simulation of sulfuric acid pressure leaching of laterite ores for the extraction of nickel and cobalt, and aqueous pressure oxidation of pyrites for the recovery of ...
Object-Oriented Simulation of Hydrometallurgical Processes Part III. Application to Leaching of Laterites and Pyrites C.T. KIRANOUDIS, N.G. VOROS, T. KRITIKOS, Z.B. MAROULIS, D. MARINOS-KOURIS, N. PAPASSIOPI, O. DIMITROPOULOU, I. PASPALIARIS, and A. KONTOPOULOS The steady-state process simulator developed by Kiranoudis et al. has been used for the detailed simulation of sulfuric acid pressure leaching of laterite ores for the extraction of nickel and cobalt, and aqueous pressure oxidation of pyrites for the recovery of gold. Advanced hydrometallurgical process models for the specific unit operations involved were developed and are appropriately described. The simulation mainly focuses on studying the overall effects of certain design parameters on the entire plant efficiency. In the case of pyrites, the autothermal performance of the pressure autoclaves can be maintained by means of the oxidized recycle stream that greatly influences the fundamental heat balances of the reactor. Flashing the reactor pulp at the exit of the autoclaves results in further precipitation of solids related to ionic equilibrium reactions. The effect of grinding is important since most reactions are facilitated by small particle diameters. The ratio of feed pyrites influences the amount of precipitation of solids in the autoclave.
I.
INTRODUCTION
PRESSURE oxidation is the most common hydrometallurgical process for treating laterites for the extraction of nickel and cobalt, and pyrites for the recovery of gold. Similar to other hydrometallurgical processes (e.g., Bayer process for the production of alumina), the flowsheets of such processes include specialized unit operations which should be accurately modeled so that meaningful results are obtained by means of a proper simulation study of the whole plant. Their importance in conjunction with the increased complexity of the entire flowsheet (compared to conventional chemical plants) makes the use of detailed simulation of such processes essential for design purposes. Simulation studies of hydrometallurgical processes have been carried out for other processes by various process simulators.[1,2] Papangelakis and Demopoulos[3,4,5] proposed a detailed mathematical model for the description of the pressure oxidation autoclaves for gold recovery from mixed pyrites. Stange[6] presented a mathematical model suitable for the simulation of the Carbon-in-Pulp (CIP) processes, where gold ions were absorbed by activated carbon. Kontopoulos and Komnitsas[7] simulated the laterite leaching process in a completely stirred tank reactor (CSTR) to study the effect on nickel and cobalt recovery and determine the optimal pulp density required. Kiranoudis et al.[1] developed an object-oriented process simulator to meet the needs of flowsheeting of complex hydrometallurgical processes, and applied it to the case of the Bayer process, for the extraction of alumina from diaspore bauxite.[8] In the current work, the simulator developed[1] has been C.T. KIRANOUDIS, N.G. VOROS, T. KRITIKOS, Research Assistants, Z.B. MAROULIS, Associate Professor, and D. MARINOSKOURIS, Professor, Department of Chemical Engineering, and N. PAPASSIOPI and O. DIMITROPOULOU, Research Assistants, I. PASPALIARIS, Assistant Professor, and A. KONTOPOULOS, Professor, Department of Mining and Metallurgical Engineering, are with the National Technical University, GR-15780, Athens, Greece. Manuscript submitted October 29, 1996. METALLURGICAL AND MATERIALS TRANSACTIONS B
used for the detailed simulation of sulfuric acid pressure leaching of laterite ores for the extraction of nickel and cobalt, and aqueous pressure oxidation of pyrites for the recovery of gold. The simulator is of a modular type regarding solution of mass and energy balance equations, while user input is carried out interactively through a specialized graphical user interface. The models used for each process unit involved in the flowsheet are appropriately described. In this way, the overall effect of certain design parameters on the entire plant efficiency could be studied and analyzed in detail. The operational parameters involved are typical for the operation of such a plant. Original performance data were supplied by LARCO SA (Larimna, Greece) and METBA SA (Athens, Greece). II.
