- Use of by-product gypsum
- Utilisation of fluorine
- Purification of phosphoric acid
- Production of phosphoric acid using acids other than sulfuric
- Phosphoric acid production by the electric furnace process
- Phosphoric acid production by the blast-furnace process
Use of by-product gypsum
In the production of wet-process phosphoric acid approximately 5 tonnes (dry basis) of by-product phosphogypsum is produced per tonne of P2O5 recovered as phosphoric acid. Only about 15% is reused. Phosphogypsum retains at least 20% of free water by weight. The principal means for using phosphogypsum are to:
- Make ammonium sulfate by reaction of the gypsum with ammonia and carbon dioxide.
- Make cement and sulfuric acid by calcining the gypsum with coke and clay or shale.
- Make plaster or plasterboard for building materials or make pressed or cast blocks for construction purposes.
- Use in cement as a set retardant.
- Use as a fertiliser filler.
- Use for direct application to farmland when the soil requires it.
Utilisation of fluorine
Most phosphate rocks of the fluoroapatite – type contain a significant quantity of fluorine, usually 3%-4% F by weight. In some cases up to 60% of this can be evolved during the manufacture of wet-process phosphoric acid. The remainder of the fluorine is retained in the gypsum (depending mainly on the rock composition), and most of the remainder is in the filter acid . The fluorine is evolved from various stages of phosphoric acid processes: from the reactor slurry surface, in the flash cooler, and in the concentration plant. In a dihydrate process the proportions are:
- Slurry surface: 3%-5%
- Flash cooler: 8%-10%
- Concentration plant: 15%-20% (up to 45% P2O5)
- Slurry surface: 12%-15%
- Flash cooler: 15%- 20%
In a typical case, about 50 kg of F is volatilised per tonne of P2O5, usually as SiF4 or HF, or some mixture of the two. The compositions of the vapours emitted from the reaction/filtration section and the evaporator are different. During acidulation of phosphate rock, the fluoride is first converted to hydrogen fluoride, which subsequently reacts with active silica to form silicon fluoride:
4HF + SiO2 = SiF4 + 2H2O
This is partially emitted as vapour in preference to hydrogen fluoride because of its higher vapour pressure. The remainder reacts further to form fluosilicic acid, which remains in the product acid, or insoluble fluosilicates, which are removed in the filter cake.
3SiF4 + 2H2O = 2H2SiF6+ SiO2
If the silicon tetrafluoride vapour leaving the reaction section is washed with water, this reaction will take place in the scrubber also and insoluble silica will form. The silica will cause scaling and eventually plug the scrubber. Therefore, the usual procedure is to scrub with a dilute solution of fluosilicic acid, in which silica remains in colloidal suspension. A scrubber design, such as cyclone venturi or void tower, that is virtually free of internal obstructions, which might become fouled, must be used in the first stage to remove the bulk of the fluorine. A second, highly efficient stage, such as a spray tower or tower with movable packing, removes most of the remaining traces of fluorine and avoids fouling (Figure 1). More dilute vapours such as air from the hood over the first filter stage and from the first filtrate separator can be introduced directly into the second-stage scrubber.
About two-thirds of the fluorine in the filter acid is volatilised when the acid is concentrated to merchant strength (50%-55% P2O5). Under the conditions of high temperature and reduced pressure obtained in vacuum evaporators, the fluosilicic acid decomposes and both silicon tetrafluoride and hydrogen fluoride are evolved.
H2SiF6 = 2HF + SiF4
However, at acid concentrations of less than 50% P2O5, more silicon tetrafluoride is evolved (creating silica in circulated fluosilicic acid) owing to the relationship between its vapour pressure and that of hydrogen fluoride.
