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A hydrometallurgical process for the dephosphorization of iron ore
HydrometaUurgy, 21 (1989) 277-292
277
Elsevier Science Publishers B.V., Amsterdam - - Printed in The Netherlands
A Hydrometallurgical Process for the Dephosphorization of Iron Ore MAMOUN MUHAMMED* and YU ZHANG**
Department of Inorganic Chemistry, The Royal Institute of Technology, S-100 44 Stockholm (Sweden) (Received October 16, 1987; revised and accepted July 30, 1988)
ABSTRACT Muhammed, M. and Zhang, Y., 1989. A hydrometallurgical process for the dephosphorization of iron ore. HydrometaUurgy, 21: 277-292. A process is proposed for the dephosphorization of iron ore. The process consists of an integrated treatment for the removal of the phosphorus from the ore by leaching and further processing of the leach solution. Phosphoric acid is extracted by isoamyl alcohol (iAmOH) and stripped by nitric acid solution. Phosphoric acid is concentrated by evaporation where most of the nitric acid is removed. The remaining nitric acid is extracted by methyl isobutyl ketone. The raffinate from the phosphoric acid extraction is treated by sulfuric acid for the regeneration of the spent nitric acid. Nitric acid is extracted by iAmOH and concentrated by distillation before re-used in further leaching. Evidence on the technical feasibility of the process was established. The economy of the process is analyzed and found to be viable.
INTRODUCTION
Some of the iron ores produced in Sweden have a relatively high phosphorus content (1-3% P). High phosphorus content in iron and steel is disadvantageous for some of their properties, e.g. material strength. However, such iron ore can still be used for the production of good quality steel if special processes that are suited to the ore quality are used. Examples are the Thomas process and LDACprocess [1 ]. However, at present, steel mills are increasingly using iron ore with low phosphorus content, so that the Swedish mining industry is having difficulties in marketing their particular ore quality. Therefore, for some time, considerable attention has been focused on developing methods for the dephosphorization of iron ores. *Author to whom all correspondence should be addressed. **Visiting Scholar from Wuhan Institute of Chemical Technology, Wuhan (China).
0304-386X/89/$03.50
© 1989 Elsevier Science Publishers B.V.
278
Principally, two different approaches can be used for decreasing the phosphorus content of the ore. Figure 1 gives a schematic representation of the two approaches. The first approach is suitable if the iron ore is to be used as pellets. In this route, the ore is extensively ground until the phosphorus bearing mineral (apatite) is completely liberated from the iron minerals. The separation is then carried out using physical methods, e.g. magnetic separation, flotation, etc. [2]. The second approach is a wet chemical method and is suitable for the treatment of iron ore sinters. In this treatment the ore is leached with a suitable solution (mineral acid) for the removal of the phosphorus by the dissolution of the apatite. Iron ore with larger particle size is more suitable to this type of treatment. The two approaches outlined above have been in industrial use in Sweden for the production of low phosphorus grade iron ore products. There are a few plants currently in operation for the production of pellets [2 ]. However, there is no report on plants either in Sweden or elsewhere, operating a leachingbased process for the production of low phosphorus iron ore. During the period 1946-1973, a plant was in operation in Sweden for the leaching of phosphorus from sinter-fines [ 3 ]. In this process the ore was leached with nitric acid and the phosphorus content was reduced from 0.2 to 0.05%. In this method no processing of the solution produced from the leaching was included. The produced acid solution was sold to another plant that manufactured mixed fertilizer. The operation costs, mainly the acid consumption, were presumably high and the plant has subsequently been shut down. Recently an increased interest has been shown for treatment methods using CRUSHING~
MINING
SCREENING~
GRINDING
LEAGNIN~_] G SEP~TIONI~_~_1 ~GNETIC~LOTATIO~-~ N FLOTATION
Phosphoric Sinter Apatite Acid Fines
Pellets
Fig. 1. Methods for the removal of phosphorus from iron ore.
279 a wet chemical approach for the production of low phosphorus sinter-fines. The economy of the process can be greatly improved by increasing the selectivity of the leaching operation and the further processing of produced leach solutions. An extensive research program was therefore initiated for developing a cost efficient process for the removal of phosphorus from the phosphatic iron ore [4-11]. The present development of this process is concerned with two main aspects: the removal of phosphorus, or dephosphorization, of iron ore; and the recovery of phosphorus as a by-product phosphoric acid of high purity. The first aspect is directly related to the quality of the iron ore, as a sinter feedstock. The second aims at reducing the cost of the treatment in order to economize the whole process. Besides, environmental consideration should be specially emphasized, where minimum emission would be permitted and maximum recycle of the process streams incorporated. Several studies were therefore undertaken in order to evaluate an alternative process with the following objectives: (1) Selective removal of phosphorus from the iron ore. This should be done without adversely affecting the quality of the ore. (2) Effective recovery of the removed phosphorus, for example, as high purity phosphoric acid. (3) The regeneration of, wherever possible, expensive reagents e.g. spent leach acid. The investigations reported in Ref. [6-11] were thus concerned with the study of the characteristics of different systems of interest for the present process development. This paper summarizes, in brief, the results obtained and is relevant to a process suggested by the authors. PROCESS PRESENTATION The proposed process is designed considering the following technological requirements: (1) Phosphorus in iron ores should be removed and the treated ores must not contain more than 0.05% of phosphorus. The loss of iron as a result of the leaching operation should be kept below 0.5%. (2) The acidity of the residual liquid in the treated ores should be as low as possible and its pH value should be 7-8 after final washing. (3) The leaching operation should be sufficiently selective so that the concentration of phosphoric acid in the leachate should be as high as possible and the mole ratio of iron to phosphoric acid in the leachate should not be higher than 0.1. (4) Phosphoric acid is to be recovered as a high purity product with as low metal impurities as possible, prior to the acid concentration. The product acid would contain 75 wt% H3PO4 preferably of food grade.
280
(5) As leaching is carried out with nitric acid, the nitrate salts in the raffinate should be used to regenerate nitric acid as completely as possible for re-use in leaching. (6) When using solvents in the different extraction processes, the entrained solvent fractions should also be recovered. OVERALLPROCESSOUTLINE
An outline of an integrated process for the dephosphorization of the ore is given by the following four main sections: ( 1 ) Phosphorus removal: nitric acid is used for selective ore leaching. (2) Phosphoric acid recovery: this could be achieved by solvent extraction. (3) Phosphoric acid concentration and purification: e.g. by distillation, evaporation or combination with solvent extraction. (4) Nitric acid regeneration: e.g. by the addition of sulfuric acid. A block diagram for the suggested process is given in Fig. 2. Two alternative conceptual flowsheets were examined. In the first flowsheet, phosphoric acid is extracted directly from the resulting leachate solution. In the second flowsheet, phosphoric acid is extracted from the solution after the addition of sulfuric acid. The first alternative is considered more attractive [5]. The second alternative offers the possibility of enhancing the phosphoric acid extraction by increasing the sulfuric acid concentration. However, the separation of sulfuric acid is considered to be more complicated. This would adversely affect the quality of the phosphoric acid produced and the quality of the leached ore as well. The quality of leached sinter-fines for the production of pig iron was evaluated. Two kinds, the sinter and metallurgical tests, were performed by the
,
)
_,
I ,t~
TT
I
V'[~-'q
I
, -I
611-I
I
ljj-I
I
- I
-
Ivs~Fo 4
Fig. 2. A conceptual flowsheet of the proposed process for the dephosphorization of iron ore.
