A CLEAN TECHNOLOGY PHOSPHORIC ACID

PROCESS

S. van der Sluis

TR diss 1550

V» T T' A C L E A N TECHNOLOGY * ' PHOSPHORIC ACID

fiy & ^ PROCES ■Tl

iu» l f S "

A CLEAN TECHNOLOGY PHOSPHORIC ACID

PROCESS

Proefschrift ter verkrijging van de graad van doctor

aan de Technische Universiteit Delft, op gezag van de Rector Magnificus,

Prof.dr. J .M. Dirken, in het openbaar te verdedigen

ten overstaan van een commissie aangewezen door het College van Dekanen

op 11 juni 1987 te 14.00 uur

door

Sierd van der Sluis

scheikundig ingenieur geboren te Oudemirdum

Delft University Press/1987

TR diss 1550

Dit proefschrift is goedgekeurd door de promotoren: Prof .dr.ir. G.M. van Rosmalen Prof.ir. J.A. Wesselingh

Copyright ©1987 by S. van der Sluis

No part of this book may be reproduced in any form by print, photoprint, microfilm or any other means without written permission from the publisher: Delft University Press, Delft, The Netherlands.

Voor Ineke en onze kinderen

VOORWOORD

Dit proefschrift is mede tot stand gekomen door de inzet van veel personeel, materieel en geld door de overheid en DSM, waarvoor ik hen hierbij bedank. De onmisbare bijdragen van vele afstudeerders, stagiairs, bijvakkers en omscholers, vaste medewerkers en collega's zijn in een aantal gevallen al terug te vinden in een co-auteurschap van een artikel of in woorden van dank, die daarbij vermeld staan. Zij, die om wat voor reden dan ook, daar niet bij genoemd zijn, wil ik hierbij alsnog bedanken voor hun inzet en bijdragen. Een speciaal woord van dank is op zijn plaats voor mijn collega Tjien T. Tjioe, want zonder zijn adviezen en hulp was dit proefschrift zeker nu nog niet tot stand gekomen. Tevens past mij een woord van dank voor de ondersteuning, die de werknemers en de diensten van de Technische Universiteit Delft, van DSM Meststoffen B.V. en van DSM Research B.V. mij gegeven hebben. Bovenal wil ik de grote groep medewerkers en medewerksters bedanken die, onbezoldigd, dag en nacht klaar stonden om, in continu dienst, proeven met het fosforzuur fabriekje op laboratorium schaal uit te voeren en hen, die daarbij weliswaar thuis bleven, maar dag en nacht oproepbaar waren. Tenslotte wil ik die mensen bedanken, die het op welke manier dan ook, mogelijk gemaakt hebben om bij verschillende fosforzuur fabrieken over de hele wereld rond te kijken en via het houden van lezingen op binnen- en buitenlandse congressen en symposia, enige internationale ervaring op te doen.

CONTENTS

1. INTRODUCTION 7 1.1. General aspects of phosphoric acid 7 1.2. The phosphate ore 7 1.3. Phosphoric acid production processes 11 1.3.1. Introduction of process routes 11 1.3.2. The dry processes 11 1.3.3. The wet processes 12 1.3.3.1. General procedure 12 1.3.3.2. Acidulation with hydrochloric acid 13 1.3.3.3. Acidulation with nitric acid 14 I.3.3.I. Acidulation with sulphuric acid 16 1.1) Impurities in relation to the applications of the calcium

sulphate byproduct 19 1.5 Aim of this study 21 1.6 Literature 21

2. SCOPE OF THIS INVESTIGATION 24

3. OVERVIEW OF THE CTPA PROCESS 28 3.1. Summary 28 3.2. Introduction 29 3.3. Objective of the study 32 3.4. Description of the CTPA process 32 3.4.1. Introduction to the process 32 3.4.2. The digestion stage 33 3.4.3. The crystallisation stages 37 3.4.4. The filtration stages 41 3.4.5. The fluoride removal 46 3.5. Simplified process flowsheet and mass balance 50 3.6. Conclusions 52 3.7. Literature 53

2

14. THE DIGESTION OF PHOSPHATE ORE IN PHOSPHORIC ACID 56 4.1. Summary 56 4 .2 . In t roduc t ion 56 4 . 3 . The d iges t ion s t a g e of the CTPA process 57 4 . 4 . Experimental 59 4 . 5 . Resul t s 60 4 . 6 . Discussion 62 4 . 6 . 1 . Inf luence of the phosphoric acid concent ra t ion 62 4 . 6 . 2 . Inf luence of the temperature 63 4 . 6 . 3 . Conclusion 63 4 .7 . A k i n e t i c model of the d iges t ion process 63 4 . 8 . Determination of the masst ransfer c o e f f i c i e n t s 66 4 .9 . Conclusive remarks 68 4 .10. Nomenclature 68 4 . 1 1 . L i t e r a t u r e 69

5 . CRYSTALLISATION OF CALCIUM SULPHATE HEMIHYDRATE 70

5 . 1 . Summary 70 5.2. Introduction - 70 5.3. Experimental 72 5.4. Results and discussion 71 5.4.1. The hemihydrate crystals 74 5.4.2. Incorporation of phosphate ions 77 5.4.3. Incorporation of cadmium ions 80 5.5. Conclusions 84 5.6. Literature 84

6. THE FILTRATION OF CALCIUM SULPHATE HEMIHYDRATE 86 6.1. Introduction 86 6.2. Theory of filtration 87 6.3. Experimental 90 6.3.1. Equipment 90 6.3-2. Chemicals 91 6.3.3. Procedure 91 6.4. Results and discussion 91 6.5. Nomenclature 97 6.6. Literature 97

FLUORIDE DISTRIBUTION COEFFICIENTS (G/L) IN WET PHOSPHORIC ACID PROCESSES 98 Summary 98 Introduction 98 Literature survey 100 Model development 102 General remarks 102 Calculation of the activity coefficients 103 Calculation of the hydrogen ion concentration in phosphoric acid 105 Determination of an expression for the fluoride distribution coefficients 107 Experimental 111 Chemicals 111 Equipment 111 Procedure 11 2 Analyses 113 Results and discussion 114 Conclusions 120 Nomenclature 121 Literature 122 Appendix: Conversion of w$ into molalities 124

MASS AND HEAT BALANCES OF THE CTPA PROCESS 125 Introduction 125 Process description 126 General approach 126 Digestion of phosphate ore 127 Cadmium removal 129 Crystallisation of calcium sulphate hemihydrate (HH) 130 Recrystallisation of HH to gypsum 131 Solid-Liquid separation 131 Fluoride removal 132 Combined mass and heat balances for each stage 132 Discussion and conclusion 147 Literature 149 Appendices 151 The solubility of HH and DH 151

4

8.6.2. Incorporation in HH and DH 151 8.6.2.1 Incorporation in calcium sulphate hemihydrate 151 8.6.2.2 Incorporation in calcium sulphate dihydrate 152 8.6.3. Vapour pressure of the PpO^-H^O system 152 8.6.1. Fluoride distribution coefficients 153 8.6.1.1. Liquid-gas 153 8.6.1.2. Liquid-solid 151 8.6.5. Density correlations 155 8.6.5.1. Phosphoric acid 155 8.6.5.2. Sulphuric acid 155 8.6.5.3. CDHP solutions 155 8.6.5.1. *V01 -^2S0H~H2° mixtures 156 8.6.5.5. Calcium sulphate hemihydrate (HH) 156 8.6.5.6. Calcium sulphate dihydrate (DH) 156 8.6.5.7. The phosphate ore 156 8.6.5.8. Water 156 8.6.6. Heat capacities and heat contents 157 8.6.6.1. Phosphoric acid 157 8.6.6.2. Sulphuric acid 157 8.6.6.3. CDHP solutions 157 8.6.6.1. H POjj-H SO^-H 0 mixtures 158 8.6.6.5. Calcium sulphate hemihydrate (HH) 158 8.6.6.6. Calcium sulphate dihydrate (DH) 158 8.6.6.7. The Phosphate ore 159 8.6.6.8. Water 159 8.6.6.9. Water vapour 159 8.6.6.10. Carbon dioxide 159 8.6.6.11. Other compounds 160 8.6.7. Enthalpies of reaction 160 8.6.7.1. Digestion of phosphate ore 160 8.6.7.2. Crystallisation of calcium sulphate hemihydrate 160 8.6.7.3. Recrystallisation of HH into gypsum 160 8.6.8. Heat of mixing in the H P0.-H SO.-H 0 system 161 8.6.8.1. General procedure 161 8.6.8.2. Binary heat of mixing in the H^O^-H-O sytem 163 8.6.8.3. Binary heat of mixing in the H^O^-H^ sytem 163 8.6.8.1. Binary heat of mixing in the H^PO^-H SO. system 163 8.6.9. Total mass and heat balances of the CTPA process 161

5

9. 9.1. 9.2. 9.2.1 9.2.2. 9.2.3. 9.2.1. 9.2.4.1. 9.2.1.2. 9.2.1.3. 9.2.1.1. 9.2.5. 9.3. 9.3.1. 9.3.2. 9.3.3. 9.3.1. 9.3.5. 9.1. 9.5. 9.6. 9.6.1. 9.6.2. 9.6.2.1. 9.6.2.2. 9.6.2.3. 9.6.3.

THE BENCH-SCALE PLANT Introduction Experimental Process conditions Chemicals Equipment Procedures Safety Startup procedure Working procedure Stop procedure Analyses Results and discussion The performance of the bench-scale run The cadmium incorporation in HH The phosphate incorporation in HH The permeability of the filter cakes Additional results Conclusions Literature Appendices Service round Pressure filter procedures The precoat procedure The filter procedure The filter change procedure The HH washing procedure

175 175 176 176 177 177 180 180 181 183 183 181 185 185 187 188 189 191 193 193 191 191 195 195 195 195 196

SUMMARY 198

SAMENVATTING 200

PUBLICATIONS 202

1. INTRODUCTION

1.1. General aspects of phosphoric acid

Phosphoric acid is an important intermediate chemical product. It is mainly used for the manufacturing of fertilisers, as demonstrated in table 1.

» 1000 tons P_0C percentage

Fertilisers Detergents Animal feed Food and beverages Surface treatment Water treatment Dentistry, tooth pastes Fire extinguisher Others

Total

32 1 1

,000 ,590 , 180 . 240 230 90 80 40

110

35,600

90 4.5 3.3 0.7 0.6 0.25 0.22 0.11 0.3

100

Table 1: World phosphate consumption and application in 1980 [1].

To meet the food needs of the increasing world population, there is a steadily growing demand for phosphate fertilisers. The phosphoric acid production is directly linked to the phosphate fertiliser consumption, which is expected to rise from 30.7 million tons P-O,- in 1982, to about 41.7 million tons

<!■ 5

P-0,- in 1990. This represents an increase of the phosphoric acid production from about 19.7 million tons P_0C in 1982 to 27.8 million tons P.O.- in 1990 [3].

£ 0 2 5 1.2. The phosphate ore

P2°5'

According to Becker [1], about 200 known minerals contain more than 1 w$ The most important minerals for the phosphoric acid industry are those

belonging to the apatite group, with the general formula Ca. 0(P°i|)6X2' in w n i c n

X can be fluoride, chloride or hydroxide [12]. Because the fluoride ion is the smallest of the three, fluoroapatite is commonly believed to be the most stable.

7

The phosphate rock normally not only contains the apatite mineral, but also several other minerals, such as calcite, dolomite, pyrite, kaolinite, quartz, feldspar and fluorite [14]. Phosphate rock deposits are located in abundance in nearly all parts of the world. The phosphate rocks can be divided in two groups, depending on their origin: - sedimentary phosphate rock - igneous phosphate rock

Igneous phosphate rock is crystallised from magma and thus found in areas with vulcanic activity. Its availability is therefore limited. The igneous phosphate rock mainly consists of fluoroapatite, in which calcium is partly replaced by barium and strontium and chloride partly substitutes fluoride.

Sedimentary phosphate rock deposits were formed by precipitation of dissolved phosphates from prehistoric seas [32]. During this process the phosphate rock was subjected to an interaction with water of varying temperature and impurity content. A large number of these impurities became included in the rock by either coprecipitation or by incorporation in the apatite crystal lattice. In this way phosphate rocks with a large variety in impurity composition and concentration were created. Most sedimentary phosphate rocks are found to be more than a million years old [27]. The sedimentary phosphate rock mainly consists of fluoroapatite, in which part of the phosphate ions are replaced by fluoride and carbonate ions. This fluoroapatite is commonly refered to as francolite.

In 1981 more than 85 % of the world's raw phosphate came from sedimentary deposits [1]. The four major raw phosphate producers in 1981, were the USA, the USSR, Morocco and China [1]. The raw phosphate, produced by open mining, contains several other minerals next to the apatite mineral. Only the apatite mineral is needed for the production of phosphoric acid. The other minerals have to be removed prior to the phosphoric acid production. The raw phosphate is normally first washed with water to remove the water soluble minerals. Thereafter several flotation steps are applied to remove the larger part of the other minerals and to obtain the phosphate ore. This process to obtain phosphate ore from raw phosphate is often called beneficiation. The removal of in particular the siliceous and carbonate gangue is necessary to avoid serious problems during the phosphoric acid production. The siliceous gangue consists of small particles, which can cause problems during filtration. Carbonate gangue removal is a prerequisite prior to the digestion of the phosphate rock by an inorganic acid. Otherwise the development of carbon dioxide in combination with the release of organic material from the rock gives rise to excessive foaming.

The current techniques for beneficiation of phosphate rock are reviewed by Lawver e.a. [21]. The procedure selected for a specific phosphate rock is determined by its carbonate content as well as by its origin. Sedimentary and igneous deposits are treated differently. One of the most important processes for the beneficiation of phosphate rock, which is applied on all rocks, is flotation. An overview of the most frequently used flotation techniques is given by Houot [16]. Flotation is used mainly to remove the siliceous gangue, but in some cases also the carbonate gangue can be removed by this technique as described by Dufour [13] for Florida phosphate rock and by Vaman RaO [39] for Mussorie phosphate rock.

The phosphate ores obtained after beneficiation of sedimentary phosphate rocks, consist mainly of francolite. McClellan and Lehr [23] derived a formula for the mean composition of francolite by averaging the composition of 110 commercially available sedimentary phosphate ores:

CaiO-aVVVPV6-x(C03)xF2+0.<,X with x = 0 - 1.5 and a = 0.6 * x. The 2H sedimentary phosphate ores described by Gremillion e.a [1t] conformed very closely to this mean francolite composition.

The chemical composition of phosphate ores varies widely with their mining location [11]. McClellan and Lehr [23] found the weight ratio of Ca0/P205 to vary from 1.32 in pure fluoroapatlte, to about 1.62 in highly substituted francolite, while according to Caro e.a. [7] the weight ratio of F/P-O- varies from 0.09 to 0.15 respectively. The influence of the weight ratio FVP2°5 on t n e

reactivity of the ore with acid is unclear. The reactivity of the ore with acid was found to be higher at increasing C0./P„0c weight ratio in the phosphate ore.

2 2 5 This observation was confirmed by Chien [10,11], who came to the same conclusion from a thermodynamic approach. i The crystallographic structure of fluoroapatlte has been described in detail by Beevers and Mclntyre [2]. Montel e.a. [2*4] report the fluoroapatlte crystal lattice to have a hexagonal structure with two series of parallel channels. One series of narrow channels has a diameter of about 2 A, while the other series consists of wide channels with a diameter of about 3.5 A, each channel containing the fluoride ions centered on its axis. The fluoroapatite lattice therefore has a very open structure. According to Kreidler [20] almost all ions seem to fit to some extent in the apatite lattice, due to its open structure.

9

Not only the fluoroapatite lattice itself has an open structure on an atomic scale, but also the fluoroapatite as a mineral has an open structure on a microscopic scale. Caro e.a. [6] investigated the pore structure of phosphate ore with mercury porosimetry and nitrogen desorption. By comparing the pore size distributions of large particle size fractions with small particle size fractions, they found the fine pore structure to be uniformly distributed throughout the material, irrespective of particle size. Hill e.a. [15] measured specific surface areas and found, for particle size fractions smaller than 150

2 um, these areas to be varying between about 1 m /g for Virginia phosphate ore 2

and 37 m /g for Tunisian phosphate ore. This relatively large surface areas are attributed to pore space within the particles. These particles are often viewed as aggregates of fine crystals. The average dimension of these elementary grains is deduced from the broadening of X-ray diffraction bands and compared with a figure calculated from surface area measurements of several suitably fine samples (10 ym). The results obtained by both methods are within one order of magnitude. The elementary grain size for Moroccan and Tunesian phosphate ore was found to be about 100 - 500 A, while for Virginian phosphate ore, it was approximately 1000 - 3000 A. The average dimension of the elementary grains is one of the most important factors determining the reactivity of the ore [15].

S I C » • COM • H U T * PHOSPHORUS

ncm SUM ACE •vntnm

Q

VBCOSITT COITML ABUTS

Figure 1: Uses of phosphoric ac id produced by dry and wet processes [ 3 2 ] .

10

1.3 Phosphoric acid production processes

1.3.1. Introduction of process routes

Phosphoric acid can be produced from phosphate ore via two major process routes: the so-called wet processes, using strong mineral acids for digestion of the ore and the dry processes, producing elemental phosphorus as an intermediate by burning of the ore in an electric furnace or in a rotary kiln [3]

An outline of the field of applications of phosphoric acid produced by dry processes as well as by wet processes is given in figure 1 [32]. Since phosphoric acid produced through dry routes contains less impurities its application lies mainly in the "high added value" areas such as detergents and food additives, while phosphoric acid from wet processes is mainly further processed into fertilisers.

1.3.2. The dry processes

In the dry processes, the phosphate ore is reduced, by addition of silica and carbon, to phosphorus and slag at about 1500 °C in a furnace:

Ca1Q(POl))6F2 + 20 Si02 + 28 C — > 20 CaSiO + 28 CO + 3 Pj, + 2 ?2

AH = 27.9 MJ/kg of ?k produced [22]

The heat, AH needed for this reaction is substantial and could until 1985 only be achieved by internal resistence heating of the molten charge in an electric furnace. In 1985, however, a direct fired rotary kiln was found to be able to supply the heat [22]. A flowsheet of the KPA process, based on this way of supplying the necessary heat is shown in figure 2.

After the combustion of the evaporating phosphorus, the obtained phosphorus pentoxide is absorbed in water or in moderately concentrated phosphoric acid, to obtain concentrated phosphoric acid (> 60 w$ P-0,-).

The direct production of clean and concentrated phosphoric acid is possible with the dry processes, because most impurities from the ore, remain in the slag [3].

Although the heat of combustion of the phosphorus and the carbon monoxide is recovered in the KPA process, these processes are still too expensive for use in the fertiliser industry, because of their high energy consumption. Moreover,

11

fine ore particles cannot be used directly in these processes, since they hamper the interstitial gas flow through the ore bed.