PROCESS MODELING
A. Aqueous Pressure Oxidation of Pyrites Pressure oxidation of refractory pyritic gold ores and concentrates involves the oxidative dissolution of the arsenic and sulfur elements at high temperatures using dissolved oxygen. Refractory materials are the ones in which gold is trapped within inclusions in the sulfide ore matrices, and is, thus, not easily leached by conventional cyanidation. Multicompartment horizontal autoclaves are the preferred reactors used in these processes. A crucial aspect in the design of pressure-leach autoclaves is to enable autogeneous operation of the reactor for maximum energy efficiency. In this case, the heat produced from the exothermic sulfide oxidation reactions is sufficient to maintain the operating temperature of the autoclave. Depending on the sulfide content of the feed, preheating or cooling by the addition of water may be required.[3] Pressure oxidation is a very complex hydrometallurgical process, in which conversions take place in three phases. Oxygen mass transfer occurs from the gas phase to the aqueous phase. Solid sulfide minerals, the most important of which are pyrite, marcasite, pyrrhotite, and arsenopyrite, are oxidized by the dissolved oxygen. The surface area and, VOLUME 28B, OCTOBER 1997—795
thus, the particle size distribution of the feed solids greatly influences the rate of reaction.[4] In addition, the ferrous ion in solution is further oxidized to ferric ion. The autoclaves operate with oxygen partial pressures that provide sufficient driving force for mass transfer of oxygen into the aqueous phase.[5] To prevent inert gases from building up inside the reactor, gas is bled continuously from the autoclave. The amount depends on the purity of oxygen supply and the quantity of carbon dioxide released due to the dissolution of carbonates in the feed. The gas bleed contributes to lower oxygen utilization and major energy losses. The multicompartment horizontal autoclave feed passes through a carbonate dissolution reactor and its output is flashed in an atmospheric expander. Sometimes, a cascade of decreasing pressure flashes is used, in order to produce steam of various pressure levels for the process. Carbonates are dissolved according to the following stoichiometric reactions. Dissolution of carbonates: CaCO3 (s) 1 H2SO4 (l)
[R1]
5 CaSO4z2H2O (s) 1 CO2 (g) MgCO3 (s) 1 H2SO4 (l)
Fe2(SO4)3 (l) 1 14/3H2O (l) 5 2/3(H3O)Fe3(SO4)2(OH)6 (s) 1 5/3H2SO4 (l) Fe2(SO4)3 (l) 1 1/3PbSO4 (s) 1 4H2O (l) 5 1/3PbFe6(SO4)(OH)12 (s) 1 2H2SO4 (l) Au(FeS2) 5 Au (s)
[R12]
Au(FeAsS) 5 Au (s)
[R13]
Au(FeS) 5 Au (s)
[R14]
Vapor/liquid equilibrium for water: H2O (g) 5 H2O (l)
O2 (g) 5 O2 (l)
QL
Mineral oxidation: FeS2 (s) 1 7/2O2 (l) 1 H2O (l)
[R4]
5 FeSO4 (l) 1 H2SO4 (l) FeAsS (s) 1 13/4O2 (l) 1 3/2H2O (l)
[R5]
5 H3AsO4 (l) 1 FeSO4 (l) FeS (s) 1 2O2 (l) 5 FeSO4 (l)
[R6]
PbS (s) 1 2O2 (l) 5 PbSO4 (s)
[R7]
Homogeneous Fe2+/Fe3+ oxidation: FeSO4 (l) 1 1/4O2 (l) 1 1/2H2SO4 (l)
[R8]
5 1/2Fe2(SO4)3 (l) 1 1/2H2O (l)
5 2FeAsO4 (s) 1 3H2SO4 (l) 796—VOLUME 28B, OCTOBER 1997
[R9]
iPS
iPL
i
i
i
i
i
L
[1] [2]
where the total liquid density of the pulp is given by an empirical equation:
rL 5 r H2O (1 1 0.8W)
[3]
The pulp residence time within each chamber, and the corresponding solid and liquid volumes are calculated as follows: t 5 VR /(QS 1 QL)
[4]
VL 5 VRQL /(QL 1 QS)
[5]
VS 5 VRQS /(QL 1 QS)
[6]
Bulk oxygen partial pressure of the incoming air stream, as a function of its purity and availability, is given by the following equation: PO2 5
p(1 2 b)(P 2 P 0) p(1 2 b) 1 (1 2 p)
[7]
Gas/liquid-phase oxygen equilibrium is given by Henry’s equation as follows: PO2 5 kHC*O2
[8]
Pyrite/arsenopyrite kinetics have been studied in depth by various researchers.[9,10,11] On the basis of controlling the surface reaction, first-order kinetics is involved. Assuming spherical particles following the shrinking-core model, the corresponding mass balances for reacting pyrites are given by the following equations:
* E * E 1`
FFeS2 5 F 0FeS2
Precipitation of reaction products: Fe2(SO4)3 (l) 1 2H3AsO4 (l)
Σ (F m /r ) 5 Σ (F m )/r
[R2]
[R3]
[R15]
The mathematical model of the cascade will be formed as a repetition of the one adopted for the corresponding oxidation chamber. If mass and energy balance calculations are performed on a molar basis, the total mass flow rates of solid and liquid phases are given by the following equations:
5 MgSO4z2H2O (s) 1 CO2 (g)
Oxygen gas/liquid transfer:
[R11]
Liberation of gold ores:
QS 5
The amount of sulfuric acid required for the complete dissolution of feed carbonates, as well as the amount of carbon dioxide produced, can be evaluated through stoichiometric mass calculations. The feed to the autoclaves is a mineral concentrate that consists of solid substances (set S) and liquid substances (set L). Pyrite and arsenopyrite are the only solid ores involved in kinetic reactions and, therefore, their particle size distribution is important for modeling the process. The principal reactions that occur during the pressure oxidation of the autoclave feed (set R) are as follows.
