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Modernisation of the leaching circuit at Pasminco metals—EZ
hydrometallurgy ELSEVIER
Hydrometallurgy39 (1995) 129-145
Modernisation of the leaching circuit at Pasminco Metals EZ M.G. Kershaw Senior Development Metallurgist Pasminco Metals --EZ, Risdon, Tasmania, Australia
Received 4 April 1995; accepted 8 June 1995
Abstract
The calcine leaching circuit at the Pasminco Metals - - E Z plant has recently been extensively upgraded, with the pachuca leaching and iron purification stages being replaced. The modernised leaching circuit is a conventional counter-current plant, composed of two low-acid calcine leaching stages, (neutral leach and weak-acid leach), an iron precipitation stage (jarosite precipitation) and three high-acid leaching stages (preneutralisation, hot-acid leach and strong-acid leach). Integration of the new leaching circuit and closure of the old pachuca leaching section commenced in March 1990. On completion in September 1991, the leaching capacity increased to produce 220 000 t zinc per annum. The thickening and filtration sections, in the high-acid leaching stages of the modernised leach circuit, did not perform satisfactorily from the outset, because of the presence of very fine particulate material containing high concentrations of silica. The silica-beating particles formed a recirculating load within the high-acid leaching stages, increasing the silica content of thickener feeds. The resultant low density thickener underflows, containing less than 200 g solids/l, increased recycle flows through the plant and severely limited the capacity of the filters used for filtering the secondary leach residue from the high-acid leaching stages. Zinc production was only maintained by reducing the lead and silica content of zinc concentrates and utilising all available filter capacity for secondary leach residue. Silica in the preneutralisation stage liquors is partially precipitated in the jarosite stage. The residual silica in the jarosite precipitation stage liquors was found to be a significant contributor to the particulate silica in the high-acid leaching stages. Various options were explored to increase silica precipitation in the jarosite precipitation stage, with the addition of a coagulant to this stage being the most cost effective. Thickeners previously full of low-density pulps are now devoid of pulp, with underfiows containing up to 400 g solids/1. The secondary leach residue filters now have surplus capacity, and the lead content of the residue has increased from 26 to 36%. The restriction on lead and silica in zinc concentrates has been removed.
0304-386X/95/$09.50 © 1995 Elsevier Science B.V. All rights reserved SSDI O304- 386 X ( 9 5 ) 0 0 0 2 6 - 7
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M.G. Kershaw / Hydrometallurgy 39 (1995) 129-145
1. Introduction The leaching circuit in the Pasminco Metals - - E Z (PMEZ) zinc plant has been extensively modernised. Commissioning and integration of the new circuit commenced in March 1990 and was completed by September 1991. The modernisation involved replacement of the pachuca leaching and iron purification sections by a conventional neutral leach operation. As well, an extensive upgrade of the jarosite precipitation stage and the high-acid leaching sections was undertaken to increase recovery of valuable metals. The low-acid leaching stages of the plant and the jarosite precipitation stage performed as expected from design. However, operations in the high-acid leaching stages performed poorly mainly due to the above design quantities ofjarosite in these stages and the effects of very fine silica-bearing particulates. The modernised PMEZ circuit, which has been described previously [1] [2], has a capacity of 220 000 t zinc per annum. It consists of a roasting stage converting zinc sulphide concentrates to a zinc calcine (calcine). The calcine is leached in dilute sulphuric acid to recover the contained zinc. The neutral solution from leaching the calcine is further purified to remove elements such as cadmium, copper, cobalt and nickel by cementation with zinc dust. The purified solution is cooled, and electrolysed to recover the zinc as zinc metal. The spent electrolyte, containing 165 to 175 g H2SO4/1, is returned to the leaching circuit as dilute sulphuric acid. Two residues are produced, an iron-bearing residue, in which the iron is present as jarosite and a lead-silver bearing residue or secondary leach residue.
2. The leaching circuit The leaching circuit is shown in Fig. 1. The tankage and equipment have been described previously [ 1 ] [2]. The principal aims of the new circuit are to increase zinc extraction from calcine and to reduce the heavy metal levels in jarosite, thereby increasing zinc recovery to at least 96%. These aims, coupled with the need to roast concentrate feed blends containing greater than 9% iron and to recover zinc from a stockpiled zinc ferrite residue, resulted in a relatively complex counter-current leach circuit with six stages, consisting of : 1. Two stages of low-acid leaching called neutral leach (NL) and weak-acid leach (WAL) stages; 2. Ajarositeprecipitation (JAR) stage including an improved jarosite filtration and washing step; 3. Three stages of high-acid leaching called preneutralisation (PN) (including a preneutralisation repulp step), hot-acid leach (HAL) and strong-acid leach (SAL) stages. 2.1. Calcine and spent electrolyte
Zinc sulphide concentrates are blended and roasted in two fluid bed roasters to produce a calcine containing zinc oxide, lead sulphate, zinc ferrite and small amounts of other compounds. Typical roaster feed and calcine analyses are given in Table 1.
LEGEND
CONC, HzSO
SPENT ELECTROLYTE STOCKPILED FERRITE RESIDUE
:
I--I
OPERATION
l
THICKENING
~
UN"
STEP
CONC.H2SO~
CONCENTRATES
J CALCINE STORAGE J CALCINE SECONDARY LEACHRESIDUE J
CALCINE
SPENT ELECTROLYTE
ZINC DUST
LIMESTONE'~IM
',o
SPENT ELECTROLYTE RETURNEDTO LEACHING
CATHODEZINC JAROSITE
~
ALLOY METALS
• ZINC DUST TO PURIFICATION ZINC METAL AND ALLOYS Fig. 1. Flowsheet for Pasminco Metals - - E Z modemised circuit.
M. G. Kershaw / Hydrometallurgy 39 (1995) 129-145
132
Table 1 Roaster feed and calcine compositions ( % )
Zn Pb Fe SiO2 ZnF~O4 ZnO PbSO4
Roaster feed
Calcine
51.2 2.35 9.7 1.3 -
57.6 2.6 11.0 1.4 24.3 60.9 3.8
The calcine is wet ground at a rate of up to 50 t/h in one of two 2.1 m × 3.05 m ball mills in NL solution to produce a slurry containing 500 g calcine/l, with 80% passing 90 microns. The slurry is pumped to various stages within the leaching plant. Typical calcine and spent electrolyte distributions are: Stage Calcine (%) Spent electrolyte (%) Neutral Leach 41 18 Weak-acid Leach 25 6 Jarosite Precipitation 18 6 Preneutralisation 16 0 Hot-acid Leach 0 0 Strong-acid Leach 0 70
2.2. Low-acid leaching stages The objectives of the low-acid leaching stages are to produce a NL thickener liquor with low concentrations of iron, arsenic, etc., for the zinc dust purification stage and a residue from the WAL stage, containing low levels of zinc oxide, to be further leached in the preneutralisation stage.
Process description of low-acid leaching stages Feed streams to the NL stage are spent electrolyte and solution from the WAL stage. Calcine is added, under pH control, to achieve a pH of between 5.0 and 5.3. The principal reactions in the NL stage are: (A) Neutralisation of sulphuric acid: ZnO +
H2SO 4 = ZnSO 4 +
H20
(B) Oxidation of ferrous iron to ferric iron using air sparging, followed by precipitation of the ferric iron as ferric hydroxide: 4Fe 2 + + 02 + 2H20
=
4Fe 3+ + 4 O H -
Fe 3+ + O H - + H z O + Z n O = F e ( O H ) 3 + Z n 2+ (C) Co-precipitation of a number of impurities such as arsenic and antimony, which are toxic to zinc electrowinning.
M. G. Kershaw / Hydrometallurgy 39 (1995) 129-145
133
The feeds to the WAL stage are liquor from the JAR stage and NL thickener underflow pulp, with the pH being controlled by the addition of either spent electrolyte or calcine slurry. The WAL stage is designed to dissolve most of the zinc oxide in the NL solids and added calcine and give a preliminary purification of JAR solution by partially precipitating elements such as arsenic and antimony.
Design criteria for the low-acid leaching stages The design of the NL stage was based on the need to produce a solution containing < 1 mg Fe 2÷/1, < 10 mg total iron/1 and < 0.5 mg As/1 for the subsequent zinc dust purification stages. Oxidation of ferrous iron is the rate controlling step and the residence time required was based on previously developed models [ 3,4], in which the oxidation rate is a function of the ferrous iron concentration and the liquor pH. Tankage was installed to oxidise up to 0.5 g Fe 2÷/1 in the combined feed stream of spent electrolyte plus liquor from the WAL stage. Air at 200 kPa was supplied to the tanks to oxidise the ferrous iron. The WAL stage was designed to precipitate ferric iron from JAR solution as ferric hydroxide at a pH between 3.0 and 3.5. Under these conditions, the extent of arsenic and antimony precipitation was expected to be greater than 97% and 69%, respectively, and the overall zinc oxide extraction from calcine to be 93%.