The product of the dry processes can, however, be used in applications, where the product from the wet processes is not sufficiently pure. This does, sometimes, require additional removal of arsenic and fluoride.

vent

phosphate ore

coke ^

silica

f raw feed preparation

1 i i

wafer

t balling preheating

A i i

fuel

---1

— —

I

air

— 1

rotary kiln reactor

'

heat recovery

* j

spent

gas

scrubbing

i ,

acid absorption

acid cleaning

solids »-

70 w % P 2 0 5

Figure 2: Flowsheet of the KPA process [22].

1.3.3. The wet processes

1.3.3.1. General procedure

Phosphoric acid can be released from phosphate ore by the action of strong mineral acids, such as nitric acid, hydrochloric acid and sulphuric acid. Sulphuric acid is the only acid, which forms an insoluble precipitate with the calcium from the phosphate ore, thus allowing the phosphoric acid to be separated directly by filtration. The chlorides and nitrates of calcium are both soluble, so special techniques, like solvent extraction [28], ion exchange [19] or cooling crystallisation [1,36] are required to recover the phosphoric acid. In spite of this phosphoric acid is being commercially produced using nitric acid as well as hydrochloric acid. The economical feasibility of these processes is mainly determined by the availability and price of the mineral acid and by local production facilities and by the desired products.

12

1.3.3.2. Acidulation with hydrochloric acid

The acidulation of phosphate ore in hydrochloric acid can be represented by the following reaction [28]:

Caln(P0„),Fo + 20 HC1 — > 6 H_PO.. + 10 CaCl. + 2 HF 1 0 4 O £ j 1 2

phosphate ore

m gas

i r dissolution

recyc le wash

solids separation

to s extr

solids washing

dissolution liquor

olvent action to » aste

dissolution liquor

recycle

stripped liquor aqueous acids t o " l»ent to distillation recovery

water ' reflux i J

J1"""J-L •^extraction)—■»] pufification] —) washing] —| stripping}-

extracted aqueous liquid

circulating solvent

- solvent streams - aqueous streams

Figure 3= Flowsheet of the ore digestion section [28].

Figure 4: Flowsheet of the solvent extraction section [28].

The hydrogen chloride can be supplied, either as a gas entering an aqueous suspension of phosphate ore or as hydrochloric acid. The phosphoric acid can be separated from the calcium chloride solution by solvent extraction with C,. and C_ alcohols or with a mixture of these solvents.

In figure 3 a flowsheet of the ore digestion section of the IMI process is shown and in figure 4 the corresponding flowsheet of the solvent extraction section. The phosphoric acid and the quantity of hydrochloric acid dissolved in the solvent are recovered from the solvent by washing with water. Thereafter the solvent is distilled off and subsequently the phosphoric acid and hydrochloric acid are separated by evaporation.

An additional advantage of this method is the reduced impurity level of the obtained concentrated phosphoric acid (> 50 w$ P_Oc), because the impurities are less easily soluble in the solvent than in phosphoric acid.

A disadvantage, however, is the severe corrosion.

13

1.3.3.3. Acidulation with nitric acid

There are three main processes in which nitric acid is used for acidulation of phosphate ore: the Odda process, the DSM sulphate-recycle process [5,36] and the Superfos process [19].

In the Odda process, the phosphate ore is digested in 60 w% HNO. at 50-70 °C. Thereafter, a large part of the calcium nitrate is removed by cooling crystallisation and separation of the crystals. The calcium nitrate crystal slurry is further processed to obtain a fertiliser, while the filtrate is partly recycled. This is shown in figure 5.

phosphate HNO,

dissolution -L^-N°x.F

-■ - iner ts

I crystallisation |

filter . . .

— HN03

1 Ca(N03)2AH20 ' Ito conversion)

neutralisation

— NH3 r » H 2 0 ■L—NH2NO3

(from conversion)

evaporation |

granulation

Figure 5: Flowsheet of the Odda process [5].

In the DSM process, the phosphate ore is digested with nitric acid at 65 °C, whereafter the solution, containing insoluble phosphate ore particles, reacts with ammonium sulphate at 55 °C. The calcium ions are precipitated as gypsum crystals and filtered off. Subsequently the moisture content of the filtrate is reduced after neutralisation by steam evaporation to obtain an ammonium phosphate fertiliser. The gypsum cake is fed, together with a 52-53 w$ ammonium carbonate solution into a crystalliser, in which calcium carbonate is precipitated and ammonium sulphate is recovered for recycling. A flowsheet of this process is shown in figure 6.

The third process is the so-called Superfos process, in which the phosphate ore is digested with nitric acid. The so obtained calcium nitrate and phosphoric acid containing solution is fed into an ion exchange section, where the calcium

11

ions are exchanged against potassium ions. The ion exchange res in , loaded with calcium, is regenerated with a potassium chloride solution, giving an CaCl effluent stream. The product, consisting of a KNO and H PO solution is used as an intermediate in the manufacturing of f e r t i l i s e r s . In a process modification, where not al l the calcium ions are removed, these residual calcium ions are precipitated with the major part of the phosphate as CaHPO.. A feed grade quality i s claimed, which i s only possible if the impurities were removed before or in the ion exchange sect ion. This process is i l lus t ra ted in figure 7.

Phosphate OPEJ I " ™ *

I V. \ ,. h * N0X.F | dissolution [_~ i n e r t s

JGypsum precipitation}—— INHJ3SQ1, (From conversion)

granulation

Figure 6: Flowsheet of the DSM process [5].

potassium chloride water phosphate ore

nitric acid

phosphoric acid

ammonia I

1 l_.

♦ '

L_ 1 __K 1 ion exchanqe ]

L

CaCl2

| concentration!

"I * ^neutralisation!

| granulation ]

chlorid

1

>-freefi PK

pre-neutrs

separation |

r

| crystallisation f-«—

| separation | 1

■

| drying |

dicalcium

ammonia

to NPK plant

phosphate

Figure 7: Flowsheet of the Superfos ion exchange NPK process with CaHPO. coproduction [19].

15

1.3-3.1). Acidulation with sulphuric acid

Although all acidulation processes are refered to as wet processes, the term wet process is mainly reserved for processes in which sulphuric acid is applied.

During the production of phosphoric acid from phosphate ore by acidulation with sulphuric acid, huge amounts of calcium sulphate are precipitated as a byproduct. Depending on the temperature as well as on the phosphate and sulphate content of the solution, either calcium sulphate dihydrate (DH), hemihydrate (HH) or anhydrite (AH) is formed, as shown in figure 8 [1,35].

The solid lines in figure 8 represent quasi equilibrium curves, indicating which phase will initially precipitate under the given conditions. The broken line in figure 8, is the border of the regions, where either AH or DH is stable. The HH phase only exists as a metastable phase. The influence of sulphuric acid can be taken into account by assuming one mole of sulphuric acid to be equivalent with about 1.5 mole of phosphoric acid [32].

100

temp°[ 80

60-

40'

20 20 30 40 50

w%P 20 5 ^

Figure 8: Phase diagram of calcium sulphate in phosphoric acid.

Commercial processes have been developed, producing either DH, HH or AH as a byproduct. An overview of these processes is given in table 2 [31], where the process is named after its byproduct.

AM precipitation

HH precipitation

DH precipitation

- - C A H stable

/ D H stable"^N

16

process | temperature | concentration | company/name

| w* P2°5

DH

HH

DH/HH double f i l t e r

HH/DH single f i l te r

HH/DH double f i l t e r

70 -85

90 - 100

stage 1 : 65 -stage 2 : 95 -

stage 1 : 90 -

stage 2 : 55 -

9 0 - 1 0 0

70 100

100

65

28

35

35

30 -

30

- 32

- 50

- 38

■32

- 50

Prayon, Jacobs Dorco, SIAPE Kellog-Lopker, Norsk Hydro, IITPIC, Rhone Poulenc

Norsk Hydro, Jacobs Dorco, Occidental

Prayon

Nippon Kokan, Nissan Mitsubishi

Singmaster and Breyer,

Norsk Hydro, Nissan, Nippon Kokan

AH | 100 -240 | 4 0 - 5 0 | (Nordengren)

Table 2: Wet phosphoric a c i d p roces ses .

The advantages and d isadvantages of the various processes are summarized in d e t a i l by Becker [ 1 ] . The main disadvantage of the DH, the s i n g l e f i l t e r HH/DH and the DH/HH p roces ses , i s the production of r e l a t i v e l y d i l u t e (about 30 wj P-0,.) phosphoric ac id , which has t o be concentrated for use in f e r t i l i s e r a p p l i c a t i o n s . Due t o the energy consuming concent ra t ion s t e p , the HH and the two f i l t e r HH/DH processes , a l lowing d i r e c t production of 40 w$ P 0 or even higher concentra ted acid from the f i l t e r , a r e r ap id ly gaining f i e l d .

17

AH processes are not operational at this moment [1], due to the serious problems observed in small scale commercial batch plants [25], such as enhanced corrosion at the required high temperature as well as a prolonged residence time following from the low growth rate of AH.

Each process has its own requirements regarding raw materials, utilities, product and byproduct quality and last but not least, overall phosphate efficiency. An important disadvantage of the HH processes for instance, is the relatively low (< 95$) phosphate efficiency. An advantage of the HH/DH and DH/HH processes is the production of a relatively clean byproduct especially with regard to P 0 incorporation.

According to Becker [1], the ultimate process route in phosphoric acid production processes is the double filter HH/DH process route, because: - concentrated phosphoric acid is directly produced from the filter - the highest overall phosphate efficiency is obtained, - the lowest sulphuric acid consumption is needed, - relatively clean phosphoric acid is produced, due to the reduced solubility

of impurities in the concentrated acid and - phosphogypsum with a low phosphate content is produced.

ISULPHURIC ACI0|-

WASTE GAS

N°1 H°2 COOLING PUMP DIGESTER TANK TANK

N°1 N"2 N"3 HYDRATION TANKS

Figure 9: The Nissan C process .

18

In figure 9, one of the commercially available double filter HH/DH processes is shown [1].

Additional aspects, which have to be taken into account before a proper process selection can be made are: capital cost, maintenance cost and on-line time, raw material and utility cost and also product as well as byproduct quality. These aspects depend largely on the local situation.

1.1. Impurities in relation to the applications of the calcium sulphate byproduct

The phosphate ore contains a lot of impurities, as is shown in table 3 [to].

In all wet processes impurities, such as radium and heavy metal ions, like cadmium, originally present in the phosphate ore, are distibuted between the phosphoric acid and the byproduct. Several methods have been developed to remove impurities from the phosphoric acid, such as flotation [33], ion exchange [9,37] and solvent extraction [29]. One of the recently published solvent extraction processes is the BESA-2 process [30], which claims to produce a phosphoric acid quality comparable with the acid produced by the dry process for almost half the price of the dry process acid.

One of the most important environmental problems in phosphoric acid production is the cadmium content of the phosphoric acid and of the calcium sulphate byproduct. Cadmium concentrations as low as 1 mg/1 in water have led to-a painful bone disease, known as Itai-Itai [1].

18 ppm 250 ppm 10 ppm 30 ppm 20 ppm 50 ppm

<0.1 ppm

Table 3: Indicative concentrations of heavy metal ions and radio-active elements in phosphate ores from sedimentary deposits.

40 K 238(J 2 3 2Th 2 2 6Ra Ra-equi valence

6 pCi/g 40 pCi/g 1.2 pCi/g

38 pCi/g to pCi/g

Cd Zn Cu Pb As Ni Hg

19

In the last years, there is a growing concern, particularly in Europe, that this toxic metal ion could enter the food chain in increasing amounts. Some countries intend to adopt a cadmium limitation of no more than 90 mg cadmium per kg P?0(- in the fertiliser product, as a preventive measure [17,38]. The Netherlands plans legislation restricting the levels of heavy metal ions in phosphate fertilisers as well as in other products used in agriculture, and also the levels in the byproduct hydrated calcium sulphate. The disposal of this byproduct is limited by its cadmium content, while its application as an indoor buiding material is also hampered by its radium content. In phosphoric acid processes operating nowadays the calcium sulphate byproduct is either disposed in water or stacked on land [3]. In urban countries, however, the land stacking is no good alternative. Another possibility is to use the so-called phosphogypsum [40].

Applications already operational or still under investigation are [8,40]: - in the building industry for the production of gypsum board, blocks etc. - as a settling retarder for cement - in super sulphated cement - as a filling material for paper, plastics, paint, etc. - for conversion into sulphuric acid and cement additives - for conversion into sulphur - for conversion into ammonium sulphate - for agricultural use - as a substitute for sand in road works - in glass production - as new building materials - in the production of gypsum ammonium nitrate as a substite for calcium ammonium nitrate

In most applications, the cadmium content of the phosphogypsum remains a problem, because leaching of heavy metal ions cannot be totally prevented [18]. If the phosphogypsum is used for manufacturing of building materials, the radium content is also a major problem, because the radio-active radon gas, a decay product of radium, can build up in the atmosphere. Methods to reduce the radium content of the phosphogypsum, however, have already been developed [26,34,40].

20

1.5. Aim of this study

In the foregoing paragraphs the importance of the impurity content of both phosphoric acid and its byproduct, phosphogypsum, has been elucidated. From the viewpoint of environmental pollution action is required.

The aim of this study is the development of a new commercially competitive process for the production of concentrated phosphoric acid, where the following requirements must be fulfilled: - a low cadmium content and a low fluoride content of the phosphoric acid and - a low cadmium content and a low phosphate content of the calcium sulphate

byproduct as well as a low radium content to make the phosphogypsum suitable for building purposes.

1.6. Literature

1. Becker, P., Phosphates and Phosphoric Acid, Fertiliser Science and Technology Series, Vol. 3, New York, Marcel Dekker Inc. (1983).

2. Beevers, C.A. and Mclntyre, P.B., Min. Mag. 27 (1947) 254. 3. British Sulphur Corporation, Phosphoric Acid, Outline of the industry,

British Sulphur Corporation Limited, London, 4th. Ed. (1984). 4. Burova, M.S. and Kazak, V.G., Khim. Prom., 1 (1985) 29. 5. Calis, G.H.M., "The role of impurities in nitrophosphate fertilizer

production," 192th A.C.S. national meeting, september 7-12 (1986) Anaheim, U.S.A.

6. Caro, J.H. and Freeman, H.P., J. Agr. Food Chem., 9, 3 (1961) 182. 7. Caro, J.H. and Hill, W.L., J. Agr. Food Chem., 4 (1956) 684. 8. Chang, W.F., Ed. "Condensed Papers of the Second International Symposium

on Phosphogypsum, December 1986, University of Miami, Florida, USA. 9. Chemische Fabrik Budenheim, DE 3218599A1 (1-2-1983),

DE 3327394A1 (14-2-1985). 10. Chien, S.H., Soil Science 123, 2 (1977) 117. 11. Chien, S.H. and Black, C.A.,

Soil Science Soc. Am. J., 40 1976) 234. 12. Deer, W.A., Howie, R.A. and Zussman, J.,

Rock Forming Minerals, 5 (1962) 324.

21

13. Dufour, P., Predali, J.J. and Ranchin, G., US. patent 4.436.616., mar. 13, 1984.

14. Gremillion, L.R. and McClellan, G.H., Trans. Soc. Mining Engineers of AIME, 270 (1981) 1975.

15. Hill, W.L., Caro, J.H. and Wieczorek, G.A., J. Agr. Food Chem., 2, 25 (1954) 1273.

16. Houot, R., Int. J. of Min. Processing, 9 (1982) 353-17. Hunter, D., "Low-Cadmium H-PO. for fertiliser use,"

Chemical Week, 136, 8 (1985) 25. 18. Jansen, M., Waller, A., Verbiest, J., Van Landschoot, R.C. and Van Rosmalen,

G.M., Industrial Crystallisation v84, The Hague, pg 171. Ed. Jancic and De Jong, Elsevier, Amsterdam (1984).

19. Knudsen, K.C., Proc. of the Fertiliser Society of London* 3 oktober 1985. 20. Kreidler, E.R., "Stoichiometry and Crystal Chemistry of Apatite,"

PhD Thesis, Department of Ceramic Science, Pennsylvania State University, Michigan, USA (1967).

21. Lawver, J.E., McClintock, W.0. and Snow, R.E., Minerals Sci. Engng., 10, 4 (1978) 278.

22. Leder, F., Park, W.C., Chang, P.W., Ellis, J.D., Megy, J.A., Hard, R.A., Kyle, H.E., Mu, J. and Shaw, B.W., Ind. Eng. Chem. Process Des. Dev., 24 (1985) 688.

23. McClellan, G.H. and Lehr, J.R., Am. Mineralogist, 54 (1969) 1374. 24. Montel, G., Bonel, G., Trombe, J.C., Heughebaert, J.C. and Rey, C ,

er Proc. 1 Congr. Int. des Composes Phosphores, Rabat, 17-21 oct., (1977). 25. Nordengren, S., Francia, I. and Nordengren, R.,

Proc. of the Fertiliser Society of London, nr. 33 (1955). 26. Olin Corporation, US patent 4,146,568, Mar. 27, (1979). 27. Philipson, T., Lantbrukshogskolans Annaler, 29 (1963) 267.' 28. Phosphoric Acid Manufacturing using Hydrochloric Acid,

Phosphorus and Potassium, 125 (1983) 29. 29. Purifying wet-process Phosphoric Acid,

Phosphorus and Potassium, 139 (1985) 34. 30. Rubin, A.G., "The BESA-2 Process,"

Phosphorus and Potassium, 137 (1985) 28..

22

31. Scott, W.C., Patterson, G.C. and Hodge, C.A., Fert Solutions, 18 (1974) 62.

32. Slack, A.V., "Phosphoric Acid", Fertiliser Science and Technology Series, Vol. 1, Marcel Dekker, New York, (1968).

33. Societe Uranium Pechiney, NL patent 8103788 (26-06-1985). 34. Stamicarbon B.V., NL patent 8006946 (16-06-1982). 35. Taperova, A.A. and Shulgina, M.N., J. Appl. Chem. USSR, 23 (1950) 27. 36. The nitro-phosphates Alternative, Fertiliser International, 209 (1985) 8. 37. Tjioe, T.T., PhD Thesis, To be published,

Technical University of Delft. 38. UKF/DSM, Private Communication with D.C. Oosterwijk. (1986). 39. Vaman RaO, D., Narayanan, M.K., Nayak, U.B-., Anantnapadmanashan, K. and

Somasundaraan, P., Int. J. of Mineral Processing, 11 (1985) 57. 40. Weterings, K., "The Utilisation of Phosphogypsum,"

Proc. of the Fertiliser Society of London, nr. 208 (1982).

23

2. SCOPE OF THIS INVESTIGATION

Especially in Europe their is a growing concern about the influence of heavy metals on human health. One of the results of extensive discussions between the members of the European Community is to prohibit the use of cadmium containing materials and to forbid the disposal of cadmium in the environment.

Some companies, however, have a license for disposal of their cadmium containing byproducts which expires in the near future. Due to the new environmental restrictions their licences will not be renewed under the same terms. So they have to look for solutions of their problems. One of these companies is DSM which license for disposal of phosphogypsum into the river ends in 1988.