281 R&D Laboratories of Luossavaara-Kiirunavaara Aktiebolag (LKAB). The results of these tests indicated that the leached ore is a good sinter feed, the quality of which is not adversely affected by the leaching treatment. In the following a description of the main sections of the proposed process is presented. PHOSPHORUS REMOVALSECTION The ore is leached with nitric acid for the removal of phosphorus in a percolation operation. Nitric acid is preferred because of its low reactivity with magnetite (the main component of the ore type studied). The characteristics of the leaching operation have been investigated in detail [6]. It was found that a properly leached ore contains phosphorus less than 0.05 % P (in some runs as low as 0.027% ). The dissolution of iron could easily be kept under 0.5%. Apatite dissolution is found to be diffusion-controlled which is sensitive to the flow-rate variation. The dissolution of magnetite is sensitive to the concentration of the leach acid used which thus limits the use of concentrated solutions. Alkali concentration in the ore was also reduced. The rate equations for the dissolution of apatite and iron were obtained. The operating conditions for the percolation leaching of the KDF ore with nitric acid were deduced. Based on the results reported above, the part of the process that is designed for the treatment of the iron ore, would include the following sequence of operations: ( 1 ) Leaching: acidulation of apatite in the ore with a leach solution of maximum 6 M HNO3. The produced leachate is used as a feed to the H3PO4 recovery section. Total residence time of the ore, including draining the solution, is approximately 24 h for a percolation operation. (2) Washing: the removal (or dilution) of the acids retained in the leached ore by recycling a processing solution. The raffinate from the H N Q regeneration section may be used. The total concentration of HNO3 and H3PO4 in this solution is less than 0.01 M. The total residence time is 4 h. (3) Neutralization: neutralizing the residual acids in the washed ore with lime solution. The pH value of the liquid remaining in the ore will be 7-8. The residence time is 4 h, including the removal of the solution (dry dripping). A general flow scheme of this section is given in Fig. 3. The leaching and washing operation is performed through a multistage, semi-countercurrent/ semi-batch operation. In this case, the ore is actually maintained in a column during the three different operating steps, while the different fractions of the processing streams are continuously pumped, in up-flow, through the fixed bed of ore. Two stages will be sufficient for leaching and four stages for washing. The flow-rate of the different solutions should be as high as possible in order to reduce the diffusion resistance.
282 Iron Ore
M
I Leachate IM H3PO4 IM HNO 3 3M Ca(N03)2
(~.01MHNO~]
(3~S~es)
0.5M HNO 3 IM Ca (NO3)2
IM Ca(N03)2
©
I
I
g
~
pH=7
- 8
(D
Treated Ore Fig. 3. A flowsheetof the phosphorus removalsection. PHOSPHORIC ACID RECOVERYSECTION Several extraction systems were studied in order to select a suitable solvent for the extraction of phosphoric acid. Synthetic solutions with different composition were used to simulate different process streams. Both batch and continuous tests using mixer settler equipment were undertaken to evaluate the extraction process using process solution. Nine different extractants were considered and four of them were investigated in more detail [7,8 ]. The extraction of H3P04 by all solvents is generally low. Most of these solvents have a selectivity for the extraction of the acids against salts. However, the selectivity of these solvents for the different acids will depend on the composition of the aqueous and organic phases. Two extractants were considered for further detailed investigation: isoamyl alcohol (iAmOH) and tributyl phosphate (TBP). The characteristics of these two extraction systems have been investigated [7,8]. Both solvents extract nitric acid preferentially to phosphoric acid. However, they can both still be used for the extraction of phosphoric acid from the leach solutions. A comparison between TBP and iAmOH showed that the extraction of H3PO4 with iAmOH can be increased if the concentration of H N Q and/or Ca(NO3)2 in
283
the feed is increased. The extraction of phosphoric acid with iAmOH is more effective than that with TBP. The selectivity of iAmOH for phosphoric acid against salts, e.g. Ca 2+ and Fe ~+, is high ( > 10) and much more selective than that of TBP.The continuous operations have indicated that the selectivity of TBP for H3P04 is so poor that no effective separation of H3P04 from Ca can be achieved. Moreover, iAmOH systems have much better phase separation characteristics. The price of iAmOH is much lower than that of TBP. Therefore, the use of iAmOH is preferred. Based on the results obtained, this section has been designed for recovery of the phosphoric acid by using an organic solvent recycled in three different operating steps: ( 1 ) Extraction (I): selective separation of H~P04 from salt impurities, mainly Ca, Mg, Fe, and A1, in the leachate. The solvent used is iAmOH, containing ca. 1-2 M H N Q and < 0.01 M H3P04. The H N Q concentration of the aqueous phase is maintained at its initial level (1-2 M H N Q ) , to ensure an efficient extraction of H~P04 without further co-extraction of H N Q that would otherwise result in the precipitation of phosphates. The H3P04 concentration in the raffinate will be as low as 0.1 M. (2) Scrubbing: removal of the co-extracted impurities from acids in the loaded solvent using a side stream of the stripping product. The resulting aqueous solution will be recycled back to the feed leachate. (3) Stripping: recovery of H3P04 and regeneration of the solvent by using a
Feed (Leachate) IM H3PO4 IM HNO3 3M Ca(NO3)z
~ Raffinate : <0.1M H3PO4 ~ IM HNO3 r I EXT~CTION I R=8 3M Ca(NOz)2 1(10-13~Stages)
1 SCRUBBING I (2-3 Stages)
R=20
S°luti°n ~ i ~ Organic OH I~ 1.5-2M HNO3
I
(3-4 Stages) I ~
0.8M H3PO 4 2-3M HNO3
I
I Strip Solution IM HNO3
HNO3+HzO
Fig. 4. A flowsheet of phosphoric acid recovery section.
284
0.20
[H3P04]o (M) Extraction/
sf
0.15
S/ / "" "
/
if" I .y .';tr,p O.lO
..