[R10]
0 FFeAsS 5 F FeAsS
0
0 FeS2
1`
0
0 FeAsS
[0,1 2 rFeS2 /D]3 dD
[9]
[0,1 2 rFeAsS /D]3 dD
[10]
where the total molar flow rates of the remaining pyrites METALLURGICAL AND MATERIALS TRANSACTIONS B
were evaluated by integration of the total particle size distribution for still-reaction diameters. The rate of pyrite firstorder reactions are given by: rFeS2 5 kFeS2 CO2 t
[11]
rFeAsS 5 kFeAsS CO2 t
[12]
The corresponding pyrite particle size distribution density functions at the exit of the autoclave chamber are given as functions of the ones in the entrance, by the following equations: EFeS2 5 E 0FeS2 [0,1 2 rFeS2 /D]3 FFeS2 /F 0FeS2
[13]
0 0 EFeAsS 5 E FeAsS [0,1 2 rFeAsS /D]3 FFeAsS /F FeAsS
[14]
The extent of oxidation reactions can now be evaluated as follows: 0 j4 5 F FeS 2 FFeS2 2
[15]
0 j5 5 F FeAsS 2 FFeAsS
[16]
0 j6 5 F FeS
[17]
0 j7 5 F PbS
[18]
2 j8 5 kFeSO4 CO2 C FeSO VL 4
[19]
The concentration of ferrous ions involved in the second order kinetics of Eq. [19] is given by the following equation: CFeSO4 5 (F 0FeSO4 1 j4 1 j5 1 j6 2 j8)t/VL
[20]
The oxygen transfer rate coming from the gas phase must meet the demand by the consumption for the oxidation reactions. Therefore, kL (C*O2 2 CO2 ) VL 5 7/2j4 1 13/4j5
[21]
1 2j6 1 1/4j8 The extent of oxygen transfer and the consumption of impure oxygen can be evaluated as follows:
j3 5 7/2j4 1 13/4j5 1 2j6 1 1/4j8
[22]
RO2 5 j3 /(pb)
[23]
The flow rates of gaseous components of the system are given by the following equations: FO2 5 F 0O2 1 j3 (1 2 b)
[24]
FN2 5 F 0N2 1 RO2 (1 2 p)
[25]
FH2O(g) 5 (FO2 1 FN2 )P 0/(P 2 P 0)
[26]
Precipitation reactions are modeled according to the ionic equilibrium of the system that is described by means of the following empirical equations: b1 C*AsO432 5 a1C*SO22 4
[27]
b2 C*Fe31 5 a2C*SO22 4
[28]
The heat balance of the reactor chamber is given by the following equation: Q5
Σ
iPS,L
Fi miCpiT 2
Σ
iPS,L
F 0i mi CpiT 1
Σ j DH
jPR
METALLURGICAL AND MATERIALS TRANSACTIONS B
j
j
[29]
Equations [1] through [29] constitute the mathematical model of the pressure oxidation autoclave chamber. The system is highly nonlinear and special attention must be taken so that the system of equations converges in robust paths. When the pulp is flashed after the autoclaves, the ionic equilibrium may differ due to the evaporation of water, resulting in more dense concentrates. Normally, the quantity of precipitated solids resulting from Reactions [R9] to [R11] by means of Eqs. [27] and [28] must be reevaluated, as well as the conventional flash equilibrium equations. B. Cyanidation and CIP Gold Recovery In the cyanidation section of a gold recovery plant, the pulp is treated with an aqueous solution of NaCN that actually dissolves gold particles held in the ground ore of the slurry stream, in the form of aurocyanide complex ions located in the solution phase. The CIP section of the same plant consists of several agitated tanks placed in series. The slurry stream coming out of the cyanidation section of the plant is introduced to the circuit as an input stream to the first tank. The solution is contacted with activated carbon and gold is adsorbed on the active sites of its surface. Each tank is equipped with screens which allow overflowing of pulp to the next tank, while the carbon particles are retained in solution. Carbon is transferred by means of centrifugal pumps in countercurrent flow. In the first tank, pumped carbon solution is fed to screens retaining carbon, while the remaining solution returns to its origin. In the last tank, only regenerated carbon is fed to the agitated solution. Carbon retained in the first tank is fed to the elution-regeneration section, while output solution from the last tank that is free of aurocyanide ions is recycled.[12] Each cyanidation tank is treated in every simulation run as a CSTR where second-order kinetics for the formation of the aurocyanide complex ion are assumed:[13,14] rAu 5 kP(y 2 y`)2
[30]
In the case of CIP reactors, the mass balances for the adsorption of gold are given by the following equation: FS (x0 2 x) 5 FC (y 2 y0) 5 CC VR rAu
[31]
where the rate of gold ion adsorption is given by the following semiempirical equation:[6,15–17] rAu 5 kb x(y + 2 y) 2 kd y
[32]
All model parameters concerning reaction kinetics and thermodynamic equilibrium are evaluated by considering experimental data obtained from the continuous operation of the process. III.
