Process Selection and Design for the Palmarejo Silver Mine J.P. Errey ABSTRACT Silver and gold bearing ore from the Palmarejo mine, located within the Mexican state of Chihuahua, contains a complex mix of electrum, native silver, silver sulphides and silver chlorides. This paper presents a case study that describes the selection, development and engineering of a tailored process plant for the recovery of silver and gold from this ore. Feasibility study modelling, metallurgical testwork and engineering studies are outlined. Options considered to optimise the recovery of the contained silver and gold are compared, based upon cost, operability, recovery and engineering aspects. Concentration by flotation was found to be the preferred primary recovery route. Intensive leaching of the produced concentrate resulted in solution grades sufficient to allow direct electrowinning to be evaluated against the more traditional Merrill Crowe process. However a flotation tail of sufficiently low economic value was unable to be produced requiring the addition of stages to allow metal extraction from this stream. The processing flow sheet, which is designed to maximise the recovery from both the flotation concentrate and tailings, and to utilise a common metal production facility, is presented. INTRODUCTION The Palmarejo Silver Gold Project is being developed by Coeur d’Alene Mines through their wholly owned subsidiary, Coeur Mexicanna S.A. de C.V. The project consists of an open pit and underground mine along with associated milling and processing facilities with a Run of Mine (ROM) throughput of 2,000,000 tonnes per annum. The project is located 265 km southwest of Chihuahua, in the southwest corner of Chihuahua State deep in the Sierra Madre Occidental Mountain Range of Mexico. It is 420 km by road from Chihuahua City. Most of the route is along the sealed Highway 127 to the township of Creel, followed by a gravel road to Palmarejo via Temoris. The Chihuahua-Pacifico rail service operates between Chihuahua and Los Mochis on the south west coast. The railway passes through the town on Temoris, which is 35 km from Palmarejo by the government maintained gravel road. This road is an extension of Highway 127 and continues through to the town of Chinipas. Airstrips for light aircraft are located at Temoris and Chinipas. A map showing the location of the Palmerjo project relative to other projects in the region is shown in Figure 1. Mining operations commenced in Palmarejo in 1818. Later in the 19th century the British constructed a railway between Palmarejo and Chinipas for the transportation of concentrate from the Palmarejo mine. Both the British mining and railway operations ceased in 1910, around the time of the Mexican revolution.

Small scale mining was intermittent throughout the 20th century with more continuous efforts made in the 1980’s. No modern exploration programs were undertaken in the district until drilling commenced in late 2003. A small milling facility capable of processing 120 tonnes per day was utilised to produce flotation concentrate, this mill last operated in 1989. Intermet Engineering Limited commenced metallurgical testwork supervision in early 2005, with engineering of the process plant beginning in late 2005. Engineering was undertaken in parallel with ongoing exploration drilling and project development and as such resulted in changes to the initial design plant throughput during 2006. This paper describes the process selection and engineering of a process plant with the emphasis on treating high sulphide silver ores by utilising direct electrowinning. Electrometals Technology Limited provided the direct electrowinning technology via the use of their EMEW powder cells. PALMAREJO MINERALISATION Mineralisation at the Palmarejo Project occurs at the intersection of the 'La Prieta' (the black one) and 'La Blanca' (the white one) structures. La Prieta trends WNW-ESE and dips 48 degrees to the SW. The La Blanca structure trends NW-SE and dips 60 degrees to the SW. The structures are filled by a quartz-vein breccia unit as a result of several phases of hydrothermal activity. Intense stockwork has developed within the rifted block at the intersection of the two faults. The veining within the La Prieta structure is considered the richer due to a high content in silver represented by bands of fine grained black sulphides (acanthite and galena). The majority of earlier underground mining was concentrated within the La Prieta structure. The La Prieta and La Blanca veins occur as polymetallic Au-Ag vein/breccias with higher grade epithermal Au-Ag overprint that adds to the formation of ore shoots within steeper dipping portions of a normal (listric) fault. This is similar to the structural setting and controls on mineralisation at other deposits in the Sierra Madre Occidental Mountains. Testwork carried out has concentrated on drill holes that have intersected either the La Prieta or La Blanca structures or both. Samples were combined to create master composites from a number of different drill holes to include both structures as well as testing the variability of individual holes. Mineralogy was completed on five drill hole head samples and on one concentrate sample produced from the master composite. Analysis identified that the contained silver predominantly occurred as either electrum or silver sulphide (acanthite) with sparodic occurrences of aurorite (Mn2+,Ag,Ca)Mn3O7·3H2O), native silver and copper-silver sulphides. Gold occurs mainly as electrum.