0.05
0 ~ 0
/ //--/'/
/
I 0.2
"
"-'" / / /
0.4
I 0.6
I 0.8
I 1.0 1.2 [H3PO4]a (M)
Fig. 5. Calculated number of theoretical stages for the extraction and stripping of phosphoric acid by iAmOH.
solution containing 2-3 M H N Q . The stripping solution is made of two processing streams, one out-flowing the H N Q regeneration section and the second from the HsPO4 purification section. The final stripping stream, containing 0.5-1 M H3PO4 and 2-3 M HNOs, is sent into the H3PO4 concentration section. The stripped solvent contains < 0.01 M H3P04 and 1-2 M H N Q . No significant re-extraction of HNOs takes place. The stripped solvent is recycled back to extraction I. A general flow scheme of this section is given in Fig. 4. In practice, a continuous, countercurrent, multistage operation is performed in each step. The suitable flow ratio of two phases and number of stages in each step will be more accurately estimated in the specific design work, which largely depends on the extractor selected. However, 15-20 stages of mixer-settlers will be sufficient for this section as the flow ratio of the organic to aqueous phase is 6-10 in the extraction step. There are 10-13 stages for extraction, 2-3 stages for scrubbing and 3-4 stages for stripping. Figure 5 gives an example of the graphical estimation of stages of extraction and stripping. PHOSPHORIC ACID CONCENTRATION AND PURIFICATION SECTION
The phosphoric acid (10%) solution from the stripping section contains a high concentration of nitric acid ( 12% ). Phosphoric acid is to be concentrated, by evaporation to 55% or 75%. During evaporation, the separation of the two
285 acids is achieved as the bulk of HNO3 is removed into the vapor phase together with water [12-14]. However, the separation of the acids is not complete. The concentrated phosphoric acid produced (about 75 wt% H3PO4) contains approximately 5 wt% H N Q . The removal of H N Q requires a high energy input. Methyl isobutyl ketone (MIBK) dissolved in aromatic diluents, e.g. toluene, can be used for the removal of nitric acid from phosphoric acid solutions. The characteristics of the extraction of nitric acid from very concentrated phosphoric acid solutions (50-75 wt% ) were studied [9]. The following section is designed for the further processing of the H3PO4 solution, in three subsections.
Primary concentration (multi-effect evaporation) (a) Removal of the organic solvent as well as evaporation of water from the H:~PO4 solution. The feed stream is produced in the stripping step of H3PO4 recovery section. The organic solvent (iAmOH) dissolved in the aqueous solution is readily evaporated together with a fraction of water at the first stage. After cooling and phase separation, the recovered solvent is recycled back to the extraction step. (b) The reduction of H N Q content as well as the further concentration of the H3P04 solution in the following two stages. The major part of HN03 is removed with the evaporation of the major part of the water. The solution is concentrated to 60-70 wt% H3PO4, which may contain about 5 wt% HNO:~. It is then pumped to the following H N Q removal subsection.
Nitric acid removal (extraction III) (a) Selective extraction of H N Q from the concentrated H~PO4 solution. The extractant used is MIBK in toluene (50 vol% (MIBK). The concentration of HN03 in the raffinate is to be reduced to lower than 0.01 wt%, while H3PO4 concentration has a level similar to or higher than that in the feed (60-70 wt%). The raffinate is then sent for final concentration. This sub-section is operated under ambient temperature, 20 ° C. (b) Stripping of H N Q loaded in the organic phase. The loaded solvent is stripped with distilled water. The stripped solvent is then recycled back to the extraction step, while the stripping solution is sent to the H3PO4 recovery section.
Final concentration (distillation) The final removal of the organic solvents in the phosphoric acid, as well as some water, is achieved. The purified HaPO4 solution is concentrated to 75 wt% H3PO4. The organic solvents dissolved, MIBK and toluene, are evaporated together with a fraction of the water during the distillation. The recovered solvent and distilled water are then re-used in the previous extraction III. A general flow scheme of this section is given in Fig. 6. The specific operating
286
L'
iAmOHHNO3~ H20
i L"~I~A'~a',,T
(2-3 Effects) Feed 0.8M ~ 3 4
STRIppING (2-3 Stages)
EXTRACTION (5-8 Stage~)
I
LIATION
ct
2-3M
60-70%H.9~4/5-1O-%~o3 Fig. 6. A flowsheet of the phosphoric acid concentration and purification section.
1.6
[HNO3]o (M)
12
/
j
0.4 ~
Stripping
0 0
0.2
0.4
I 0.6
I 0.8
I 1.0 1.2 [HNO3]~ (M)
Fig. 7. Calculated number of theoretical stages for the extraction and stripping of nitric acid from 65% phosphoric acid by 50% MIBKdissolved in toluene.
modes for each subsection can be estimated in further design work. Figure 7 gives an example of the basic calculations of the nitric removal subsection (extraction III). As seen, for the given operating conditions, the number of theoretical stages is four for the extraction as well as for the stripping operation. NITRIC ACID REGENERATION SECTION
Since the solubility of gypsum, C a S O 4. 2H20, in nitric acid is low, maximum 2.5 wt% at 25 oC [ 15 ], the regeneration of HNO3 can be easily achieved by the precipitation of CaSO4-2H20. The nitric acid solution can be recycled for further leaching. During the leaching, however, precipitation of calcium sulfate occurs as HNO3 concentration decreases and Ca(NO3)2 concentration in-
287 creases, This will result in an increase of the sulfur content in the ore, which is undesirable with respect to the quality of the iron ore. The determination of the solubility of CaS04 in the aqueous systems H+-Ca2+-NO~]--S04, related to the composition of the present process solutions is reported in [10]. The results indicate that the direct reuse of the completely generated nitric acid at about 5 M will introduce ca. 7.9 kg CaS04 ( 1.7 kg S ) into 1 ton of the leached ore. In order to avoid the precipitation of gypsum during the use of regenerated nitric acid, a partial regeneration of the acid is suggested. In this way a certain amount (a steady-state concentration) of calcium nitrate will be maintained in the solution thus reducing the solubility of gypsum due to the common ion effect. This section is designed to consist of three subsections.
Calcium sulphate precipitation The extraction raffinate resulting from H3PO4 recovery section is mixed with 97 wt% H2804 in a stirred tank, and the solid CaSO 4- nH20, or gypsum, is formed. The stoichiometric requirement for H2SO4 addition is that the total sulphate concentration should be 0.7-0.8 M lower than the final calcium concentration. The resultant suspension, which has an acidity of 5-6 M HNO:~ in the solution, is then sent to the following H N Q recovery subsection.
Nitric acid recovery (extraction II/crystallization) (a) Selective extraction of H N Q against sulfate in the solution. The suspension is fed into first stage of the three-phase extractors (mixer-settlers or columns), iAmOH is used as a solvent. In this way, a complete extraction of H N Q into the organic phase results in the efficient crystallization of the gypsum as the acidity of the aqueous solution is greatly reduced. The gypsum precipitates at the bottom of the extractor where the aqueous raffinate contains less than 0.01 M of acids, including HN0:~. This suspension is sent to gypsum removal subsection. (b) Scrubbing of the loaded solvent for the removal of the sulphate. A side stream of the stripping solution is used for scrubbing, and then added to the extraction feed (suspension). (c) Stripping of the scrubbed solvent for primary recovery of H N Q . The strip feed is a mixture of the washing streams, containing 0.5-1 M HNO~. The stripping product may contain 3-4 M H N Q , and is to be concentrated to 5-6 M by distillation in order to be re-used in leaching. (d) Washing of the stripped solvent for further recovery of HNO~. Distilled water is used for washing. The washed solvent would contain > 0.01 M HNO~ and is then recycled for extraction. The washing product is added to the strip feed.
288
H2S04 (98~)
Q
<0.1M H3PO4 [ IM HNO 3 1 3M Ca(NO3)z t
<0.01M HNO 3 IM Ca(NO3)2
~_~..~o~
~
_J ~ o ~
~_~
~
'!IF[(5"7RSt ~ =ages) 2
~l ~
Ill (4-6 Stages) ~
~
.
i
~s~,~ I
-
(2-3R=10 IStages) ~7 1 (2-3~Stages) --
iAmOU
.2o,
~
.%,
[ U20
. . . . . .