CASE STUDIES
PRISMA has been successfully used to simulate with accuracy sulfuric acid pressure leaching of laterite ores for the extraction of nickel and cobalt, and gold recovery through aqueous pressure oxidation of arsenopyrites. Both case studies are analyzed in detail in the current section, where simulation results are also presented. All process flowsheets were initially solved using a standard set of input data (standard conditions), and then several simulation VOLUME 28B, OCTOBER 1997—797
NiO (s) 1 H2SO4 (l) 5 NiSO4 (l) 1 H2O (l)
[R20]
CoO (s) 1 H2SO4 (l) 5 CoSO4 (l) 1 H2O (l)
[R21]
Hydrolysis of aluminium and iron also takes place at the beginning, when acidity is relatively high: 2FeO(OH) (s) 1 3H2SO4 (l) 5 2Fe2(SO4)3 (l)
[R22]
1 4H2O (l) 6AlO(OH) (s) 1 9H2SO4 (l) 5 3Al2(SO4)3 (l)
[R23]
1 12H2O (l) while at higher temperatures, the produced ions of iron and aluminium precipitate according to the following chemical equations: Fe2(SO4)3 (l) 1 3H2O (l) 5 Fe2O3zSO3zH2O (s)
[R24]
1 2H2SO4 (l) Fe2O3zSO3zH2O (s) 5 Fe2O3 1 H2SO4 (l)
[R25]
3Al2(SO4)3 (l) 1 14H2O 5 3Al2O3z4SO3z9H2O (s)
[R26]
1 5H2SO4 (l) Fig. 1—Flowsheet of the laterite leaching process.
runs were performed in order to investigate the behavior of the processes under different operating conditions, feed characteristics, and/or flowsheet arrangements. A. Sulfuric Acid Pressure Leaching of Laterite Ores for the Extraction of Nickel and Cobalt The flowsheet presented in Figure 1 has been proposed as an alternative method for nickel and cobalt extraction from low-grade laterite ores.[7] All data used for the simulation of this flowsheet were based on results of batchleaching experiments performed by the same researchers, coupled with data obtained by bibliographic research. Figure 1 shows that before being fed to the leaching autoclave, the ground laterite ore is mixed with sulfuric acid and the solution is recycled through the washing section. Steam heats the autoclave in order to keep the temperature at the required level (260 7C). Slurry leaving the autoclave enters a flash unit and then it is fed to a filter where the leached solid residue is separated from the liquid phase. The filtrate, which constitutes the pregnant solution, is the main product of the circuit. Filtered solids (cake with approximately 30 pct water content) are washed in a countercurrent system comprised of five washing units. Wash water is fed to the last washer, the underflow of which constitutes the final residue which is lead to the tailings sink, while the overflow of the first washing unit—enriched in nickel, cobalt, and sulfuric ions—is recycled to the autoclave. The leaching autoclave is modeled as a stoichiometric CSTR heated with steam. The user specifies the conversion of each reaction involved and the desired operation temperature. Component flow rates in the product stream, extent of reactions, and the required steam flow rate are calculated by the model. Nickel and cobalt dissolution are described by the following chemical reactions: 798—VOLUME 28B, OCTOBER 1997
Al2(SO4)3 (l) 1 2H2O 5 3Al2O3z2SO3zH2O (s) 1 2H2O (l)
[R27]
In the washing section, the specified data needed are the desired solids concentration of the underflow (dense) slurry, the fraction of solids carried to the overflow, and the washing efficiency. The performance of the sedimentation units is studied by means of the Whilhem–Naide model. The filtration unit is modeled as a simple stream separator, where the user is asked to specify the desired pulp density of the cake, while a zero content of suspended solid particles is assumed for the filtrate. Flash units are modeled by taking into consideration the boiling-point rise of the liquid phase. The simulator prompts the user to enter the operation pressure and applies built-in algorithms to calculate the corresponding temperature of the output streams. Feed ore is considered to have the average composition of the Greek laterites—extracted from Ag. Ioannis and Euboia