Mineralogical analysis of the flotation concentrate confirmed that the sample is predominately a pyrite concentrate with other base metal sulphides including galena, sphalerite, chalcopyrite, chalcocite, covellite, bornite and marcasite. Comminution testwork identified three main rock types to be treated, namely quartz vein breccia, amygdaloidal andersite and footwall sediments. Each rock type has been tested separately, allowing models for each type to be established. This ensured adequate information was available for the comminution circuit design. TREATMENT OPTIONS Palmarejo ore contains significant amounts of silver sulphide and electrum which are slow leaching, resulting in poor recoveries in typical cyanide leaching practices. Pre-treatment of the ore indicated high recoveries could be achieved from flotation, producing a high grade concentrate. A number of treatment options for the resulting flotation concentrate were investigated including:

• sale of flotation concentrate, • leaching of flotation concentrate only at site, • leaching of both flotation concentrate and tailings streams, and • direct electrowinning recovery versus Merrill Crowe recovery.

Flotation Concentrate Sale Preliminary terms for the sale of flotation concentrate were discussed with existing smelters in Mexico. These terms suggested downstream treatment of the flotation concentrate at site would be more economical. Additionally, treatment of flotation tailings was also economic and as such the option to sell concentrate was not pursued further. Flotation Concentrate Leaching Only Treatment of a flotation concentrate at site had the advantage of requiring a small leaching and recovery facility, greatly reducing the project capital cost. Despite devoting considerable effort in refining the flotation techniques, the flotation route was unable to produce a tailings stream containing grades that were uneconomic to leach. Subsequent risk analysis indicated this route was a high risk, particularly when treating oxide ores, as it relied heavily on the flotation response of the ore. Leaching of Both Flotation Concentrate and Tailings Streams Leaching of both flotation streams offered the most robust solution, being able to cope with variations in flotation response, offering the highest recovery from the resource and ultimately the greatest return to the project. However, the required leaching conditions for the two separate flotation streams are different, requiring the installation of two leaching circuits. Direct Electrowinning Technology versus Merrill Crowe Technology

The use of direct electrowinning to treat solutions produced from the leaching circuits was considered early in the project engineering phase. Electrometals Technology Limited (Electrometals) was involved in early discussions to determine the treatment options and direction for testwork programs. As testwork progressed and plant selection and sizing became available, an engineering study was completed to compare direct electrowinning versus Merrill Crowe process options. The study involved developing comparative nominal capital and operating cost estimates. Capital costs were very similar, with the direct electrowinning option having a slightly lower cost. There was a significant operating cost saving with direct electrowinning, with the estimated operating cost approximately half that for a Merrill Crowe circuit. Readers should note that these were comparative costs and do not represent current market costs. METALLURGICAL TESTWORK AND PROCESS DESIGN Six testwork campaigns were conducted over a two year period, between 2003 and 2005. These campaigns included standard bench scale testing through to two pilot plant campaigns. A summary of the results is given below. Comminution Comminution testwork was undertaken on samples from all ore types and comprised:

• unconfined compressive strength tests, • autogenous media competency, • crushing work index, • rod mill work index, • ball mill work index, • JK drop weight test parameters, • JK SMC testing, • ore SG, and • abrasion index.

A summary of the comminution data is shown in Table 1. It was concluded that Palmarejo ores would be amenable to any one of the three circuits:

• SAB circuit, semi autogenous grinding (SAG) and ball milling circuit, • SABC circuit, SAG, ball and crushing milling circuit, or • three stage crushing and ball milling.

Selection of the comminution circuit for the plant considered the following:

• capital and operating costs, • risk of critical size build-up in the SAG mill circuit with some ores, • possibility of future plant capacity expansion, and • the availability of potential second hand mills.