"-!
V"
oypsu~
(4-6 S t a g e s ) ~
T
©
'
R-I
56M.No,
~a(~O~®
Fig. 8. A flowsheet for the nitric acid regeneration section.
[HN03] o (M) 4.0
Strip
3.0
e = 1
4---
0=2
/ / In/ /
2.0
Extraction
iv
1.0
0
1.0
I
I
I
2.0
3.0
4.0
I
5.0 6.0 [HNO3] a (U)
Fig. 9. Calculated number of theoretical stages for the extraction and stripping of nitric acid by iAmOH.
289
Gypsum separation (a) Further crystallization of gypsum, as well as gravity sedimentation, in the raffinate solution: the suspension from the raffinate is fed into a continuous thickener. The clear overflow stream is sent to the P-removal section where it will be used for washing the leached ore, while the thickened underflow is fed to the filtration step. (b) Filtration of gypsum: The filtrate is pumped back to the thickener, while the gypsum is to be washed with a washing stream of p H 7-8 from the Premoval section. A general flow scheme of this section is given in Fig. 8. The specific operating modes for each subsection should be developed in further design work. An example of the calculation is given for the extraction II, as shown in Fig. 9. ASSESSMENT OF PROCESS ECONOMY We have examined the economics for the process described above. The cost estimates were made for a plant that processes 4 million tons of iron ore per year and produces high purity phosphoric acid (75% H3P04) [ 11 ].
Cost estimates The capital investment costs were estimated. The estimates calculated for the different items, e.g. equipment, building, etc, are given in Table 1. As seen, the total capital investment is estimated to be about 444 million Swedish crowns TABLE1 The capital investment costs (in millions of Swedish crowns)
Production units: Phosphorus removal section H3P04 recovery section {extraction I) HN03 regeneration section Acid concentration section Solvent inventory Subtotal Contingency (10%) Auxiliary: Material storage Gypsum disposal Land and preparation Subtotal Total
80.5 34.8 40.5 150.6 8.8 315.2 31.5 61.5 32.7 2.6 96.8 443.5
290 TABLE 2 The operation costs (in Swedish crownsper ton ore) Materials: Sulfuric acid ( > 97% ), 76 kg Nitric acid (> 63%, 2 kg (make-up) Others (water, lime, solvent make-up, etc. )
19.0 1.5 1.0
Subtotal
21.5
Energies: Steam Electricity, 26 kWh
12.0 5.9
Subtotal
17.9
Labor Maintenance, 4% of total investment Others
Total
2.5 4.4 5.6 51.9
(61 million US$) at 1986 values. The operation costs are estimated from the process parameters. The material costs were calculated from the process material balance mainly concerning H2804 and H N Q consumptions. The energy costs were calculated from the energy balance mainly for the evaporation and distillation. These costs are specified in Table 2. As seen, the total operatingcost is estimated to be 52 Swedish crowns per ton ore or 208 million Swedish crowns (28.5 million US$) per year. Revenue and profitability estimation
The increase of the ore value, as a result of dephosphorization, was set by the LKABcompany to 17-20 Swedish crowns per ton ore processed (1980). The same value is used also in the present estimation. From the material balance, the production of phosphoric acid (75 wt%) is calculated to be more t h a n 32.0 kg per ton ore (considering a total of 80% recovery). The present price is taken to be 350-400 US$ per ton of 75% H:~PO4 for technical grade (ex factory; Supra AB, Sweden). The higher purity acid produced by thermal process is at least 50% higher in price. The purity grade of product acid of this process may be sufficiently high to meet the requirement of food grade acid. The revenue and profit are estimated in Table 3. As seen, the revenue is to be 100-160 Swedish crowns per ton ore or 400-640 million Swedish crowns (55-88 million US$) per year at a full production capacity, while the operating profit 28-90 Swedish crowns per ton ore or 110-
291 TABLE 3 The revenue, income and profit (in Swedish crowns per ton ore) Quality of acid:
Technical
Estimation level:
low
Ore value increase Value of 75% H:~PO4
High purity high
low
high
17.0 81.8
20.0 93.4
17.0 122.6
20.0 140.2
Revenue Operating cost
98.8 - 51.9
113.4 - 51.9
139.6 - 51.9
160.2 - 51.9
Income Interest on loan Depreciation
46.9 - 11.1 - 7.9
61.5 - 11.1 -- 7.9
87.7 - i 1.1 - 7.9
108.3 - 11.1 - 7.9
27.9
42.5
68.7
89.3
Profit
360 million Swedish c r o w n s ( 1 5 - 5 0 million U S $ ) per year. T h i s indicates a high p r o f i t a b i l i t y of t h e process u n d e r t h e s e conditions. CONCLUSIONS F r o m the results r e p o r t e d t h e p r o p o s e d process is s h o w n to be t e c h n i c a l l y feasible. T h e quality of the leached ore has i m p r o v e d u p o n leaching. P h o s p h o r i c acid could be r e c o v e r e d as h i g h - p u r i t y grade. T h e e s t i m a t e s of the process e c o n o m y indicate t h a t the process can be o p e r a t e d with profit. However, the e c o n o m y of the process is very m u c h d e p e n d e n t on: (a) price of p h o s p h o r i c acid; (b) e n e r g y costs; a n d (c) costs of sulfuric acid. ACKNOWLEDGEMENTS T h e work has b e e n f i n a n c e d b y the Swedish B o a r d for T e c h n i c a l Developm e n t (STU). T h e a s s i s t a n c e a n d c o o p e r a t i o n of several colleagues f r o m t h e d e p a r t m e n t s of C h e m i c a l E n g i n e e r i n g at the R o y a l I n s t i t u t e of T e c h n o l o g y a n d at the U n i v e r s i t y of L u n d are gratefully acknowledged.
REFERENCES 1 Fagerberg,B. and Sandgren, P., 1980. German -British-Swedish Symposium Conserving our Resources. Hamburg (May). 2 EdstrSm, J., Engelbert, T., Selin, R., Werme, A. and Wijk, O., , 1981. St~lframst~illning ur fosforrik j~irnmalm, STU Inf. No. 256-1981, Stockholm, Sweden.