mines—which are treated in the GMM LARCO SA plant with a pyrometallurgical method for the production of ferronickel alloy. Standard composition is 1.2 pct NiO, 0.08 pct CoO, 50 pct FeO(OH) and 6.5 pct A1O(OH). The flowsheet was designed to treat 150 tons of dry ore on a daily basis. Make-up sulfuric acid solution concentration is estimated to satisfy the process requirement regarding the acid-to-ore ratio (A/O), whose optimum value has been found to be 0.3. The corresponding pulp density in the autoclave at the optimum level was 30 pct. Operating conditions for the process units involved are given on Table I. A detailed mass and energy balance of the flowsheet was carried out using the PRISMA simulator. Some indicative results obtained, which concern several critical parameters of the process (A/O ratio in the CSTR, pulp density of the reacting slurry, losses of nickel and cobalt in the tailings, etc.) are presented in Table II. Figure 2(a) shows the profile of the diluted components’ concentration along the washer’s overflow streams. As indicated, the recovery of diluted components increases from bottom to top of the METALLURGICAL AND MATERIALS TRANSACTIONS B
Table I.
Operating Conditions of Process Units for Ni,Co Extraction
Leaching autoclave Operating temperature Flash unit Operating pressure Filter Pulp density of the cake Solids percentage in the filtrate Washing units Underflow pulp density Efficiency
260 7C 1 bar 70 pct 0 pct 50 pct 90 pct
Table II. Indicative Simulation Results for Ni,Co Extraction (Standard Conditions) Wash water flow rate Nickel production Cobalt production Solid residue
200 m3/h 1.266 ton/h or 8.434 kg/ton solids 0.065 ton/h or 0.431 kg/ton solids 150.8 ton/h
Fig. 3—Flowsheet of the gold production process.
washing cascade. Product recovery is given in Figure 2(b), as a function of the number of washing steps used. In this way, the plant performance is investigated in terms of alternative operating schemes (i.e., washing-cascade malfunction and structural design cases). B. Gold Production from Refractory Pyrite Concentrates with the Pressure Oxidation Process
(a)
(b) Fig. 2—Simulation results for the laterite leaching process. (a) Concentration profile through the CCW steps. (b) Nickel and cobalt content in the pregnant solution vs the number of washing steps. METALLURGICAL AND MATERIALS TRANSACTIONS B
Pressure oxidation is used as an oxidative pretreatment step, aiming at liberating the refractory gold which is encapsulated in the sulfide lattice. Liberated gold is subsequently leached by cyanide and recovered by the CIP method. The flowsheet of the whole process, as simulated by PRISMA, is presented in Figure 3. It consists of the following sections: (1) pressure oxidation of the sulfides and washing of the oxidized residue; (2) gold recovery from the oxidized residue; and (3) neutralization of the acidic solution. After mixing the concentrate mixed with water, it is directed to the grinding unit. Ground concentrate is then mixed with recycled oxidized solids derived from the countercurrent washing (CCW) circuit. Recycling is a crucial parameter for the performance of the whole process for the following reasons. (1) Recycled slurry contains sulfuric acid, which decomposes the carbonates contained in the feed material (carbonates decomposition reactor). (2) Reacting sulfides (FeS2, FeAsS) exhibit a propensity to form elemental sulfur as an intermediate oxidation product. This sulfur can coat the unreacted sulfide minerals. Coated particles can then agglomerate, resulting in lower oxidation rates. Agglomeration is avoided by ‘‘diluting’’ the sulfides of the feed material with inert oxidized solids and by maintaining high pulp densities in the autoclave. (3) Oxidation of sulfides are highly exothermic reactions and cooling water must be injected to the autoclave in order to maintain the operating temperature. Recycled slurry, which is cooled during the flashing and washing VOLUME 28B, OCTOBER 1997—799
Table III.