Single stage crushing followed by an open circuit SAG mill and a ball mill closed by hydrocyclones was selected as the preferred circuit configuration. The top ball size and charge levels for the SAG mill is anticipated to eliminate any likelihood of critical size build up in the mill. However, provision was made in the plant layout for a recycle crusher should it be required.

Flotation Initial testwork programs investigated the different circuit configurations to maximise silver and gold recovery and included:

• whole ore leaching, • gravity concentration followed by leaching of the concentrate and tailings, • flotation followed by leaching of the flotation concentrate and tailings.

The results indicated that the best recovery was achieved by flotation followed by separate leaching of both product streams. A summary of the different process routes tested is provided in Table 2. Subsequent testwork campaigns concentrated on optimising the flotation circuit configuration and reagent selection. Early flotation testwork indicated high flotation recoveries into a rougher concentrate with a mass pull to concentrate of between 6% to 10% (average 9.5%) with recoveries of between 77% to 89% for both silver and gold. Due to these high recoveries, flotation tests were carried out to try to obtain a ‘throw away tail’ from flotation, i.e. the flotation tailings could be sent straight to the tailings dam. Other tests included the use of gravity separation, followed by flotation and also controlled potential sulphidisation flotation. These tests failed to produce a tail that was of a low enough grade to be discarded, so these approaches were not pursued further. Flotation testwork then concentrated on the best option for handling the rougher concentrate following cyanide leaching. It was found with the early rougher concentrates that a considerable quantity of fine material was present that gave difficulties with settling and filtration. The addition of cleaner flotation stages indicated that the mass pull could be reduced to between 1.1% and 5.3%, (average of 3.6%), with considerable improvement in obtained grade. While the addition of cleaning lowered the recovery of both silver and gold, (averaging 80.9% and 79.1% respectively) compared to rougher flotation, the settling properties of the concentrate were found to have significantly improved. Locked-cycle testing was carried out following the batch flotation testwork to allow optimisation of the circuit configuration and reagent selection and dosage. A master composite sample prepared from three drill holes was used for this test. The test was carried out for seven cycles and gave a mass pull of 5.3% to concentrate with gold and silver recoveries of 92.0 and 85.1% respectively. Two pilot plant flotation runs were then conducted. The first was undertaken on four RC drill hole samples. The primary objective of the pilot plant run was to produce a flotation concentrate that could be leached, with the resulting leach liquor separated from the solids and sent for electrowinning testing. The flotation circuit included only roughing and scavenging, with no cleaner stage. This pilot produced a high mass pull to concentrate (17.2%) with relative low gold and silver grades, at 25 g/t and 2,297 g/t respectively. The second pilot test was carried out on combined samples from diamond drill hole and bulk underground samples. The flotation circuit for this run consisted of rougher

flotation followed by a cleaner stage. The flotation circuit initially ran well with control samples indicating high grade concentrate and low tailings grade, however after a number of hours it became apparent that a circulating load of fine gangue material had built up in the circuit which finally reported to the cleaner concentrate. This resulted in a high mass pull of cleaner concentrate which was of low grade. This highlighted the need to monitor the circulating load in the cleaner circuit at site. Additionally, the plant design includes options to feed the cleaner tail to different points within the rougher circuit. As the primary objective of the pilot program was to again produce a leached solution for electrowinning testing, it was decided to re-clean the cleaner concentrate by pumping the concentrate through the cleaner cell again, after the pilot trial was complete. This cleaning increased the concentrate grades to 23.7 g/t gold and 3,170 g/t silver. This concentrate was leached, the solids removed and the clear solution sent for electrowinning testing. A summary of results for the batch rougher tests, cleaner flotation tests, locked cycle testing and pilot plant tests are given in Table 3. Cyanide leach Leaching testwork has been carried out on all flotation concentrate and tailings samples resulting from the major testwork campaigns. Leaching of the concentrates indicated that high cyanide levels (initially 5%) were required to ensure high silver and gold recoveries. The addition of air to the leaching slurry was seen to have little impact on the cyanide requirement, however oxygen sparging was found to lower the cyanide concentration requirement to 1%. In general, the recovery of silver and gold by leaching was in excess of 97%, although one variability test on drill hole 078D only achieved concentrate leach recoveries of 44.4% for silver and 94.4 % for gold. This was a high grade sample and the leach curves from this test indicate that silver leaching was not complete at the end of the test. As such a higher cyanide concentration may be required to increase the leaching rate for this sample. This is supported as the master composite, made up from drill holes 078D, 115D and 125D achieved 97.9% silver recovery on the concentrate leach. A summary of the leaching results for the flotation concentrate and tailings is given in Table 4. The leaching of the flotation tailings samples generally yielded lower leach recoveries than achieved from the concentrate samples, these averaging 84% for gold and 72.3% for silver. Recoveries were quite variable which is partly due to the head grades treated and varied between 55% and 96.6%. Overall, combined leach recoveries from both the concentrate and tailings were 93.2% for gold and 87.9% for silver. This was seen to improve to 94.0% for gold and 91.7% for silver if the 078D variability sample is excluded. For flotation concentrate leaching, the average reagent consumption was 19.9 kg/t cyanide and 0.59 kg/t lime. Cyanide consumption was seen to vary between 0.31 kg/t