Performance of the low-acid leaching stages The ferrous iron oxidation in the NL tanks is better than design (Table 2), which has been attributed to the more efficient gas dispersion properties of the installed agitators, compared to the Rushton turbine impellers used in the pilot plant studies. The liquor, after solids/liquid separation in two stages of thickeners, contains 146 g Zn/1, 26 mg total iron/ 1, 1.1 mg soluble Fe/l, 1.5 mg As/1 and 100 mg particulates/1 and is forwarded to the zinc dust purification stages. In the NL stage, the extraction of the zinc oxide component from calcine has averaged 83%, compared to design of 85%. The performance of the WAL stage is measured on the overall zinc oxide extraction from calcine. The extraction has averaged 87%, which is lower than design as a result of operating at a higher pH. The operating pH was increased by 0.5 due to higher than design flows of JAR liquor and poor agitation in the WAL tanks. The flow of JAR stage liquor to the WAL Table 2 Ferrous iron concentrations in the NL stage Tank
Premix Tank NL Tank 1 NL Tank 2 NL Tank 3 NL Tank 4 NL Tank 5 *From pilot plant data.
pH
5.1 5.3 5.3 5.3 5.3
Fe 2+ (mg/l) Actual
Expected *
800 40 l 1 1 < 1
800 335 102 31 9 3
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M.G. Kershaw/ Hydrometallurgy39 (1995) 129-145
stage is nearly twice design, due to the need to clarify jarosite filtrate (designed to bypass the WAL stage) and the higher than design flow of spent electrolyte to the SAL stage. Arsenic and antimony precipitation levels of 98% and 90%, respectively, are better than design. Improvements in NL stage clarification and agitation in the WAL tanks are expected to bring the performance of these two stages up to design expectations. 2.3. Jarosite precipitation stage
Clarified PN liquor is pumped to the first JAR tank, where aqueous ammonia solution and jarosite seed are added. Calcine slurry is added to the second, third and fourth tanks to control the acidity. The ferric iron (dissolved in the high-acid leach stages) is precipitated at 97°C, predominantly as ammonium jarosite with minor amounts of other jarosites also being precipitated (e.g. sodium and potassium), by the reaction: M + +3Fe 3+ +2SOZ4- + 6 H 2 0 = MFe3(SOn)2(OH)6+6H + where M = NH4, K, Na, H30. The acidity is increased in the sixth JAR tank to 20 g HzSO4/I with spent electrolyte. The acidification reduces scale formation in the thickeners and pumping lines and dissolves species, which interfere with flocculation. The solids are removed by thickening and the thickener underflow slurry is neutralised with limestone prior to filtration on three 512 m 2 Rittershaus and Blecher recessed chamber filter presses. JAR stage design criteria The pilot plant studies showed that the rate of iron precipitation is a function of temperature and concentration of ferric iron, acidity and ammonium ion concentration [5]. For the design conditions of 10 g H2SO4/1, 3 g NH~-/1 and 97°C a residence time of 4.4 h was required to precipitate 80% of the ferric iron as jarosite from the PN liquor. JAR performance About 75% of the ferric iron in the PN liquor is precipitated, giving a solution containing between 4 and 7 g Fe 3 ÷/1. The precipitation occurs at an acidity of 3 to 6 g H2SO4/1 and an ammonium ion concentration of 2 to 2.5 g NH~-/1. The actual residence time in the JAR stage is one hour less than design. The effects of the lower residence time and lower ammonium ion concentration have been small and have been offset by operating the JAR stage at a lower acidity. Typical compositions of the jarosite filter cake and thickener overflow are given in Table 3 and Table 4. 2.4. High-acid leaching stages
The high-acid leaching stages are the PN, HAL and SAL stages in which the acidity ranges from 5 to 160 g H2SO4]1.
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135
Table 3 Jarosite filter cake analyses Component
Zn Fe Pb AI SiO2 Ca
Composition (%) Design
Plant
3.2 28.6 1.1 0.45 1.8 5.9
3.5 24.9 1.3 0.42 2.4 6.3
Table 4 JAR stage thickener overflowliquor composition Component
Zn Fe3+ Fe2+ H2SO4 NH4 SiO2
Composition (g/l) Pilot plant
Plant
150.5 5.7 1.5 11.8 3.5 0.85
125.7 6 0.5 20 2.1 0.88
Process description of the high-acid leaching stages The operating acidity in the PN stage is the lowest of the high-acid leach stages. The feeds to the PN stage are liquor from the H A L stage and W A L thickener underflow pulp. The main purpose of the PN stage is to reduce the acidity of the H A L stage thickener liquor as low as possible without precipitating ferric iron as jarosite. The zinc oxide and ferric hydroxide components in the W A L solids are leached with preneutralisation repulp (PN repulp) thickener overflow liquor at 20 to 25 g HzSO4/1. The acidity of the slurry is further reduced by calcine additions to an acidity of 5 g H2SO4/1 before being thickened in a high capacity thickener. The liquor from the thickener is further clarified to minimise carry-over of calcine residue to the JAR stage. Both PN thickener underflow and PN clarifier underflow streams are pumped to the PN repulp thickener where they are mixed with H A L liquor, to effect an acid wash of the solids. The H A L stage is the second of the two high-acid leaching stages, designed primarily for the dissolution of the zinc ferrite from calcine residue and stockpiled zinc ferrite residue, according to: ZnFe/O4 + 4H2SO4 = ZnSO4 + Fe2(504) 3 df.4H20 PN repulp solids are leached in liquor from the SAL stage at 85°C and an acidity ranging from 120 g H2SO4/1 in the first tank to 70 g H2SO4/1 in the fifth tank. The stockpiled zinc ferrite can be added to the second leach tank for acidity control. After thickening, the
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M.G. Kershaw/ HydrometaUurgy39 (1995)129-145
clarified liquor is pumped to the PN repulp step and the thickened pulp is pumped to the SAL stage. The thickened slurry from the HAL stage is leached in the SAL stage in a mixture of contact sulphuric acid and hot spent electrolyte (heated to 85°C by steam in plate heat exchangers) to dissolve zinc ferrite and jarosite from the HAL solids, increasing the zinc ferrite extraction from the zinc residues to 95%. The thickened SAL solids are filtered and washed on three 35 m 2 Hoesch pressure filters. The filters have excess capacity and can easily filter the 75 t/d of secondary leach residue. This residue is a valuable by-product containing greater than 80% of the lead and silver in the calcine added to the leaching circuit.
Design criteria for the high-acid leaching stages The design criteria for the PN stage were, 100% dissolution of the ferric hydroxide and zinc oxide from the WAL solids and calcine with a final acidity in the PN liquor of 5 g H2SO4/1 and with less than 5% precipitation of iron as jarosite, so that the PN solids contain less than 8.5% iron as jarosite. Pilot plant studies indicated that jarosite formation in the PN stage would be of concern, with scale formation rates of 1 mm/day being measured. A spent electrolyte cleaning system was installed to descale the PN tanks and thickener on a weekly cycle. The dissolution of zinc ferrite is relatively slow. To maximise the amount of zinc ferrite dissolved in the HAL stage, the acidity in SAL liquor must be reduced as low as possible. However, jarosite precipitation is favoured by low acidity and long residence times. The pilot plant studies showed that the optimum minimum acidity in the HAL tanks was 60 g H2SO4/1. At this acidity, 80% of the zinc ferrite in the PN repulp solids and in 100 t/d of stockpiled zinc ferrite could be dissolved in the five 250 m 3 tanks. Under these conditions limited dissolution of jarosite from the PN repulp solids occurred in the first HAL tank followed by some jarosite precipitation in the later tanks and thickener. The design of the HAL plant assumed no net jarosite formation. The SAL stage was designed to dissolve at least 90% of the iron contained in the HAL solids, from both jarosite and zinc ferrite. The minimum acidity and temperature from the pilot plant studies was 140 g HzSO4/1 and 95°C, respectively, with a residence time of 3 hours. Two 35 m 2 Hoesch pressure filters were initially installed to filter up to 100 t/d of secondary leach residue. Performance of the high-acid leaching stages It has proven to be very difficult to control the acidity in the second part of the PN stage at 5 g H2SO4/1 due to the residence time in the reactors being too short, compared to the frequency of acidity determinations by the on-stream analysers. The actual acidity is in the range of 5 to 10 g H2SO4/1. Small quantities of zinc oxide remain in the PN solids with the amount increasing as the quantity of calcine required in the PN stage for acidity control increases. The main causes of this problem are short circuiting of the calcine in the tanks and the residence time being too low for complete reaction of the zinc oxide contained in the calcine. The zinc oxide in the PN solids is dissolved in the PN repulp step.
M.G. Kershaw / Hydrometallurgy 39 (1995) 129-145
137
Table 5 PN stage solid and liquor analyses PN solids (%)
Component
Component
Pilot plant
Plant
Zn Fe Pb NH4 K SiO2
17.1 28.7 7.7 0.28 0.18 3.6
17.4 31.9 6.0 0.33 0.22 3.6
Zn Fe 3÷ H2SO4 NIL, K SiO2
PN liquor (g/l) Pilot Plant
Plant
117.6 36.4 14.2 3.9 0.13 4.6
117.0 25.0 9.6 2.5 0.03 2.6
"includes particulate silica. Table 6 HAL stage solid and liquor analyses Component
Zn Fe Pb SiO2 K NH4 ZnFe2Oa K jarosite NH4 jarosite
Solid composition (%) Pilot plant
Plant
9.5 17.5 17.2 9.1 0.75 0.35 25.7 9.2 9.3
6.8 19.9 18.2 13.0 0.63 0.50 23.1 7.7 13.3
Component
Liquor composition (g/l) Pilot plant
Plant
K NFL,
91.2 42.6 1.8 56.1 2.9 0.24 3.8
70 35.7 0.4 79.6 2.0 0.24 2.8
Component
Liquor composition ( g / 1)
Zn Fe 3+ f e 2+ H2504
*includes particulate silica. Table 7 SAL stage solid and liquor analyses Component
Zn Fe Pb SiO2 K NIL, ZnF~O4 K jarosite NI-~ jarosite
Solid composition (%) Pilot plant
Plant 1994
6.4 4.3 37.0 11.1 0.07 0.08 9.0 < 1 <2
4.0 2.9 38.3 13.4 0.07 0.09 6.0 < 1 <2
* Si02 analysis includes particulates.