Pressure Oxidation and Washing Circuit Operational Parameters for Gold Extraction Plant
Input Streams
Operational Parameters Pressure Oxidation Circuit
Concentrate Flow rate: 13.5 ton/h Composition FeAsS: 26 pct, FeS2: 64 pct, PbS: 2 pct, CaCO3: 1.0 pct, MgCO3: 1 pct, SiO2: 6 pct Gold distribution Au(FeAsS): 50 g/ton, Au(FeS2): 21 g/ton Water mixed with concentrate 4.5 ton/h, 35 7C Wash water (stage 3): 100 ton/h, 35 7C Wash water (stage 2): 200 ton/h, 35 7C
Oxidized residue Third thickener underflow Carbon feed Flow rate: 40 kg/h Gold concentration: 25 g/ton
Acid solution First thickener overflow
Grinding unit PSD function: Rosin–Rammler Parameters characteristic diameter 5 20 mm dispersion degree 5 1.18 Autoclave compartments Operating volume first compartment: 50 m3 compartments 2 through 6: 25 m3 Pressure Temperature O2 purity O2 utilization Flash tank Pressure Thickeners Solids in underflow
101 kPa 55 wt pct
Gold Recovery Circuit Cyanidation reactors Operating volume: 70 m3 CIP tanks Operating volume: 50 m3 Carbon inventory: 25 kg/m3 Thickener Solids concentration: 55 wt pct Neutralization Circuit First reactor Neutralization reagent: CaCO3, excess 20 pct Second reactor Neutralization reagent: Ca(OH)2, excess 20 pct Thickener Solids concentration: 35 pct
operations, provides a heat sink. In this way, the amount of cooling water is minimized and the pulp density is maintained at the required high levels. Following the mixing of the fresh concentrate with the recycled solids and the carbonate’s decomposition, the slurry is fed to the autoclave. The latter is simulated as a series of six CSTRs. Cooling water is introduced in each compartment. Oxidized slurry is discharged from the autoclave through a flash tank and water vapor is released. Slurry is then washed in a three-stage countercurrent decantation system. From the second wash thickener underflow, a part of the oxidized solids is recycled to the autoclave feed tank and mixed with the fresh concentrate. Due to the recycling of solids, the first two thickeners operate with higher load compared to the third. For this reason, additional washing water is fed countercurrently to the second thickener. Oxidized residue is fed into two cyanidation reactors, where the gold is leached from the solids with a sodium cyanide solution. Pulp is then directed to a cascade of six CIP reactors. Carbon is fed to the last stage and transferred countercurrently through the CIP cascade. Gold is transferred from the cyanide solution to the carbon. Output streams of the cyanidation-CIP circuit are the loaded carbon from the first CIP stage and the barren solution and solids from the last CIP stage. Overflow of the first thickener, containing the metals dis800—VOLUME 28B, OCTOBER 1997
1800 kPa 190 7C 99 pct 85 pct
solved in the pressure oxidation, is directed to the neutralization circuit to be treated in the neutralization tanks, first with limestone and then with lime, to precipitate ferric arsenate and metal hydroxides; the associated sulfate is removed as gypsum. Resulting sludge is thickened and the underflow is mixed with the CIP tailings and rejected. A case study used as a reference for the simulation runs corresponds to the treatment of a refractory pyrite concentrate produced at the Olympias beneficiation plant in Eastern Chalkidiki, Greece. This concentrate assays on the average 40 pct Fe, 40 pct S, 12 pct As, and 25 g/t Au. A gold plant designed to treat 13.5 ton/h is simulated by PRISMA. The recycled stream is controlled to return 106 ton/h of oxidized solids in the autoclave. In this way, the sulfur content of the solids fed to the autoclave is reduced from the initial value of 40 pct (fresh concentrate) to the recommended operating value of 4.5 pct (concentrate and recycled solids). Operating parameters for all the units included in the flowsheet are presented in Table III. Percent sulfides oxidation through the autoclave compartments is presented in Figure 4(a). Beginning with a 75 pct oxidation in the first compartment, a 99.5 pct oxidation is achieved at the exit of the autoclave. Ferric and arsenate ions, produced from the oxidation reactions, precipitate in the form of ferric arsenate and jarosite hydrolysis products. Solids precipitation in the six compartments is shown in Figure 4(b). FeAsO4 precipitation amounts to 3.5 ton/h at METALLURGICAL AND MATERIALS TRANSACTIONS B
(a)
(b) Fig. 5—(a) Cooling water addition and water evaporation in the compartments of the autoclave. (b) Pulp density variation.