The use of the phosphogypsum is also limited by its impurity content. Moreover, the total quantity of phosphogypsum produced in the Netherlands by far exceeds its potential use in known applications.

Studies initiated and sponsered by the Government, were therefore started to look for new applications of phosphogypsum and for new processes to produce clean phosphogypsum, suitable for application. Additionally, the cadmium content of fertilisers, made from phosphoric acid should be reduced as well to prevent the build up of cadmium in the soil and thus an increase in cadmium in the food chain.

This study is part of the development of a new process route for the production of clean phosphoric acid and clean gypsum. There are three basically different routes.

First of all the amount of cadmium introduced into the process can be reduced by either using phosphate ore with a low cadmium content, of which the availability is limited, however, or by removal of cadmium from the phosphate ore prior to its use. This last possibility is attainable by either calcination of the ore or by selective digestion of the carbonate fraction of carbonate rich phosphate ores with a mineral acid with a relatively high halogenide content. Although the carbonate fraction of these ores is known to contain an appreciable quantity of the total amount of cadmium present in the ore, a reduction with a factor two of the total cadmium content of the ore is the best result obtained so far. The calcination procedure consumes much energy and moreover, the CdO is released into the atmosphere, because it is difficult to recover the cadmium completely from the off gasses.

24

The second method is to prevent the incorporation of cadmium in the phosphogypsum by addition of complexing agents for cadmium and to remove it from the phosphoric acid afterwards, by either solvent extraction, flotation, flocculation or ion exchange. If the cadmium content of the ore is high, the free cadmium level in the solution can only be reduced to an acceptable low level with the addition of relatively large amounts of complexing agents. The addition of complexing agents for cadmium is therefore, a good solution for phosphate ores with a relatively low cadmium content.

The third alternative is the integral method by which the cadmium is removed from the process stream and thus simultaneously reducing its incorporation in the phosphogypsum. This alternative can be obtained in two ways. The first one is the removal of cadmium, ions from the solution before the precipitation of gypsum. Therefore a predigestion step is necessary to obtain a clear calcium-di-hydrogen-phosphate solution, from which the cadmium ions can be removed by either ion exchange, extraction or another technique. The second way is based on the relatively high cadmium uptake of calcium sulphate anhydrite (chapter 5). The cadmium is removed from the solution by its incorporation in a small amount of calcium sulphate anhydrite, precipitated from the recycled phosphoric acid by addition of calcium ions.

To produce concentrated phosphoric acid and phosphogypsum both with a low cadmium content the alternative with the predigestion step has been chosen, because this process route also has the possibility to optimize the digestion step and the crystallisation step independently from each other, and moreover complexing agents can be applied for a further reduction of the cadmium content of the phosphogypsum.

This thesis covers important parts of the development of a "clean technology phosphoric acid" (CTPA) process along a predigestion route. It is composed of a series of self consistent articles, some of which have already been published in the course of the investigation. A small overlap in the contents of the various chapters as well as some inconsistency in the notations were therefore unavoidable. The various articles have been arranged according to the sequence in process steps, when possible.

25

Chapter 3 gives a brief overview of the Clean Technology Phosphoric Acid (CTPA) process, as developed at the Technical University of Delft. Also the separate process steps are indicated and briefly discussed. This chapter can be regarded as an extended summary of the investigation. It links the separate process steps together.

One of the new and essential steps of the CTPA process is treated in chapter 4. Here the digestion of phosphate ore in phosphoric acid is treated. The digestion conditions were systematically varied by changing the phosphoric acid concentration, temperature and particle size of the Zin phosphate ore.

The subject of chapter 5 concerns the second new and important step of the CTPA process, the crystallisation of calcium sulphate hemihydrate from a clear solution. The crystallisation conditions were varied over a wide range by changing the phosphoric acid concentration, the sulphuric acid concentration, the stirrer speed, the solid over liquid weight ratio and the residence time in the crystalliser. The cadmium as well as the phosphate incorporation in the precipitated calcium sulphate hemihydrate (HH) were measured.

Chapter 6 describes the filtration experiments, performed with the phosphoric acid-HH slurries obtained from the experiments described in chapter 5. The permeability of the calcium sulphate hemihydrate (HH) cake obtained by filtration of the phosphoric acid-HH slurry from the crystalliser is determined as a function of process parameters like the sulphate content and the residence time of the solution in the crystalliser.

Chapter 7 deals with the fluoride distribution between the phosphoric acid and the gas phase. Because of environmental restrictions the fluoride removal must take place in a controlled way. The fluoride distribution coefficient is studied as a function of temperature, phosphoric acid concentration, sulphuric acid concentration and fluoride content of the solution. Moreover, an attempt is made to develop an equation, which can be used to predict the fluoride distribution coefficients as a function of the process parameters.

Most of the results obtained in the foregoing chapters are used in chapter 8 to estimate the mass and heat balances of a 1000 tons P„0C per day producing plant, operating according to the CTPA process. Most attention a priori was focussed on energy and phosphate efficiencies. The flowsheeting program TISFLO, developed dy DSM is used for the calculation of the composition of all process streams and of the overall mass and heat balances.

26

To study the influence of impurities, a bench-scale plant of the more relevant parts of the CTPA process was build. The influence of impurities on the phosphate and cadmium incorporation in the HH as well as on the permeability of the HH cake obtained by filtration of the HH slurry was studied.

The results obtained in this study should be combined with the results obtained from a simultaneous study of the cadmium removal techniques [1] and with the results obtained from a reasonably sized pilot plant; operating according to the CTPA process, to give a clear view of the possibility to produce "clean" phosphoric acid and "clean phosphogypsum" with the CTPA process in an economically feasible way.

[1] Tjioe, T.T., PhD Thesis, To be published, Technical University of Delft

27

3. OVERVIEW OF THE CTPA PROCESS

3.1. Summary

Phosphoric acid for use in fertiliser applications is mainly produced by a 'wet process', i.e. by digestion of phosphate ore with sulphuric acid. In such wet processes, however, impurities like cadmium and radium, originating from the phosphate ore are distributed between the phosphoric acid and the byproduct, a calcium sulphate modification. The use or disposal of the byproduct and probably the use of the phosphoric acid in future as well is limited by its impurity content.

The aim of the currently developed process is the direct production of concentrated phosphoric acid with a low cadmium content as well as the production of a major amount of calcium sulphate as hemihydrate (HH), with a low phosphate, cadmium and radium content, in a commercially feasible way.

For this purpose the phosphate ore is first completely digested in recycled phosphoric acid, containing about 40 w? P?0,- a n d 1«8 w? H?S0j,. After separation of the insoluble ore residue together with a minor amount of calcium sulphate hemihydrate, precipitated in the digester, a clear calcium-di-hydrogen-phosphate

2+ (CDHP) solution is obtained. From this solution the Cd -ions can be removed by e.g. ion exchange. Thereafter the calcium ions are removed by adding concentrated sulphuric acid to the CDHP solution in the crystalliser at 90 °C. In this way a clean calcium sulphate hemihydrate can be obtained. In order to optimize the individual process steps, each step had to be investigated separately.

By performing HH crystallisation experiments, a linear relationship was found between the molar phosphate over sulphate ratio in the crystals and in the solution. The phosphate content of the crystals decreases with increasing sulphate concentration in the crystalliser. It was found that above 3 w$ H SO. the phosphate content of the HH was lower than 0.1 vi% P2°c:- T n e cadmium incorporation was also measured as function of the operating conditions and appears to increase significantly with raising sulphate concentrations above 2 w* H 2 S O r

Filtration studies showed, that the HH crystals obtained during digestion of the ore, are difficult to filter, while the HH crystals developed in the crystalliser, filter quite well. A maximum filtration rate was reached, if a sulphate content of about 1.8 wï H SO. was maintained in the crystalliser.

28

A study of the fluoride emissions in the process led to the development of an expression for the fluoride distribution coefficient as a function of the operating conditions.

Preliminary results from a continuously operated bench-scale plant, in which part of the process is studied, were in agreement with the results obtained from the separate studies.

A. preliminary process flowsheet and mass balance, in combination with our experimental results, show that concentrated phosphoric acid (40 wJ'P-Cv) with less than 5 ppm Cd as well as clean calcium sulphate hemihydrate with less than 1 ppm Cd and less than 0.3 w% P?0_ can be produced.

3.2. Introduction

Phosphoric acid is mainly used for the manufacturing of fertilisers. The phosphoric acid production is almost directly linked to the world phosphate fertiliser consumption, which is expected to rise from 30,7 million tons P-CL in 1982 to about 11,7 million tons P„0,. in 1990. This represents an increase of the wet phosphoric acid production from 19,7 million tons P-CL in 1982 to 27,8 million tons P.0,. in 1990 [19].

Phosphoric acid can be released from phosphate ore by the action of strong mineral acids such as nitric acid, sulphuric acid and hydrochloric acid. Sulphuric acid is the only acid which forms an insoluble precipitate with the calcium of the phosphate ore, thus allowing the phosphoric acid to be separated directly by filtration. The chlorides and nitrates of calcium are both soluble, so special techniques, like solvent extraction [18] ion exchange [14] or cooling crystallisation [2], [31] are required to recover the phosphoric acid. In spite of this, phosphoric acid is being commercially produced using nitric acid as well as hydrochloric acid. The economical feasibility of these processes is mainly determined by the availability and price of the mineral acid and the local production facilities.

Although these acidulation methods are all referred to as "wet processes", the term "wet process" is mainly reserved for processes in which sulphuric acid is applied.

During the production of phosphoric acid from phosphate ore by wet processes, e.g. by applying sulphuric acid, huge amounts of calcium sulphate are precipitated as a byproduct. Depending on the temperature as well as on the phosphate and sulphate content of the solution, either calcium sulphate

29

dihydrate (DH), hemihydrate (HH) or anhydrite (AH) will be formed, as is shown in figure 1 [1, 29].

100-

temp°C 80-

60-

40'

20 20 30 40 50

w % P 2 0 5 ^

Figure 1: Phase diagram of calcium sulphate in phosphoric acid.

The solid lines in figure 1 represent quasi-equilibrium curves, which indicate which phase will initially precipitate under the given conditions. The broken line in figure 1 bounds the regions, where either DH or AH is stable. The HH phase only exists as a metastable phase. The influence of the sulphuric acid content can be taken into account by assuming that 1 mole of sulphuric acid is equivalent to about 1.5 mole of phosphoric acid.

Commercial wet processes are developed producing either DH, HH or AH or even combinations as a byproduct. Becker [1] summarizes the advantages and disadvantages of such processes.

The main disadvantage of the DH, the single filter HH/DH and the DH/HH processes is the production of a relatively diluted (about 30 w$ P„0_)

2 D phosphoric acid, which has to be concentrated for use in fertiliser applications. Due to the expensive concentration step, the HH and the two filter HH/DH processes, allowing direct production of 40 - 50 w$ P 0 from the filter, are rapidly gaining field. AH processes are not operational at this moment [1], due to serious problems observed in small scale commercial batch plants [17], such as the severe corrosion at the required high temperature and the low growth rate of AH.

AH precipitation

HH precipitation

DH precipitation

- - C A H stable

, / D H stable""*N

30

In all wet processes, impurities, such as radium ions and heavy metal ions, like cadmium, originally present in the phosphate ore, are distributed between the phosphoric acid (up to 80$ of the Cd) and the byproduct, 'phosphogypsum'.

Cadmium is responsible for a painful bone disease known as Itai-Itai. Concentrations as low as 1 mg/1 in water have led to that illness [1], Cadmium is one of the most worrisome environmental problems, and in the last years there is growing concern, particularly in Europe, that this toxic metal, which is present in e.g. the wet phosphoric acid, could enter into the food chain in increasing amounts. Some countries intend to adopt a cadmium limitation of no more than 90 mg per kg P20_ in the fertiliser product as a preventive measure [12,33]. The Netherlands plans legislation restricting the levels of heavy metals in phosphate fertilisers as well • as in other products, used in agriculture and also the levels in phosphogypsum, which is disposed into the water. These restrictions may well be taken up by the European Economic Community in the near future.

The byproduct, calcium sulphate, also tends to incorporate phosphate ions, which lowers the overall phosphate efficiency of the process. The use of this byproduct as a building material is hindered by its radium and phosphate content, while the development of new applications for phosphogypsum is limited by the uptake of heavy metal ions, since leaching of heavy metal ions can not be totally prevented [13]. The direct production of concentrated and relatively clean phosphoric acid is so far only possible by applying the furnace process, an alternative to acidulation processes.

In the furnace process the phosphate ore is reduced, by addition of silica and carbon, to phosphorus and slag at about 1500 °C:

2 Ca1Q(P01))6F2 + 20 Si02 + 28 C -—> 20 CaSiO + 28 CO + 3 P^ + 2 F2 (1)

AH = 27.9 MJ/kg of P^ produced [16].

The heat needed for this reaction is substantial, and could, until a few

years ago only be achieved by internal resistance heating of the molten charge in an electric furnace. In 1985, however, a direct fired rotary kiln was used to supply the heat [16]. After combustion of the phosphorus, the P 0, is

2 5 absorbed into water or moderately concentrated phosphoric acid, to obtain concentrated phosphoric acid (>60 w$ P„0,-). The direct production of concentrated and clean phosphoric acid is possible with the furnace process, because most impurities from the ore remain in the slag [19]. This process,

31

however, is too expensive for use in the fertiliser industry, due to the high energy cost. Moreover, there are other possibilities to remove Cd from wet phosphoric acid, like for instance flotation [28], solvent extraction [21] or the Budenheim process [3]. In all these processes, however, the disposal of the byproduct remains a problem, mainly due to its cadmium content, while the use of the byproduct as e.g. building material, is hampered by its radium content.

In order to produce not only relatively clean and concentrated phosphoric acid for fertiliser applications, at a feasible price, but also pure HH, the wet Clean Technology Phosphoric Acid (CTPA) process is currently being developed.

3.3. Objective of the study

The aim of the CTPA process is the production of concentrated phosphoric acid with a low cadmium and fluoride content as well as the production of calcium sulphate hemihydrate with a low cadmium, radium and phosphate content in a commercially feasible way.

3.4. Description of the CTPA process

3.4.1. Introduction to the process

2+ To meet the foregoing requirements, the Cd -ions have to be removed from 2+ a large process stream. An alternative would be to prevent incorporation of Cd

2+ 2+ ' -ions in the HH or to leach the Cd -ions from the HH, in which cases the Cd -• ions only have to be removed from the much smaller product acid stream. Removal of Cd from the phosphate ore by e.g. calcination, is far too expensive to be a reasonable alternative. Selective leaching of cadmium from ground phosphate ore by a diluted mineral acid with a high chloride concentration is another possibility [4]. Due to the high chloride content, however, severe corrosion will probably occur. So far no methods are known to prevent the Cd uptake during crystallisation sufficiently, to produce HH with less than 0.5 ppm Cd, nor methods to leach the Cd from the byproduct until an acceptable low level is reached. Therefore, removal of Cd from the large process stream, before the major amount of the calcium sulphate is crystallised, is a necessity.

For this purpose the digestion of phosphate ore and the crystallisation of HH, which normally occur more or less simultaneously, have to be divided into two separate stages (see figure 2).

32

This also makes it possible to control each stage apart and to optimize each stage on its own, instead of the combination of the two. The first step of the process, the digestion of phosphate ore, occurs in a recycle stream of phosphoric acid. Since the recycled acid stream, however, unavoidably contains sulphate ions in an amount dictated by the operating conditions in the crystalliser, some of the HH, up to a maximum of 25 w$, will precipitate during digestion of the ore.

After filtration of the insoluble ore residue, together with the'relatively small amount of HH precipitated, a clear calcium-di-hydrogen-phosphate (CDHP)

2+ solution is obtained, from which the Cd -ions can be removed by ion exchange or extraction. Thereafter the HH is crystallised. The HH precipitated during digestion of the phosphate ore has to be recrystallised into DH in order to raise the overall phosphate efficiency of the process to an acceptable level (> 98 w$ P„0C recovery). Although recrystallisation of this part of the HH is necessary, concentrated phosphoric acid (> 40 w$ P„0_ ) is still produced,

c. D

because the recrystallisation step is fully integrated in the process. The scheme of the process is given in figure 2.

H2SOfc

\ digestior Cd removal

recrystallisation H H — — OH

t DH

recycle acid

HH crystallisation

1

HH

product

Figure 2: Process scheme.

The most important steps of the CTPA process will be discussed in the next sections, except for the Cd removal step, which will be presented elsewhere [32]. The principle of the Cd removal by ion exchange is the addition of a cadmium specific inorganic complexing agent, like bromide or iodide ions, followed by the removal of the Cd-halogenide complex by anion exchange with commercially available basic ion exchange resins.

3.4.2. The digestion stage

The main phosphate ores used today are sedimentary phosphates or franeolltes. These ores mainly consist of fluoroapatite with part of the phosphate ions replaced by fluoride and carbonate ions. By beneficiation of the

33

ore, large amounts of waste material are removed. The treated ore, however, still contains residual calcite and dolomite, which cause severe foaming during the digestion [1].

Fluoroapatite is dissolved in pure phosphoric acid according to the following reaction:

Ca1Q(P04)6F2 + 114 HgPO,, — > 10 Ca(H2P04)2 + 2 HF (2) fluoroapatite CDHP

after which the calcium-di-hydrogen-phosphate, further referred to as CDHP, dissolves in the phosphoric acid stream.

6

5

4

3

2

1

0 30 35 40 45 ^ w%P 20 5

Figure 3: Solubility of CaO in phosphoric acid.

The amount of CDHP which can be dissolved in phosphoric acid strongly depends on the temperature and the phosphate content of the solution, as shown in figure 3 [7].

Phosphoric acid to be used for the production of mono- and di-ammonium-phosphate must at least contain 40 w? P„0C [11]. The temperature in the

2 D

digestion stage is res t r ic ted by the HH-AH phase boundary, because the precipitated calcium sulphate has to be HH, which can be easily recrystal l ised into DH. In order to minimize the recycle acid stream, necessary for complete digestion of the phosphate ore, conditions with the highest CaO solubili ty have

3t

w%CaO

100 °C

^ \ _

to be chosen. The optimum conditions for digestion of the phosphate ore are therefore about 40 w? P-O- and approximately 90 °C.

When a mixture of phosphoric acid and sulphuric acid is used for the digestion of the ore, two processes will occur more or less simultaneously: digestion of the phosphate ore and precipitation of calcium sulphate (HH) [1]. At sulphate concentrations in the digestion stage above 2.5wJ H SO., the HH tends to precipitate not only upon the present HH crystals, but also upon the ore particles. Such coating of the ore reduces the digestion rate and is called blinding [5].