292 3 4 5 6 7 8 9
10 11 12 13 14 15
Hohenstein, P., 1967. Erzbergbau Metalhiittenswes., 217. Grenthe, I., Liem, H., Muhammed, M. and Setterwall, F., 1981. Swedish Patent 8103114-8. Muhammed, M., 1981. Defosforering av KDF j~irnmalm. Rep., STU Dnr 79-5304; TRITAOOK-2022, Stockholm, Sweden. Zhang, Y. and Muhammed, M., 1989. The removal of phosphorus from iron ore by leaching with nitric acid. Hydrometallurgy, 21 (3): 255-275. Zhang, Y. and Muhammed, M., 1988. Extraction of phosphoric acid from nitrate solutions with isoamyl alcohol. Solvent Extr. Ion Exch., 6 (6): 973. Zhang, Y., Muhammed, M. and Valiente, M., 1988. Extraction of nitric and phosphoric acid with tributyl phosphate. Solvent Extr. Ion Exch., 7 (2): in press. Zhang, Y. and Muhammed, M., 1988. Removal of nitric acid from concentrated phosphoric acid by extraction with methyl isobutyl ketone in aromatic diluents. Solvent Extr. Ion Exch., 6(6): 993. Zhang, Y. and Muhammed, M., 1989. Solubility of calcium sulphate in nitric acid solutions containing calcium nitrate and phosphoric acid. J. Chem. Eng. Data, 34(1 ): 121. Muhammed, M. and Zhang, Y., 1987. In: J. Edstr5m et al. (Editors), St~lframst~illning ur fosforrik j~irnmalm, del 2. STU Inf. No. 622-1987, 64, Stockholm, Sweden. Tichonov, A. and Korsinkina, J., 1936. Zh. Khim. Promsti., 13, 1345. Flatt, P. and Bonnet, J., 1958. Chimia, 12,343. Petrov, et al., 1969. Neftekhimya, 9,643, p. 16. Kruchenko, V. and Beremzhanov, B., 1980. Zh. Neorg. Khim., 25, p. 3076.
277
Elsevier Science Publishers B.V., Amsterdam - - Printed in The Netherlands
A Hydrometallurgical Process for the Dephosphorization of Iron Ore MAMOUN MUHAMMED* and YU ZHANG**
Department of Inorganic Chemistry, The Royal Institute of Technology, S-100 44 Stockholm (Sweden) (Received October 16, 1987; revised and accepted July 30, 1988)
ABSTRACT Muhammed, M. and Zhang, Y., 1989. A hydrometallurgical process for the dephosphorization of iron ore. HydrometaUurgy, 21: 277-292. A process is proposed for the dephosphorization of iron ore. The process consists of an integrated treatment for the removal of the phosphorus from the ore by leaching and further processing of the leach solution. Phosphoric acid is extracted by isoamyl alcohol (iAmOH) and stripped by nitric acid solution. Phosphoric acid is concentrated by evaporation where most of the nitric acid is removed. The remaining nitric acid is extracted by methyl isobutyl ketone. The raffinate from the phosphoric acid extraction is treated by sulfuric acid for the regeneration of the spent nitric acid. Nitric acid is extracted by iAmOH and concentrated by distillation before re-used in further leaching. Evidence on the technical feasibility of the process was established. The economy of the process is analyzed and found to be viable.
INTRODUCTION
Some of the iron ores produced in Sweden have a relatively high phosphorus content (1-3% P). High phosphorus content in iron and steel is disadvantageous for some of their properties, e.g. material strength. However, such iron ore can still be used for the production of good quality steel if special processes that are suited to the ore quality are used. Examples are the Thomas process and LDACprocess [1 ]. However, at present, steel mills are increasingly using iron ore with low phosphorus content, so that the Swedish mining industry is having difficulties in marketing their particular ore quality. Therefore, for some time, considerable attention has been focused on developing methods for the dephosphorization of iron ores. *Author to whom all correspondence should be addressed. **Visiting Scholar from Wuhan Institute of Chemical Technology, Wuhan (China).
0304-386X/89/$03.50
© 1989 Elsevier Science Publishers B.V.
278
Principally, two different approaches can be used for decreasing the phosphorus content of the ore. Figure 1 gives a schematic representation of the two approaches. The first approach is suitable if the iron ore is to be used as pellets. In this route, the ore is extensively ground until the phosphorus bearing mineral (apatite) is completely liberated from the iron minerals. The separation is then carried out using physical methods, e.g. magnetic separation, flotation, etc. [2]. The second approach is a wet chemical method and is suitable for the treatment of iron ore sinters. In this treatment the ore is leached with a suitable solution (mineral acid) for the removal of the phosphorus by the dissolution of the apatite. Iron ore with larger particle size is more suitable to this type of treatment. The two approaches outlined above have been in industrial use in Sweden for the production of low phosphorus grade iron ore products. There are a few plants currently in operation for the production of pellets [2 ]. However, there is no report on plants either in Sweden or elsewhere, operating a leachingbased process for the production of low phosphorus iron ore. During the period 1946-1973, a plant was in operation in Sweden for the leaching of phosphorus from sinter-fines [ 3 ]. In this process the ore was leached with nitric acid and the phosphorus content was reduced from 0.2 to 0.05%. In this method no processing of the solution produced from the leaching was included. The produced acid solution was sold to another plant that manufactured mixed fertilizer. The operation costs, mainly the acid consumption, were presumably high and the plant has subsequently been shut down. Recently an increased interest has been shown for treatment methods using CRUSHING~
MINING
SCREENING~
GRINDING
LEAGNIN~_] G SEP~TIONI~_~_1 ~GNETIC~LOTATIO~-~ N FLOTATION
Phosphoric Sinter Apatite Acid Fines
Pellets
Fig. 1. Methods for the removal of phosphorus from iron ore.
279 a wet chemical approach for the production of low phosphorus sinter-fines. The economy of the process can be greatly improved by increasing the selectivity of the leaching operation and the further processing of produced leach solutions. An extensive research program was therefore initiated for developing a cost efficient process for the removal of phosphorus from the phosphatic iron ore [4-11]. The present development of this process is concerned with two main aspects: the removal of phosphorus, or dephosphorization, of iron ore; and the recovery of phosphorus as a by-product phosphoric acid of high purity. The first aspect is directly related to the quality of the iron ore, as a sinter feedstock. The second aims at reducing the cost of the treatment in order to economize the whole process. Besides, environmental consideration should be specially emphasized, where minimum emission would be permitted and maximum recycle of the process streams incorporated. Several studies were therefore undertaken in order to evaluate an alternative process with the following objectives: (1) Selective removal of phosphorus from the iron ore. This should be done without adversely affecting the quality of the ore. (2) Effective recovery of the removed phosphorus, for example, as high purity phosphoric acid. (3) The regeneration of, wherever possible, expensive reagents e.g. spent leach acid. The investigations reported in Ref. [6-11] were thus concerned with the study of the characteristics of different systems of interest for the present process development. This paper summarizes, in brief, the results obtained and is relevant to a process suggested by the authors. PROCESS PRESENTATION The proposed process is designed considering the following technological requirements: (1) Phosphorus in iron ores should be removed and the treated ores must not contain more than 0.05% of phosphorus. The loss of iron as a result of the leaching operation should be kept below 0.5%. (2) The acidity of the residual liquid in the treated ores should be as low as possible and its pH value should be 7-8 after final washing. (3) The leaching operation should be sufficiently selective so that the concentration of phosphoric acid in the leachate should be as high as possible and the mole ratio of iron to phosphoric acid in the leachate should not be higher than 0.1. (4) Phosphoric acid is to be recovered as a high purity product with as low metal impurities as possible, prior to the acid concentration. The product acid would contain 75 wt% H3PO4 preferably of food grade.