Fig. 4—(a) Percent sulfides oxidation through the autoclave compartments. (b) Solids precipitation in the autoclave compartments. (c) Aqueous phase composition in the autoclave and in the flash tank.
the exit of the last autoclave chamber. Most of it (2.5 ton/h) precipitates in the first compartment. A slight redissolution (0.1 ton/h) occurs in the flash tank. All the lead produced from the oxidation of PbS in the first compartment precipitates as Pb-jarosite. A small quantity of the recycled METALLURGICAL AND MATERIALS TRANSACTIONS B
H3O-jarosite (1 ton/h) is dissolved in the first compartment, due to the H2SO4 generated from the oxidation of the sulfides. The H3O-jarosite is precipitated in the following compartments, reaching the amount of 3.1 ton/h at the exit of the autoclave. More than half of this amount (1.7 ton/h) is redissolved during the flash operation. Redissolution occurs due to the evaporation of 43 ton/h of water in the flash tank; this results in increasing the H2SO4 concentration from 51.7 to 61.8 g/L. Aqueous-phase composition in the autoclave and in the flash tank is shown in Figure 4(c). Addition of cooling water in the six compartments is shown in Figure 5(a). The highest quantity of cooling water is required in the second compartment (27 ton/h), though 75 pct of the exothermic reactions are completed in the first compartment. This is due to the fact that 90 pct of the heat generated in the first compartment is consumed in heating up the feed slurry from 40 7C to 190 7C. The amount of water evaporated in the autoclave and in the flash tank is also shown in Figure 5(a). Evaporation taking place inside the autoclave is limited (2.5 ton/h). It must be noted that this amount depends mainly on the N2 and O2 flow rates at the exit of the autoclave (ventilation gases), which are fixed through the definition of the autoclave operating parameters, namely the oxygen purity and VOLUME 28B, OCTOBER 1997—801
(a) (a)
(b) Fig. 6—(a) Washing circuit performance. (b) Gold content in the solution and on the carbon.
oxygen utilization parameters. All gas components are directed with the slurry in the flash tank and discharged in the vapor phase. Variation of the pulp density is presented in Figure 5(b). Slurry fed to the autoclave contains 56.7 pct solids. Pulp density drops to 49.3 pct in the first compartment, due mainly to the dissolution of sulfides. In the other compartments, the further decrease of the pulp density to 41.7 pct is due to the addition of cooling water. Evaporation in the flash tank increases the final pulp density to 48.7 pct. Washing circuit performance is shown in Figure 6(a). Sulfuric acid concentration decreases from 61.8 g/L in the feed stream to 0.24 g/L in the third thickener. The washing operation contributes also in cooling the slurry from 100 7C to 35 7C. Recycled slurry from the second thickener is cooled at 42 7C and contains 4.5 g/L H2SO4. The third thickener underflow is directed to the gold recovery circuit. It contains 6.7 ton/h of solids carrying the liberated gold. Gold is leached from the solids in the cyanidation reactors and the pulp is fed to the CIP cascade. Gold concentration in the feed solution amounts to 15 g/ton. The carbon is fed to the last stage with an initial concentration of 25 g/ton, and is countercurrently transferred at an average rate of 26 kg/h. Gold in solution and carbon profiles are presented in Figure 6(b). Gold concentration in solution is seen to decline rapidly from the first to the fourth stage, 802—VOLUME 28B, OCTOBER 1997
(b) Fig. 7—(a) The effect of grinding on the oxidation of sulfides. (b) Effect of gold losses in the barren CIP solution as a function of the initial feed gold concentration.
while it remains almost constant in the last two stages. Five or even four stages, instead of six, would be sufficient in the case examined. The barren solution assays 0.002 g/ton, which corresponds to 99.99 pct recovery in the CIP section. First thickener overflow is directed to the neutralization circuit. Neutralization of this acidic solution produces 42 ton/h of solids, consisting mainly of gypsum (74 wt pct). Final wastes of the process are the 6.7 ton/h of oxidized solids after gold recuperation and 42 ton/h neutralization solids. In the case examined, 3.6 kg of solids must be disposed for each kg of treated concentrate. The simulator was also used to evaluate the performance of the process when important operating parameters change. (1) Grinding of the concentrate: The importance of grinding on the oxidation of sulfides is illustrated in Figure 7(a). The concentrate ‘‘as received’’ has a Rosin–Rammler diameter of 105 m (nominal size from a Rosin–Rammler distribution plot). If this concentrate is fed to the autoclave, the maximum oxidation is 85 pct in the sixth compartment. However, when the concenMETALLURGICAL AND MATERIALS TRANSACTIONS B
trate is ground to a Rosin–Rammler diameter of 20 m, a 99.8 pct oxidation is achieved. (2) CIP circuit: In the CIP circuit, gold recovery is mainly dictated from the amount of gold adsorbed on the carbon feed. As presented in Figure 7(b), gold losses in the barren solution increase with increasing amounts of gold already adsorbed on the carbon feed. (3) Concentrate composition: The FeAsS/FeS2 molar ratio in the concentrate influences the amount of solids precipitated in the autoclave, as presented in Figure 8(a). During the treatment of 13.5 ton/h of a concentrate containing a FeAsS/FeS2 molar ratio of 0.4 (reference case), 6.7 ton/h of oxidized solids are produced. Increasing the FeAsS/FeS2 ratio to 5 results in the production of 14 ton/h of oxidized residue for the same quantity of fresh concentrate. On the contrary, the amount of neutralization solids decreases from 42 ton/h at FeAsS/FeS2 ranging from 0.4 to 25 ton/h at FeAsS/FeS2 equal to 5, as shown in Figure 8(b). The total amount of solid wastes, consisting of the solids precipitated in the autoclave and those produced in the neutralization circuit, is also decreased when the FeAsS/FeS2 ratio increases, as shown in Figure 8(c). IV.