and 54.2 kg/t. The average cyanide consumption decreased in the tests using oxygen to approximately 10 kg/t. Reagent consumption for the flotation tailings leaching averaged 0.71 kg/t cyanide and 1.21 kg/t lime. Cyanide consumption varied between 0.33 kg/t and 1.64 kg/t. The average combined reagent consumption for the concentrate and tailings leach was 1.26 kg/t cyanide and 1.16 kg/t lime. Direct electrowinning Two solutions were produced from the two pilot plant trials that were conducted at the SGS Lakefield Oretest laboratory. The first solution was produced from a rougher concentrate product that was leached with a 5 % cyanide concentration leach in April 2005. The leach solution contained approximately 1,000 ppm silver and 10.3 ppm gold. The electrowinning test showed silver plate was initially produced at higher solution tenors, followed by powder production as the silver concentration dropped below 300 ppm. The second leach solution was produced from a flotation cleaner sample from the second pilot trial and leached at 1% cyanide in November 2005. This solution contained approximately 1,900 ppm silver and 18.2 ppm gold. The solution produced silver powder for the full range of testing and indicated a higher production rate per cell than the first solution tested. The second solution produced is expected to be more representative of the full scale leaching circuit product, as the process used closely represented the final plant flowsheet. That being:

• The concentrate used to produce the second solution was cleaned in the flotation circuit to remove excess gangue material prior to leaching.

• The cyanide concentration of the leach was at the planned level of 1%. Electrometals Technologies used the results from the second solution test to size the required electrowinning circuit. The advantage of being able to use the powder cells is that this style of cell can be automated and the powder can be collected in a closed filter. This helps in securing the product, as well as minimising the workforce required to work in the refinery area. Detoxification The CIL residue will contain high levels of free and weak acid dissociable (WAD) cyanide when compared to plants treating oxide gold ores. Design discharge limits for the CIL residue for WAD cyanide were set at 50 ppm. Due to the high free cyanide levels contained in the CIL residue, the recycling of some of this solution to the cyanide leach circuits was investigated. A CIL tailings thickener was proposed to recover some of this solution, followed by dilution of the thickener underflow with tailings return water prior to feeding the detoxification circuit. An economic evaluation of the capital cost versus projected operating cost savings from cyanide recovery and detoxification reagent savings was undertaken.

Cyanide destruction testing evaluated the Inco SO2/air system using 100 kg of master composite sample treated through a 7 kg batch flotation cell. The results of all of the cyanide destruction testwork are summarised in Table 5. The method used to prepare the sample for detoxification testing was quite involved due to the nature of the circuit design. The method included:

• leaching a sample of flotation concentrate for 48 hours, • addition of carbon to the leach in the last 4 hours, • removal of carbon at the end of the leach, • combining of the concentrate leach tailings and flotation tailings in a ratio

representative of the final circuit, • leaching of the combined slurry for 24 hours, • addition of carbon to the leach in the last 4 hours, and • removal of carbon at the end of the leach.

The final leach tailings were then flocculated and left to settle overnight. Excess clear solution was decanted from the slurry and fresh water added to achieve the density required for detoxification testwork.

