pot , then flow t o t h e No. 1

FLUID BED ROASTING, ACID, AND FERTILIZER AT VALLEY FIELD 171

reac tor . Sulphuric ac id i s d i l u t ed from 93 t o 74 percent s t rength and cooled t o 70°C i n a Karbate f a l l i n g f i l m cooler. The acid i s then mixed with 25 m3/hr of 20 percent phosphoric acid recycled from the f i l t e r sec t ion and flows i n to No. 1 reactor .

To ensure maximum P205 recovery, t h e phosphate rock t o sulphuric acid feed r a t i o must be c lose ly control led. An excess of f r e e sul- phuric acid causes an impe,rvious l aye r of calcium sulphate t o be formed on t h e sur face of the phosphate rock p a r t i c l e s , stopping t he react ion. This e f f e c t becomes evident when t h e sulphuric acid con- cen t ra t ion i n t he s l u r r y exceeds 3.5 percent and unless it i s cor- rected quickly, reac t ion ceases. Conversely, i f t h e sulphuric acid content of the reac t ion s lu r ry f a l l s belo:$ 1 percent t h e formation of insoluble dicalcium phosphate i l l a ssoc ia t ion with gypsum i s pro- moted. Dicalcium phosphate (CaHPOq.2H20) is s imi l a r i n c r y s t a l s t r uc tu re t o dihydrate gypsum (CaSOq.2H20) m d it is believed t h a t groups ,-)f phosphoric ac id ions replace some of t he sulphate ions i n t h e gypsum c r y s t a l l a t t i c e .

T h e s l u r ry from No. 4 reac tor conraining about 28% P205 phosphoric ac id and 40 percent gypsum s o l i d s i s f i l t e r e d on a Model 13 Bird- Prayon t i l t i n g pan f i l t e r having 38.7 m2 of f i l t r a t i o n area. After t he separat ion of the 28% P205 product ac id , t he gypsum cake i 5

given t w o washes with progressively weaker phosphoric ac id and a f i n a l wash with hot water. The s t ronges t wash f i l t r a t e , containing about 20% P205 i s recycled t o t h e d iges t ion system. The 28 percent product acid i s pumped t o two storage tanks having a capaci ty of 284 m3. The gypsum cake dumped from the f i l t e r i s s l u r r i e d with water a d pumped t o a ?40,000 m2 disposal pond.

Evaporation: To provide phosphoric acid of t he appropriate s t rength f o r t h e production of phosphate f e r t i l i z e r s t he p lan t has two forced c i r cu l a t i on vacuum evaporators. The u n i t s cdn be operated i~ s e r i e s o r p a r d l l e l t o concentrate 135 T/da t o P2O5 i n 28 percent acid t o 54 percent acid. Normally t h e un i t s make 45% P2O5 acid f o r t r i p l e superphosphate and 39% P2O5 f o r t h e manufacture of diarnmonium phos- phate.

The heat f o r evaporation i s supplied by c i r cu l a t i ng t he acid through a ve r t i c a l hea t exchanger heated by 0.7 atm steam or1 t h e s h e l l s i d e and containing 178 impervious graphi te (Karbate) tubes, 5 i n long. Evaporate is withdrawn snd condensed by 100 m3/hr of cold water and t he non-condensable gas i s removed by a two-stage e jec tor . The evaporator operates a t a pressure of 100 mrn Hg ab- so lu t e with a so lu t ion temperature of 650C. Concentrated acid i s pumped t '3 four rubber-lined s t e e l s to rage tanks each capable of s t o r ing 170 m3 of acid.