280
(5) As leaching is carried out with nitric acid, the nitrate salts in the raffinate should be used to regenerate nitric acid as completely as possible for re-use in leaching. (6) When using solvents in the different extraction processes, the entrained solvent fractions should also be recovered. OVERALLPROCESSOUTLINE
An outline of an integrated process for the dephosphorization of the ore is given by the following four main sections: ( 1 ) Phosphorus removal: nitric acid is used for selective ore leaching. (2) Phosphoric acid recovery: this could be achieved by solvent extraction. (3) Phosphoric acid concentration and purification: e.g. by distillation, evaporation or combination with solvent extraction. (4) Nitric acid regeneration: e.g. by the addition of sulfuric acid. A block diagram for the suggested process is given in Fig. 2. Two alternative conceptual flowsheets were examined. In the first flowsheet, phosphoric acid is extracted directly from the resulting leachate solution. In the second flowsheet, phosphoric acid is extracted from the solution after the addition of sulfuric acid. The first alternative is considered more attractive [5]. The second alternative offers the possibility of enhancing the phosphoric acid extraction by increasing the sulfuric acid concentration. However, the separation of sulfuric acid is considered to be more complicated. This would adversely affect the quality of the phosphoric acid produced and the quality of the leached ore as well. The quality of leached sinter-fines for the production of pig iron was evaluated. Two kinds, the sinter and metallurgical tests, were performed by the
,
)
_,
I ,t~
TT
I
V'[~-'q
I
, -I
611-I
I
ljj-I
I
- I
-
Ivs~Fo 4
Fig. 2. A conceptual flowsheet of the proposed process for the dephosphorization of iron ore.
281 R&D Laboratories of Luossavaara-Kiirunavaara Aktiebolag (LKAB). The results of these tests indicated that the leached ore is a good sinter feed, the quality of which is not adversely affected by the leaching treatment. In the following a description of the main sections of the proposed process is presented. PHOSPHORUS REMOVALSECTION The ore is leached with nitric acid for the removal of phosphorus in a percolation operation. Nitric acid is preferred because of its low reactivity with magnetite (the main component of the ore type studied). The characteristics of the leaching operation have been investigated in detail [6]. It was found that a properly leached ore contains phosphorus less than 0.05 % P (in some runs as low as 0.027% ). The dissolution of iron could easily be kept under 0.5%. Apatite dissolution is found to be diffusion-controlled which is sensitive to the flow-rate variation. The dissolution of magnetite is sensitive to the concentration of the leach acid used which thus limits the use of concentrated solutions. Alkali concentration in the ore was also reduced. The rate equations for the dissolution of apatite and iron were obtained. The operating conditions for the percolation leaching of the KDF ore with nitric acid were deduced. Based on the results reported above, the part of the process that is designed for the treatment of the iron ore, would include the following sequence of operations: ( 1 ) Leaching: acidulation of apatite in the ore with a leach solution of maximum 6 M HNO3. The produced leachate is used as a feed to the H3PO4 recovery section. Total residence time of the ore, including draining the solution, is approximately 24 h for a percolation operation. (2) Washing: the removal (or dilution) of the acids retained in the leached ore by recycling a processing solution. The raffinate from the H N Q regeneration section may be used. The total concentration of HNO3 and H3PO4 in this solution is less than 0.01 M. The total residence time is 4 h. (3) Neutralization: neutralizing the residual acids in the washed ore with lime solution. The pH value of the liquid remaining in the ore will be 7-8. The residence time is 4 h, including the removal of the solution (dry dripping). A general flow scheme of this section is given in Fig. 3. The leaching and washing operation is performed through a multistage, semi-countercurrent/ semi-batch operation. In this case, the ore is actually maintained in a column during the three different operating steps, while the different fractions of the processing streams are continuously pumped, in up-flow, through the fixed bed of ore. Two stages will be sufficient for leaching and four stages for washing. The flow-rate of the different solutions should be as high as possible in order to reduce the diffusion resistance.
282 Iron Ore
M
I Leachate IM H3PO4 IM HNO 3 3M Ca(N03)2
(~.01MHNO~]
(3~S~es)
0.5M HNO 3 IM Ca (NO3)2
IM Ca(N03)2
©
I
I
g
~
pH=7
- 8
(D
Treated Ore Fig. 3. A flowsheetof the phosphorus removalsection. PHOSPHORIC ACID RECOVERYSECTION Several extraction systems were studied in order to select a suitable solvent for the extraction of phosphoric acid. Synthetic solutions with different composition were used to simulate different process streams. Both batch and continuous tests using mixer settler equipment were undertaken to evaluate the extraction process using process solution. Nine different extractants were considered and four of them were investigated in more detail [7,8 ]. The extraction of H3P04 by all solvents is generally low. Most of these solvents have a selectivity for the extraction of the acids against salts. However, the selectivity of these solvents for the different acids will depend on the composition of the aqueous and organic phases. Two extractants were considered for further detailed investigation: isoamyl alcohol (iAmOH) and tributyl phosphate (TBP). The characteristics of these two extraction systems have been investigated [7,8]. Both solvents extract nitric acid preferentially to phosphoric acid. However, they can both still be used for the extraction of phosphoric acid from the leach solutions. A comparison between TBP and iAmOH showed that the extraction of H3PO4 with iAmOH can be increased if the concentration of H N Q and/or Ca(NO3)2 in
283
the feed is increased. The extraction of phosphoric acid with iAmOH is more effective than that with TBP. The selectivity of iAmOH for phosphoric acid against salts, e.g. Ca 2+ and Fe ~+, is high ( > 10) and much more selective than that of TBP.The continuous operations have indicated that the selectivity of TBP for H3P04 is so poor that no effective separation of H3P04 from Ca can be achieved. Moreover, iAmOH systems have much better phase separation characteristics. The price of iAmOH is much lower than that of TBP. Therefore, the use of iAmOH is preferred. Based on the results obtained, this section has been designed for recovery of the phosphoric acid by using an organic solvent recycled in three different operating steps: ( 1 ) Extraction (I): selective separation of H~P04 from salt impurities, mainly Ca, Mg, Fe, and A1, in the leachate. The solvent used is iAmOH, containing ca. 1-2 M H N Q and < 0.01 M H3P04. The H N Q concentration of the aqueous phase is maintained at its initial level (1-2 M H N Q ) , to ensure an efficient extraction of H~P04 without further co-extraction of H N Q that would otherwise result in the precipitation of phosphates. The H3P04 concentration in the raffinate will be as low as 0.1 M. (2) Scrubbing: removal of the co-extracted impurities from acids in the loaded solvent using a side stream of the stripping product. The resulting aqueous solution will be recycled back to the feed leachate. (3) Stripping: recovery of H3P04 and regeneration of the solvent by using a
Feed (Leachate) IM H3PO4 IM HNO3 3M Ca(NO3)z
~ Raffinate : <0.1M H3PO4 ~ IM HNO3 r I EXT~CTION I R=8 3M Ca(NOz)2 1(10-13~Stages)
1 SCRUBBING I (2-3 Stages)
R=20
S°luti°n ~ i ~ Organic OH I~ 1.5-2M HNO3
I
(3-4 Stages) I ~
0.8M H3PO 4 2-3M HNO3
I
I Strip Solution IM HNO3
HNO3+HzO
Fig. 4. A flowsheet of phosphoric acid recovery section.