Cyanide detoxification testwork included six batch tests and two semi-continuous tests. The batch testing was done in two separate phases, with the aim being to define the conditions required for the semi-continuous tests. The first phase of batch testing was conducted with conditions simulating a cyanide concentration in the concentrate leach of 5%, which resulted in a detoxification feed containing 870 mg/l WAD CN and a total CN of 1,100 mg/l. The testing of this slurry was undertaken at 45% w/w solids, with the pH maintained at 9.0 and metabisulphite additions of 115%, 150% and 165% of the stoichiometric requirement. These tests reduced the WAD cyanide level to 126 mg/l and as such did not achieve the target level of 10 mg/l. It was noted during the tests that the viscosity of the slurry increased dramatically during testing. It is believed that this had a significant impact on the effectiveness of the detoxification process. It was also felt that a lower slurry pH would aid the process. Subsequent to the first phase of testing, cyanide optimisation testwork had been carried out on the flotation concentrate sample. This optimisation resulted in the addition of oxygen to the concentrate and a reduction in the cyanide concentration to 1%. The second phase of detoxification testing was done using the lower cyanide level in the concentrate leach and also was planned to study the affect of pulp density on cyanide destruction. Feed slurry to the detoxification testwork therefore set at a cyanide concentration of 240 mg/l WAD CN. Three tests were conducted, with two tests at a target density of 40% solids and one at 35% solids. The two tests at 40% solids used a metabisulphite target of 150% and 200% of the stoichiometric requirement, while the test at 35% solids used 150%. The density of the tests was found to be higher than planned, at 42.5% and 37.5%, however all three tests decreased the WAD cyanide level below the target value of 10 mg/l.

A final large slurry sample was produced to allow two semi-continuous tests to be carried out. These tests were done to confirm the conditions determined from the batch testwork and to confirm final cyanide levels in the resultant slurry. Both tests were conducted at 42% solids, with the pH maintained between 8.0 and 8.5, had a residence time of 70 minutes and tested the metabisulphite stoichiometry at 150% and 165%. Both tests resulted in WAD cyanide levels of less than 2 mg/l. Thickening Two settling testwork campaigns were done by Outokumpu Technologies which concentrated on flocculant screening and dynamic settling tests. The unit rate for the flotation tailings was found to be best at 0.76 t/m²/h and for the concentrate 0.28 t/m²/h. A summary of results for the settling testwork is given in Table 6. Both campaigns indicated that the targeted underflow densities should be achieved, however overflow clarity may not be as good as expected. This is especially true for the concentrate sample that is at a high pH and has high sodium content due to the high cyanide levels. Further testwork on optimum flocculants will be done on site during commissioning. Miscellaneous A number of other metallurgical tests were carried out to collect data required for the plant design and as alternative process route. These have included the following:

• Rheology Testwork. • Oxygen Uptake Testwork. • Merrill Crowe Testwork. • Tailings Testwork

Rheology testwork has been carried out on flotation concentrate and tailings samples at different densities. This data has been used in pump and agitator designs. Oxygen uptake tests were done on both flotation concentrate and tailings samples. This was done to determine the oxygen requirement for each slurry type and in turn allow for sizing of the oxygen plant. The concentrate slurry has a high oxygen demand, 0.3207 mg/l/min for the first six hours and then slowly reduces over the next 18 hours. The tailings sample oxygen demand was quite low with a peak of 0.0286 mg/l/min. A summary of results for the oxygen uptake testwork is given in Table 7. Merrill Crowe testwork was done on high grade flotation concentrate leach liquors as an alternative process route to direct electrowinning. The testwork indicated high silver and gold precipitation rates from the solution, although some minor re-dissolution was seen over one hour. Solution tenors dropped from 2,360 ppm silver to approximately 7 ppm in 15 minutes at a zinc stoichiometry of 1:1, 0.85:1 and 1.15:1. Final tailings samples following cyanide detoxification were sent to the consulting geotechnical laboratories for testing. The results from these tests will be used for the final tailings dam design by the consultant.