Since the highest possible digestion rate should be obtained, the sulphate concentration in the digestion stage must be carefully selected in order to avoid blinding. In a continuous process, blinding can only be avoided by keeping a low sulphate concentration in the digester, which implies maintaining a high calcium concentration. Under such conditions, the continuously added sulphate ions precipitate upon the hemihydrate crystals and the digestion can be treated as if proceeding in pure phosphoric acid. The digestion experiments were thus performed in pure phosphoric acid [25]. At 90 °C, however, the digestion proceeds very fast and is accompanied by severe foaming. Most experiments were thus performed at lower temperatures, with pretreated ore.

The phosphate ore used was a mixture of Khouribga/Zin in a 40/60 \>% ratio. In order to reduce the foaming, the ore was pretreated to remove calcite and dolomite with a 0.5 N tri-ammonium-citrate solution at pH = 8.1 for a few days [22]. The remaining fluoroapatite was washed with water and dried at 60 °C for a few days. Thereafter, the ore was divided in twelve fractions by sieving through "Twente" sieves. The ore particle sizes range from 150 to 2000 ym. The 1.5 liter reaction vessel was thermostated. About 50 grams of ore were suspended in about 1 kg of acid. To monitor the digestion process with time, periodically samples were taken from the reactor by vacuum withdrawal, filtered over a G^ glass filter, coated with perlite, and analysed for the CaO content.

From the results obtained, it can be deduced that in a first approach the digestion of phosphate ore in phosphoric acid depends on the concentration of the phosphoric acid as well as on the temperature and the particle size of the phosphate ore. In figures 4 and 5 it is shown that, as expected, the smaller the particle size of the ore and the higher the temperature, the higher the digestion rate of the calcium present in the ore.

If the rate determining step in the digestion process were a chemical reaction, an increase of the digestion rate with increasing acid concentration is expected, due to the increase of the hydrogen ion concentration. If, however,

35

diffusion is the rate determining step, a decrease of the digestion rate is expected, due to the increase of the viscosity with increasing acid concentration.

conversion

0.00129-0,00151 0,000655-0.000780 0.000462 0,000181-0,000256

time (mini

Figure 4: Influence of the particle size on the digestion rate of the phosphate ore at 50 °C in 40 u% Po0r.

2 D

conversion

emp.PC) particle size (ml

50 0,00185 70 0.00185 50 0,00078 60 0,00078 70 0,00078

10 time (min) -

Figure 5: Influence of the temperature at different particle sizes on the digestion rate of the ore in 40 w% P„0,_. 2 5

In figure 6 the influence of the concentration of phosphoric acid on the digestion rate shows that from 35 up to 45 w? P_0 the digestion rate decreases, while for 30 up to 35 w$ PT) the digestion rate seems to increase.

2 O

This can be explained by assuming that at 50 °C the rate determining step changes from chemical reaction at low phosphoric acid concentrations to diffusion for higher phosphoric acid concentrations. At 90 °C diffusion will be the rate determing step also for 30-35 wj P-O... This study has been continued in

2 5 order to determine the rate controlling mechanism at different operating conditions [23, chapter 4]. Preliminary results from our bench-scale plant showed that a residence time of about 1 hour is enough, for ore particles up to 2000 vim, to reach more than 99 % conversion.

36

t conversion

1-

50°C 0.000655m

f I

i u 5 10 time (min) • -

Figure 6: Influence of the phosphoric acid concentration on the digestion rate of phosphate ore at 50 °C.

3.4.3. The crystallisation stages

In the CTPA process several crystallisation stages exist. The first occurs in the digestion stage. The recycled acid, used for digestion of the phosphate ore, contains approximately 1.8 w$ H_SOü, which almost completely precipitates as HH. From previous batch experiments [13] the phosphate incorporation in this HH is expected to be high (above 2 w$ P„0_). In our continuously operated bench-scale plant also a phosphate incorporation of approximately 3 w? P_0C was found in this HH. Because about 25 wj of the total calcium present in the phosphate ore is precipitated during digestion, the overall phosphate efficiency of the process would be unacceptably low. Therefore, this HH must be recrystallised to reduce the phosphate incorporated in the calcium sulphate crystal lattice. It is known [10] that due to recrystallisation of HH into DH, a reduction in the phosphate incorporation of about 90 % can be achieved.

Batch recrystallisation experiments of HH into DH showed that a phosphate incorporation of 0.3 w$ P2°5 in D H '-35 i 3 a reasonable estimate of the value to be obtained in a production plant. A residence time of approximately 5 hrs [10], however, is needed for complete conversion of HH into DH. In order to reduce

temperature = particle size =

w% P ;

30 35 40

37

this residence time, a more detailed investigation of the recrystallisation process is required and will be reported in the near future [35].

The most important crystallisation step in the CTPA process is the direct production of clean HH from a clean and clear CDHP solution and concentrated sulphuric acid. This crystallisation was extensively investigated in order to determine the optimal operating conditions. The influence of the w$ P

20 R ' the

w$ H9SOj. and the residence time in the crystalliser on the phosphate and cadmium incorporation in the HH as well as on the size and shape of the HH crystals was studied. The size and shape of the HH crystals have been correlated with their filterability.

The crystallisation experiments were performed with chemically pure CDHP and sulphuric acid solutions of varying concentrations in a separate, 1 liter, continuous crystalliser at 90 °C. A more detailed description of this study is given elsewhere [27]. Experiments were performed with residence times of 20 up to 80 minutes. Each experiment lasted for 8 to 10 residence times. After each residence time, samples were taken and analysed for the cadmium as well as for the phosphate content of the HH crystals. The cadmium and phosphate content of the crystals remained constant after the 6 residence time. The samples taken after the 8 residence time were considered to be representative for the "steady state" of this system.

In phosphoric acid containing 40 u% P_Oc and a low sulphate concentration (less than 0.7 w% H?S0n) aggregates of HH crystals were obtained. At raising sulphate concentrations single needles were formed in addition to aggregates. At still higher sulphate concentrations (above 2 w$ H-SOj.) mostly single needles were grown which were longer (up to 2 mm) and relatively thinner. Also many tiny needles were observed. The influence of the HH crystals on their filterability is quantified in section 3.1.3 (see also figure 10, 11 and 12). By performing experiments in more concentrated phosphoric acid (47 w$ PpOc) mainly aggregates were obtained, possibly due to the higher supersaturation caused by the lower solubility of HH, which stimulates primary nucleation.

The phosphate incorporation in the HH was found to be a competition between the phosphate and the sulphate ions in the solution. In figure 7 the molar phosphate over sulphate ratio in the HH crystals is given versus the same ratio in the solution.

No significant influence of stirring rate and residence time on the phosphate uptake was found. Small variations in supersaturation therefore do not appear to have any influence on the uptake. In batch experiments, however, where

38

larger variations in the supersaturation were used the phosphate incorporation appeared to increase with enhanced supersaturation [13].

The relationship between the molar ratios presented in figure 7 for sulphate concentrations between 0.1 and 2.5 w$ is given by:

[ HPO^2 "] [ H3P04 ] — — ( crystals ) = 1.4 * 10~4 * ( solution) (3) [ SO,,2" ] [ H2S04 ]

From equation (3) it can be calculated that at sulphate concentrations above 1.8 w% in phosphoric acid containing HO w$ P-0_ a phosphate incorporation in the HH crystals of less than 0.3 w$ P-0C is obtained. This incorporation is acceptable for our process and is in agreement with the preliminary results obtained from our bench-scale plant.

0.04

0.03—

0.02-

0.01-

0 100 200 300

[H3P0 J -(solution)

(H2SO4] ^

Figure 7: Phosphate incorporation expressed as the molar phosphate over sulphate ratio in the crystals versus the same ratio in the solution.

2+ The incorporation of cadmium in the HH, expressed as the molar Cd over 2+ Ca ratio in the HH crystals is plotted in figure 8 as a function of both the

2+ 2+ sulphate content of the solution and the molar Cd over Ca ratio in the

i

(HPO^]

| s o 2 - | install 4 '

250 rpm 7°

fo ° / ■ ^ 1570 rpm

•

residence time 20 min. 40 min. 60 min. 80 min.

w%P205

40 47 • ■ O D

0

39

solution. In this figure, only results are given for experiments in phosphoric acid containing 40 w$ P„Cv.

At higher phosphate concentrations the cadmium uptake was found to be higher, possibly due to the reduced solubility of HH. The curve presented in figure 8 shows that, up to about 2.5 w$ H^O^, the cadmium incorporation remains almost constant for experiments with a residence time of 10 minutes. An increase in uptake occurs with increasing sulphate concentrations for experiments with a residence time of 20 minutes. No explanation for this difference can be given.

3

2

1

500 1000 1500

5 («I |Ca|

2+ 2+ Figure 8: Cadmium incorporation expressed as the molar Cd over Ca ratio in the HH crystals versus the same ratio and the sulphate concentration in the solution.

During continuous operation of the bench-scale plant, however, it was observed that in the digester, at very high calcium concentrations (see table 1) and thus at low sulphate concentrations (below 0.05 w$ H,SCv) the precipitated

2+ 2+ HH crystals contained a very small amount of Cd. The molar Cd over Ca ratio in the crystals was less than 4 * 10~ . Therefore, it can not be excluded that below 0.1 wj H2S0. the actual Cd uptake will decrease. Variations of the residence time and the stirring rate had no significant influence on the Cd uptake either. From the results mentioned above and from the bench-scale plant, it can be deduced that the cadmium incorporation in the HH obtained in the HH crystalliser, can only be reduced to less than 1 ppm by preventing the cadmium

I ,

1 0 5 x l " l I Ca J

0

1 i

hal)

Q ^

r 'sidence fime • 20 min. o 40 min.

» w% H SO 2 3 1 1

HO

incorporation through the use of complexing agents or by the previous removal of the cadmium from the process stream.

An additional result from the bench-scale plant experiments was that a large quantity of the radium present in the ore was precipitated with the HH during digestion. The majority of the HH, obtained in the HH crystalliser, thus contained less Ra than HH obtained from commercial processes. This result is in good agreement with those described in a DSM patent [6].

3.1.1. The filtration stages

Three filtration stages are encountered in our CTPA process as is shown in figure 13. The first, located after the digestion of phosphate ore, separates the ore residue particles together with a minor amount of HH precipitated in the digestion stage from the CDHP solution. The second filter is placed after recrystallisation of the HH into DH. The third filter is installed after crystallisation of the major part of the HH, which should be clean enough to be used as e.g. a building material.

In order to calculate the filter area needed for each stage, the calcium sulphate cakes obtained in all three stages, should be characterized. This can be achieved by measuring the permeability of the formed cake, which can be used to calculate the linear filtration velocity, U, from the equation:

1 dV K*AP U = — * (4)

A dt n*h

in which A = filter area V = volume of filtrate t = filtration time K = cake permeability AP = pressure drop over the cake n = liquid viscosity h = cake height

The permeability, K, can be obtained by substituting the specific cake resistance, o, the solid density, p and the cake porosity, e into the equation:

[m ] [m3] [s] [m2] [N*m~2] [N*s*m~2] [m]

11

where a = specific cake resistance [m*kg ] e = cake porosity

-•3 p = solids density [kg*m ]

The cake porosity follows from a mass balance over the filter cake:

G * s = (1-e) * a * h » p (6) s

in which G «■ mass of slurry f i l tered [kg] s = weight fraction solids in the slurry

The specific cake resistance can be calculated from the slope, B, of the At/AV versus V curve, obtained from f i l t ra t ion experiments at constant pressure by the equation [34]:

B * 2 * A * AP

a (7) c * n

in which c = (s * p . ) / (A * (1 - m * s ) [kg»nf5] e * Px

m = 1 +

( 1 - e ) * p -3 p. = liquid density [kg*m ]

To obtain uniform crystals and reproducable results, all crystallisation experiments have been performed in a continous crystalliser, with an unclassifying slurry withdrawal.

First the filtration of the small HH particles (mean diameter = 100 ym), precipitated during digestion, mixed with the very tiny ore residue particles (mean diameter = 5 pm) was studied. This filtration is obviously difficult, because of the small particles involved. Only by using a pressure filter we were

-1 3 able to calculate with equation (H) an approximate permeability, K, of 1*10 -ill p to 1*10 m for the HH/ore residue cake. This is very low, especially since a

large throughput of the huge process stream is necessary. In practice, a

42

sedimentary system is therefore recommended to separate the solids from the liquid. Washing of the cake is not needed, because the crystals have to be recrystallized.

The CDHP solution must be almost free of solids to avoid blockage of the packed columns, which might be used for removal of cadmium by anion exchange. The filtration of the DH obtained after recrystallisation of the HH precipitated during digestion, is assumed to be equal to the filtration of DH obtained after recrystallisation of HH produced in common HH/DH processes. The permeability of such a DH cake was calculated with equation (4) and with data given by Becker

-12 2 [1], to be about 1*10 m . This permeability was also found by filtration of a DH slurry from a commercial HH/DH process. Because in practice the quantity of acid to be filtered here is relatively small, the industrial filter area needed (about - 20 operational l imitat ion.

30 m for a 1000 tons P_0_ per day plant) will never be an

70 -

50-

3 0 -

10 -

1012xKlm2) residence time —•- 20 min --o-- 40min

60 min

- * • w%H2S0<,

Figure 9: Influence of the residence time and the sulphate concentration on the permeability of the HH cake obtained in the crysta l l iser at 40 w% Po0lr.

The third f i l t r a t i on stage, of the majority of the HH, was studied more intensively [21], part icularly because the capacity of this f i l t e r was expected to be an important item in the capital cost of the plant.

43

In figure 9 the influence of the sulphuric acid content in the crystalliser on the permeability of the HH cake is shown. A maximum permeability is observed between 1.5 and 2 u% H SO. as was also found by Dahlgren [5] for a single stage HH process. This result can be explained by the shape and size of the crystals. At low sulphate concentrations very small crystals are formed, while at high sulphate concentrations both large (up to 2 mm) and small crystals were observed. These different crystals are shown in figure 10, 11 and 12.

Figure 10: HH crystals obtained at 90 °C in 40 u% ?20 and 0.5 wj H2S0

The influence of the residence time on the permeability of the HH cake is also demonstrated in figure 9. Above about 1.5 u% H_S0j. an increase in residence time seems to lead to a reduction in the permeability of the HH cake. This can be caused by the formation of small crystals due to the breakage of the longer HH needles obtained at longer residence times. By maintaining about 1.8 w? H SO. in the crystalliser and a residence time of 20 minutes, the permeability

-12 2 of the HH cake is about 65*10 m . From preliminary results, obtained in our bench-scale plant, this permeability is expected to be in practice one fifth of the value obtained, due to the presence of impurities. The filter area needed

2 can then be calculated to be approximately 75 m for a 1000 tons P 0~ per day producing plant. This value is satisfactory small. More information about this subject will be gathered in the near future [23].

44

Figure 11 : HH crystals obtained at 90 °C in 40 w% P_0,- and 2.4 w$ H„S0„. do i n

Figure 12: HH crystals obtained at 90 °C in 40 w$ P.O.. and 3.2 wj H„S0, . 2 5 2 4

45

3.4.5. The fluoride removal

Fluoroapatite contains approximately 1 w$ F [1]. During its digestion by phosphoric acid, the fluoride is released as hydrogen fluoride. The hydrogen fluoride reacts with the silica present in the ore or added as a clay, to form a fluorosilicic acid solution in phosphoric acid. Some of the fluorosilicic acid precipitates with sodium or potassium ions as Na SiF,, K SiF,, or NaKSiF,. The remaining fluorosilicic acid partly decomposes in SiFü and HF, which on their turn partly evade in an unknown ratio. The residual fluoride in solution is distributed between the phosphoric acid and the byproduct during its

2-precipitation. In the byproduct, fluoride is assumed to be present as A1F_ , 2- 5

which can substitute for SOj. [26]. The presence of fluoride in phosphoric acid not only makes this acid more corrosive, but also makes it unsuitable for fertiliser applications, if the concentration is too high. It also causes post precipitation of compounds like chuckrovite. The fluoride content of the byproduct may also hamper the use of the "phosphogypsum". Discharge of fluoride in the air is limited in the Netherlands to about 2 ppm for environmental reasons [1, 33].

So the fluoride discharge has to be carefully controlled. In commercial phosphoric acid plants, the fluoride mainly evades during digestion and crystallisation, while the residual fluoride is stripped from the product acid stream. Half of the fluoride present in phosphate ore, however, is removed as solid material, either incorporated in the byproduct or as a sodium or potassium fluorosilicate precipitate. To calculate the amount of fluoride released during digestion and crystallisation and to design the fluoride stripping unit, fluoride distribution coefficients between the phosphoric acid and the ambient air are needed. Due to lack of available data for our system, fluoride distribution coefficients were measured between the acid and the air as a function of the operating conditions like: w? P„0r, w$ H„SO„ and w% H„SiF, in the acid and the temperature. Also the influence of impurities was taken into account [26].

Experiments were performed in which a N ? gas stream was first led through three bottles containing phosphoric acid of the desired concentration and temperature, without fluoride, to saturate the N~ with water. Thereafter, the saturated N. gas stream was led through four bottles containing the same phosphoric acid, but this time with fluoride added as H„SiF,, in order to saturate the N„ gas with fluoride. To be able to measure the fluoride concentration in the gas, it was led through two absorption bottles filled with

16

water for a given time. The absorbed fluoride was determined with a fluoride specific electrode, after addition of a buffer solution to keep the ionic strength constant.

A theoretical model used to correlate the measured distribution coefficients is based on a simplified model of the H..PO.-H SiFg-H 0 system [26]. In this model the following reactions are assumed to take place:

SiFg2" + 2 H+ < — > SiF4(l) + 2 HF(1) (8)

SiF^l) <---> SiF^Cg) (9)

2 HF(1) < — > 2 HF(g) (10)

The fluoride distribution coefficient H, can now be defined as:

total F concentration in the gas phase H (11)

total F concentration in the liquid phase

If it is further assumed that the concentration of SiF^ and HF in the liquid are negligible compared to the concentration of SiF, , equation (11) can be

o rewritten as:

( 4 * [SiF^] + [HF] ) gas phase H — (12)

6 * [SiF62_]

Because the silica concentration in the gas phase at 90 °C is too small to be measured with any accuracy, the separate SiFü and HF concentrations in the gas phase are unknown. These concentrations are, however, related to the SiFü and HF activities in the solution by Henry's law. If the gases are assumed to behave ideal, these relationships are given by:

[HF] (gas phase) h(HF) - (13)

a(HF)

and:

47

[SiFjj] (gas phase) hCSiF^) = - - - - - (14)

a (S iF 4 )

in which a(HF) = activity of HF a(SiF.) = activity of SiFy h(HF) = Henry coefficient of HF h(SiF.) = Henry coefficient of SiFn

If furthermore, a constant ratio q = a(HF)/a(SiF^) (15) is assumed and the equilibrium constant, K, of equation (8), is defined as:

a(HF)2 * atSiF^) K = (16)

a(SiF62") * a(H+)2

where a(H ) = activity of H 2- 2-

a(SiFg ) = activity of SiFg

equation (12) can be rewritten as:

2 h(SiF„) a(SiF,2") * K * a(H+)2 6 1 V?