284
0.20
[H3P04]o (M) Extraction/
sf
0.15
S/ / "" "
/
if" I .y .';tr,p O.lO
..
0.05
0 ~ 0
/ //--/'/
/
I 0.2
"
"-'" / / /
0.4
I 0.6
I 0.8
I 1.0 1.2 [H3PO4]a (M)
Fig. 5. Calculated number of theoretical stages for the extraction and stripping of phosphoric acid by iAmOH.
solution containing 2-3 M H N Q . The stripping solution is made of two processing streams, one out-flowing the H N Q regeneration section and the second from the HsPO4 purification section. The final stripping stream, containing 0.5-1 M H3PO4 and 2-3 M HNOs, is sent into the H3PO4 concentration section. The stripped solvent contains < 0.01 M H3P04 and 1-2 M H N Q . No significant re-extraction of HNOs takes place. The stripped solvent is recycled back to extraction I. A general flow scheme of this section is given in Fig. 4. In practice, a continuous, countercurrent, multistage operation is performed in each step. The suitable flow ratio of two phases and number of stages in each step will be more accurately estimated in the specific design work, which largely depends on the extractor selected. However, 15-20 stages of mixer-settlers will be sufficient for this section as the flow ratio of the organic to aqueous phase is 6-10 in the extraction step. There are 10-13 stages for extraction, 2-3 stages for scrubbing and 3-4 stages for stripping. Figure 5 gives an example of the graphical estimation of stages of extraction and stripping. PHOSPHORIC ACID CONCENTRATION AND PURIFICATION SECTION
The phosphoric acid (10%) solution from the stripping section contains a high concentration of nitric acid ( 12% ). Phosphoric acid is to be concentrated, by evaporation to 55% or 75%. During evaporation, the separation of the two
285 acids is achieved as the bulk of HNO3 is removed into the vapor phase together with water [12-14]. However, the separation of the acids is not complete. The concentrated phosphoric acid produced (about 75 wt% H3PO4) contains approximately 5 wt% H N Q . The removal of H N Q requires a high energy input. Methyl isobutyl ketone (MIBK) dissolved in aromatic diluents, e.g. toluene, can be used for the removal of nitric acid from phosphoric acid solutions. The characteristics of the extraction of nitric acid from very concentrated phosphoric acid solutions (50-75 wt% ) were studied [9]. The following section is designed for the further processing of the H3PO4 solution, in three subsections.
Primary concentration (multi-effect evaporation) (a) Removal of the organic solvent as well as evaporation of water from the H:~PO4 solution. The feed stream is produced in the stripping step of H3PO4 recovery section. The organic solvent (iAmOH) dissolved in the aqueous solution is readily evaporated together with a fraction of water at the first stage. After cooling and phase separation, the recovered solvent is recycled back to the extraction step. (b) The reduction of H N Q content as well as the further concentration of the H3P04 solution in the following two stages. The major part of HN03 is removed with the evaporation of the major part of the water. The solution is concentrated to 60-70 wt% H3PO4, which may contain about 5 wt% HNO:~. It is then pumped to the following H N Q removal subsection.
Nitric acid removal (extraction III) (a) Selective extraction of H N Q from the concentrated H~PO4 solution. The extractant used is MIBK in toluene (50 vol% (MIBK). The concentration of HN03 in the raffinate is to be reduced to lower than 0.01 wt%, while H3PO4 concentration has a level similar to or higher than that in the feed (60-70 wt%). The raffinate is then sent for final concentration. This sub-section is operated under ambient temperature, 20 ° C. (b) Stripping of H N Q loaded in the organic phase. The loaded solvent is stripped with distilled water. The stripped solvent is then recycled back to the extraction step, while the stripping solution is sent to the H3PO4 recovery section.
Final concentration (distillation) The final removal of the organic solvents in the phosphoric acid, as well as some water, is achieved. The purified HaPO4 solution is concentrated to 75 wt% H3PO4. The organic solvents dissolved, MIBK and toluene, are evaporated together with a fraction of the water during the distillation. The recovered solvent and distilled water are then re-used in the previous extraction III. A general flow scheme of this section is given in Fig. 6. The specific operating
286
L'
iAmOHHNO3~ H20
i L"~I~A'~a',,T
(2-3 Effects) Feed 0.8M ~ 3 4
STRIppING (2-3 Stages)
EXTRACTION (5-8 Stage~)
I
LIATION
ct
2-3M
60-70%H.9~4/5-1O-%~o3 Fig. 6. A flowsheet of the phosphoric acid concentration and purification section.
1.6
[HNO3]o (M)
12
/
j
0.4 ~
Stripping
0 0
0.2
0.4
I 0.6
I 0.8
I 1.0 1.2 [HNO3]~ (M)
Fig. 7. Calculated number of theoretical stages for the extraction and stripping of nitric acid from 65% phosphoric acid by 50% MIBKdissolved in toluene.
modes for each subsection can be estimated in further design work. Figure 7 gives an example of the basic calculations of the nitric removal subsection (extraction III). As seen, for the given operating conditions, the number of theoretical stages is four for the extraction as well as for the stripping operation. NITRIC ACID REGENERATION SECTION
Since the solubility of gypsum, C a S O 4. 2H20, in nitric acid is low, maximum 2.5 wt% at 25 oC [ 15 ], the regeneration of HNO3 can be easily achieved by the precipitation of CaSO4-2H20. The nitric acid solution can be recycled for further leaching. During the leaching, however, precipitation of calcium sulfate occurs as HNO3 concentration decreases and Ca(NO3)2 concentration in-
287 creases, This will result in an increase of the sulfur content in the ore, which is undesirable with respect to the quality of the iron ore. The determination of the solubility of CaS04 in the aqueous systems H+-Ca2+-NO~]--S04, related to the composition of the present process solutions is reported in [10]. The results indicate that the direct reuse of the completely generated nitric acid at about 5 M will introduce ca. 7.9 kg CaS04 ( 1.7 kg S ) into 1 ton of the leached ore. In order to avoid the precipitation of gypsum during the use of regenerated nitric acid, a partial regeneration of the acid is suggested. In this way a certain amount (a steady-state concentration) of calcium nitrate will be maintained in the solution thus reducing the solubility of gypsum due to the common ion effect. This section is designed to consist of three subsections.
Calcium sulphate precipitation The extraction raffinate resulting from H3PO4 recovery section is mixed with 97 wt% H2804 in a stirred tank, and the solid CaSO 4- nH20, or gypsum, is formed. The stoichiometric requirement for H2SO4 addition is that the total sulphate concentration should be 0.7-0.8 M lower than the final calcium concentration. The resultant suspension, which has an acidity of 5-6 M HNO:~ in the solution, is then sent to the following H N Q recovery subsection.