Flowsheet The process developed for the Palmarejo ore is shown in Figure 2. The plant is to operate on a continuous 24 hour per day, seven days per week basis. The flowsheet has utilised conventional mineral processing techniques, but has incorporated them into novel flow sheet that utilises direct electrowinning as the primary recovery of silver in a large scale process plant. PLANT DESIGN AND ENGINEERING CONSIDERATIONS Process Plant The testwork carried out has allowed the design of the process plant for Palmarejo to proceed. The circuit design utilises conventional mineral processing techniques and consists of SAG and ball milling, flotation, leaching, CCD washing, CIL, electrowinning, cyanide destruction and disposal in a tailings facility. The Palmarejo processing plant is required to process a range of ores with varying hardness. The milling circuit has to handle a blend of hard competent amygdaloidal andesite, friable quartz vein breccia and competent footwall sediments. The comminution testwork indicates that the ore is amenable to SAG milling; however a two stage milling circuit is required to process the ore. Blending of the ore types will be important to maintain a steady throughput as well as to ensure that the mills are not damaged during changes in ore hardness being fed to the plant. Blending will also be important to ensure that silver and gold feed grades are maintained at a relative stable grade to minimise process disruptions through the flotation and electrowinning stages of the plant. The process plant was designed with a flotation circuit to concentrate the majority of the silver and gold minerals into a low weight concentrate. This gives the advantage of:

• allowing a high cyanide concentration to be utilised to treat the slow leaching silver and gold minerals in the concentrate,

• allows the production of a high grade solution that can be sent directly to electrowinning, and

• minimises the tonnage of material that must be processed through a solids – liquids separation process.

The flotation concentrate will be thickened to remove excess water and so minimise the amount of flotation reagents going forward to the leaching circuit. The concentrate will then be diluted to 50% solids w/w before it is leached at 1% cyanide concentration for 48 hours. The slurry will then be sent to a three stage CCD washing circuit to wash the high grade solution from the leached solids. The overflow solution from the CCD circuit is pumped to the electrowinning circuit. The underflow from the final CCD thickener is pumped to the head of the CIL circuit. This is done to allow additional leaching of this slurry to ensure complete leaching of the concentrate. The flotation tailings still contain sufficient silver and gold grades to warrant the leaching of this product. As such, the flotation tails and final CCD underflow are

combined in a conventional CIL circuit. The circuit consists of one leach tank with a residence time of 4 hours and seven carbon adsorption tanks with a combined residence time of 20 hours. The carbon stripping circuit consists of a 10 tonne split Anglo American (AARL) stripping circuit. The design of this circuit allows for three strips per day to be completed, however it is only planned to do 13 strips per week. The incorporation of a CIL circuit into the plant design adds considerable flexibility to the plant operation. This circuit acts as the surge control within the circuit and will smooth process upsets from the flotation, concentrate leach and electrowinning circuits. It also can handle oxide ores that may have lower flotation recoveries. Tailings from the CIL circuit will be pumped to the tailings thickener to recover excess water and the contained cyanide. The underflow solution is then diluted with tailings return water and sent to the cyanide detoxification circuit incorporating the Inco SO2/air system. The pregnant solution produced by the carbon stripping circuit is sent to the electrowinning circuit along with the overflow from the CCD circuit. The Electrometals electrowinning circuit consists of 380 EMEW Powder Electrowinning Cells. These are a fully automatic cell that produces a silver and gold powder from the feed solution. The feed grade of solution to this circuit is expected to vary between 1,500 ppm and 2,000 ppm precious metal (silver plus gold). The EMEW circuit is designed to produce 1,250 kg/t of precious metal powder, while the barren solution tenor is 50 ppm precious metal. This solution will be returned to both the concentrate leach circuit and leach circuits as dilution water. Second hand equipment The use of high quality second hand equipment was considered a priority for the project to help with the development schedule. An extensive search in 2006 for available equipment found two 5 MW autogenous grinding (AG) mills in Spain. An inspection of these mills indicated that they were in extremely good condition and comminution analysis indicated they would be suitable for the project, with a possible production throughput of 2.5 million tonnes per annum. Additionally, flotation cells were available at the same location and could be transported at the same time as the mills. Flotation rougher and cleaner banks were purchased along with the two mills, plus numerous mill spares, including pinion, girth gear and gearbox. No other second hand equipment was sourced for the project. The mills were fitted with 50 Hz motors and, as the power supply in Mexico is 60 Hz, a study was carried out to determine the most effective solution for modifying the drives to meet the power supply requirements. Additionally, the comminution study had indicated that the SAG mill would need to be variable speed to handle the variation in ore hardness. The study indicated the lowest cost and most effective solution for the SAG mill was to install variable speed drives for the two SAG mill motors. The best solution for the ball mill was to purchase two new 60 Hz motors. These were selected with a slower speed so that the mill would operate at 68% of critical speed. This was done as the configuration of the mill was not ideal for ball milling.