T r ip l e S;~perphosph_> Production

T r ip l e superphosphate i s msde by the reac t ion of phospho-. L i ~ acid

and f i n e l y ground phosphate rock t o y i e l d water-soluble monocalcium phosphate. The p r i n c i p a l r e a c t i o n i s represented by t h e equation:

F i g w e 10 i s t h e flowsheet of t h e Klrhlman process employed a t SLF. The same grade of rock a s i s used t o make phosphoric a c i d i s ground t o 85-90 percen t minus 7.4 micron i n a 1.33 m diameter Raymond r o l l e r m i l l equipped with a Whizzer separa to r . The ground rock i s de l ive red by a weighing f e e d e r t o a smal l , a g i t a t e d mix po t where it is tho- roughly mixed with a metered stream of 45% P205 phosphoric ac id . The s l u r r y then .claws onto a ruS5er conveyor b e l t 1.2 m wide and 42.5 m long where i t r a p i d l y s o l i d i f i e s . A t t h e end o f t h e b o l t a r o t a t i n g drum d i s i n t e g r a t o r breaks up t h e r e a c t i o n mass before it passes t o a r o t a r y k i l n dryer . The production b e l t i s v e n t i l a t e d by a f i b e r - g l a s s re in forced p l a s t i c f an moving 10,000 m3/hr through a Model 362 Aeromix scrubber.

The product i s forced cured and d r i e d i n a co-current d ryer 2.15 m i:l diameter by 24 rn long equipped with a fu rnace burning No. 6 o i l t o g ive a h e a t r e l e a s e of about 1 x lo6 kcal /hr . The d ryer has t h e capac i ty t o d ry about 9 T/hr of product from 17 t o 8 pe rcen t mois- t u r e . The d r i e d product i s t r anspor ted 65 m t o t h e s t o r a g e bu i ld ing on a s e r i e s of conveyors. Dryer gases a r e f i r s t cleaned i n a cyclone separa to r t h e n blown through a rubber-lined s t e e l Model 400 Aeromix scrubber by a s t a i n l e s s s t e e l f a n having s capac i ty of 13,500 m3/hr again:st z pressure drop of 15 nun Hg.

!./hen t h e product i s discharged from t h e d r y e r , t h e r e a c t i o n ts produce s o l u b l e monocalcium phosphate i s 85-95 percen t complete arid proceeds t o completion over a cur ing pe r iod of about 14 days i n t h e s t o r a g e bui lding.

Diammonim~ Phosphate Production

F e r t i l i z e r grade "diammonium" phosphate (18-46-0) i s a mixture of t h e clmpounds diammonium phosphate, mono ammonim phosphate p l u s impur i t i e s p r e c i p i t a t e d from t h e wet process phosphoric acid . The p r i n c i p a l r e a c t i o n is represejl ted by t h e fol lowing equat iox:

The vapour p ressure of ammonia over a s a t u r a t e d , b o i l i n g s o l u t i o n of diammo.~ium phosphate i s such t h a t ammonia vapour l o s s e s would be excess ive un less s p e c i a l precaut ions a r e t aken t o recover t h i s am- monia. Most p l a n t s i n North America producing t h i s product avoid t h e d i f f i c u l t y by p a r t i a l l y armnoniating t h e phosphoric acid i n a p re -neu t ra l i ze r and running t h e s l u r r y onto a d r i e d bed of recycled product I n a g ranu la t ion drum o r blunger and f i n i s h t h e ammoniation d t t h i s point . The excess ammonia i s captured i n a scrubber by a c i r c u l a t i n g s o h t i o i of phosphoric ,3cid.

i-2 T S P To Storage

Figure 10 - Trip le Superphosphate Production

The Kuhlman process used a t SLF i s shown i n Figure 11. It is unique i n t h a t t he phosphoric acid i s ammoniated t o diaminonium phos- phate i n a s i ng l e s t age reactor . The problem of excess ammonia is overcome by passing t he off-gas from the reac tor through a heat-ex- changer where t he grea te r p a r t of t h e steam i s condensed and flows as condensate t o a s t r i pp ing tower and t h e ammonia recovered by steam s t r ipp ing . The ammonia passing through the condenser is re-compres- sed by ~3 small r o t a ry compressor, conwined with t he ammonia from the s t r i p p e r and re-injected i n t o t he reactor . New ammonia f o r t he pro- cess i s received as l i qu id from the s torage sphere, converted t o vapour i n a small s t e e l hea t exchanger, combined with t h e recycle stream and i n j ec t ed i n t o t h e r eac to r through e i g h t spxcgers a t a pressure of .5 atm. Phosphoric ac id containing about 38% P2O5 flows i n t o ' t h e reac tor v i a a s m s l l t a i l gas abssrber through which t h e in- e r t gases a r e purged from t h e reactor .