At acid concentrations of 50%-55% P2O5, the molar ratio HF:SiF4 is approximately 2, while at higher acid concentrations hydrogen fluoride predominates . There is least danger of problems with silica deposition in the scrubber if the HF:SiF4 ratio is at least 2, since the only product of scrubbing will be fluosilicic acid, while it is undesirable to allow it to exceed 2 by too much, since the product will then contain free hydrogen fluoride. Because the scrubbing solution used is fluosilicic acid, it is desirable to maintain a high concentration from the point of view of transport costs and use of the product. To avoid dilution by the water from the evaporator, the scrubber, which is interposed between the evaporator and the condenser, must be operated at a high temperature at which water will not significantly condense. In practice, this sets a maximum of about 18%-22% H2SiF6 on the concentration of the scrubbing medium because at higher concentrations the equilibrium vapour pressures of hydrogen fluoride and silicon tetrafluoride are too high for efficient fluorine removal.
The void spray tower type of scrubber, originally developed by Swift and Co. and further developed by Fisons and others, is often used. The scrubber is conveniently mounted atop the evaporator and receives vapours from the entrainment separator, which returns any droplets or mist particles of phosphoric acid to the acid system. Fluorine scrubbers typically operate at an efficiency of 90%-93%; the remainder of the fluorine vapours is largely removed in the condenser into which the vapours subsequently pass. The condenser is usually of the barometric direct-contact type, but after cooling the water must be directed to a wastewater treatment installation to remove the contamination.
Numerous processes for recovering saleable fluorine compounds have been proposed and developed experimentally; some of these processes are used on an industrial scale. A description of all of these processes is beyond the scope of this resource.
Fluorine is usually recovered in the form of an aqueous solution of fluosilicic acid, H2SiF6. The concentration may be as high as 20%-25%. In some countries the fluosilicic acid is used directly for ‘fluoridation’ of municipal water supplies to prevent tooth decay. The acid is shipped to various municipalities in rubber-lined railcars. Salts of fluosilicic acid, such as sodium, potassium, and ammonium fluosilicates, have various uses and can be readily produced from the acid. Sodium fluosilicate is used for fluoridation of municipal water supplies, but fluosilicic acid is generally preferred.
Processes have been developed for making aluminium fluoride(AlF3) and cryolite(Na3AIF6) from fluosilicic acid, which are used by the aluminium industry in substantial quantities. However, these materials must be quite pure for this purpose; in particular, the silicon and phosphorus must be very low. This requirement complicates production from by-products of the phosphate industry. Calcium fluoride may be produced and used instead of natural fluorspar for production of hydrogen fluoride (HF), which is the basic material for producing numerous organic and inorganic fluorine compounds or for metallurgical purposes . There are also processes to produce hydrofluoric acid solutions or anhydrous hydrogen fluoride from fluosilicic acid solutions.
In general, the economics of producing chemical-grade fluorine compounds from by-product fluosilicic acid is marginal to unfavourable for small phosphoric acid plants. However, large plants may find recovery profitable. Where several phosphoric acid plants are operating within an economical shipping distance, the crude fluosilicic acid may be shipped to a central point for production of refined fluorine compounds. For more details on fluorine use, see reference .
Purification of phosphoric acid
For most fertiliser production processes, purification of wet-process phosphoric acid is not necessary. However, there are two common fertiliser uses that may call for partial purification:
‘Merchant-grade” acid that is shipped by rail, barge, or ocean vessels and is often stored at shipping and receiving terminals should be purified sufficiently so that formation of insoluble precipitates (sludge) during shipping and storage is minimised.
Phosphoric acid to be used in the production of liquid fertilisers, such as ammonium polyphosphate solution, sometimes requires partial purification to prevent formation of precipitates upon ammoniation or during storage of the ammoniated solution.
Although ammonium polyphosphate sequesters most of the common impurities, excessive amounts of some impurities (especially magnesium and organic matter) cause precipitate formation. Superphosphoric acids usually do not form sludge, but magnesium and titanium have been known to cause sludge-forming precipitates.
A major fraction of sludge in most merchant-sludge acids is the compound (Fe,Al)3KH14(PO4)8 • 4H2O. It precipitates slowly over a period of several weeks; therefore, long storage periods are required to ensure reasonable completion of the precipitation reaction.