3 [SiFg2"] q2

1 h(HF) + - * — * (K * q * a(SiF 2") * a(H+)2)1"33 (17)

2- 6

6 [SiF^ ]

Since K, q, h(HF) and h(SiF.) are constants, only the activity coefficients and the hydrogen ion molality are needed to calculate the fluoride distribution

2-coefficient, because the SiF, concentration follows from the total amount of fluoride added as H.SiF,. The ion specific interaction model of Pitzer [20] can be used to calculate the activity coefficients, while the hydrogen ion molality can be calculated from Elmore's model of phosphoric acid [8], In Pitzer*s as well as in Elmore's model molalities instead of concentrations are used.

18

Therefore, the fluoride distribution coefficient, H is redefined as:

total F concentration in the gas phase H = - — (18)

H SiF, molality in the liquid phase

Equation (17) now becomes:

K * a(H+)2

H = 1 * h(SiF„) * ( —)°-33 * Y(siF,2") + 2 2-2 6

q^ * atSiF^ ) d

K * q * a(H+)2

+ h(HF) * (— ) 0 , 3 3 * Y(SiF 2~) (19) 2-2 6

in which Y(SiFg ") = the activity coefficient of SiFg

Taking the logaritm of equation (19) and rearranging the terms leads to:

in H = C. + |* ln(m(H+)) + §* ln(Y(H+)) - ;* m(SiF,2") + \* ln(Y(SiF,2")) (20) 1 3 3 3 . o 3 o

where C. = a constant Y(H+) = activity coefficient of H+

m = molality

Also the temperature influence on the fluoride distribution coefficient can be taken into account by applying a Clausius Clapeyron equation for the relationship between the temperature and the gas pressure. Moreover, if the contribution of the sulphuric acid to the hydrogen ion molality in the solution is calculated with Pitzer's model, and assuming the sulphuric acid to be completely dissociated, equation (20) becomes:

In H = C, + | * ln(m(H+)) - | » ln(m(SiF62")) + C2 * ind^PO^) + C3 * m(H2SiF6) +

+ C4 * m(H2S0^) + C5/T (21)

49

The values of the constants, derived from our experimental data, were:

C - 32,n C„ = 0.156 C. = 1.0 C„ = 0.157 C = -9550 \ d i n D

The above equation was found to describe the experimental results in the following area:

30 < w? P205 < 50 70 < temperature °C < 95 0 < wj H SO,, < 6 w$ H2SiF6 < 4

and a molar F/Si ratio less than 10.

The influence of impurities on the distribution coefficient has hardly been studied. Acid from a commercial wet process, however, was used to check equation (21) for an impure system. The measured fluoride distribution coefficients were in fairly good agreement with the values, calculated with equation (21), except for "merchant grade" acid, where the molar F over Si ratio was much higher than 10. In this case, equation (21) cannot be used, because the fluoride is not only present as fluorosilicic acid. It is present also as hydrogen fluoride and mono-fluoro-phosphoric acid. The existance of mono-fluoro-phosphoric acid was already suggested by Kopylov e.a. [15].

With equation (21) the amount of fluoride which will release in each process stage can be estimated. The fluoride recuperated as fluorosilicic acid can be used to produce hydrogen fluoride as well as pyrogenic SiO_ [93. The hydrogen fluoride can be used in the production of engineering plastics like PTFE and PVDF, while the Si0? can be used for the production of ceramic materials [30].

3.5. Simplified process flowsheet and mass balance

Most of the foregoing experiments were performed with chemically pure reagents. In order to check the influence of impurities on these results, a bench-scale plant has been built, covering the essential parts of the process. This is necessary, because only by continuously operating under steady state conditions, the impurity levels are completely built up. The digestion of the phosphate ore and the crystallisation of HH as well as their respective filtrations, are investigated by performing continuous runs in the bench-scale

50

plant. Not all steps are investigated, owing to the complexity of the necessary equipment and the difficulty in controlling the process on a bench-scale.

Preliminary results obtained with this equipment, showed no essential differences with those achieved with chemically pure reagents, only the filtration of HH was five times worse. Because all process steps are still under investigation only a simplified process flowsheet and mass balance can be presented. A heat balance is being studied. Basic information about the process conditions is given in table 1, while the composition of the phosphate ore used, is given in table 2.

digestion | recrystallisation | crystallisation |

w% P205

wj HgSO,, w$ CaO wj s o l i d s temperature

40.6 -5 4

90

28 5.0 -

23 60

40 |

1.7 | |

10 | 90 |

Table 1: Process conditions.

w$ P205

wj CaO

w% SOjj

w$ F

w$ H20

31.55

50.80

2.40

4.24

1.0

Table 2: Composition of the phosphate ore.

The simplified process flowsheet and mass balance for a 1000 tons P-0,. per day plant operating 24 hours per day, is given in figure 13, where the mass streams are given in tons per hour. The overall phosphate efficiency of the process can be estimated to be about 99 %, if the filtration losses are neglected.

Accurate residence times can not yet be given. For the digestion the residence time, however, will be about 1 hour, for the crystallisation about 30

51

minutes, and for the recrystallisation approximately 5 hours. The number of tanks needed for digestion of phosphate ore will be two or three, just as the number of tanks required for the recrystallisation of HH into DH. For the crystallisation of HH one tank will be sufficient.

The HH as well as the DH filter can be vacuum filters. After the digestion, the solids and the CDHP solution may be separated by sedimentation. Hl°JW , , , ^

113

H2S04 , ) 9 108

36

ore 132 vapour Yi "11

digestion * - '■

1052

• * 60

Cd removal 1113

vapour 5__

recrystallisation

. i , ,

203

vapour 4

<— HH crystallisation

137

33 " " 46 DH 23 water

104

recycle acid 1052

Figure 13: Simplified process flowsheet and mass balance.

3 . 6 . Conc lus ions

1354

product acid

138 HH 69 water

1156

104

A new process for the production of clean phosphoric acid (10 w$ P„0,-) as 2 5

well as clean calcium sulphate hemihydrate is currently being developed. This process has the following features: * The rate limiting step in the digestion of phosphate ore by phosphoric acid at 50 °C changes from chemical reaction at low v% P„0C (about 30 wj) to diffusion beyond 35 wj P ^ . At 90 °C diffusion will be the rate limiting step also for 30 vi$ P205.

* Due to the extremely low sulphate concentration in the digester, unground phosphate ore can be used in this process, because blinding will not occur.

52

* The phosphate incorporation in HH can be expressed as a linear relationship between the molar phosphate over sulphate ratio in the solution and in the crystals between 40 - 48 w$ P^Cv at 90 °C.

* The phosphate content of the HH crystals decreases with increasing sulphate concentration in the crystalliser. It was found that above 3 w$ H.S0L in 40 - 48 w$ PpCv» the phosphate content of the HH was lower then 0.1 w$ P 0 .

* The Cd incorporation increases with increasing sulphate concentration in the crystalliser for residence times of 20 minutes.

* No influence of small variations in stirring rate and residence time on the cadmium and phosphate uptake has been observed.

* In the Cd removal step, the Cd concentration has to be lowered until about 3 ppm in solution, to obtain HH with less then 1 ppm Cd incorporated, unless complexing agents for Cd are used.

* Due to the precipitation of a quantity of radium in the digestion stage, the HH obtained from the crystalliser contains less radium than HH from commercial processes.

* Separation of the solids after the digestion stage is difficult due to the presence of fine ore residue particles.

* A maximum filtration rate has been observed with HH crystals obtained from phosphoric acid containing 40 w$ P.0_ and about 1.8 w% H„S0„.

* The filtration rate of the HH crystals decreases with increasing residence time, above 20 minutes, in the crystalliser.

* It is recommended that, in the crystallisation stage, the sulphate concentration should not exceed 2 w? H.SO,,, because of the steep increase in Cd incorporation and the relatively slow decrease in phosphate incorporation with further raising sulphate concentrations. Moreover, the filtration rate is decreasing with increasing sulphate concentrations above 2 w$ H.SO..

* An equation has been derived for the calculation of the fluoride distribution coefficient as a function of the operating conditions.

3.7. Literature

1. Becker, P., (1983), Phosphate and Phosphoric Acid, Fertiliser Science and Technology Series, Vol. 3, New York, Marcel Dekker, Inc.

2. Burova, M.S. and Kazak, V.G.,0985), Crystallization of calcium nitrate from nitric-phosphoric acid solutions in the presence of an inert coolant. Khim. Prom. 1 (1985) 29.

53

3. Chemische Fabrik Budenheim, DE 3218599 A1 1-2-1983, DE 3327391 A1 11-2-1985. I. Chemische Fabrik Budenheim, DE 3332698 21-3-1985. 5. Dahlgren, S.E. and Hakansson, R., (1966), Manufacture of 10-45 per cent P-Cv

d D

phosphoric acid by the semihydrate process. Acta Polytech. Scand. Chem. Met. Ser. 52 (1966) 3.

6. DSM NL-Patent NOS 8006916 20-12-80. 7. Elmore, K.L. and Farr, T.D., (1910), Ind. Eng. Chem. 32, 1 (1910) 580. 8. Elmore, K.L., Hatfield, J.D., Dunn, R.L. and Jones, A.D. (1965),

Dissociation of phosphoric acid solutions at 25 °C, J. Phys. Chem. 69 (1965) 3520.

9. Flemmert, G.L. (1977), Hydrogen fluoride and pyrogenic silica from fluosilicic acid. The Fertiliser Society of London, proceedings No 163.

10. Frochen, J., Laine, J., Allyot, H. and Gosse, G. (1972), Concentrated phosphoric acid manufacturing process by double crystallization. Bulletin d'informations HEURTEY No 57 (1972).

II. Hignett, T.P. and Mudahar, M.S. (1985), Energy consumption in the world fertilizer sector. Fertilizer Technology 1, 1 (1985) 39.

12. Hunter, D. (1985), Low-Cadmium H.-PO for fertilizer use. Chemical Week 136, 8 (1985) 25.

13. Jansen, M., Waller, A., Verbiest, J., Van Landschoot, R.C. and Van Rosmalen G.M. (1981), Industrial Crystallization '81 The Hague, Jancic and De Jong, ed., pg 171.

11. Knudsen, K.C., (1985), The Fertiliser Society of London, Proceedings of 3 oct. 1985.

15. Kopylov, V.A., Senotova, G.I. and Pozin M.E. (1975), Equilibrium partial vapor pressure of HF, SIFj. and H„0 over solutions of the system H.PO.,-H_SiF,-H_0, Zh. Prikl. Khim. 18, 11 (1975) 2155. 3 H d 0 d

16. Leder, F., Park, W.C., Chang, P.W., Ellis, J.D., Megy, J.A., Hard, R.A., Kyle, H.E., Mu, J. and Shaw, B.W. (1985), New process for technical-grade phosphoric acid, Ind. Eng. Chem. Process. Des. Dev. 21 (1985) 688.

17. Nordengren, S., Francia, I. and Nordengren R. (1955), The firs.t installation of a phosphoric acid plant according to the anhydrate method at Verticelli, Italy. The Fertiliser Society of London, Proceedings No 33.

18. Phosphoric acid manufactoring using hydrochloric acid. Phosphorus and Potassium 125 (1983) 29.

19. Phosphoric Acid, Outline of the industry. British Sulphur Corporation Limited, London, 1981.

51

20. Pitzer, K.S. (1973), Thermodynamics of electrolytes. 1: Theoretical basis and ground equations, J. Phys. Chem. 77 (1973) 268.

21. Purifying wet-process phosphoric acid. Phosphorus and Potassium 139 (1985) 34.

22. Silverman, S.R., Fuyat, F.R. and Weiser, J.D. (1952), Am. Mineralogist 37 (1952) 211.

23. Sluis, S. van der, (1987), PhD thesis. To be published. Technical University of Delft.

24. Sluis, S. van der, Leenhouts, W., Wesselingh, J.A. and Rosmalen, G.M. van, (1987), Filtration and washing of calcium sulphate-phosphoric acid slurries. To be published.

25. Sluis, S. van der, Murach, Z., Wesselingh, J.A. and Rosmalen, G.M van (1986), The predigestion stage of a new phosphoric acid process, Procedings of World Congrss 111 of Chemical Engineering, Tokyo, Japan, Vol. 4, (1986) 96.

26. Sluis, S. van der, Schrijver, J.H.M., Baak, F.P.C, and Rosmalen, G.M. van, Fluoride distribution coefficients in wet phosphoric acid processes, to be published.

27. Sluis, S. van der, Witkamp, G.J. and Rosmalen, G.M. van (1986), The crystallization of calsiumsulfate in concentrated phosphoric acid. J. Cryst. Growth, 79, (1986) 620.

28. Societe Uranium Pechiney. patent NL 8403788, 16-06-1985. 29. Taperova A.A. and Shulgina, M.N. J. Appl. Chem., (USSR), 23, (1950), 27. 30. Technisch Keramiek, Technieuws W-84-09, okt. 1984.

Ministerie van Economische zaken, Direktie R. en 0., 's Gravenhage. 31. The nitro-phosphates alternative.

Fertilizer international 209, (1985) 8. 32. Tjioe, T.T., Weij, P. and Rosmalen, G.M. van (1986),

"Removal of cadmium by anion exchange in a wet phosphoric acid process", Proceedings of World Congress 111 of Chemical Engineering, Tokyo, Japan, Vol. 2, (1986) 925.

33. UKF, private communications (1986). 34. Willis, M.S., Tosun. I. and Collins, R.M. (1985),

Filtration mechanisms. Chem. Eng. Res. Des., 63, (1985), 175. 35. Witkamp, G.J., Schuit, S.P.J. and Rosmalen, G.M. van (1986),

Recrystallization of calcium sulfate modifications in phosphoric acid, Condensed papers of the Second International Symposium on Phosphogypsum, University of Miami, W.F. Chang ed., Miami, Florida, USA, (1986) 106.

55

1. THE DIGESTION OF PHOSPHATE ORE IN PHOSPHORIC ACID

1.1 . Summary

Phosphoric acid is used for fertiliser applications. To minimize environmental pollution by heavy metals, a new process is being designed. One of the important steps in this process is the complete digestion of phosphate ore in phosphoric acid. The rates of digestion of fluoroapatite particles in phosphoric acid were determined. The particle size was varied from 150 to 2000 ym, the phosphoric acid concentration from 30 to 50 w$ P2

0c anci t h e temperature between 60 and 90 °C. The rate of digestion was found to be controlled by transport of calcium ions to the solution. A model for the digestion process was developed, based on the diffusion of calcium ions as the rate limiting step. The masstransfer coefficients, calculated by applying this model on the experimental data, are slightly increasing with temperature and

-5 -1 have values of 5 to 10 * 10 i « 3

1.2. Introduction

Phosphoric acid is a major component of many fertilisers. It is mainly produced by digestion of phosphate ore (fluoroapatite) with sulphuric acid. This yields phosphoric acid as a product and hydrated calcium sulphate as a byproduct. The ore also contains traces of heavy metal ions, of which cadmium and radium are the most worrisome. They can cause serious environmental problems (Becker,1983). A new "clean technology" phosphoric acid (CTPA) process is therefore being developed to reduce these problems (Van der Sluis e.a.,1986).

This process aims at producing a 10 w$ P?0_ solution, with less than 5 ppm cadmium and hydrated calcium sulphate with less than 1 ppm cadmium.

Contrary to the conventional processes, the ore is not digested in a mixture of sulphuric acid and recycled phosphoric acid, but only in recycled phosphoric acid. This allows almost complete digestion of the ore, without the simultaneous precipitation of all calcium sulphate. After separation of the insoluble ore residue and a minor amount of calcium sulphate a clear solution is obtained, from which the cadmium can be removed by ion exchange (Tjioe e.a., 1986).

After removal of the cadmium ions, the precipitation of the calcium ions is performed by adding sulphuric acid.

56

Large equipment is required, due to the very large recycle stream of phosphoric acid needed for complete digestion. For its design reliable data on the digestion rates are a necessity.

Several publications have been devoted to this subject, which are unfortunately contradictory. Three different rate determining steps have been proposed: - diffusion of calcium ions away from the particle.

(Serdyuk e.a., 1982 and Hufmann e.a., 1957) - diffusion of hydrogen ions towards the particle.

(Bloise e.a., 198H) - the chemical reaction of the acid with the ore particle.

(Ivanov e.a., 1977) This inconsistency explains, why the subject was reinvestigated under the

conditions expected in the CTPA process.

t.3. The digestion stage of the CTPA process

The phosphate ores used in wet processes today are mainly sedimentary phosphates or francolites. These ores consist of fluoroapatite, with part of the phosphate ions replaced by fluoride and carbonate ions. By beneficiation of the phosphate rock, large amounts of waste material are removed. The phosphate ore, however, still contains residual waste components.

If phosphate ore is digested by phosphoric acid, the following reactions can take place (Elmore and Farr, 1940):

Ca1Q(P0lt)6F2 + U H-POjj > 10 CaHPO^ + 2 HF (1)

C aiO ( P 04 )6 F2 + 14 H3P°t > 1 0 C a ( H2P 0H )2 + 2 HF <2>

The amount of calcium ions, which can be dissolved strongly depends on the temperature as well as on the phosphoric acid concentration of the solution. These parameters also determine, which type of calcium phosphate is formed as shown in figure 1 (Elmore and Farr, 1940).

The phosphoric acid used for the digestion is recycled product acid. This recycle stream unavoidably contains sulphate ions in an amount dictated by the operating conditions in the crystalliser. So part of the calcium sulphate, up to a maximum of 25 w%, will precipitate during digestion.

57

When a mixture of phosphoric acid and sulphuric acid is used for the digestion of the ore, two processes occur more or less simultaneously: digestion of the ore and precipitation of calcium sulphate hemihydrate. At high sulphate concentrations the calcium sulphate tends to precipitate not only upon the hemihydrate crystals, but also upon the ore particles. Such a coating is often called blinding (Becker, 1983).

6_i w%CaO

5 -

4-

3-

2-

1-

100°C

V 30 35 40 45 w%P205

Figure 1: Solubility of CaO in phosphoric acid

Since a high digestion rate, as well as complete digestion should be obtained, the sulphuric acid concentration in the digestion stage has to be carefully selected to avoid blinding. In a continuous process blinding can only be avoided by maintaining a high calcium concentration in the digestion stage to keep the sulphuric acid concentration sufficiently low (< 2 w$ H-SO^MVan der Sluis e.a., 1986). Under such conditions, sulphate ions precipitate upon the hemihydrate crystals and the digestion process can be treated as if proceeding in pure phosphoric acid.