Site terrain The Palmarejo mine site is located in an extremely mountainous region of the Sierra Madre Occidental Mountain Range. Whilst the altitude of the final plant site is not extremely high, approximately 1000 m, the actual site consists of steep mountain slopes with very little flat ground. Due to this restriction, the plant design is extremely compact in an attempt to minimise the amount of earthworks required to provide suitable flat areas for the plant layout. Limits on the largest individual equipment piece were imposed by restricted access to the site caused by narrow and steep mountain roads. Considerable effort was expended on improving the road access from Creel to site to allow the delivery to site of the large mechanical equipment. Even with the road improvements, the mills required additional modification to enable delivery of the mill shells to site. Water supply The original plan for the process plant water supply was to build an emergency containment dam (ECD) in one of the valleys below the mine site. The original schedule was to have this complete before the last wet season prior to plant commissioning. Due to difficulties in expected dam construction and variability in rainfall, it was also decided to secure water from the Chinipas River, alongside the town of Chinipas. The Chinipas River is at an altitude of 440 metres, while the ECD is at an altitude of 790 metres. The pipeline route between the two is 17 km long and an engineering study was done to investigate the most effective pumping system for this duty. A single stage pumping station was found to offer the lowest capital and operating costs. Water from the ECD will then be pumped in two stages to supply water to the process plant. The first pumping station will deliver water to a lower water tank farm, some 5 km from the ECD with a static lift of 220 metres. The second pump station will deliver water to the main process plant area which is 1 km further up the valley with a static lift of 140 metres. Tailings return water will be pumped back from the tailings dam to a separate storage tank so this water can be distributed as dilution water for cyanide detoxification and also as make-up water in the milling circuit. This pumping system consists of a two stage pumping station that delivers water from the tailings dam 3 km to the process plant. The static lift required for this duty is 255 metres. Due to the high static head of returning tailings water to the plant, an economic analysis was undertaken into adding a final thickener after the detoxification circuit. This would reduce the amount of water required to be returned to the plant. The study indicated that the pay back period was less than 12 months, however ongoing investigations were required to determine if sufficient room was available at the process plant site for the addition of another 22 metre thickener. Power supply

Power will initially be supplied from an onsite 12 MW diesel power house, consisting of 12 x 1 MW caterpillar generators. Diesel will be trucked to site for the power house as well as for use by the mining fleet. An overland power supply will be constructed once the process plant is completed to supply power from the Mexican power grid. Provision for future expansion An integral component of the Palmarejo project strategy is further exploration at the Palmarejo site as well as exploration on leases in the immediate area around Palmarejo. It is expected that additional resources will be identified that may justify an expansion in capacity. Provision has been included in the design of the plant as follows:

• crushing circuit has been sized for throughput increase to 2.5 Mt/a, • Milling circuit has been sized for throughput increase to 2.5 Mt/a, • Milling circuit has been designed for the addition of a pebble crushing circuit

to be added, • Flotation circuit has been sized for throughput increase to 2.5 Mt/a, • Thickeners have been sized for throughput increase to 2.5 Mt/a, and • Electrometals EMEW circuit has additional area available to add 26% more

cells. The rectifiers have also been oversized for potential high grade solution treatment.

REFERENCES Ammtec Campaign – January 2004. SGS Lakefield Oretest Campaign – 9609, December 2004. SGS Lakefield Oretest Campaign – 9632, May 2005. Ammtec Ltd Campaign – A9848, September 2005. SGS Lakefield Oretest Campaign – 9745, December 2005. SGS Lakefield Oretest Campaign – 9772, December 2005. Outokumpu Technologies Pty Ltd – S559TA, July 2005. Outokumpu Technologies Pty Ltd – December 2005. Electrometals Technologies Ltd – November 2005. ACKNOWLEDGEMENTS Coeur D’Alene