Zn Fe 3+ f e 2+
H2SO4 K NH4
Pilot plant
Plant 1994
61.0 16.3 0.3 175.6 3.2 0.55 3.8
49.0 9.1 0.1 168.0 1.7 0.15 2.8
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M.G. Kershaw/ HydrometaUurgy39 (1995) 129-145
Iron, as jarosite, in the PN solids averages 11.2%. Half of this iron is from jarosite in the JAR stage thickener overflow liquor which is settled in the WAL stage thickeners. The seeding effect of this jarosite in the PN stage is unknown. The other half of the jarosite is precipitated in the PN stage. In general, the operations of the PN stage are similar to design, albeit at a higher acidity. If the acidity in the PN stage could be controlled at 5 g H2SO4/1,the calcine to the JAR stage could be reduced by 10%, increasing the overall zinc recovery from calcine in the leaching circuit by 0.3%. Solid and liquor analyses from the PN stage are compared to the pilot plant data in Table 5. The HAL and SAL stages in the two years following commissioning operated well below design expectations. The problems experienced are discussed in Section 3. The performance of the HAL stage is now within design expectations, with the zinc ferrite extraction averaging 84% and a net dissolution of both ammonium jarosite (50%) and potassium jarosite (6%). Dissolution ofjarosite occurs mainly in the first HAL tank where the acidity is usually greater than 100 g H2SO4/1. Typical compositions of HAL solids and liquors are given in Table 6. Since October 1993, the performance of the SAL stage has been close to design, with 85% iron dissolution from the HAL solids at a minimum acidity of 150 g H2SO4/1. Additional leaching occurs in the thickening step prior to filtration, increasing iron extraction to 87%. The lower than expected iron extraction has been attributed to a reduced residence time of 2 hours, because of the need to increase the acidity in the SAL stage by increasing the flowrate of spent electrolyte (see Section 3). Typical solid and liquor compositions for the SAL stage are given in Table 7.
3. Problems since commissioning The major problems since commissioning which limited the capacity of the leaching circuit were, the excessive quantity of solids to be filtered on the Hoesch filters and the poor settling and filtering characteristics of the HAL and SAL solids. 3.1. Poor settling and filtration
Settling in the SAL, HAL and PN stages of the plant steadily deteriorated after the plant was commissioned. Settling and filtration were so bad that all the SAL thickeners (two 6.5m diameter high capacity and one 15-m diameter conventional thickeners) and the HAL thickeners (two conventional 15-m diameter thickeners) were full of low-density pulps containing < 200 g solids/1 in the thickener underflow slurry. (Design was based on > 300 g solids/l in the thickener underflow slurry and > 400 g solids/1 in the Hoesch filter feed). The overflow from the thickeners contained unsettled solids which increased the thickening problems in the following stage. To break the recycle of solids via thickener overflows, the PN repulp thickener was used as a polishing thickener for the HAL thickener overflow, with the underflow being campaign filtered on one of the jarosite filters. As well, the old secondary leach residue filtration equipment was recommissioned and 2000 m 3 of SAL pulp
M.G. Kershaw / Hydrometallurgy 39 (1995) 129-145
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Table 8 Composition of PN, HAL and SAL solids, June 1992 Component
Zn Fe Pb SiO2 K NH 4 ZnFe204 NH4 jarosite K jarosite
Solids composition (%) PN
HAL
SAL
12.8 24.5 4.2 7.3 0.22 0.90 34.3 24.0 2.8
4.7 25.0 8.0 8.8 0.34 1.41 16.1 37.5 4.4
3.5 10.6 23.4 13.4 0.20 0.60 7.0 16.0 2.6
was stockpiled in redundant leach tankage. Even with these measures in place, the plant was not capable of sustaining zinc production because of limited Hoesch filter capacity, with 75 t/d being the maximum sustainable filtration capacity, whereas the actual secondary leach residue production was in excess of 130 t/d. Higher grade concentrates, containing less lead and silica, were purchased to partly offset the lack of filtration capacity. The analysis of the secondary leach residue being produced during this period had high residual concentrations of zinc ferrite and jarosites. Typical analyses of the solids from the PN, HAL and SAL stages at this time (June 1992) are given in Table 8. Increased leaching of zinc ferrite and dissolution of jarosite in the HAL and SAL stages were achieved by increasing the spent electrolyte flow to the SAL stage, thereby increasing the acidity in the SAL and HAL stages. Jarosite formation in the PN and HAL stages was reduced by lowering the temperature 5°C. The changes (summarised in Table 9) resulted in a 25% reduction in tonnage of material being produced (Table 10), with the extraction of zinc ferrite and jarosite approaching design levels. Table 9 Modifications to the operating parameters ofPN, HAL, SAL Parameters Design
Modified
SAL
Temperature (°C) Spent flow (m3ph) ( % of total) Acidity (g H2SO4/L)
95 < 135 60 > 140
90 < 175 70 > 155
> 60 85-90
>70 80-85
85-90
80-85
HAL
Acidity (g H2SO4/L) Temperature (°C) PN
Temperature (°C)
M.G. Kershaw /Hydrometallurgy 39 (1995) 129-145
140
Table 10 Tonnage of solids in the HAL and SAL stages before and after the changes Tonnage (tpd)
HAL solids
Total Contained jarosite zinc ferrite SiO2 PbSO4 Accounted (%)
SAL solids
Before
After
Before
389
183
133
163 63 34 46 79
55 24 29 46 84
31 9.5 18 46 79
After 93 <5 11 18 46 86
Despite these changes, the filtration rate of the secondary leach residue on the Hoesch filters continued to decline. 3.2. Silica
The silica content of the solids in the PN, HAL and SAL stages remained the only major difference between the plant in September 1992 and the pilot plant study (Table 11). A detailed mass balance for silica showed that there was a recirculating load of 12 t SiO2/d moving in the liquor stream between the SAL, HAL and PN stages (Table 12) in the thickener overflow streams and returning in the thickener underflow streams. A further analysis of the data showed that silica precipitated from the JAR stage thickener overflow solution in the WAL stage did not redissolve in the PN stage and is the source ofrecirculating silica. In the pilot plant the silica dissolved and returned to the JAR stage, where 80% precipitated, equating to 50% of the silica contained in the added calcine. However, in the plant, only 25% of the silica in calcine reported to jarosite. By increasing the deportment of silica in calcine to jarosite to 50%, the tonnage of secondary leach residue to be filtered would be reduced to less than 83 t/d. The aim of this phase of the post-commissioning optimisation was to increase silica rejection with jarosite by increasing silica precipitation in the JAR stage. Table 11 Silica deportment to solids Solids
Plant September 1992
Pilot plant
Source
% SiO2
%Pb
% SiO2
% Pb
Calcine PN HAL SAL
1.9 7.0 18.6 13.8
3.1 6.9 13.2 30.6
2.0 3.6 9.1 11.1
3.5 7.7 17.2 37.0
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141
Table 12 Silica in thickener overflow liquors (Sept. 1992) Liquor
SAI O/F HAL O/F PN repulp O/F PN O/F JAR O/F WAL O/F NL O/F
Silica concentration (g/l)
SiO2 in overflow
Unfiltered
filtered
% in solids
t/d
2.1 4.0 4.2 2.3 1.6 0.13 0.15
0.25 0.07 0.24 1.34 1.6
67.4 25.0 16.6 13.3
2.1 7.5 0.5 1.9
Silica precipitation in the JAR stage A number of text book options to increase silica precipitation involving modifications to design operating parameters were examined (e.g. increasing/decreasing the pH at which ferric iron was precipitated as jarosite, the use of silica seed material, etc.) These options were all rejected, either as being incompatible with the objectives of the modernised leaching circuit or just did not work. Thus an alternative strategy was required. In the early commissioning phase, cationic flocculants were found to settle the fine silicabearing particles in the SAL and HAL stages, but resulted in the thickener underflows having unacceptably low solids loading. The use of cationic flocculants was discontinued in November 1991, with nonionic flocculants being used instead. When the silica problem was first identified in September 1992, cationic flocculants were trialed as potential silica precipitants in the JAR stage, without success. However, lower molecular weight epichlorohydrin polyamine type condensates, with cationic behaviour, from Allied Colloids Australia Pty Ltd (e.g. Magnafloc 1597) and from Maxwell Chemicals Pty Ltd (e.g. Proset 2830) when added to material from the fifth JAR tank did result in silica precipitation. In these scoping tests, the condensates or coagulants were added at 50 mg/l of slurry and were conducted using one litre of slurry maintained at 95°C in an agitated vessel. After 30 min the treated slurry was flocculated with 5 mg of nonionic flocculant/l and settled. The silica content of the clear decant was generally less than 0.5 g/1. Similar levels of silica precipitation occurred when using polyethylene glycol supplied by ICI Australia Operations Pty Ltd. A series of pilot plant tests simulating the operations of the JAR stage were conducted using the most promising coagulants (Magnafloc 1597 and Proset 2830) to confirm the one-litre tests. The clarified liquor was used in further tests to assess if there were any adverse down-stream effects and to confirm if the benefits from increased silica removal in the JAR stage would be realised. In continuous pilot plant tests, freshly collected PN clarifier overflow liquor from the plant was pumped into the first of six 6-1 tanks, simulating the plant situation. The pilot plant results were validated against plant performance to confirm that silica behaviour was being duplicated (Table 13). The cationic coagulants were also added to the sixth tank at 50 mg/l. Both coagulants reduced the silica concentration in the sixth tank by about 60% (Table 13), at an efficiency of 20 g SiO2 precipitated per gram of coagulant. The pilot
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Table 13 SIO2 Concentrations (g/l) in the JAR stage tanks Stream
PN O / F JAR 1 JAR 2 JAR 3 JAR 4 JAR 5 JAR 6 O/F U/F
Plant ( 2 / 9 / 9 2 )
2.2 1.3 1.7 1.8 1.7 1.7 1.6 1.6 0.5
Coagulant added in 6-1 pilot plant tests NIL
50 mg/1 M 1597"
1.8 1.1 1.2 1.3 1.5 1.4 1.4 1.5 0.29
1.7 1.4 1.6 1.6 1.8 1.6 0.68 0.8 0.24
50 mg/1 P 2830" * 1.8 1.0 1.2 1.2 1.2 1.4 0.47 0.31
"M 1597 = Magnafloc 1597. ""P 2830 = Proset 2830. Table 14 Silica distribution in the leach plant after coagulant additions Stream SiO2 in J A R O/F (g/l) % SlO2 in PN solids HAL solids SLR solids
Before coagulant
After coagulant
1.6
0.9
7.0 15.7 13.8
3.6 10.6 9.3
2.3 4.2 4.0 2.1
2.2 1.7 1.6 1.2
SiO2 in O/F liquors (g/l) PN PN repulp HAL SAL O / F = thickener overflow liquor. SLR = secondary leach residue.