In the production of sodium tripolyphosphate and other condensed phosphates for use as builders in detergents, the whiteness of the product and all metal ions (without sodium and potassium) are the principal concern. For use in human food processing and animal feed supplement production, the content of toxic impurities must be diminished to insignificant levels. These include fluorine, arsenic and any heavy metals that are present in the phosphate rock and sulfuric acid. When purified acid is required only for the production of phosphates, impurity removal is facilitated by neutralisation, since many impurities are rapidly precipitated as insoluble phosphates. When it is not permissible to neutralise the acid, it is necessary to rely on physical treatment such as solvent extraction or ion exchange, with or without chemical treatment. Phosphoric acid manufacturers purifying phosphoric acid for sale have inclined towards this type of treatment, usually based on solvent extraction. In most cases when chemicals are used (e.g. in the case of removing arsenic), they react with the impurity alone and do not chemically alter the phosphoric acid.
Removal of Organic Materials
In the oxidising and dehydrating environment of the attack section, organic impurities tend to be decomposed to carbon; if the rock has a relatively high organic content, the resultant acid will be black. Two methods of removing this discoloration include calcining the rock, which destroys a good proportion of the organic matter, and treating the acid with active carbon, with or without a flocculating agent. Conducting the latter treatment at a moderate temperature (60°-80°C) increases the efficiency of organics removal to the extent that it may be possible to avoid calcining some high-organic rocks. Other techniques that have been proposed rely on using a special flocculating agent or hydrogen peroxide in the presence of a mixed, metal oxides catalyst.
Even if wet-process phosphoric acid is concentrated to the superphosphoric acid stage, it still contains too much fluorine for production of feed phosphates although the other impurity levels may be acceptable for animal feed. For animal feed preparation, it is desirable that the phosphorus:fluorine ratio should be at least 100 P : 1F, but normal wet-process phosphoric acid has a P: F ratio of between 15: 1 and 54 : 1, depending on its concentration, the composition of the rock from which it is made, and its production conditions.
Several procedures have been proposed to remove fluorine from phosphoric acid based on volatilisation or chemical precipitation, in the latter case with or without other impurities.
Precipitation – In this type of process, a large amount of water-miscible organic solvent is added to the phosphoric acid, which causes settling of many of the dissolved impurities out of solution and enables the acid-solvent mixture to be separated.
Particularly, if methanol is used, alkali metal or ammonium ions must be present in sufficient quantities to convert the sludge impurities into less soluble double salts. Other solvents that could be used include acetone, methyl ethyl ketone, other lower alcohols, and dioxane.
The inconvenience of this type of process is that the phosphoric acid must be separated from a large volume of solvent, usually by stripping, which is expensive.
Extraction – The alternative method of purifying phosphoric acid by means of organic solvents is liquid/ liquid extraction; crude phosphoric acid is brought into countercurrent contact with a partially or substantially water-immiscible solvent in a series of mixer-settlers or columns. In general, the partition coefficient of phosphoric acid between the aqueous solution and the solvent is very unfavourable; thus the crude acid is usually concentrated before processing and only a relatively small proportion of the acid is extracted. The remaining partially depleted raffinate, containing most of the impurities, is usually disposed of in fertiliser manufacture.
Depending on the degree of purification required, the extract may be washed with purified phosphoric acid and is then usually contacted with water into which the majority of the acid passes. The end product is a somewhat weaker but still purer acid than the original crude acid and a depleted solvent phase, which is recycled. ·
When there are no possibilities to use raffinate in fertiliser production, it must be strongly acidified with sulfuric acid before additional extraction; most of the P2O5 is driven into the solvent phase. Impurities remain in the acidified aqueous phase, which is neutralised with lime before disposal as waste.
Production of phosphoric acid using acids other than sulfuric
Phosphate rock can be dissolved by several organic and inorganic acids to produce phosphoric acid. Commercial nitrophosphate processes produce phosphoric acid that contains nitrates· hence these processes are used to produce compound NP or NPK fertilisers. It is technically feasible to produce phosphoric acid that is substantially free of calcium or nitrates by separation methods involving solvent extraction. One such process using tertiary amyl alcohol as the solvent was developed in Finland and described by Lounamaa ; however, no commercial use has been reported.