The s l u r r y of diamm0ni.m phosphate i s bled from the reac tor th- rough a var iab le o r i f i c e d i r e c t l y t o a twin-shaft blunger dr iven by a 55 kW motor. I n t h e blunger t he s l u r r y coa ts a recycle steam of 80-90 T/hr of d r i ed , granulated prod.uct.

The wet p e l l e t s drop d i r e c t l y i n t o a co-current dryer 2.75 m i n diameter- by 24.5 m long heated by a furnace burning No. 6 . o i l t o give a hea t r e l ea se of about 2 x 106 kcal/hr. The d r i ed product i s ele-

vated by a continuous chain bucket e leva tor t o t h r ee screens i n ser- i e s . The f i r s t 1 m x 2.5 m mechanically vibrated screen removes t h e plus 12 mm material and r e j e c t s it to a s i ng l e r o l l crusher followed by two p a r a l l e l double r o l l crushers.

The undersize from the lump screen i s screened on a 1.25 m x 4.6 m Tyler H m e r screen a t 1.00 mm with t he undersize going d i r e c t l y t o recycle. The oversize from t h e s e a n d screen i s f i n a l l y screened on a 1.25 m x 3.0 m Tyler Hummer screen w i t h a 3.36 rn deck from which the oversize goes t o the crusher syscem and t h e undersize t o a 2.45 m diameter x 15.5 m long a i r cooled product cooler and then t o storage.

Table V I I I summarizes SLF operat ing s t a t i s t i c s f o r a t yp i ca l month.

Evaporator Amrnon~a + From-- - 4 Storage Recyc le Ammon~a

I i ! - I

Condensor I

c

i compressor - - - - - - - - J

S t e a m

j Ammonia 1 ' Stripper

Dryer Scrubber

Reactor

\

.- 4

Figure 11 - Diamrnonium Phosphate Production

FLUID BED ROASTING, ACID, AND FERTILIZER AT VALLEY FIELD

Table VIII

SLF Operating Statistics - February 1970

perating Time 0% P2O5 Acid Production (P205) oncentrated Acid Production (P205) verall gypsum efficiency otal P2O5 losses jn gypsum verall recovery acid plant hospha.te rock consumed ulphuric acid (100%) consumed

uel oil consumption

roduct size -3.36 mm +1.41 mrn

onia consumption hosphor:ic acid (P2O5) consumption uel oil consumption

otal steam consumption otal operating labour otal ma.intenance labour

REFERENCES

(1) Jephson, A.C., A.Y. Bethune and R.C. Kelahan, llZinc Recovery a t Canadian E lec t ro ly t i c Zinc's New Valleyfield Plant", Journal of Metals, Vol. 18, NO. 8 , Aug. 1966, pp 947-956.

(2) Spira , P., llA Radioactive Tracer Test a t t he Orchan Concentrator 24 September 1969", In t e rna l Report No. 188, Noranda Research Centre, February 1970.

( 3 ) Spira, P., "A Radioactive Tracer Test i n No. 1 Roaster a t Cana- dian E lec t ro ly t i c Zinc", In t e rna l Report No. 187, Noranda Research Centre, February 1970.

(4 ) Donovan, J . R . and P.J. Stuber, l lSulfuric Acid Production from Ore Roaster Gases", Journal of Metals, Vol. 19, No. 11, November 1967, 9

Canadian E

lectrolytic Zinc Ltd.

Valley field, Q

ue., Canada

Chapter 5, page 144

S.A. V

ieille-Montagw

B

alen, Belgium

C

hapter 6, page 178

Mitsubishi m

etal Mining C

o., Ltd.

Alcita, Japan

Chapter 7, page 198

Ruhr-Zink G

MBH

Datteln, F

ederal Republic of

Germ

any C

hapter 9, page 247

Am

ericun Zinc C

o. East St. Louis, Illinois, U

SA

C

hapter 11, page 308

Com

inco, Ltd.

Trail, B.C., Canada

Chapter 12, page 330

Nisso Sm

elting Co. L

td. A

izu, Japan C

hapter 15, page 409