thickener supernatant clarity and the settling rate were greatly improved by the addition of the coagulants. The decant liquor from the pilot plant was used in batch simulations of the low-acid leaching and zinc dust purification stages. The simulations of the high-acid leaching stages all showed improved thickening and clarification characteristics with the respective solids containing less silica. Also, there were no adverse effects from the coagulants in the zinc dust purification steps. The tested coagulants were used in confirmatory plant trials, the results of which were so outstanding that within four weeks of commencing the trials, additions of the coagulant Magnafloc 1597 at a rate of 40 mg/1 had become routine. Every stage in the leaching circuit was beneficially affected.
M.G. Kershaw / Hydrometallurgy 39 (1995) 129-145
143
The effect on silica concentrations in the leach circuit before and after the trial is shown in Table 14. Silica behaviour Silica precipitation from PN liquor occurs in the early JAR tanks equivalent to 25% of the silica in calcine. There is normally a slight increase in silica concentration in the second, third and fourth JAR tanks due to dissolution of silica from the added calcine, equivalent to 0.3 to 0.5 g SIO2/1. With no coagulant addition there is no significant silica precipitation between the fifth tank and the thickener overflow liquor (Table 14). Silica size profiles through the PN and JAR stages showed the majority of silica ( > 90%) in the PN liquors and the early JAR tanks to pass a 0.45 micron filter. By the fifth JAR tank less than 50% of the silica passed a 0.45 micron filter and there had been a reduction in the silica passing a 0.003 micron filter (Table 15). The coagulant "precipitates" the majority of that silica coarser than 0.45 microns, reducing the silica concentration to 0.31 g/l, of which 80% has a size less than 0.003 micron. It is concluded that silica must grow to about 0.45 micron before it can be "precipitated" by the coagulant. However, no silica precipitation occurs by adding the coagulant to an acidified solution taken from the fifth JAR tank. Thus the presence of the jarosite solids and the nonionic flocculant are also important for silica "precipitation". When the residence time between the PN thickener and the addition of the coagulant was reduced by 40% the efficiency of the coagulant was reduced to less than half, resulting in a return to high recirculating loads of silica within four weeks. Silica levels in the high-acid leaching stages of the plant returned to low levels within two weeks, once the residence time was returned to normal. To effectively "precipitate" the silica with the coagulants, a residence time of at least 4.5 h is required for the silica dissolved in the PN stage to polymerise and agglomerate to greater than 0.45 microns.
Table 15 Solution silica sizing Liquor stream
pH
SiO2 concentration (g/l) Total
PN thickener O / F PN clarifier O / F JAR 1 JAR4 JAR5 JAR6 *
0.92 0.85 0.94 1.24 1.14 0.62
• After addition of 50 m g / L of Magnafloc 1597.
1.82 1.73 1.48 0.91 0.70 0.31
Passing filters sized 0.45/zm
0.003/~m
1.76 1.58 1.40 0.53 0.36 0.26
0.77 0.53 0.42 0.36 0.34 0.25
144
M.G. Kershaw/ Hydrometallurgy39 (1995) 129-145
Benefits from the use o f coagulant
The benefits from using the coagulant were very extensive, affecting nearly every stage, and included: 1. Thickening in the SAL, HAL and PN stages is now at design with all underflows containing greater than 300 g solids/1. The HAL stage is operating on only one thickener and the PN repulp step has returned to its design duty. 2. Filtration of the secondary leach residue is no longer the plant bottleneck. (All the stockpiled SAL pulp has now been retreated and filtered). 3. The lead and silver contents of the secondary leach residue are significantly higher as a result of reducing the weight of gangue material in the residue (ie, silica, jarosites, etc.) and markedly increased its commercial value as a high quality source of lead and silver. 4. Flocculation and thickening in the JAR, WAL and high-acid leach stages had greatly improved. As a result, flocculant consumption has decreased plant-wide by at least 50%. 5. A marked reduction in the jarosite content of the WAL stage solids due to lower solids content of the JAR stage thickener overflow, again as a result of improved flocculation characteristics of the JAR solids. 6. The purchase of high grade concentrates is no longer required for solids handling in the SAL and HAL stages. The major benefit has been the increase in the plant leaching capacity by 25,000 t/a of zinc to be at least 5% above design.
4. Conclusions
The lower than design leaching of zinc ferrite and jarosites in the high-acid leaching stages of the modernised leaching circuit at Pasminco Metals - - E Z has been corrected by increasing the hot spent electrolyte flow to these stages. The resultant increased flow through the subsequent stages has had a minor impact on their respective performances. However, the operations of these stages required some modification to account for the extra flow. A coagulant addition to the JAR stage was required to increase silica precipitation to reduce the build up of silica in the high-acid leaching stages. The reduced silica concentrations in both solid and liquor streams have improved all stages of thickening and clarification. Flocculant consumption has halved, with the savings of flocculant more than compensating for the cost of the coagulant. Thickening in the HAL and SAL stages is no longer a plant bottleneck. The secondary leach residue at 83 t/d is at design and can be easily filtered on the three Hoesch filters (a third was commissioned during 1992/93) with mostly only two now being required. An unexpected benefit was that the lower flocculant usage virtually eliminated flocculant break-through from the NL stage which has improved the performance of the zinc dust purification stages. The modernised leach circuit at Pasminco Metals - - E Z is now operating to design expectations.
M.G. Kershaw / Hydrometallurgy 39 (1995) 129-145
145
Acknowledgements The author wishes to thank the m a n a g e m e n t of P a s m i n c o Metals - - - E Z for permission to publish this paper, and those people i n v o l v e d with the project, the success of which is a testimony to their endeavours. T h a n k s are also extended to Dr Ian Matthew ( C o n s u l t a n t ) and Dr R. R i n g of A N S T O , for their assistance in solving the silica problem.
References [ 1] Pammenter, R.V., Gilmour, R. and Salmon, P.T., Production of zinc in the modernised Pasminco Metals - EZ Plant, Hobart, Tasmania. Int. Syrup. World Zinc '93, Hobart, Tasmania, pp. 287-293 (1993). 121 Gilmour, R., Electrolytic zinc production by Pasminco Metals --EZ. In: J.T. Woodcock and J.K. Hamilton (Editors), The Sir Maurice Mawby Memorial Volume, Australasian Mining and Metallurgy, 1:573-578 (1993). [ 3 ] Higuchi, B., Kawaguchi, Y., Asaki, Z. and Kondo, Y., Oxygenation of neutral ferrous sulphate solution by gas bubbling. Denki Kagaku Oyabi Kogyo Butsui Kagaku, (46)12:648-655 (1978). [4] Minegishi, T., Asaki, Z., Higuchi, B. and Kondo, Y., Oxidation of ferrous sulphate in weakly acidic solution by gas bubbling. Metall. Trans., B, 14B: 17-24 (1983). [5] Pammenter, R.V., Kershaw, M.G. and Horsham, T.R., The low-contaminant jarosite process --further developments and the implementation of the process. Ellis Horwood, Chichester (1986), Iron Control in Hydrometallurgy, In: J.E. Dutrizac and A.J. Monhemius (Editors), pp. 603-617.