Several processes using hydrochloric acid have been developed or patented, but few have been used commercially. The main stages of such a process is.
Dissolution of phosphate rock by hydrochloric acid, which results in an aqueous solution of calcium chloride and phosphoric acid;
Liquid-liquid contacting in a number of solvent extraction steps to obtain a solution of substantially pure phosphoric acid; and
Acid concentration to obtain 95% H3PO4 (69% P2O5).
The raw materials and reagents used are as follows:
- Phosphate rock (any commercial grade). The P2O5 recovery is more than 98%.
- Hydrochloric acid for acidulation can be used as a solution of 20% HCl or higher or in gaseous form by combining absorption with reaction. For economic reasons, concentrated acid is preferred because most of the water accompanying the acid must be evaporated in a later stage of the process. HCl consumption is dependent on the composition of the rock.
- Solvent. Several solvents can be used for extraction. Those preferred are technical isoamyl alcohol (IM), n-butanol, or a mixture of both. Solvent makeup is 4 kg/tonne P2O5.
- Process water.
- Auxiliary reagents. Depending on the type of rock and the method of separation of insoluble residue from dissolution liquor, minor quantities of filter aids or flocculating agents may be required.
Dissolution and Mechanical Separation of Insoluble Residue
The dissolution of phosphate rock is essentially decomposition of fluorapatite by HCI according to the following equation:
Ca10F2(PO4)6 + 20 HCl -> 2HF + 6H3PO4 + 10CaCl2
Other acid-soluble components of the rock, such CaCO3, decompose simultaneously. The rock is dissolved by hydrochloric acid. The insoluble residue amounts to a small percentage of the rock feed and consists mainly of silica, silicates, and insoluble organic matter. The insoluble matter can be separated from the dissolution liquor by filtration, followed by washing of the cake or by sedimentation in a thickener, followed by countercurrent decantation washing of the sediment The choice of the proper method of separating solids depends on the character of the insoluble residue and on economic considerations. The dissolution liquor is fed to the subsequent section.
Figure 2 shows a typical flowsheet of the HCl phosphoric acid process.
This stage consists of several operations: extraction, purification, washing, and stripping.
Extraction – Extraction is performed by a countercurrent contact of dissolution liquor with the selected solvent. Phosphoric acid transfers selectively from the aqueous dissolution liquor to the organic solvent phase; the resultant extract and calcium-chloride brine (raffinate) contain substantially all the impurities, such as fluorine and iron.
Purification -The solvent extract, which contains small amounts of Ca++ and some other impurities, is purified further by countercurrent contact with an aqueous phase.
Washing- The acid of the purified extract is transferred into water. The solvent that leaves this stage is virtually free of acid.
Stripping- The acid-free solvent stream extracts the residual acids present in the raffinate and is recycled to extraction. The spent calcium-chloride brine is stripped by steam to recover any dissolved solvent.
The dilute aqueous acid emerging from washing consists of an aqueous solution of H3PO4, HCl, and some dissolved solvent. This solution is concentrated to 95% H3PO4, which is the end product. The separation of H3PO4 from other components of the solution is essentially a distillation operation; this permits a full recovery of the minor quantity of solvent that is dissolved in the aqueous phase on washing and of the HCl, both maintained in closed cycle in the process.
The main requirement of this operation is heat economy and a multiple-effect evaporator is used to achieve this. The amount of steam used is less than 0.5 tonne/tonne of water evaporated. The complete absence of dissolved solids in the solution being concentrated permits the maintenance of high heat-transfer coefficients. All of the volatile streams from the process are recycled to the previous steps of the process.
Solvent Recovery from Spent Calcium Chloride Brine
The residual brine leaving the stripping stage contains a small amount of dissolved solvent, which must be recovered for economic reasons. The solvents used form an azeotrope with water on rectification so that the simplest system to be applied is steam stripping. The costs of this operation are reduced by the recovery of heat from the brine leaving the system. The recovered solvent is recycled to the liquid-liquid contacting section and the brine is discarded.