In most of the work on the digestion process, as presented in literature, the solid product is a coprecipitate of hydrated calcium sulphate and solid calcium-(mono or di)-hydrogen-phosphate, since this product is often directly used as a fertiliser. The complexity of this combined precipitation, however, hardens the unraveling of the mechanisms of the individual steps. The calcium-(mono or di)-hydrogen-phosphate can, however, also be produced as a solution by digestion of the phosphate ore in a large amount of phosphoric acid, at low

58

sulphate concentration. In this case the rate determining step was found to be calcium diffusion both, in diluted phosphoric acid (< 5 w$ P.0C) at 25 °C (Hufmann e.a., 1957) and in 47 wj P ^ , between 15 and 85 °C (Serdyuk e.a, 1982). It was, however, also reported to be a fourth order chemical reaction in 24-47 w$ P-O^, between 80 and 120 °C (Ivanov e.a., 1977). The confusion became

£ D . . . even larger, when in 1984 a mathematical model was presented, in which the rate limiting step of the digestion step at 70 °C was assumed to be diffusion of hydrogen ions to the particle surface (Bloise e.a., 1984). No unambiquous conclusion could therefore be drawn about the rate determining step of the digestion process.

1 = thermostated reactor, 2 = polypropene baffles, 3 = six-bladed polyvinyl-

idene fluoride starrer, 4 = thermometer, 5 = polypropene lid with

silicon-rubber seals, 6 = syringe, 7 = reflux cooler.

The sampling hole is not shown.

Figure 2: Experimental reactor used for digestion of phosphate ore in phosphoric acid.

4.4. Experimental

The phosphate ore used was Zin ore from Israel. Four different size fractions (150 - 212 ym, 425 - 500 ym, 1000 - 1180 ym and 1700 - 2000 ym) were obtained by sieving through "Twente" sieves. To remove agglommerates of smaller ore particles from these fractions, 500 grams of each size fraction were gently stirred twice with about 2 kg of distilled water and sieved again. Then the

59

fractions were dried at 60 °C for at least 21 hours. Chemically pure phosphoric acid (85 wj H^POj.) and distilled water were used to prepare the solutions containing 30,40 and 50 w$ P0CL (i.e. 11, 55 and 69 wj H.PO,,).

d 3 3 1 Figure 2 shows the reactor used, a double walled glass vessel with baffles.

The temperature was kept at a constant value of 60, 75 or 90 °C by a thermostat. To prevent evaporation of water during the experiment, the vessel was

equiped with a polypropene lid with a reflux cooler. A 50 ml plastic syringe was used to feed the ore suspension in water to the reactor. The power input,

3 provided by the six-bladed turbine stirrer, was maintained at about 5 kW/m (1000 rpm). This power input is sufficient to keep even the largest particles suspended.

An amount of 20 or 10 grams of an ore fraction was suspended in 8 or 16 grams of water and 1 to 3 grams of antifoaming agent in the plastic syringe. This suspension was added to 500 or 1000 grams of phosphoric acid of the selected concentration and temperature. In this way an instantaneous suspending of the ore particles in solution was obtained.

The digestion process was followed by taking 3 ml samples from the reactor with a "Finn"-pipette. Sampling intervals varied between 15 seconds at the start to 15 minutes towards the end of the experiment. The samples were immediately and quickly filtered over Teflon filters, provided with "Millipore AP 2002500" filter cloth. The filtrate was accurately weighted.

The CaO content of the filtrate was analysed by ICP (Klok e.a., 1985) and/or a potentiometric titration (Jordan and Monn, 1967; Pribil and Vesely, 1966; Szekeres e.a., 1965).

1.5. Results

The results are shown in the figures 3, 1 and 5 for various concentrations of the phosphoric acid (30, 10 and 50 wj P2°5^* t h e temperature (6o, 75 and 90 °C) and the mean particle sizes (181, 163, 1090 and 1850 ym).

Since the Zin phosphate ore used is a mineral, it is not homogeneous. As a consequence the CaO content of each particle size fraction is different as shown in table 1.

The final CaO concentration reached in the solution is therefore used to estimate the CaO content of the ore particle size fraction, used in that particular experiment.

60

Figure 3: Digestion curves for phosphate ore in phosphoric acid containing 30 w$ P»0_ at three temperatures; mean ore particle diameter: 2 5 a = 181 ym, b = 463 ym, c = 1090 ym and d = 1850 ym.

Figure 4: Digestion curves for phosphate ore in phosphoric acid containing 40 w$ P.0_ at three temperatures; mean ore particle diameter: a = 181 ym, b = 463 ym, c = 1090 ym and d = 1850 ym.

Figure 5: Digestion curves for phosphate ore in phosphoric acid containing 50 w$ Pj^c; a^ three temperatures; mean ore particle diameter: a = 181 ym, b = 463 ym, c = 1090 ym and d - 1850 ym.

61

I Analysis of w? CaO (ICP) mean particle size (urn) | 1 2 3

181 | 55.3 52.9 53-0 463 | 53.0 1090 | 50.8 50.6 51.3 1850 I 52.2

Table 1: CaO content of Zin phosphate ore.

4.6. Discussion

4,6.1. Influence of the phosphoric acid concentration

The effect of the phosphoric acid concentration on the digestion rate of the phosphate ore was studied at constant values of: the stirrer speed, the particle size, the temperature and the weight ratio of phosphoric acid and phosphate ore. Comparison of the figures 3a, 4a and 5a up to 3d, 4d and 5d shows, that the time needed for complete digestion increases from 30 to 50 w? P Cv. This is probably due to the raise in viscosity with increasing phosphoric acid concentration as shown in table 2.

Temperature | Phosphoric acid concentration °C I 30 wjt Po0c 40 w* P,0C 50 wj Po0c

do do d O ^ ^

60 | 1.69 2.88 5.47 75 | 1.37 2.30 4.21 90 | 1.14 1.88 3.31

Table 2: Viscosity data (n * 1O3 Ns*m ) for phosphoric acid (Slack, 1968)

The viscosity is an important parameter, because of its influence on the diffusion rate. The decreasing digestion rate with increasing acid concentration, which is most clearly revealed by the curves for the largest particles, is a strong argument against a hydrogen ion diffusion limitation. Moreover, diffusion of hydrogen ions is commonly considered to proceed very

62

fast. Either the diffusion of calcium ions or the chemical reaction therefore seem more likely to be the rate controlling step in the digestion process.

4.6.2. Influence of the temperature

The effect of the temperature can also be observed from the figures 3, 4 and 5, and in particular from the curves for the largest particle sizes. In all cases, the time needed to reach complete digestion decreases with increasing temperature. The differences between the digestion rates for 60, 75 and 90 °C is, however, rather small. By using only the slope at t = 0 of the digestion curves presented in the figures 3» '4 and 5, an almost constant energy of activation of the digestion process, Ea, has been calculated for each curve. Its value was about 15 kJ per mol, independent of particle size and phosphoric acid concentration. This rough value of Ea, indicates that the digestion rate is not controlled by a chemical reaction, which normally is more sensitive to the temperature.

4.6.3. Conclusion

From the results obtained, it is reasonable to assume, that the diffusion of calcium ions from the surface of the reacting ore particles into the bulk of the solution is the rate determinig step of the digestion process.

In the following sections a kinetic model of the digestion process, based on this assumption, will be derived and the experimental data will be used to calculate the masstransfer coefficients and thus to verify the validity of the model.

4.7. A kinetic model of the digestion process

In this section a model wil be developed for the digestion process of ore particles of one size, which is based on the following assumptions: - the transport of acid to the surface of the particles proceeds fast. - at the surface, the acid rapidly reacts, until the saturation concentration

of Ca(HoP0„)„ is reached.

63

- the transport of this calcium phosphate into the solution is rate limiting. The masstransfer coefficient, k for its transport has a constant value.

- the particles of one size fraction are spherical and have the same initial radius, RQ. They do not desintergrate during digestion.

- the solution is well mixed. The rate at which calcium ions (taken as CaO) diffuse from a single

particle into the solution is then given by:

f * ~ = k * A * (c - cj (1) dt s t

Since the particles are spherical, their mass, M and external surface area, A are given by:

M = p * 5 * IT * R3 (2)

A - n * TT * R 2 (3)

This leads to:

f * »s * o1 = k * (cs - V <"> It is convenient to simplify this equation by introducing the ratio's:

R - g-, o - -- and t = T r 5 - g - - S - ; - j .

This yields:

— ; - 1 - o (5) dt

A second relation is obtained from the total CaO balance:

M0 = Mfc + L * cfc (6)

64

After complete digestion the final CaO concentration in the solution, c is obtained:

MQ = L * cf (7)

Eq. (6) and (7) are easily combined to give:

M c M0 Cf

or in a dimenslonless form:

R*3 - 1 - Si (9)

cf with p = — cs

So o = p « (1 - R 3) (10)

Together eq. (5) and (10) yield:

----- ;= = dt* (11) a + p * R

with a = (1 - p). Integration of the left hand side of eg. (11) (Weast, 1976) leads to:

F(R*> ■ 3-5" * [\ * I" (^~-5-2-«) ♦ /3 * arctan (-2--X---"-)) (12)

with Y = (a/p)1/3. In a digestion experiment the following parameters are known beforehand, R

and e„. s The c values used in this paper are those determined by Elmore e.a. 3 » (Elmore e.a., 19^0). If c,. and c. are measured, R can be calculated from eq. *- r * (9) and so the corresponding value of F(R ). According to eq. (11) a plot of * F(R ) against time should give a straight line with a slope:

65

k * c f"-R--;--P ;

The masstransfer coefficient, k is then determined by the value of this slope. The density of pure fluoroapatite is 320O kg/m3 (Weast, 1976). The real

density of the Zin ore was estimated with a pyknometer to be about 2700 kg/m3. This value was used in the calculations.

4.8. Determination of the masstransfer coefficient

The function F(R ) given by eq. (12) was plotted against time for all experiments presented in the figures 3, 4 and 5. For each experiment a straight line was obtained, which could be fitted by the method of least squares with a regression coefficient larger than 0.99. Examples of these fits are given in figure 6. From the slope of the straight lines, the masstransfer coefficients were calculated. These values are presented in table 3.

FIR')

60 °Crc= 0,994 75°Crc= 0,999 90 °C re =0,991

100 «- lime Is)

Figure 6: F(R ) as a funct ion of time in 40 w$ P„0_ a t t h r e e temperatures for a mean ore p a r t i c l e diameter of 463 \m-

• 30w%P2O5

X 40w%P205 D 50w%P2O5

-1 and Figure 7: In k as a function of T the phosphoric acid concentration for a mean particle diameter of 463 ym.

66

For a given temperature, the masstransfer coefficients depend only slightly on the particle size. Their values are close to those predicted by existing correlations for the masstransfer coefficients of particles in mixed vessels (Calderbank, 1961).

The influence of the phosphoric acid concentration can best be observed from the results obtained for the largest particle size fractions, because their digestion proceeds slowest. A slight decrease of the k values with increasing phosphoric acid concentrations is observed. This decrease is caused by the increase in viscosity, given in table 2 and also follows from the existing correlations.

concentration of phosphoric acid and temperature

30 wj P205

60 75 90

40 w* P205

60 75 90

50 w* P205

60 75 90

mean

181

4 6 7

5 6 8

4 6 7

diameter of

urn 463

4 6 8

5 6 7

4 5 6

the

pm

ore particle

1090

5 7 9

5 6 8

4 6 7

urn

size fraction

1850 ym

7 9 11

6 8 9

4 7 8

Table 3: Calculated k values (« 10 m * s ).

The temperature also has a slight influence on the k values. Due to the decrease of the viscosity and the increase of the diffusion coefficient there is a small increase of the k values with temperature. The influence of the temperature on the digestion process can be related to an activation energy, which can be obtained from an Arrhenius plot of In k versus T . In figure 7

67

this is illustrated for one particle size fraction (463 um) and three phosphoric acid concentrations (30, 40 and 50 u% P 20 5).

The energy of activation appeared to be independent of the particle size fraction and nearly independent of the phosphoric acid concentration. Its value lies between 13 and 23 kj/mol, an order of magnitude expected for a diffusion controlled process.

t.9. Conclusive remarks

Although Huffman e.a. (Huffman e.a., 1957) also found that calcium ion diffusion was the rate limiting step of the digestion process, their model did not take into account the considerable change in surface area of the ore with time. Their equation describing the digestion process in diluted phosphoric acid is therefore too simple and cannot be used to calculate the masstransfer coefficients. The model presented by Serdyuk e.a. (Serdyuk e.a., 1982) was an empirical model, in which a coefficient of digestion was introduced and an undefined constant. Their model is thus only applicable within the limits of their experiments and cannot be used to calculate masstransfer coefficients, because the relationship between the rate of digestion and the masstransfer coefficient is not well defined. The model used here does not suffer from these disadvantages.

From the foregoing, it can be concluded that the digestion of Zin phosphate ore in chemically pure phosphoric acid (30 - 50 w$ P»0_) between 60 and 90 °C

d. o can be described by a model in which calcium ion diffusion is the rate limiting step. 1.10. Nomenclature

2 A external surface area of the phosphate ore m c. CaO concentration in the bulk of the solution after

-3 complete digestion of the ore kg * m

-■* c saturation concentration of CaO in phosphoric acid kg * m 3 -3

c CaO concentration in the bulk of the solution at time t kg * m f weight fraction of CaO in the phosphate ore k masstransfer coefficient m * s L mass of phosphoric acid kg M mass of the phosphate ore at time t kg R radius of an ore particle (mean) at time t m

68

t time p density of the Zin phosphate ore

1.11. Literature

1. Becker, P., Phosphate and Phosphoric Acid, Fertiliser Science and Technology Series, Vol. 3., Marcel Dekker Inc., New York (1983).

2. Bloise, R., Shakourzadeh, K. and Baratin, F., Industrie Minerale Techniques, 9 (1984) 721.

3. Calderbank, P.H. and Moo-Young, M.B., Chem. Eng. Sci., 16 (1961) 39. 4. Elmore, K.L. and Farr, T.D., Ind. Eng. Chem., 32, 4 (1940) 580. 5. Huffman, E.O., Cate, W.E., Deming, M.E. and Elmore, K.L.,

J. Agr. Food Chem., 5, 4 (1957) 266. 6. Ivanov, E.V., Zinyuk, R.Yu. and Pozin, M.E.,

Zh. Prikl. Khim., 50, 6 (1977) 1193. 7. Jordan, D.E. and Monn, D.E., Anal. Chim. Acta., 37 (1967) 42. 8. Klok, A., Tiggelman, J.J., Weij, P., Dalen, J.P.J. van and Galan, L. de,

Proc. XXIV Colloq. Spectrosc. Intern. (1985) 98. 9. Pribil, R. and Vesely, V., Talanta, 13 (1966) 233-10. Serdyuk, V.V., Tereshchenko, L.Ya., Panov, V.P. and Chekreneva, G.M.,

Zh. Prikl. Khim., 55, 10 (1982) 2190. 11. Slack, A.V., Ed., Phosphoric Acid, Fertiliser Science and Technology

Series, Vol. 1., Marcel Dekker Inc., New York (1968). 12. Sluis, S. van der, Meszaros, Y., Wesselingh, J.A. and Rosmalen, G.M. van,

Proc. of the Fertiliser Society of London, nr. 249 (1986). 13. Szekeres, V.L., Kardos, E. and Szekeres, G.L.,

J. fur Praktische Chemie, 4, 28 (1965) 113. 14. Tjioe, T.T., Weij, P. and Rosmalen, G.M. van,

Proc. of World Congress 111 of Chemical Engineering, Tokyo, Japan, Vol. 2 (1986) 925.

15. Weast, R.C., Ed., Handbook of Chemistry and Physics, 57 ed, CRC Press Inc., Cleveland, Ohio (1976) A-116.B-241.

s kg *

69

5. CRYSTALLISATION OF CALCIUM SULPHATE HEMIHYDRATE

5.1. Summary

In a continuous crystallisation procedure calcium sulphate hemihydrate (HH) crystals were formed by feeding a calcium-di-hydrogen-phosphate (CDHP) solution in concentrated phosphoric acid (5.6 and 6.6 mole H.P0u per kg), simultaneously with a sulphuric acid solution, into a 1 liter crystalliser at 92 °C. The slurry removal occurred unclassified. Residence times of 20 to 80 minutes were maintained. The solid over liquid weight ratio depended on the calcium content of the CDHP solution. By slight adjustment of the feed streams the molar calcium over sulphate ratio in the solution was varied between about 0.05 and 8. A cadmium concentration of 30 ppm was maintained in the crystallising solution. The shape of the HH crystals varied with increasing sulphate concentrations from mainly aggregates to mainly needles.

During crystallisation of the HH crystals, cadmium as well as phosphate ions are incorporated into the lattice. The degree of uptake strongly depends on the operating conditions. The phosphate uptake decreases with increasing sulphate concentrations in the crystalliser, while the cadmium uptake increases. The molar phosphate over sulphate ratio in the crystals is found to be proportional to the same ratio in the solution. A comparable linear relationship is also observed for the molar cadmium over calcium ratios at a residence time of 20 minutes. At high sulphate concentrations, long residence times and high temperatures (about 95 °C) the stable anhydrite modification (AH) tends to develop next to the metastable HH phase. The cadmium uptake in AH is at least 10 times higher than in HH, while the phosphate uptake is not influenced by a small amount of AH in the crystals. The cadmium incorporation in the HH lattice can be reduced by adding halogenides to the solution.

5.2. Introduction

During the production of phosphoric acid for fertiliser applications huge amounts of calcium sulphate are precipitated as a byproduct. This byproduct not only tends to incorporate phosphate ions, but also cadmium ions, originating from the phosphate ore [1,11]. Uptake of phosphate lowers the overall efficiency of the process [i|], while the incorporation of cadmium hampers the disposal of the byproduct due to environmental restrictions. In order to obtain hydrated

70

calcium sulphate with a low cadmium and phosphate content from concentrated phosphoric acid, a new process is being developed.

In this process [t] a clear calcium-di-hydrogen-phosphate (CDHP) solution is obtained by digestion of the phosphate ore (fluoroapatite) in recycled

2+ phosphoric acid, enabling the removal of Cd ions by e.g. ion exchange [17]. Thereafter the major amount of calcium sulphate hemihydrate is precipitated from the CDHP solution by adding sulphuric acid. The incorporation of phosphate ions and of the remaining cadmium ions during this crystallisation has to be minimized. The amount of foreign ions incorporated largely depends on the operating conditions, given by the H.PCv and H-SO. content of the solution, the temperature, and the supersaturation, which is, in our experiments, on its turn determined by the residence time and the solid over liquid weight ratio. Also impurities originating from the ore or purposely added components affect the uptake of foreign ions by calcium sulphate. Components which are able to coordinate the cadmium ions in solution, like e.g. halogenides [8,9], reduce the cadmium uptake.

1 1 5 10

^ mole h^PCykg

Figure 1: The precipitated calcium sulphate phase'as a function of the H,PCv content and temperature of the solution. Note that HH only exists as a metastable phase, and that AH is s table in a very wide range of conditions. Our working conditions are indicated with a + sign.

71

The nature of the precipitated calcium sulphate phase also influences the uptake of foreign ions. Depending on the temperature and the H^PCK and H SO, content of the solution either calcium sulphate dihydrate (DH), hemihydrate (HH) or anhydrite (AH) will be formed [16], as shown in figure 1.