Hydrometallurgy39 (1995) 129-145
Modernisation of the leaching circuit at Pasminco Metals EZ M.G. Kershaw Senior Development Metallurgist Pasminco Metals --EZ, Risdon, Tasmania, Australia
Received 4 April 1995; accepted 8 June 1995
Abstract
The calcine leaching circuit at the Pasminco Metals - - E Z plant has recently been extensively upgraded, with the pachuca leaching and iron purification stages being replaced. The modernised leaching circuit is a conventional counter-current plant, composed of two low-acid calcine leaching stages, (neutral leach and weak-acid leach), an iron precipitation stage (jarosite precipitation) and three high-acid leaching stages (preneutralisation, hot-acid leach and strong-acid leach). Integration of the new leaching circuit and closure of the old pachuca leaching section commenced in March 1990. On completion in September 1991, the leaching capacity increased to produce 220 000 t zinc per annum. The thickening and filtration sections, in the high-acid leaching stages of the modernised leach circuit, did not perform satisfactorily from the outset, because of the presence of very fine particulate material containing high concentrations of silica. The silica-beating particles formed a recirculating load within the high-acid leaching stages, increasing the silica content of thickener feeds. The resultant low density thickener underflows, containing less than 200 g solids/l, increased recycle flows through the plant and severely limited the capacity of the filters used for filtering the secondary leach residue from the high-acid leaching stages. Zinc production was only maintained by reducing the lead and silica content of zinc concentrates and utilising all available filter capacity for secondary leach residue. Silica in the preneutralisation stage liquors is partially precipitated in the jarosite stage. The residual silica in the jarosite precipitation stage liquors was found to be a significant contributor to the particulate silica in the high-acid leaching stages. Various options were explored to increase silica precipitation in the jarosite precipitation stage, with the addition of a coagulant to this stage being the most cost effective. Thickeners previously full of low-density pulps are now devoid of pulp, with underfiows containing up to 400 g solids/1. The secondary leach residue filters now have surplus capacity, and the lead content of the residue has increased from 26 to 36%. The restriction on lead and silica in zinc concentrates has been removed.
0304-386X/95/$09.50 © 1995 Elsevier Science B.V. All rights reserved SSDI O304- 386 X ( 9 5 ) 0 0 0 2 6 - 7
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M.G. Kershaw / Hydrometallurgy 39 (1995) 129-145
1. Introduction The leaching circuit in the Pasminco Metals - - E Z (PMEZ) zinc plant has been extensively modernised. Commissioning and integration of the new circuit commenced in March 1990 and was completed by September 1991. The modernisation involved replacement of the pachuca leaching and iron purification sections by a conventional neutral leach operation. As well, an extensive upgrade of the jarosite precipitation stage and the high-acid leaching sections was undertaken to increase recovery of valuable metals. The low-acid leaching stages of the plant and the jarosite precipitation stage performed as expected from design. However, operations in the high-acid leaching stages performed poorly mainly due to the above design quantities ofjarosite in these stages and the effects of very fine silica-bearing particulates. The modernised PMEZ circuit, which has been described previously [1] [2], has a capacity of 220 000 t zinc per annum. It consists of a roasting stage converting zinc sulphide concentrates to a zinc calcine (calcine). The calcine is leached in dilute sulphuric acid to recover the contained zinc. The neutral solution from leaching the calcine is further purified to remove elements such as cadmium, copper, cobalt and nickel by cementation with zinc dust. The purified solution is cooled, and electrolysed to recover the zinc as zinc metal. The spent electrolyte, containing 165 to 175 g H2SO4/1, is returned to the leaching circuit as dilute sulphuric acid. Two residues are produced, an iron-bearing residue, in which the iron is present as jarosite and a lead-silver bearing residue or secondary leach residue.
2. The leaching circuit The leaching circuit is shown in Fig. 1. The tankage and equipment have been described previously [ 1 ] [2]. The principal aims of the new circuit are to increase zinc extraction from calcine and to reduce the heavy metal levels in jarosite, thereby increasing zinc recovery to at least 96%. These aims, coupled with the need to roast concentrate feed blends containing greater than 9% iron and to recover zinc from a stockpiled zinc ferrite residue, resulted in a relatively complex counter-current leach circuit with six stages, consisting of : 1. Two stages of low-acid leaching called neutral leach (NL) and weak-acid leach (WAL) stages; 2. Ajarositeprecipitation (JAR) stage including an improved jarosite filtration and washing step; 3. Three stages of high-acid leaching called preneutralisation (PN) (including a preneutralisation repulp step), hot-acid leach (HAL) and strong-acid leach (SAL) stages. 2.1. Calcine and spent electrolyte
Zinc sulphide concentrates are blended and roasted in two fluid bed roasters to produce a calcine containing zinc oxide, lead sulphate, zinc ferrite and small amounts of other compounds. Typical roaster feed and calcine analyses are given in Table 1.
LEGEND
CONC, HzSO
SPENT ELECTROLYTE STOCKPILED FERRITE RESIDUE
:
I--I
OPERATION
l
THICKENING
~
UN"
STEP
CONC.H2SO~
CONCENTRATES
J CALCINE STORAGE J CALCINE SECONDARY LEACHRESIDUE J
CALCINE
SPENT ELECTROLYTE
ZINC DUST
LIMESTONE'~IM
',o
SPENT ELECTROLYTE RETURNEDTO LEACHING
CATHODEZINC JAROSITE
~
ALLOY METALS
• ZINC DUST TO PURIFICATION ZINC METAL AND ALLOYS Fig. 1. Flowsheet for Pasminco Metals - - E Z modemised circuit.
M. G. Kershaw / Hydrometallurgy 39 (1995) 129-145
132
Table 1 Roaster feed and calcine compositions ( % )
Zn Pb Fe SiO2 ZnF~O4 ZnO PbSO4
Roaster feed
Calcine
51.2 2.35 9.7 1.3 -
57.6 2.6 11.0 1.4 24.3 60.9 3.8
The calcine is wet ground at a rate of up to 50 t/h in one of two 2.1 m × 3.05 m ball mills in NL solution to produce a slurry containing 500 g calcine/l, with 80% passing 90 microns. The slurry is pumped to various stages within the leaching plant. Typical calcine and spent electrolyte distributions are: Stage Calcine (%) Spent electrolyte (%) Neutral Leach 41 18 Weak-acid Leach 25 6 Jarosite Precipitation 18 6 Preneutralisation 16 0 Hot-acid Leach 0 0 Strong-acid Leach 0 70
2.2. Low-acid leaching stages The objectives of the low-acid leaching stages are to produce a NL thickener liquor with low concentrations of iron, arsenic, etc., for the zinc dust purification stage and a residue from the WAL stage, containing low levels of zinc oxide, to be further leached in the preneutralisation stage.
Process description of low-acid leaching stages Feed streams to the NL stage are spent electrolyte and solution from the WAL stage. Calcine is added, under pH control, to achieve a pH of between 5.0 and 5.3. The principal reactions in the NL stage are: (A) Neutralisation of sulphuric acid: ZnO +
H2SO 4 = ZnSO 4 +
H20
(B) Oxidation of ferrous iron to ferric iron using air sparging, followed by precipitation of the ferric iron as ferric hydroxide: 4Fe 2 + + 02 + 2H20
=
4Fe 3+ + 4 O H -
Fe 3+ + O H - + H z O + Z n O = F e ( O H ) 3 + Z n 2+ (C) Co-precipitation of a number of impurities such as arsenic and antimony, which are toxic to zinc electrowinning.
M. G. Kershaw / Hydrometallurgy 39 (1995) 129-145
133
The feeds to the WAL stage are liquor from the JAR stage and NL thickener underflow pulp, with the pH being controlled by the addition of either spent electrolyte or calcine slurry. The WAL stage is designed to dissolve most of the zinc oxide in the NL solids and added calcine and give a preliminary purification of JAR solution by partially precipitating elements such as arsenic and antimony.
Design criteria for the low-acid leaching stages The design of the NL stage was based on the need to produce a solution containing < 1 mg Fe 2÷/1, < 10 mg total iron/1 and < 0.5 mg As/1 for the subsequent zinc dust purification stages. Oxidation of ferrous iron is the rate controlling step and the residence time required was based on previously developed models [ 3,4], in which the oxidation rate is a function of the ferrous iron concentration and the liquor pH. Tankage was installed to oxidise up to 0.5 g Fe 2÷/1 in the combined feed stream of spent electrolyte plus liquor from the WAL stage. Air at 200 kPa was supplied to the tanks to oxidise the ferrous iron. The WAL stage was designed to precipitate ferric iron from JAR solution as ferric hydroxide at a pH between 3.0 and 3.5. Under these conditions, the extent of arsenic and antimony precipitation was expected to be greater than 97% and 69%, respectively, and the overall zinc oxide extraction from calcine to be 93%.