In those parts of the process where solvent is present, acid-resistant construction materials that are also solvent resistant are required. These parts included the liquid-liquid contacting section and a part of the sections where acid is concentrated and where solvent is recovered from the brine.
Rubber-lined steel is the least expensive material for the dissolution and mechanical separation of insoluble residue. In the liquid-liquid contacting section, rigid polyvinyl chloride (PVC) is satisfactory. In the parts of the system operating at elevated temperatures, impervious graphite can be used for the heat exchangers. Other construction materials include thermosetting resins and lined steel.
Quality of HCl-Route Phosphoric Acid
HCl-route phosphoric acid is much cleaner than wet-process acid, and its analysis is similar to that of thermal acid (Table 1). By making slight adjustments in the process, food-grade acid can be obtained. The composition of wet-process acid is dependent on the rock as raw material; whereas, almost the opposite is true for HCl-route phosphoric acid.
The capital investment required for a plant may vary from location to location. However; for comparative purposes, it will be noted that the capital cost for the HCl process is about 35% higher than for the standard (H2SO4) wet process when production of the acids is excluded.
If HCl were available as a by-product from another process, the capital cost would be lower than for a wet process plant including H2SO4 production facilities. However, when using by-product HCl, the scale of the operation would be limited by the amount of by-product HCl available.
Operating costs may be estimated from the process requirements, which are given in Table 2 for a plant of 100-tpd capacity.
Disposal of CaCl2 may be difficult and expensive.
HCl-route phosphoric acid has certain disadvantages as compared with wet-process acid. Its production is economic only in places where HCl is available or where it can be produced at a moderate price. Transport of HCl in the form of an aqueous solution of perhaps 33% HCl is possible only in pipes or railcars lined with rubber, PVC, or similar materials.
However, HCl-route phosphoric acid has some advantages over wet-process acid. Unlike wet-process acid, it contains no scale-forming components, and its composition and quality are practically independent of the· type of phosphate rock used. Superphosphoric acid (70%-72% P2O5) can easily be produced from HCl-route phosphoric acid, without any additional purification process.
HCI can be used where it is available as a byproduct. This is important for countries producing NaOH, where there is no captive market for the chlorine that is produced simultaneously. Byproduct hydrochloric acid is sometimes available from other sources and may even create disposal problems. In such cases the production of phosphoric acid by the acidulation of phosphate rock with hydrochloric acid can be advantageous if the quantity of the byproduct is adequate for an economical scale of operation and disposal or use of the calcium chloride brine is economically feasible.
An interesting possible source of hydrochloric acid is through calcination and hydrolysis of magnesium chloride according to the following equation:
MgCl2 + H2O = 2HCl + MgO
Magnesium oxide could be useful for production of refractories.
Another possible source of hydrochloric acid is from the production of potassium phosphate from phosphoric acid and potassium chloride.
At present the only plants using the HCl process are relatively small ones, and most of the product is used to make industrial phosphates rather than fertilisers.
As an alternative route and process, dicalcium phosphate (DCP) can be produced from phosphate rock and HCl. DCP can be used for animals feed purpose or it can also be reprocessed with sulfuric acid to produce phosphoric acid.
Phosphoric acid production by the electric furnace process
The first step in the production of furnace acid is to produce elemental phosphorus in an electric furnace (Figure 3). Phosphate nodules or other lump phosphate material, silica pebble, and coke are mixed and fed to the furnace. The electric current that is supplied to the furnace through carbon or graphite electrodes fuses the rock and silica, and the carbon in the coke reduces the phosphate. A mixture of phosphorus vapour and carbon monoxide gas is withdrawn continuously from the furnace. The phosphorus is condensed to a liquid that is converted to phosphoric acid in a separate plant, often located far from the phosphorus plant. Molten calcium silicate slag and an iron-phosphorus compound known as ferrophosphorus are tapped from the furnace periodically. The following- equation represents the principal reaction in the furnace:
Ca10F2(PO4)6 + 15C + 6SiO2 à 1.5P4 + 15CO + 3(3CaO•2SiO2) + CaF2
One advantage of the furnace process is its ability to use low-grade phosphate rock provided that the principal impurity is silica. Iron oxide and alumina are much less objectionable in the furnace process than in the wet process. Siliceous phosphate rock containing about 24% P2O5 is used in several plants; such rock may be obtained at a very low cost in some locations. Rock containing as much as 7% Al2O3 is acceptable.