Nitric acid recovery (extraction II/crystallization) (a) Selective extraction of H N Q against sulfate in the solution. The suspension is fed into first stage of the three-phase extractors (mixer-settlers or columns), iAmOH is used as a solvent. In this way, a complete extraction of H N Q into the organic phase results in the efficient crystallization of the gypsum as the acidity of the aqueous solution is greatly reduced. The gypsum precipitates at the bottom of the extractor where the aqueous raffinate contains less than 0.01 M of acids, including HN0:~. This suspension is sent to gypsum removal subsection. (b) Scrubbing of the loaded solvent for the removal of the sulphate. A side stream of the stripping solution is used for scrubbing, and then added to the extraction feed (suspension). (c) Stripping of the scrubbed solvent for primary recovery of H N Q . The strip feed is a mixture of the washing streams, containing 0.5-1 M HNO~. The stripping product may contain 3-4 M H N Q , and is to be concentrated to 5-6 M by distillation in order to be re-used in leaching. (d) Washing of the stripped solvent for further recovery of HNO~. Distilled water is used for washing. The washed solvent would contain > 0.01 M HNO~ and is then recycled for extraction. The washing product is added to the strip feed.
288
H2S04 (98~)
Q
<0.1M H3PO4 [ IM HNO 3 1 3M Ca(NO3)z t
<0.01M HNO 3 IM Ca(NO3)2
~_~..~o~
~
_J ~ o ~
~_~
~
'!IF[(5"7RSt ~ =ages) 2
~l ~
Ill (4-6 Stages) ~
~
.
i
~s~,~ I
-
(2-3R=10 IStages) ~7 1 (2-3~Stages) --
iAmOU
.2o,
~
.%,
[ U20
. . . . . .
"-!
V"
oypsu~
(4-6 S t a g e s ) ~
T
©
'
R-I
56M.No,
~a(~O~®
Fig. 8. A flowsheet for the nitric acid regeneration section.
[HN03] o (M) 4.0
Strip
3.0
e = 1
4---
0=2
/ / In/ /
2.0
Extraction
iv
1.0
0
1.0
I
I
I
2.0
3.0
4.0
I
5.0 6.0 [HNO3] a (U)
Fig. 9. Calculated number of theoretical stages for the extraction and stripping of nitric acid by iAmOH.
289
Gypsum separation (a) Further crystallization of gypsum, as well as gravity sedimentation, in the raffinate solution: the suspension from the raffinate is fed into a continuous thickener. The clear overflow stream is sent to the P-removal section where it will be used for washing the leached ore, while the thickened underflow is fed to the filtration step. (b) Filtration of gypsum: The filtrate is pumped back to the thickener, while the gypsum is to be washed with a washing stream of p H 7-8 from the Premoval section. A general flow scheme of this section is given in Fig. 8. The specific operating modes for each subsection should be developed in further design work. An example of the calculation is given for the extraction II, as shown in Fig. 9. ASSESSMENT OF PROCESS ECONOMY We have examined the economics for the process described above. The cost estimates were made for a plant that processes 4 million tons of iron ore per year and produces high purity phosphoric acid (75% H3P04) [ 11 ].
Cost estimates The capital investment costs were estimated. The estimates calculated for the different items, e.g. equipment, building, etc, are given in Table 1. As seen, the total capital investment is estimated to be about 444 million Swedish crowns TABLE1 The capital investment costs (in millions of Swedish crowns)
Production units: Phosphorus removal section H3P04 recovery section {extraction I) HN03 regeneration section Acid concentration section Solvent inventory Subtotal Contingency (10%) Auxiliary: Material storage Gypsum disposal Land and preparation Subtotal Total
80.5 34.8 40.5 150.6 8.8 315.2 31.5 61.5 32.7 2.6 96.8 443.5
290 TABLE 2 The operation costs (in Swedish crownsper ton ore) Materials: Sulfuric acid ( > 97% ), 76 kg Nitric acid (> 63%, 2 kg (make-up) Others (water, lime, solvent make-up, etc. )
19.0 1.5 1.0
Subtotal
21.5
Energies: Steam Electricity, 26 kWh
12.0 5.9
Subtotal
17.9
Labor Maintenance, 4% of total investment Others
Total
2.5 4.4 5.6 51.9
(61 million US$) at 1986 values. The operation costs are estimated from the process parameters. The material costs were calculated from the process material balance mainly concerning H2804 and H N Q consumptions. The energy costs were calculated from the energy balance mainly for the evaporation and distillation. These costs are specified in Table 2. As seen, the total operatingcost is estimated to be 52 Swedish crowns per ton ore or 208 million Swedish crowns (28.5 million US$) per year. Revenue and profitability estimation
The increase of the ore value, as a result of dephosphorization, was set by the LKABcompany to 17-20 Swedish crowns per ton ore processed (1980). The same value is used also in the present estimation. From the material balance, the production of phosphoric acid (75 wt%) is calculated to be more t h a n 32.0 kg per ton ore (considering a total of 80% recovery). The present price is taken to be 350-400 US$ per ton of 75% H:~PO4 for technical grade (ex factory; Supra AB, Sweden). The higher purity acid produced by thermal process is at least 50% higher in price. The purity grade of product acid of this process may be sufficiently high to meet the requirement of food grade acid. The revenue and profit are estimated in Table 3. As seen, the revenue is to be 100-160 Swedish crowns per ton ore or 400-640 million Swedish crowns (55-88 million US$) per year at a full production capacity, while the operating profit 28-90 Swedish crowns per ton ore or 110-
291 TABLE 3 The revenue, income and profit (in Swedish crowns per ton ore) Quality of acid:
Technical
Estimation level:
low
Ore value increase Value of 75% H:~PO4
High purity high
low
high
17.0 81.8
20.0 93.4
17.0 122.6
20.0 140.2
Revenue Operating cost
98.8 - 51.9
113.4 - 51.9
139.6 - 51.9
160.2 - 51.9
Income Interest on loan Depreciation
46.9 - 11.1 - 7.9
61.5 - 11.1 -- 7.9
87.7 - i 1.1 - 7.9
108.3 - 11.1 - 7.9
27.9
42.5
68.7
89.3
Profit
360 million Swedish c r o w n s ( 1 5 - 5 0 million U S $ ) per year. T h i s indicates a high p r o f i t a b i l i t y of t h e process u n d e r t h e s e conditions. CONCLUSIONS F r o m the results r e p o r t e d t h e p r o p o s e d process is s h o w n to be t e c h n i c a l l y feasible. T h e quality of the leached ore has i m p r o v e d u p o n leaching. P h o s p h o r i c acid could be r e c o v e r e d as h i g h - p u r i t y grade. T h e e s t i m a t e s of the process e c o n o m y indicate t h a t the process can be o p e r a t e d with profit. However, the e c o n o m y of the process is very m u c h d e p e n d e n t on: (a) price of p h o s p h o r i c acid; (b) e n e r g y costs; a n d (c) costs of sulfuric acid. ACKNOWLEDGEMENTS T h e work has b e e n f i n a n c e d b y the Swedish B o a r d for T e c h n i c a l Developm e n t (STU). T h e a s s i s t a n c e a n d c o o p e r a t i o n of several colleagues f r o m t h e d e p a r t m e n t s of C h e m i c a l E n g i n e e r i n g at the R o y a l I n s t i t u t e of T e c h n o l o g y a n d at the U n i v e r s i t y of L u n d are gratefully acknowledged.
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