Performance of the low-acid leaching stages The ferrous iron oxidation in the NL tanks is better than design (Table 2), which has been attributed to the more efficient gas dispersion properties of the installed agitators, compared to the Rushton turbine impellers used in the pilot plant studies. The liquor, after solids/liquid separation in two stages of thickeners, contains 146 g Zn/1, 26 mg total iron/ 1, 1.1 mg soluble Fe/l, 1.5 mg As/1 and 100 mg particulates/1 and is forwarded to the zinc dust purification stages. In the NL stage, the extraction of the zinc oxide component from calcine has averaged 83%, compared to design of 85%. The performance of the WAL stage is measured on the overall zinc oxide extraction from calcine. The extraction has averaged 87%, which is lower than design as a result of operating at a higher pH. The operating pH was increased by 0.5 due to higher than design flows of JAR liquor and poor agitation in the WAL tanks. The flow of JAR stage liquor to the WAL Table 2 Ferrous iron concentrations in the NL stage Tank
Premix Tank NL Tank 1 NL Tank 2 NL Tank 3 NL Tank 4 NL Tank 5 *From pilot plant data.
pH
5.1 5.3 5.3 5.3 5.3
Fe 2+ (mg/l) Actual
Expected *
800 40 l 1 1 < 1
800 335 102 31 9 3
134
M.G. Kershaw/ Hydrometallurgy39 (1995) 129-145
stage is nearly twice design, due to the need to clarify jarosite filtrate (designed to bypass the WAL stage) and the higher than design flow of spent electrolyte to the SAL stage. Arsenic and antimony precipitation levels of 98% and 90%, respectively, are better than design. Improvements in NL stage clarification and agitation in the WAL tanks are expected to bring the performance of these two stages up to design expectations. 2.3. Jarosite precipitation stage
Clarified PN liquor is pumped to the first JAR tank, where aqueous ammonia solution and jarosite seed are added. Calcine slurry is added to the second, third and fourth tanks to control the acidity. The ferric iron (dissolved in the high-acid leach stages) is precipitated at 97°C, predominantly as ammonium jarosite with minor amounts of other jarosites also being precipitated (e.g. sodium and potassium), by the reaction: M + +3Fe 3+ +2SOZ4- + 6 H 2 0 = MFe3(SOn)2(OH)6+6H + where M = NH4, K, Na, H30. The acidity is increased in the sixth JAR tank to 20 g HzSO4/I with spent electrolyte. The acidification reduces scale formation in the thickeners and pumping lines and dissolves species, which interfere with flocculation. The solids are removed by thickening and the thickener underflow slurry is neutralised with limestone prior to filtration on three 512 m 2 Rittershaus and Blecher recessed chamber filter presses. JAR stage design criteria The pilot plant studies showed that the rate of iron precipitation is a function of temperature and concentration of ferric iron, acidity and ammonium ion concentration [5]. For the design conditions of 10 g H2SO4/1, 3 g NH~-/1 and 97°C a residence time of 4.4 h was required to precipitate 80% of the ferric iron as jarosite from the PN liquor. JAR performance About 75% of the ferric iron in the PN liquor is precipitated, giving a solution containing between 4 and 7 g Fe 3 ÷/1. The precipitation occurs at an acidity of 3 to 6 g H2SO4/1 and an ammonium ion concentration of 2 to 2.5 g NH~-/1. The actual residence time in the JAR stage is one hour less than design. The effects of the lower residence time and lower ammonium ion concentration have been small and have been offset by operating the JAR stage at a lower acidity. Typical compositions of the jarosite filter cake and thickener overflow are given in Table 3 and Table 4. 2.4. High-acid leaching stages
The high-acid leaching stages are the PN, HAL and SAL stages in which the acidity ranges from 5 to 160 g H2SO4]1.
M.G. Kershaw/ Hydrometallurgy39 (1995) 129-145
135
Table 3 Jarosite filter cake analyses Component
Zn Fe Pb AI SiO2 Ca
Composition (%) Design
Plant
3.2 28.6 1.1 0.45 1.8 5.9
3.5 24.9 1.3 0.42 2.4 6.3
Table 4 JAR stage thickener overflowliquor composition Component
Zn Fe3+ Fe2+ H2SO4 NH4 SiO2
Composition (g/l) Pilot plant
Plant
150.5 5.7 1.5 11.8 3.5 0.85
125.7 6 0.5 20 2.1 0.88
Process description of the high-acid leaching stages The operating acidity in the PN stage is the lowest of the high-acid leach stages. The feeds to the PN stage are liquor from the H A L stage and W A L thickener underflow pulp. The main purpose of the PN stage is to reduce the acidity of the H A L stage thickener liquor as low as possible without precipitating ferric iron as jarosite. The zinc oxide and ferric hydroxide components in the W A L solids are leached with preneutralisation repulp (PN repulp) thickener overflow liquor at 20 to 25 g HzSO4/1. The acidity of the slurry is further reduced by calcine additions to an acidity of 5 g H2SO4/1 before being thickened in a high capacity thickener. The liquor from the thickener is further clarified to minimise carry-over of calcine residue to the JAR stage. Both PN thickener underflow and PN clarifier underflow streams are pumped to the PN repulp thickener where they are mixed with H A L liquor, to effect an acid wash of the solids. The H A L stage is the second of the two high-acid leaching stages, designed primarily for the dissolution of the zinc ferrite from calcine residue and stockpiled zinc ferrite residue, according to: ZnFe/O4 + 4H2SO4 = ZnSO4 + Fe2(504) 3 df.4H20 PN repulp solids are leached in liquor from the SAL stage at 85°C and an acidity ranging from 120 g H2SO4/1 in the first tank to 70 g H2SO4/1 in the fifth tank. The stockpiled zinc ferrite can be added to the second leach tank for acidity control. After thickening, the
136
M.G. Kershaw/ HydrometaUurgy39 (1995)129-145
clarified liquor is pumped to the PN repulp step and the thickened pulp is pumped to the SAL stage. The thickened slurry from the HAL stage is leached in the SAL stage in a mixture of contact sulphuric acid and hot spent electrolyte (heated to 85°C by steam in plate heat exchangers) to dissolve zinc ferrite and jarosite from the HAL solids, increasing the zinc ferrite extraction from the zinc residues to 95%. The thickened SAL solids are filtered and washed on three 35 m 2 Hoesch pressure filters. The filters have excess capacity and can easily filter the 75 t/d of secondary leach residue. This residue is a valuable by-product containing greater than 80% of the lead and silver in the calcine added to the leaching circuit.
Design criteria for the high-acid leaching stages The design criteria for the PN stage were, 100% dissolution of the ferric hydroxide and zinc oxide from the WAL solids and calcine with a final acidity in the PN liquor of 5 g H2SO4/1 and with less than 5% precipitation of iron as jarosite, so that the PN solids contain less than 8.5% iron as jarosite. Pilot plant studies indicated that jarosite formation in the PN stage would be of concern, with scale formation rates of 1 mm/day being measured. A spent electrolyte cleaning system was installed to descale the PN tanks and thickener on a weekly cycle. The dissolution of zinc ferrite is relatively slow. To maximise the amount of zinc ferrite dissolved in the HAL stage, the acidity in SAL liquor must be reduced as low as possible. However, jarosite precipitation is favoured by low acidity and long residence times. The pilot plant studies showed that the optimum minimum acidity in the HAL tanks was 60 g H2SO4/1. At this acidity, 80% of the zinc ferrite in the PN repulp solids and in 100 t/d of stockpiled zinc ferrite could be dissolved in the five 250 m 3 tanks. Under these conditions limited dissolution of jarosite from the PN repulp solids occurred in the first HAL tank followed by some jarosite precipitation in the later tanks and thickener. The design of the HAL plant assumed no net jarosite formation. The SAL stage was designed to dissolve at least 90% of the iron contained in the HAL solids, from both jarosite and zinc ferrite. The minimum acidity and temperature from the pilot plant studies was 140 g HzSO4/1 and 95°C, respectively, with a residence time of 3 hours. Two 35 m 2 Hoesch pressure filters were initially installed to filter up to 100 t/d of secondary leach residue. Performance of the high-acid leaching stages It has proven to be very difficult to control the acidity in the second part of the PN stage at 5 g H2SO4/1 due to the residence time in the reactors being too short, compared to the frequency of acidity determinations by the on-stream analysers. The actual acidity is in the range of 5 to 10 g H2SO4/1. Small quantities of zinc oxide remain in the PN solids with the amount increasing as the quantity of calcine required in the PN stage for acidity control increases. The main causes of this problem are short circuiting of the calcine in the tanks and the residence time being too low for complete reaction of the zinc oxide contained in the calcine. The zinc oxide in the PN solids is dissolved in the PN repulp step.
M.G. Kershaw / Hydrometallurgy 39 (1995) 129-145
137
Table 5 PN stage solid and liquor analyses PN solids (%)
Component
Component
Pilot plant
Plant
Zn Fe Pb NH4 K SiO2
17.1 28.7 7.7 0.28 0.18 3.6
17.4 31.9 6.0 0.33 0.22 3.6
Zn Fe 3÷ H2SO4 NIL, K SiO2
PN liquor (g/l) Pilot Plant
Plant
117.6 36.4 14.2 3.9 0.13 4.6
117.0 25.0 9.6 2.5 0.03 2.6
"includes particulate silica. Table 6 HAL stage solid and liquor analyses Component
Zn Fe Pb SiO2 K NH4 ZnFe2Oa K jarosite NH4 jarosite
Solid composition (%) Pilot plant
Plant
9.5 17.5 17.2 9.1 0.75 0.35 25.7 9.2 9.3
6.8 19.9 18.2 13.0 0.63 0.50 23.1 7.7 13.3
Component
Liquor composition (g/l) Pilot plant
Plant
K NFL,
91.2 42.6 1.8 56.1 2.9 0.24 3.8
70 35.7 0.4 79.6 2.0 0.24 2.8
Component
Liquor composition ( g / 1)
Zn Fe 3+ f e 2+ H2504
*includes particulate silica. Table 7 SAL stage solid and liquor analyses Component
Zn Fe Pb SiO2 K NIL, ZnF~O4 K jarosite NI-~ jarosite
Solid composition (%) Pilot plant
Plant 1994
6.4 4.3 37.0 11.1 0.07 0.08 9.0 < 1 <2
4.0 2.9 38.3 13.4 0.07 0.09 6.0 < 1 <2
* Si02 analysis includes particulates.