lf lump or pebble rock of suitable size (about 0.6-4.0 cm) that is resistant to decrepitation on heating is available, the cost of agglomerating the charge may be avoided. However, such rock is seldom available; therefore, the rock usually is agglomerated and calcined or sintered before charging it to the furnace. Carbon monoxide gas, which is a by-product from the furnace, is the usual fuel for the calcination. Even so, this step is expensive.
The recovery of P2O5 as elemental phosphorus usually is in the range of 86%-92% of that charged to the furnace. The loss of P2O5 in the slag is about 3%. From 2%-8% of the P2O5 charged is recovered as ferrophosphorus, which contains about 23% phosphorus, 70% iron, and small amounts of manganese, silicon, and other metallic elements, depending on the charge composition. The amount of ferrophosphorus formed depends on the iron oxide content of the charge. The ferrophosphorus is sold to the steel industry, but the income only partially compensates for the loss of phosphorus production.
Of the phosphorus recovered as elemental phosphorus, about 5% is in the form of sludge, even after a series of settling steps to separate sludge from clean phosphorus. This sludge phosphorus may be recovered by burning it separately to produce impure phosphoric acid, by distillation or by dewatering and returning it to the furnace.
The main disadvantages of the furnace process are the relatively high capital cost of the plant and the scarcity of locations where low-cost electricity is available. For this reason the electric furnace process is used almost exclusively to produce phosphorus and phosphoric acid for industrial chemicals, insecticides, detergents, and food or animal feed additives.
Production of phosphoric acid from elemental phosphorus is relatively simple. It is carried out by burning liquid elemental phosphorus in air and hydrating the resulting P2O5 to H3PO4. A diagram of a typical plant is shown in Figure 4. All the process equipment is made of stainless steel, usually type 316. The overall reaction is:
P4 + 5O2 +6H2O à 4H3PO4
Typical process requirements per tonne of P2O5 produced as phosphoric acid, assuming 86% overall recovery for a plant of about 100,000 tonnes P2O5/year capacity, are given in Table 3.
Fuel consumption is negligible – about 1.2 million kcal, depending on efficiency of use of by-product carbon monoxide gas.
If phosphate rock is available that is satisfactory for use without agglomeration and calcining, the plant cost would be 25%-30% lower, maintenance and labour costs would also be 25%-30% lower, and fuel requirements would be virtually eliminated. The by-product carbon monoxide gas from the furnace would be more than sufficient for drying coke, silica and rock. Some rocks that have been used successfully in electric furnaces without agglomeration, calcining, or sintering are screened Florida pebble (plus 6 mm), Ronda hard rock, and Montana hard rock (crushed and screened). Use of uncalcined rock may increase the electric power consumption in the furnace by as much as 10%, depending on the CO2 combined water content.
The Tennessee Valley Authority (TVA) began development of the electric-furnace process for producing phosphate fertilisers in 1933 and produced phosphorus and phosphoric acid from 1934 to 1977. At one time five furnaces were in operation. ln 1977 operation of all furnaces was discontinued by TVA since the process could no longer compete with the wet process for fertiliser production.
Phosphoric acid production by the blast-furnace process
A flow diagram of a TVA pilot plant for producing phosphoric acid by the blast-furnace process is shown in Figure 5. The scale of the pilot plant was about 1 tonne of P2O5 per day .
In general, the blast-furnace process differs from the electric furnace in the following aspects:
Coke is used for both fuel and phosphorus reduction. The estimated coke requirement for a large scale unit is 2.5 tonnes/tonne of P2O5 recovered as phosphoric acid (allowing for ferrophosphorus losses). About 0.6 tonne of coke is consumed in reduction of P2O5 to phosphorus, and the remainder generates heat by combustion with preheated air to form carbon monoxide.