CONCLUSIONS
The steady-state simulator developed by Kiranoudis et al.[1,8] has been successfully used for the detailed simulation of sulfuric acid pressure leaching of laterite ores for the extraction of nickel and cobalt, and pyrites for the recovery of gold. The models used for each process unit involved in the flowsheet were analyzed in detail. The overall effect on the entire plant efficiency of certain design parameters has been studied and analyzed, and specific operational problems were underlined. In the case of pyrites, the amount of feed pyrites dictates the thermal performance of the pressure autoclaves. For minerals poor in pyrites, heating is required. On the other hand, autothermal performance can be maintained by means of the oxidized recycle stream, which greatly influences the fundamental heat balances of the reactor. Flashing the reactor pulp at the exit of the autoclaves results in further precipitation of solids related to ionic equilibrium reactions. A CIP process can be safely used for gold recovery from particles coming from the cyanidation section. Grinding of feed pyrites is important, since most reactions are facilitated by small particle diameters. The ratio of feed pyrites influences the amount of precipitation of solids in the autoclave. ACKNOWLEDGMENTS This work has been supported by the Greek General Secretariat of Research and Technology, Aluminium de Gre´ce, GMM LARCO SA, and METBA Ltd. Fig. 8—(a) Effect of the arsenopyrite/pyrite ratio on the amount of neutralization solids. (b) Effect of the arsenopyrite/pyrite ratio on the total amount of solid wastes. (c) Effect of the arsenopyrite/pyrite ratio on the amount of solids precipitated.
METALLURGICAL AND MATERIALS TRANSACTIONS B
[a,b] a 1,a 2,b 1,b 2 b DHj mi
NOMENCLATURE maximum of a and b ionic equilibrium constants oxygen availability heat of reaction j molecular weight of species i VOLUME 28B, OCTOBER 1997—803
jj ri rL rS CC CCW Ci C*i CIP Cpi CSTR D Ei E0i FC Fi F0i FS kFeAsS,kFeS2,kFeSO4 kH kL kb,kd,y+ kP p P P0 PRISMA Pi Q QL Q0L QS Q0S ri RO2
extent of reaction j density of component i density of liquid phase density of solid phase carbon concentration countercurrent washing solution concentration of species i equilibrium solution concentration of species i carbon-in-pulp specific heat of species i completely stirred tank reactor particle diameter particle distribution of a slurry leaving an autoclave chamber particle distribution of a slurry entering an autoclave chamber carbon flow rate molar flow rate of species i leaving a process molar flow rate of species i entering a process solution flow rate kinetic constants for pyrite oxidation Henry equilibrium constant mass-transfer volumetric coefficient gold adsorption kinetic constants cyanidation kinetic constant oxygen purity total pressure vapor pressure Process Integrated Simulator for Metallurgical Applications partial pressure of species i heat produced liquid phase volumetric flow rate leaving the process liquid phase volumetric flow rate entering the process solid phase volumetric flow rate leaving the process solid phase volumetric flow rate entering the process rate of reaction consumption of impure oxygen
804—VOLUME 28B, OCTOBER 1997
t VL VR VS x x0 y y0 y`
retention time liquid phase volume reactor volume solid phase volume gold concentration of the pregnant leaving the CIP reactor gold concentration of the pregnant entering the CIP reactor gold concentration of carbon particles leaving the CIP reactor gold concentration of carbon particles entering the CIP reactor equilibrium gold concentration of solids REFERENCES
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