Zn Fe 3+ f e 2+
H2SO4 K NH4
Pilot plant
Plant 1994
61.0 16.3 0.3 175.6 3.2 0.55 3.8
49.0 9.1 0.1 168.0 1.7 0.15 2.8
138
M.G. Kershaw/ HydrometaUurgy39 (1995) 129-145
Iron, as jarosite, in the PN solids averages 11.2%. Half of this iron is from jarosite in the JAR stage thickener overflow liquor which is settled in the WAL stage thickeners. The seeding effect of this jarosite in the PN stage is unknown. The other half of the jarosite is precipitated in the PN stage. In general, the operations of the PN stage are similar to design, albeit at a higher acidity. If the acidity in the PN stage could be controlled at 5 g H2SO4/1,the calcine to the JAR stage could be reduced by 10%, increasing the overall zinc recovery from calcine in the leaching circuit by 0.3%. Solid and liquor analyses from the PN stage are compared to the pilot plant data in Table 5. The HAL and SAL stages in the two years following commissioning operated well below design expectations. The problems experienced are discussed in Section 3. The performance of the HAL stage is now within design expectations, with the zinc ferrite extraction averaging 84% and a net dissolution of both ammonium jarosite (50%) and potassium jarosite (6%). Dissolution ofjarosite occurs mainly in the first HAL tank where the acidity is usually greater than 100 g H2SO4/1. Typical compositions of HAL solids and liquors are given in Table 6. Since October 1993, the performance of the SAL stage has been close to design, with 85% iron dissolution from the HAL solids at a minimum acidity of 150 g H2SO4/1. Additional leaching occurs in the thickening step prior to filtration, increasing iron extraction to 87%. The lower than expected iron extraction has been attributed to a reduced residence time of 2 hours, because of the need to increase the acidity in the SAL stage by increasing the flowrate of spent electrolyte (see Section 3). Typical solid and liquor compositions for the SAL stage are given in Table 7.
3. Problems since commissioning The major problems since commissioning which limited the capacity of the leaching circuit were, the excessive quantity of solids to be filtered on the Hoesch filters and the poor settling and filtering characteristics of the HAL and SAL solids. 3.1. Poor settling and filtration
Settling in the SAL, HAL and PN stages of the plant steadily deteriorated after the plant was commissioned. Settling and filtration were so bad that all the SAL thickeners (two 6.5m diameter high capacity and one 15-m diameter conventional thickeners) and the HAL thickeners (two conventional 15-m diameter thickeners) were full of low-density pulps containing < 200 g solids/1 in the thickener underflow slurry. (Design was based on > 300 g solids/l in the thickener underflow slurry and > 400 g solids/1 in the Hoesch filter feed). The overflow from the thickeners contained unsettled solids which increased the thickening problems in the following stage. To break the recycle of solids via thickener overflows, the PN repulp thickener was used as a polishing thickener for the HAL thickener overflow, with the underflow being campaign filtered on one of the jarosite filters. As well, the old secondary leach residue filtration equipment was recommissioned and 2000 m 3 of SAL pulp
M.G. Kershaw / Hydrometallurgy 39 (1995) 129-145
139
Table 8 Composition of PN, HAL and SAL solids, June 1992 Component
Zn Fe Pb SiO2 K NH 4 ZnFe204 NH4 jarosite K jarosite
Solids composition (%) PN
HAL
SAL
12.8 24.5 4.2 7.3 0.22 0.90 34.3 24.0 2.8
4.7 25.0 8.0 8.8 0.34 1.41 16.1 37.5 4.4
3.5 10.6 23.4 13.4 0.20 0.60 7.0 16.0 2.6
was stockpiled in redundant leach tankage. Even with these measures in place, the plant was not capable of sustaining zinc production because of limited Hoesch filter capacity, with 75 t/d being the maximum sustainable filtration capacity, whereas the actual secondary leach residue production was in excess of 130 t/d. Higher grade concentrates, containing less lead and silica, were purchased to partly offset the lack of filtration capacity. The analysis of the secondary leach residue being produced during this period had high residual concentrations of zinc ferrite and jarosites. Typical analyses of the solids from the PN, HAL and SAL stages at this time (June 1992) are given in Table 8. Increased leaching of zinc ferrite and dissolution of jarosite in the HAL and SAL stages were achieved by increasing the spent electrolyte flow to the SAL stage, thereby increasing the acidity in the SAL and HAL stages. Jarosite formation in the PN and HAL stages was reduced by lowering the temperature 5°C. The changes (summarised in Table 9) resulted in a 25% reduction in tonnage of material being produced (Table 10), with the extraction of zinc ferrite and jarosite approaching design levels. Table 9 Modifications to the operating parameters ofPN, HAL, SAL Parameters Design
Modified
SAL
Temperature (°C) Spent flow (m3ph) ( % of total) Acidity (g H2SO4/L)
95 < 135 60 > 140
90 < 175 70 > 155
> 60 85-90
>70 80-85
85-90
80-85
HAL
Acidity (g H2SO4/L) Temperature (°C) PN
Temperature (°C)
M.G. Kershaw /Hydrometallurgy 39 (1995) 129-145
140
Table 10 Tonnage of solids in the HAL and SAL stages before and after the changes Tonnage (tpd)
HAL solids
Total Contained jarosite zinc ferrite SiO2 PbSO4 Accounted (%)
SAL solids
Before
After
Before
389
183
133
163 63 34 46 79
55 24 29 46 84
31 9.5 18 46 79
After 93 <5 11 18 46 86
Despite these changes, the filtration rate of the secondary leach residue on the Hoesch filters continued to decline. 3.2. Silica
The silica content of the solids in the PN, HAL and SAL stages remained the only major difference between the plant in September 1992 and the pilot plant study (Table 11). A detailed mass balance for silica showed that there was a recirculating load of 12 t SiO2/d moving in the liquor stream between the SAL, HAL and PN stages (Table 12) in the thickener overflow streams and returning in the thickener underflow streams. A further analysis of the data showed that silica precipitated from the JAR stage thickener overflow solution in the WAL stage did not redissolve in the PN stage and is the source ofrecirculating silica. In the pilot plant the silica dissolved and returned to the JAR stage, where 80% precipitated, equating to 50% of the silica contained in the added calcine. However, in the plant, only 25% of the silica in calcine reported to jarosite. By increasing the deportment of silica in calcine to jarosite to 50%, the tonnage of secondary leach residue to be filtered would be reduced to less than 83 t/d. The aim of this phase of the post-commissioning optimisation was to increase silica rejection with jarosite by increasing silica precipitation in the JAR stage. Table 11 Silica deportment to solids Solids
Plant September 1992
Pilot plant
Source
% SiO2
%Pb
% SiO2
% Pb
Calcine PN HAL SAL
1.9 7.0 18.6 13.8
3.1 6.9 13.2 30.6
2.0 3.6 9.1 11.1
3.5 7.7 17.2 37.0
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141
Table 12 Silica in thickener overflow liquors (Sept. 1992) Liquor
SAI O/F HAL O/F PN repulp O/F PN O/F JAR O/F WAL O/F NL O/F
Silica concentration (g/l)
SiO2 in overflow
Unfiltered
filtered
% in solids
t/d
2.1 4.0 4.2 2.3 1.6 0.13 0.15
0.25 0.07 0.24 1.34 1.6
67.4 25.0 16.6 13.3
2.1 7.5 0.5 1.9
Silica precipitation in the JAR stage A number of text book options to increase silica precipitation involving modifications to design operating parameters were examined (e.g. increasing/decreasing the pH at which ferric iron was precipitated as jarosite, the use of silica seed material, etc.) These options were all rejected, either as being incompatible with the objectives of the modernised leaching circuit or just did not work. Thus an alternative strategy was required. In the early commissioning phase, cationic flocculants were found to settle the fine silicabearing particles in the SAL and HAL stages, but resulted in the thickener underflows having unacceptably low solids loading. The use of cationic flocculants was discontinued in November 1991, with nonionic flocculants being used instead. When the silica problem was first identified in September 1992, cationic flocculants were trialed as potential silica precipitants in the JAR stage, without success. However, lower molecular weight epichlorohydrin polyamine type condensates, with cationic behaviour, from Allied Colloids Australia Pty Ltd (e.g. Magnafloc 1597) and from Maxwell Chemicals Pty Ltd (e.g. Proset 2830) when added to material from the fifth JAR tank did result in silica precipitation. In these scoping tests, the condensates or coagulants were added at 50 mg/l of slurry and were conducted using one litre of slurry maintained at 95°C in an agitated vessel. After 30 min the treated slurry was flocculated with 5 mg of nonionic flocculant/l and settled. The silica content of the clear decant was generally less than 0.5 g/1. Similar levels of silica precipitation occurred when using polyethylene glycol supplied by ICI Australia Operations Pty Ltd. A series of pilot plant tests simulating the operations of the JAR stage were conducted using the most promising coagulants (Magnafloc 1597 and Proset 2830) to confirm the one-litre tests. The clarified liquor was used in further tests to assess if there were any adverse down-stream effects and to confirm if the benefits from increased silica removal in the JAR stage would be realised. In continuous pilot plant tests, freshly collected PN clarifier overflow liquor from the plant was pumped into the first of six 6-1 tanks, simulating the plant situation. The pilot plant results were validated against plant performance to confirm that silica behaviour was being duplicated (Table 13). The cationic coagulants were also added to the sixth tank at 50 mg/l. Both coagulants reduced the silica concentration in the sixth tank by about 60% (Table 13), at an efficiency of 20 g SiO2 precipitated per gram of coagulant. The pilot
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Table 13 SIO2 Concentrations (g/l) in the JAR stage tanks Stream
PN O / F JAR 1 JAR 2 JAR 3 JAR 4 JAR 5 JAR 6 O/F U/F
Plant ( 2 / 9 / 9 2 )
2.2 1.3 1.7 1.8 1.7 1.7 1.6 1.6 0.5
Coagulant added in 6-1 pilot plant tests NIL
50 mg/1 M 1597"
1.8 1.1 1.2 1.3 1.5 1.4 1.4 1.5 0.29
1.7 1.4 1.6 1.6 1.8 1.6 0.68 0.8 0.24
50 mg/1 P 2830" * 1.8 1.0 1.2 1.2 1.2 1.4 0.47 0.31
"M 1597 = Magnafloc 1597. ""P 2830 = Proset 2830. Table 14 Silica distribution in the leach plant after coagulant additions Stream SiO2 in J A R O/F (g/l) % SlO2 in PN solids HAL solids SLR solids
Before coagulant
After coagulant
1.6
0.9
7.0 15.7 13.8
3.6 10.6 9.3
2.3 4.2 4.0 2.1
2.2 1.7 1.6 1.2
SiO2 in O/F liquors (g/l) PN PN repulp HAL SAL O / F = thickener overflow liquor. SLR = secondary leach residue.