The solid lines in figure 1 represent equilibrium curves, but since some of the modifications are metastable under the given conditions, these lines merely indicate which phase will initially precipitate. The broken line in figure 1 separates the regions where either DH or AH is stable. The HH phase only exists as a metastable phase. The influence of H_SOü on the position of the curves in figure 1 can be taken into account by assuming 1 mole of H?SOj, to be equivalent with about 1.5 mole of H-POn [1]. The working conditions in our process, where HH is produced, are indicated in figure 1.

The aim of the present study is to investigate the incorporation of both phosphate and cadmium ions in HH in dependance on the above mentioned operating conditions during continuous crystallisation. In addition a start has been made with the investigation of the influence of halogenides on the incorporation of cadmium ions.

5.3. Experimental

In all experiments chemically pure H PCv, H-SCL, 3CdSCL.8H_0 and ammonium halogenides were used. The CDHP solution was prepared by dissolving CaCO.. in a concentrated phosphoric acid solution at about 92 °C. The calcium content of the CDHP solution is restricted by its solubility and was about 1 mole/kg in a solution containing 5.6 mole H.PCv/kg (55 w$ H_PCL) and approximately 0.6 mole/kg in a solution containing 6.6 mole H.PCv/kg (65 w% H-PCL). The CDHP

-4 2+ solutions contained 3*10 mole Cd /kg, which is the same as in the crystalliser.

Concentrated sulphuric acid was added in the experiments where a solution containing 5.6 mole H-POjVkg was feeded into the crystalliser, while a solution containing 1.7 mole HpSCv/kg and 5.4 mole H,P0^/kg was added in experiments with 6.6 mole H_P0j. per kg solution in the crystalliser.

In figure 2 the experimental set-up is given. The 1 liter double-walled crystalliser of glass was provided with a six bladed polyvinylidene fluoride (PVDF) turbine stirrer. The CDHP and the sulphuric acid inlet tubes as well as the drain tube served as baffles. The power input, generated by the stirrer, was calculated from the stirrer speed, the slurry density, and the geometry of both vessel and stirrer [1], and was normally approximately 1 kW/m (stirrerspeed 640

72

rpm). The supply vessels as well as the crystalliser were thermostated at about 92 °C except for the concentrated H2

S0n suPP!y vessel, which was kept at room temperature. Peristaltic pumps were used, equipped with marprene or viton tubes.

Figure 2: Scheme of the crystallisation apparatus.

At the start of each experiment the crystalliser was filled with a solution composed in accordance with the selected experimental conditions and an adapted amount of HH from a former experiment. The continuous crystallisation was performed by simultaneous addition of the CDHP solution and the sulphuric acid solution to the crystalliser. The slurry volume in the reactor was kept constant by continuous, unclassified, removal of slurry from the crystalliser. The residence time of the slurry, defined as the reactor volume divided by the slurry volume stream, was varied in the experiments from 20 to 80 minutes.

In each experiment a steady state was obtained after 8 to 10 residence times. For experiments were no HH was initially added, a longer runtime appeared to be necessary to obtain a steady state. In experiments where solutions were used containing 5.6 mole H P0./kg solution, the solid over liquid weight ratio was about 0.11, while in case of 6.6 mole H POjj/kg solution this ratio was 0.05. To vary the H SO^ content of the solution from 0.02 to 0.35 mole/kg at a given H-POjj concentration, corresponding with molar calcium over sulphate ratios between about 0.1 and 8, only small changes in the CDHP and H2S01) feed streams

73

were needed. The influence of such small variations on the solid over liquid ratio was negligible. Also the residence time had hardly any effect on it.

After each residence time unclassified samples were taken from the crystalliser with a sample tube. This tube was connected with a G.-glass filter, placed on a reception tube under vacuum. The liquid was diluted with water and

2+ analysed for its phosphate and sulphate content by X-ray fluorescence, its Ca 2+ content by complexometric titration and its Cd content by inductively coupled

plasma spectroscopy (ICP) [7]. The crystals were washed two times with water of about 90 °C, saturated with gypsum and then three times with acetone and dried

2-at 60 °C. The HPCv content of the crystals was determined by colorimetry, 2 + while the Cd content was measured by ICP. X-ray diffraction was used to

identify the phase of the crystals.

Figure 3a: HH crystals obtained at 92 °C in 5.6 mole H.POj,/ kg and 0.05 mole f^SO^/kg.

5.4. Results and discussion

5.4.1. The hemihydrate crystals

When hexagonal calcium sulphate hemihydrate crystals are precipitated from a solution containing 5.6 or 6.6 mole H,P0j,/kg, where the calcium ions are in

74

excess of the sulphate ions, aggregates are formed. Each aggregate consists of numerous needles originating from one point and outstretched like the radii of a sphere (see figure 3a) •

At increasing sulphate concentration in the solution, further referred to as [H„SO.], separate needles start to be formed beside the aggregates. At still higher [H SOJ mostly separate needles are grown, which are longer and relatively thinner. In addition some bundles of two or more needles are formed, which are often oriented in twin positions (see figure 3b).

Figure 3b: HH crystals obtained at 92 °C in 5.6 mole H,P0,,/kg and 0.24 mole H S0^ /kg.

At [H SOJ beyond about 0.25 mole/kg many tiny needles are formed, beside needles with a length up to 2000 ym (see figure 3c).

In 6.6 mole H P0n/kg aggregate formation still occurs at much higher [H-SO.] than in solutions containing 5.6 mole H.PO./kg. Also the needles are smaller than those obtained in 5.6 mole H_P0u/kg. This might be caused by a hampered volume diffusion to the tops of the needles, due to the higher viscosity of the solution containing 6.6 mole H,P0j./kg.

In many of the long needles, hollow cores were observed in the direction of the c-axis (see figure 3d).

75

Figure 3c: HH crystals obtained at 92 °C in 5.6 mole H-POj/kg and 0.32 mole H S0^ /kg.

Figure 3d: SEM view of hollow cores in HH crystals.

76

Examination of the obtained c r y s t a l s by X-ray d i f f r a c t i o n revealed t h a t some of the batches of c r y s t a l s contained anhydr i t e . The de tec t ion l i m i t for AH in the c r y s t a l s i s approximately 0.5 - 1 w%. The amount of AH varied even a t f ixed opera t ing cond i t i ons . In s p i t e of t h i s i r r e p r o d u c e b i l i t y , i t can be concluded tha t a combination of a prolonged res idence t ime, dosage of concentrated su lphur ic ac id and a high [H„S0,J in the so lu t ion promotes the formation of AH.

The conversion r a t e of HH in to AH i s enhanced, when the operat ing condi t ions are approaching the HH-AH equi l ibr ium l i n e given in f igure 1, i . e . at high [HpSOn] and/or [H POJ and a t temperatures c lose t o 95 °C. The dosage of concentra ted su lphur ic acid might l o c a l l y induce AH n u c l e i . Once AH nuclei a re p r e s e n t , t h e res idence time becomes very important .

5 . 1 . 2 . Incorpora t ion of phosphate ions

The predominant ions in the phosphoric ac id so lu t i on a r e , according t o Elmore [ 3 ] , the H , H_P_0Q and HS0„ ions . In the c r y s t a l s , however, more

2- 5 2 8 2 - ^ 2-l i k e l y HPO,. ions rep lace SO,, i ons . The HPO,. ion can e a s i l y s u b s t i t u t e for a su lphate ion, s ince these two ions are almost s imi l a r in s i ze and share the a f f i n i t y towards calcium i o n s . The s i m i l a r i t y of these anions i s a l s o r e f l ec t ed by the ex i s t ence of two comparable s a l t s : CaHP0ü.2H.0 and CaS0ü.2H 0, which are both spar ing ly so lub le in water and have the same molar volume [ 1 9 ] .

F i r s t c r y s t a l l i s a t i o n experiments were performed in a s o l u t i o n containing 5.6 mole H.-PO^/kg, with res idence times of 20 and 40 minutes, and [H SO,,] between 0.02 and 0.35 mole/kg. These condi t ions yielded a s o l i d over l iqu id

2-weight r a t i o of 0 .11 . In f igure 4 the HPO,, incorpora t ion in the c r y s t a l s , ?- 2 -

expressed as the molar [HPO,. ] over [SO,, ] r a t i o in the c r y s t a l s i s given versus the molar [H-P0J over [H S O j r a t i o in the s o l u t i o n .

The phosphate uptake decreases with r a i s i n g [HpSO,] and a l i n e a r dependence between t h e molar phosphate over su lpha te r a t i o s in the c r y s t a l s and in the so lu t ion appears to e x i s t . The r e l a t i o n s h i p between the molar r a t i o s presented by the drawn l i n e in f i g u r e 1 for 0.02 < [H SOJ < 0.35 mole/kg i s given by equation ( 1 ) :

[HPO,2"] [ H , P 0 J - — ( c r y s t a l ) = 1 . 4 * 1 0 * --*---- ( s o l u t i o n ) (1)

CSO^ ] [H SO ]

77

where 1.U*1o" denotes the partition coefficient, D(HPOa ).

100 200 300 Figure 4: The phosphate incorporation, expressed as the molar phosphate over

sulphate ratio in the crystals, as a function of the same ratio in the solution for various residence times. The slope of the curve,

2- -4 D(HPCL )» is approximately 1. 4 * 1 0 .

This linearity points at a competition between the two anions for the available adsorption sites at the crystal surface. An influence of the linear growth rate upon the phosphate uptake can therefore be expected, since at a higher growth rate more phosphate ions will be entrapped by the propagating steps. This growth rate is determined by the actual supersaturation. Measurement of the supersaturation is impossible because the differences between the calcium and sulphate concentrations in the solution and their respective saturation values are too small to be determined. Therefore, the production rate, defined as the amount of HH obtained per kg suspension per second, has been used as a yardstick for the linear growth rate, under the assumption that the specific surface area of the crystals remains constant. Although the specific surface area tends to increase at higher supersaturation, it will never obscure the higher supersaturation inherent to the higher production rate.

In case of unclassified slurry withdrawal the production rate is directly correlated with the solid over liquid weight ratio and the residence time. In

78

all experiments the solid over liquid weight ratio is about constant, and is determined by the calcium content of the CDHP solution and by the sulphate concentration of the sulphuric acid solution. This is only true because the amount of not precipitated calcium ions is extremely low, due to the slight solubility of HH in phosphoric acid (about 0.07 mole per kg) [1]. At shorter residence times, the production rate is higher and therefore the supersaturation will be higher, and an enhanced phosphate incorporation is expected. Comparison of the results obtained for various residence times, as given in figure 4, does not show this tendency, so the differences in supersaturation are apparently too small to cause a visible effect in uptake.

Additional experiments were performed, with a power input of either 0.04 kW/m or 11 kW/m (stirrer speed 250 and 1570 rpm respectively). No effect of these variations on the phosphate uptake was noticed. This shows that either the film layer thickness is hardly effected by the change in power input or that volume diffusion is not a predominant growth barrier in our experiments.

Besides the experiments in 5.6 mole H.PCL/kg also experiments in 6.6 mole hLPOj/kg were performed to examine the effect of the H-.P0. content. For 5.6 mole H PO,, /kg it was found that for changing [H?S0j.] in the solution, the phosphate

2- 2-uptake in the crystals, given by the molar [HPO^ ] over [SO. ] ratio, is ruled by the same ratio in the bulk of the solution. If this rule remains valid for varying phosphate concentrations in solution, no shift of the curve in figure H is expected for different H..P0,. concentrations. Experiments performed in a solution containing 6.6 mole H PO./kg yielded values, lying on thesame line as obtained for solutions containing 5.6 mole H-PO^/kg. The crystallising solution, however, is quite different. A lower solid over liquid weight ratio and a lower solubility of calcium sulphate prevail, while due to the higher viscosity of the solution the diffusion coefficient is also lower. No estimation can therefore be made of the differences in supersaturation for experiments performed in either 5.6 or 6.6 mole H-PO^/kg. Since, however, all results obtained in 6.6 or in 5.6 mole H.POjVkg for various residence times contribute to the same line in figure 4, small changes in supersaturation, due to variations in residence time and phosphoric acid concentration, apparently do not influence the uptake significantly.

Results formerly obtained from batch experiments [6] in 6.6 mole H,P0j/kg, where the production rate was varied over a much wider range, showed that the influence of larger increases in feed rates and thus in supersaturation can no longer be neglected. Therefore, if much higher feed rates were applied, a higher uptake can be expected in our experiments.

79

In some of our experiments AH was formed. In the experiments where the AH percentage of the crystals remained below approximately 10 w}, no significant increase of the phosphate uptake was observed. A visible higher uptake was noticed only for AH percentages above 15 w$.

Finally a few experiments were performed where under further similar conditions either Cl , Br or I was added. Addition of halogenides did not influence the phosphate uptake.

5.4.3. Incorporation of cadmium ions

2+ 2+ 2+ Since the ionic radii of Ca and Cd are almost equal( r(Ca ) = 1.12 A

and r(Cd2+) = 1.07 A [5], both with an 8-fold coordination [2,12]), the Cd2+ 2+ ions can replace the Ca ions in the crystals, without causing much lattice

strain. The chemical properties of these ions, on the contrary, are quite different. Cadmium sulphate as well as its hydrates are very soluble in aqueous

2+ 2-solutions [11], so the interaction between the Cd ions and the SOj. ions at

2+ 2-the crystal surface must be much weaker than between the Ca ions and the S0Ü

ions. This also counts for a solution containing 5.6 mole H-POn/kg at 92 °C, where the molar solubility of cadmium sulphate has been measured to be about 0.7 mole per kg solution, which is ten times higher than the molar solubility of HH. For the experiments in 5.6 mole H P0./kg with sulphuric acid concentrations ranging from 0.02 to 0.35 mole/kg and residence times of 20 and 40 minutes, the cadmium concentration in the HH crystals was measured. In figure 5 the cadmium uptake, expressed as the molar cadmium over calcium ratio in the crystals, is presented as a function of the sulphate concentration in the solution.

The [HpSOj,] in the solution can be converted to a molar cadmium over 2+ calcium ratio, because the product of [Ca ] and [H_S0.] is about constant, and

-3 2 2 2+ equal to 5*10 mole /kg in 5.6 mole H,P0,./kg, while [Cd ] in the solution was 3*10 mole/ kg in all experiments. In figure 5 the molar cadmium over calcium ratio in the solution is also given on the horizontal axis. If, in analogy with the phosphate uptake, a competition between cadmium and calcium ions for the available adsorption sites at the interface takes place, a linear relationship between the cadmium over calcium ratio in the crystals and in the solution is expected. For experiments performed with a residence time of 20 minutes this linearity indeed exists for sulphate concentrations of 0.1 up to 0.35 mole/kg.

Experiments with 40 minutes residence time could not be performed at sulphate concentrations beyond about 0.25 mole/kg, without excessive AH formation. 80

The partition coefficient for cadmium, expressed as:

D (Cd2+) ( [Cd2+] / [Ca2+] ) ( crystal )

( [Cd2+] / [Ca2+] ) ( solution ) (2)

equals 1*10 J for HH.

3.10 5-

2.10~5H

1.10"5H

residence time Q12%AH

[Cd] — - (crystal) ICa]

• 20 min O M)min

O 11%AH

.0-' O

0.1 0.2 °,3 I H 2 ^ % ) 0.5-10" 1.10" 1.5-10"'

^ (solution)-[Ca] Figure 5: The Cd incorporation expressed as the molar Cd + over Ca ratio

in the crystals as a function of the same ratio and of the sulphate concentration in the solution. The star indicates the reduced level of incorporation that could be obtained by addition of halogenldes

-2 to the solution (see text), e.g. 0.7*10 mole I per kg solution.

For sulphate contents below 0.15 mole/kg the cadmium uptake is higher for 40 than for 20 minutes residence time. For the moment no explanation for this phenomenon can be given. An indication, that at sulphate concentrations below 0.05 mole/kg the cadmium incorporation is lower than predicted by our experiments with 40 minutes residence time, is given by the results of continuous operation of the phosphoric acid process on a bench-scale. During this operation a small amount of HH is precipitated during the digestion of the ore. The solution contains in that case about 0.005 mole H SO./kg, and in the crystals only a very small amount of cadmium, corresponding to a molar cadmium

81

over calcium ratio below H * 10 is found. This difference in cadmium uptake can hardly be resulting from a difference in residence time (120 versus 10 minutes) and in solid over liquid ratio (0.01 versus 0.11 or 0.05) only. It is therefore not excluded that for [H-SO^] below 0.05 mole/kg the actual cadmium uptake will decrease. The tendency of increasing uptake with raising [H_S0^] is also found in batch experiments [17], although in that case the extent of the incorporation was higher [12],

A variation in power input from 0.01 to 11 kW/m did not have a visible effect on the cadmium uptake. In figure 5 the strong influence of AH formation on the cadmium uptake is elucidated. The cadmium uptake in the AH at various sulphate concentrations could be estimated from experiments, where AH was formed, after correction for the amount of cadmium incorporated in the HH. These preliminary data make a linear relationship between the molar cadmium over calcium ratio in AH and the solution likely. The corresponding partition

-2 coefficient D is about 1*10 . This linearity points at a competition between the cations for the available adsorption sites at the interface, which is also assumed to occur for the phosphate and sulphate ions. Comparison of the partition coefficient for cadmium ions in the HH phase, which for a residence

-3 time of 20 minutes equals 1*10 , with the partition coefficient for phosphate -it 2+

uptake, which equals 1.1*10 in HH, demonstrates that substitution of the Ca 2+ 2-

ions by Cd ions occurs about 7 times more frequently than replacement of SOj, 2-ions by HPOj. ions. 2+ 2-

Regarding the weak affinity of the Cd ions towards the S0U ions at the 2- 2+

interface and the much stronger affinity of HPO^ ions towards the Ca ions at 2+ the interface, this high Cd uptake is rather surprising. The preliminary 2+ -2

partition coefficient for Cd in AH, D = 1*10 , is even substantly higher than 2+ 2-

D(Cd )HH and D(HP0Ü )HH. A possible reason for this enhanced uptake follows 2+ from the lower solubility (about two times) of AH, [Ca ] * [H.SOj is about -3 2 2 2*10 mole /kg in 5.6 mole H P0L/kg, which causes the supersaturation for AH S * 2* to be much higher than for HH, resulting in a higher Cd uptake in AH due to a kinetic overgrowth effect. Also thermodynamic effects have to be considered as a possible reason for the enhanced cadmium uptake in AH compared to the uptake in HH. These thermodynamic effects on the partition coefficient can be taken into account by writing D, according to Mclntire [13], and assuming the activity

2+ 2+ coefficients of Cd and Ca to be about equal, as: K(CaSCv) f . A l l / R T x

■>«"2+> = K(cds5;)* *( i u / R T ) ( 3 )

82

2 + 2 + Here the difference in chemical potential between Cd and Ca in the solution is assumed to be taken into account mainly by the quotient of the solubility products K, while An represents the difference in chemical potential

2+ 2+ between Cd and Ca in the crystal lattice. 2+ If Ay were similar for Cd uptake in AH and in HH, the partition

2+ coefficient for Cd in HH is expected to be twice the value in AH, due to the 2+ lower solubility of AH. The observed D value for Cd in AH is at least ten

times higher than its value in HH. Therefore, if we assume eq.(3) to be valid, 2+ 2+ Ay cannot be the same, and replacement of Ca by Cd possibly requires less

energy in AH than in HH. This assumption is not unlikely because CdSO. and CaSCv are almost identical salts, while this is not true for their hydrates.