As with the electric furnace, the charge – phosphate rock, coke and silica – must be in lump or agglomerated form, but it is not necessary to calcine or dry the charge since there is sufficient heat for this purpose in the ascending gases in the furnace shaft.
The gas from the furnace contains about 37% CO and 1.0%-1.5% P4 by volume. The remainder is mainly nitrogen. Although recovery of elemental phosphorus by cooling and condensation is feasible, it would be difficult to recover a high percentage because of the low concentration in the gas. In the TVA pilot plant, phosphoric acid was recovered after preferential oxidation of the phosphorus in the gas with air.
The gas remaining after phosphoric acid recovery contains about 34% CO, 1%-2%O2, and the remainder N2 (dry basis). About 40% of this gas can be used advantageously for preheating the air to the blast furnace. The remainder would be available for other uses.
The use of the blast furnace to produce phosphoric acid for fertiliser purposes seems unpromising at present due to the high cost of coke. However, with some improvements it might be considered in certain circumstances . As does the electric furnace it can use low-grade siliceous ore with moderately high alumina and iron oxide content.
30. ‘Sanders, M.D. 1968. ‘Recovery of Fluorides as By-products’. IN Phosphoric acid, A.V. Slack (Ed.) pp765-778, Marcel Dekker Inc., New York, NY, USA.
31. ‘Production of Synthetic Fluorspar from Waste Fluosilic Acid’. 1976. Paper No 22. ISMA Technical Conference, The Hague, Netherlands.
40. Lounamaa, N. and L. Niinimaka. 1971. ‘Typpi Oy’s Solvent Extraction Process for Producing Compound Fertilizers’, Paper No. ID/WG99/20, Presented at the Second Interregional Fertilizer Symposium (UNIDO), Kiev, USSR.
41. ‘Phosphoric Acid Production from Hydrochloric Acid.’ 1987. Negev Phosphates Ltd. Marketing Division, Israel.
43. Hignett, T. P. 1948. ‘Development of Blast Furnace Process for Production of Phosphoric Acid,’ Chemical Engineering Progress, 44(10): 753- 764; 44(11):821-832; and 44(12):895-904.
44. Hignett, T. P. 1968. ‘Use of the Blast Furnace for Production of Phosphoric Acid,’ Phosphorus and Potassium, 36:20-21,24.
Links to Related IFS Proceedings
187, (1979), Uranium Recovery from Phosphoric Acid (A Process Engineering Review), A P Kouloheris
201, (1981), From Wet Crude Phosphoric Acid to High Purity Products, A Davister, G Martin
587, (2006), Phosphogypsum Management and Opportunities for Use, J Hilton
668, (2010), The Phosphate Life-Cycle: Rethinking the Options for a Finite Resource, J Hilton, A E Johnston, C J Dawson
744, (2014), History, Development, Status and Opportunities for Kiln Phosphoric Acid, T P Fowler
748, (2014), Management of Fluorine in Phosphoric Acid Production, B Van Massenhove, M G J A Collin, T Theys
752, (2014), Life Cycle Management of Phosphogypsum Stacks, G R Albarelli, B K Birky
766, (2015), Uranium Extraction from Phosphoric Acid: The Experience of Prayon, M G J A Collin
804, (2017), Phosphogypsum stacking: A new approach and case study, V Dardenne, J Peret and S Plainchamp
821, (2018), Approaches to improving the quality of phosphoric acid, T Henry
Links to external sources
Becker, P. (1989) Phosphates and Phosphoric Acid: Raw Materials: Technology, and Economics of the Wet Process. Marcel Dekker, Inc., New York, NY, U.S.A.
Havelange, S. et al. (2022). Phosphoric Acid and Phosphates in Ullmann’s Encyclopaedia of Industrial Chemistry.
Slack, A.V. (1968). Phosphoric Acid (Part I and II). Marcel Dekker, Inc., New York, NY, U.S.A.
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