thickener supernatant clarity and the settling rate were greatly improved by the addition of the coagulants. The decant liquor from the pilot plant was used in batch simulations of the low-acid leaching and zinc dust purification stages. The simulations of the high-acid leaching stages all showed improved thickening and clarification characteristics with the respective solids containing less silica. Also, there were no adverse effects from the coagulants in the zinc dust purification steps. The tested coagulants were used in confirmatory plant trials, the results of which were so outstanding that within four weeks of commencing the trials, additions of the coagulant Magnafloc 1597 at a rate of 40 mg/1 had become routine. Every stage in the leaching circuit was beneficially affected.
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143
The effect on silica concentrations in the leach circuit before and after the trial is shown in Table 14. Silica behaviour Silica precipitation from PN liquor occurs in the early JAR tanks equivalent to 25% of the silica in calcine. There is normally a slight increase in silica concentration in the second, third and fourth JAR tanks due to dissolution of silica from the added calcine, equivalent to 0.3 to 0.5 g SIO2/1. With no coagulant addition there is no significant silica precipitation between the fifth tank and the thickener overflow liquor (Table 14). Silica size profiles through the PN and JAR stages showed the majority of silica ( > 90%) in the PN liquors and the early JAR tanks to pass a 0.45 micron filter. By the fifth JAR tank less than 50% of the silica passed a 0.45 micron filter and there had been a reduction in the silica passing a 0.003 micron filter (Table 15). The coagulant "precipitates" the majority of that silica coarser than 0.45 microns, reducing the silica concentration to 0.31 g/l, of which 80% has a size less than 0.003 micron. It is concluded that silica must grow to about 0.45 micron before it can be "precipitated" by the coagulant. However, no silica precipitation occurs by adding the coagulant to an acidified solution taken from the fifth JAR tank. Thus the presence of the jarosite solids and the nonionic flocculant are also important for silica "precipitation". When the residence time between the PN thickener and the addition of the coagulant was reduced by 40% the efficiency of the coagulant was reduced to less than half, resulting in a return to high recirculating loads of silica within four weeks. Silica levels in the high-acid leaching stages of the plant returned to low levels within two weeks, once the residence time was returned to normal. To effectively "precipitate" the silica with the coagulants, a residence time of at least 4.5 h is required for the silica dissolved in the PN stage to polymerise and agglomerate to greater than 0.45 microns.
Table 15 Solution silica sizing Liquor stream
pH
SiO2 concentration (g/l) Total
PN thickener O / F PN clarifier O / F JAR 1 JAR4 JAR5 JAR6 *
0.92 0.85 0.94 1.24 1.14 0.62
• After addition of 50 m g / L of Magnafloc 1597.
1.82 1.73 1.48 0.91 0.70 0.31
Passing filters sized 0.45/zm
0.003/~m
1.76 1.58 1.40 0.53 0.36 0.26
0.77 0.53 0.42 0.36 0.34 0.25
144
M.G. Kershaw/ Hydrometallurgy39 (1995) 129-145
Benefits from the use o f coagulant
The benefits from using the coagulant were very extensive, affecting nearly every stage, and included: 1. Thickening in the SAL, HAL and PN stages is now at design with all underflows containing greater than 300 g solids/1. The HAL stage is operating on only one thickener and the PN repulp step has returned to its design duty. 2. Filtration of the secondary leach residue is no longer the plant bottleneck. (All the stockpiled SAL pulp has now been retreated and filtered). 3. The lead and silver contents of the secondary leach residue are significantly higher as a result of reducing the weight of gangue material in the residue (ie, silica, jarosites, etc.) and markedly increased its commercial value as a high quality source of lead and silver. 4. Flocculation and thickening in the JAR, WAL and high-acid leach stages had greatly improved. As a result, flocculant consumption has decreased plant-wide by at least 50%. 5. A marked reduction in the jarosite content of the WAL stage solids due to lower solids content of the JAR stage thickener overflow, again as a result of improved flocculation characteristics of the JAR solids. 6. The purchase of high grade concentrates is no longer required for solids handling in the SAL and HAL stages. The major benefit has been the increase in the plant leaching capacity by 25,000 t/a of zinc to be at least 5% above design.
4. Conclusions
The lower than design leaching of zinc ferrite and jarosites in the high-acid leaching stages of the modernised leaching circuit at Pasminco Metals - - E Z has been corrected by increasing the hot spent electrolyte flow to these stages. The resultant increased flow through the subsequent stages has had a minor impact on their respective performances. However, the operations of these stages required some modification to account for the extra flow. A coagulant addition to the JAR stage was required to increase silica precipitation to reduce the build up of silica in the high-acid leaching stages. The reduced silica concentrations in both solid and liquor streams have improved all stages of thickening and clarification. Flocculant consumption has halved, with the savings of flocculant more than compensating for the cost of the coagulant. Thickening in the HAL and SAL stages is no longer a plant bottleneck. The secondary leach residue at 83 t/d is at design and can be easily filtered on the three Hoesch filters (a third was commissioned during 1992/93) with mostly only two now being required. An unexpected benefit was that the lower flocculant usage virtually eliminated flocculant break-through from the NL stage which has improved the performance of the zinc dust purification stages. The modernised leach circuit at Pasminco Metals - - E Z is now operating to design expectations.
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Acknowledgements The author wishes to thank the m a n a g e m e n t of P a s m i n c o Metals - - - E Z for permission to publish this paper, and those people i n v o l v e d with the project, the success of which is a testimony to their endeavours. T h a n k s are also extended to Dr Ian Matthew ( C o n s u l t a n t ) and Dr R. R i n g of A N S T O , for their assistance in solving the silica problem.
References [ 1] Pammenter, R.V., Gilmour, R. and Salmon, P.T., Production of zinc in the modernised Pasminco Metals - EZ Plant, Hobart, Tasmania. Int. Syrup. World Zinc '93, Hobart, Tasmania, pp. 287-293 (1993). 121 Gilmour, R., Electrolytic zinc production by Pasminco Metals --EZ. In: J.T. Woodcock and J.K. Hamilton (Editors), The Sir Maurice Mawby Memorial Volume, Australasian Mining and Metallurgy, 1:573-578 (1993). [ 3 ] Higuchi, B., Kawaguchi, Y., Asaki, Z. and Kondo, Y., Oxygenation of neutral ferrous sulphate solution by gas bubbling. Denki Kagaku Oyabi Kogyo Butsui Kagaku, (46)12:648-655 (1978). [4] Minegishi, T., Asaki, Z., Higuchi, B. and Kondo, Y., Oxidation of ferrous sulphate in weakly acidic solution by gas bubbling. Metall. Trans., B, 14B: 17-24 (1983). [5] Pammenter, R.V., Kershaw, M.G. and Horsham, T.R., The low-contaminant jarosite process --further developments and the implementation of the process. Ellis Horwood, Chichester (1986), Iron Control in Hydrometallurgy, In: J.E. Dutrizac and A.J. Monhemius (Editors), pp. 603-617.