Partition coefficients were also determined by Kushnir, who investigated the incorporation of many foreign ions in DH, HH and AH, precipitated from or recrystallised in seawater [10,11]. A striking difference between our results

2+ 2+ 2+* for Cd uptake and Kushnir's results for Mg uptake is, that D(Mg )AH is two

2+ 2+ times smaller than D(Mg )HH, while we found D(Cd )AH to be more than ten times

2+, 2+ larger than D(Cd )HH. The partition coefficient for Mg in AH from seawater is 2+ smaller than for Cd in AH from phosphoric acid, although the solubilities of CdSCv and MgSCv in aqueous solutions are comparable. The apparantly existing

2 + 2 + differences in Ay values of Mg and Cd in CaSCv are probably caused by the 2+ 2+ 2+

larger differences, e.g. in ionic radii, between Mg and Ca compared to Cd 2+ 2+ 2+

and Ca , leading to a larger lattice mismatch for Mg . The Mg ions probably 2+ fit equally bad in the HH as in the AH lattice, while Cd ions fit better in 2+ 2+

the AH than in the HH lattice and apparently better than Mg ions. Our Cd incorporation results will be verified in the near future by recrystallisation experiments.

Finally a few experiments were performed where either Cl , Br or I was added. The preliminary results presented in figure 5 demonstrate that when 2.7*1 o" mole Cl" , 1.3*10~2 mole Br" or 0.7*10-2 mole i" is added to 1 kg

-14 2+ 2+ solution containing 3*10 mole of Cd , a reduction in Cd incorporation in the crystals of about HQ% is achieved. These results are in agreement with literature data concerning cadmium complexation by halogenides in aqueous solutions [6,7]. In these data a decrease in complexing activity is mentioned in the order I~ > Br" > Cl~.

83

5.5. Conclusions

- The habit of the HH crystals changes with raising [H SO.] from mainly aggregates to mainly needles.

2-- The HPCv incorporation decreases with raising [H~SOv] and is directly proportional to the phosphate over sulphate ratio in the bulk of the solution.

2+ - The Cd incorporation increases with raising [H.SCv] and is directly proportional to the molar cadmium over calcium ratio in the solution for residence times of 20 minutes and between about 0.15 and 0.35 mole H„S0. per kg solution.

- High HpSOj, concentrations, high temperatures and long residence times stimulate the formation of AH.

2+ - Under identical operating conditions, the Cd uptake in AH is at least 10 times higher than in HH.

2+ - Halogenides reduce the Cd uptake. 5.6. Literature

[I] Becker, P., "Phosphates and Phosphoric Acid," Fertiliser Science and Technology Series, Vol. 3, Marcel Dekker N.Y. (1983).

[2] Bushuer, N.N., Russian Journal of Inorganic Chemistry, 27, (1982) 341. [3] Elmore, K.L., Hatfield, J.D., Dunn R.L. and Jones, A.D.,

J. Phys. Chem. 69 (1965) 3520. [4] Gorecki, H., Indian Journal of Technology, 23 (1985) 51. [5] Huheey, J.E., "Inorganic Chemistry," Harper and Row (1975), p. 74. [6] Jansen, M., Waller, A., Verbiest, J., Landschoot, van R.C. and Rosmalen,

G.M. van, Industrial Cryst. (1984), 171, ed. Jancic and De Jong, Elsevier Amsterdam.

[7] Klok, A., Tiggelman, J.J., Weij, P., Dalen, J.P.J. van and Galan, L. de, Proc. XXIV Colloq. Spectrosc. Intern. (1985) 98.

[8] Kolthoff, I.M. and Elving, P.J., "Treatise on analytical chemistry," Vol. 3, Part 2, Interscience N.Y. (1961).

[9] Kostelnik, R.J. and Bothner-By, A.A., J. Magn. Resonance 14 (1974) 141. [10] Kushnir, J., Geochimica et Cosmochimica Acta, 44, (1980) 1471. [II] Kushnir, J., Geochimica et Cosmochimica Acta, 46, (1982) 433. [12] Lager, G.A., Armbruster, Th., Rotella, R.J., Jorgensen J.D. and Hinks, D.G.

Amer. Mineralogist, 69, (1984) 910. [13] Mclntire, W.L., Geochimica et Cosmochimica Acta, 27. (1963) 1209.

84

[1H] Slack, V.A., "Phosphoric Acid," Fertiliser Science and Technology Series, Vol. 1, Marcel Dekker N.Y. (1968).

[15] Sluis, van der S., Murach, Z., Wesselingh J.A. and Rosmalen, G.M. van, "The predigestion stage of a new phosphoric acid process," Proc. of World Congress III of Chemical Engineering, Vol JJ (1986) 96.

[16] Taperova, A.A. and Shulgina, M.N., J. Applied Chem. (USSR) 23 (1950) 27. [17] TJioe, T.T., Weij, P. and Rosmalen, G.M. van, "Removal of cadmium by anion

exchange in a wet-phosphoric acid process," Proc. of World Congress III of Chemical Engineering, Vol. 2 (1986) 925.

[18] Tjioe, T.T., Woude, van der H., Verbiest, J., Durville P.F.M. and Rosmalen, G.M. van, Proceedings of the international conference "Heavy metals in the environment," Vol. 1, 292, Athens 1985, Editor Lekkas, T.D.

[19] Weast, R.C., "Handbook of Chemistry and Physics," 59th edition, CRC press, Florida, (1978).

85

6. FILTRATION OF CALCIUM SULPHATE HEMIHYDRATE

6.1. Introduction

The separation of calcium sulphate from its slurry in phosphoric acid is always done by filtration.

The main objective of the filtration is to obtain a clear phosphoric acid, which is not diluted with wash water, by means of simple and economical equipment, which is also suitable for a thorough cake washing, to prevent discharge of phosphoric acid with the calcium sulphate cake.

The washing of the cake is extremely important, because about one million US dollars a year are lost, if 0.5 % of the P 0 produced in a 1000 tons P 0 per day plant is discharged with the cake. The amount of water available for cake washing strongly depends on the acid strength to be produced, on the type of hydrate of calcium sulphate formed as well as on the concentration of sulphuric acid used. It also depends on the moisture content of the cake discharged from the filter. Phosphoric acid processes operating nowadays, nearly always use three washing steps. This is certainly necessary for the processes, in which a high acid strength is produced, because the quantity of wash water available is then very small.

The wash water must be supplied, when the liquid level reaches the surface of the cake, because otherwise the cake tends to detach itself from the walls of the filter compartment, lowering the efficiency of the washing sometimes even from about 98 to 80 % [4],

Some of the most important parameters in the filtration and washing procedure are [2]: the physical properties of the liquid, the particle size, shape, size distribution and packing characteristics and the solids content of the slurry to be filtered.

The description of the filtration and washing of the calcium sulphate cakes is relatively easy, because the resistance of the filter cloth and support medium can usually be neglected and because the calcium sulphate cakes are found to be incompressible [4].

The total area needed for filtration and washing is found to be about four times the area needed for filtration only, if three wash steps are used [1]. To obtain a clear phosphoric acid, most of the time a so-called prefiltration section is incorporated. In general the pores of the filter cloth will be larger than the particles, which are to be removed and the filter will work efficiently only after an initial deposit has been trapped in the medium [1,2].

86

This first part of the product acid is normally recycled to the digestion stage of the plant. In some cases anionic polyelectrlytes are used as flocculants to improve the filtration characteristics [1].

Although filters are always selected from experience with the same type of ore in the same process or from pilot plant studies [1], laboratory studies of the filtration can give an estimation of the filter area needed. The reason, why plant data are always preferred, is the large influence of impurities on the size and shape of the crystals obtained as well as the occurrence of scaling, which is hardly observed in laboratory studies. Scaling can be caused by the formation of chukhrovite [1] or alkaline fluoro compounds, and is frequently encountered in phosphoric acid plants [1].

The first continuous filter used in 1930, was a horizontal belt filter. Around 19t5 horizontal table and tilting pan filters came into use [4].

The types of filters used at this moment are quite well described by Slack [1]; they have not changed greatly over the last twenty years. The most important parameters for selecting a filter are the capacity, the total investment, maintanance cost and susceptibility to corrosion, which is severe in phosphoric acid media.

The type of filter cloth used varies widely and depends on the acid strength to be produced, the temperature, the type of phosphate ore used and the required clarity of the phosphoric acid. Some of the materials used are : saran, polyethylene polypropylene, terylene and dacron.

The aim of this study is to determine the influence of the sulphate concentration and the residence time in the crystalllser on the filterabllity of the HH-phosphoric acid slurry, expressed as the permeability of the HH cake.

6.2. Theory of filtration

The principle of filtration is shown in figure 1. The filtrate flows through the cake and the filter medium, due to a pressure gradient. The resistance against the flow is the combined resistance of the cake and the filter medium.

In the experiments to be described, the pressure drop over the cake and filter medium is held constant. The cake formed is assumed to be incompressible [4] and homogeneous. Furthermore, the cake and filter medium are assumed to form two independent resistances in series. The filtrate flow rate per unit filter area is then given by:

87

AP "*~(Rc~ Rm) (1)

where Re and Rm are the cake and filter medium resistances respectively, AP the pressure drop over the cake and filter medium and n the viscosity of the filtrate. This can also be written as:

(Re + Rm) * 1 * 5Y A dt n (2)

1 = slurry 2 = filter cake 3 = filter cloth 1 = support medium 5 = filtrate

Figure 1: Filtration principle

The cake resistance can be expressed as a relationship between the specific cake resistance, a and the mass of solids deposited per unit filter area, w as:

Re * w = a * ps * (1 - e) * (3)

This proportionality between the cake resistance and the cake height can also be written as:

Rc= - (4)

where K is the cake permeability. If slip between the particles and the fluid in the feed is neglected, the

superficial velocities of the solid and the liquid are equal and the following relationship yields:

88

us Ul f 2 — = -i > u = i * u (5)

(1 - ef) ef s ef 1 The volume of solids deposited can be calculated from the following expression:

V = A * y u * dt (6) 3 O S and the cake height, h can be obtained from eq. (6) and the cake porosity, e, as:

h ■ r b * oK "dt - u-h) * ~--iy~ *o* ui •dt (7)

This can be written as:

(1 - e ) h - s--.-(rh) * i- (8)

in which V = volume of filtrate. After substitution of eq. (8) in eq. (1) , the result can be inserted in eq.

(2) to give:

(1 - E ) * V ( „___ff..-.I— + Rm) , 1 * 3V = AP (9) ^e„ * (1 - e) * A * K "m; A dt n ^'

which can be written as:

-Jj - o, * c2 * V (10)

« » . , * . 1 * Rm „ (1 " e ) « n in which c = -=-»-j and c .

ef * (1 - e) * A * K * AP So a plot of -- versus V should give a straight line. The slope of this line then yields, together with the values of e and ef, the cake permeability, K.

89

6.3. Experimental

6.3.1. Equipment

The equipment used for the filtration experiments is shown in figure 3. The top part of the filter consists of a double walled pipe of 0.4 m height, with a diameter of 0.031 m. This tube is thermostated. The liquid level in the tube can be measured with a measuring-tape on the outside. The lower part of the filter is shown in more detail in figure 4. It can be detached from the upper part by a screw fitting for easy cleaning. The connection between the two parts, as well as the connections between the glass and the metal are sealed with viton rings to prevent leakages. The liquid flowing through the filter medium passes a glass valve (inner diameter 0.01 m) and is collected in a round flask of 0.5 litre.

Figure 3= Filtration apparatus. Figure 4: Lower part of the filter with filter medium.

One outlet of this flask is connected to a vacuum pump and a manometer. The other outlet is also connected to the manometer and a 50 litre vessel, which acts as a buffer to maintain a constant vacuum. The volume of filtrate was registered by the lowering of the slurry level in the glass tube. After each cm, the time was noted down. The filter medium was filter cloth made of terylene 23-42 and glued between two PVDF rings.

90

6.3.2. Chemicals

The phosphoric acid-calcium sulphate hemihydrate slurries, produced from chemically pure reagents, were obtained from the crystallisation experiments described in chapter 5. These slurries were used for filtration experiments immediately after their formation, because otherwise the HH might be recrystallised into DH before the filtration and erroneous results would have been obtained.

6.3-3« Procedure

About 300 grams of the freshly formed calcium sulphate hemihydrate-phosphoric acid slurry is weighted and suspended in the' glass tube, which is thermostated at 90 °C. The slurry is homogenised by gently stirring for a few minutes. In the meantime the pressure below the filter medium is reduced to about 25 kPa. So a mean pressure drop over the cake and filter medium of 29 kPa will result during the filtration. The stirrer is quickly removed, just before the start of the filtration experiment. At time t = 0 the glass valve is completely opened and the level of the slurry in the glass tube is monitored as a function of time. The experiment is stopped, when the slurry level almost reaches the cake level. After measuring the cake height, the filter is opened and air is blown through the tube to remove the cake. The solids are three times washed with water saturated with gypsum, and thereafter three times with acetone. The weight of the solids is determined after overnight drying.

6.4. Results and discussion

The reproducebility of the results obtained with our equipment was first checked by performing filtration experiments after each residence time of the crystals in the crystalliser. From the results obtained in the crystallisation experiments as well as from literature, it is known that at least five residence times are required, before the crystalliser attains a steady state. It is also believed, that the crystal size, shape and size distribution hardly change after five residence times under the same conditions. The crystallisation experiment was performed at 90 °C in 40 w$ P n and 1.8 wj H»S0„. The mean residence time

dO d. H

in the crystalliser was 40 minutes. After 2, 3, 5, 6, 7 and 8 residence times a dt/dV versus V curve was made; these are shown in figure 5.

91

Figure 5 shows, that after 5 residence times the slope of the dt/dV versus V curve does not change anymore. This means that the permeability of the HH cake obtained is indeed constant and the results obtained with our equipment are reproduceble.

Important process parameters are the sulphate concentration in the solution and the residence time of the crystals in the crystalliser. The influence of both parameters on the permeability is shown in figure 6.

A maximum permeability of the HH cake is observed between a sulphate content of 1.5 and 2.0 w$ H SO. in the crystalliser. This was also found by Dahlgren [3] in a single stage HH process, without a separate digestion stage.

1 -

"-i-x10"5s/m3

10

number of residence times o = 2, • = 3 , D = 5 1 ■ = 6.7,8

50 -*- Vx106m3

100

Figure 5: Influence of the number of residence times of the crystals in the crystalliser on the dt/dV versus V curve.

The influence of the sulphate content in the crystalliser on the permeability of the HH cake can be explained by the size, shape and size distribution of the crystals.

At low sulphate concentrations small crystals and clusters of crystals are formed as shown in figure 7. At higher sulphate concentrations the crystals are larger and more uniform, as shown in figure 8, while at very high sulphate concentrations both large and small crystals are obtained, as shown in figure 9.

The influence of the residence time is also illustrated in figure 6. Above 1.5 w$ HpSO,, in the solution in the crystalliser, an increase in residence time

92

from 20 to 60 minutes seems to lead to a reduction in the permeability of the HH cake. This might be caused by the formation of small crystals, due to the breakage of the longer HH needless, obtained at longer residence times and sulphate concentrations above about 1.5 w$ H-SOn. The permeability reaches its maximum value with a residence time of 20 minutes and about 1 .8 w$ H?S0,. in the

-12 2 solution in the crystalliser. The maximum value is about 65 * 10 m .

70-

50-

30-

10-

residence time —•- 20 min --o-- 40min

60 min

I 3

w%H 2S0 4

Figure 6: Influence of the residence time and the sulphate content of the solution in the crystalliser on the permeability of the HH cake obtained at 90 °C in 40 w$ P 0_ with a stirrer speed of 650 rpm.

<i 5 x. = 11 w% AH, x 2 = 1560 rpm, x, 280 rpm.

Furthermore, the influence of a small amount of AH in the crystals on the permeability of the crystal cake was checked. The AH crystals are normally smaller than HH crystals and have a different size, shape and size distribution [1,4], The influence of only 11 w$ AH in the crystals on the permeability of the crystal cake is large indeed. The permeability obtained was about 50 % of the normal value, as illustrated in figure 6 by point x .

The influence of halogenides, added to reduce the cadmium incorporation in the crystals (see chapter 5 ) , is not shown in figure 6. No significant change of the permeability of the HH cake, however, was noticed.

93

Figure 7: HH crystals obtained at 90 °C in MO w$ P ^ and 0.5 w% H^O^.

Figure 8: HH crystals obtained at 90 °C in 10 w* P20 and 2.4 w$ H^O^.

94

A process parameter, which did influence the permeability was the stirring speed in the crystalliser. Under otherwise similar conditions the stirring speed of about 650 rpm was increased in one experiment to about 1560 rpm (point x_ in figure 6) and in another experiment lowered to about 280 rpm (point x_ in figure 6). Lowering of the stirring speed seems to increase the permeability of the cake. This might be caused by less breakage of the HH needles. Good stirring is, however, required to keep the crystals in suspension and to mix the CDHP and sulphuric acid solution thoroughly and quickly. The mixing of these solutions must be fast enough to prevent the formation of small crystals, due to locally high supersaturations. These small crystals will inevitably lower the permeability of the HH cake.

Figure 9: HH crystals obtained at 90 °C in HO w$ P Cv and 3.2 w$ H-SO^.

The results obtained here are for HH from chemically pure reagents. In reality, impurities will influence the size, shape and size distribution of the HH crystals. This influence has been checked in the bench-scale plant and will be presented in chapter 9.

Last but not least a comparison was made between the results predicted by our equipment and the results obtained with a commercial filter. The linear filtration velocity of the commercial filter can be obtained from the amount of Pp0_ produced per square meter filter area per day, which is given as 7 tons

2 P20,-/m per day. Because 50 % of the acid is recycled and 50 % is actually

95

produced, the amount of acid filtered is thus 14 tons P„0,_/m /day. The data 2 5

for the density and concentration of the commercial acid given in table 1 have -4 been used to calculate a linear filtration velocity of about 5 * 1 0 m/s. The slurry of the commercial plant was also used in our equipment, in which

-12 2 a permeability of this DH cake of 1 * 10 m was obtained (the slurry was filtered at 60 °C). The linear filtration velocity calculated from this permeability and the values of the cake height, viscosity and pressure drop

-4 given in table 1, is also equal to about 5 * 10 m/s. So simular results are obtained in our laboratory equipment and on a commercial filter for the same material.

acid concentration viscosity density cake height pressure drop permeability

over (chapt

cake er 9)

26 1.5 * 10~3

1250 5 * 10~2

50 -1 2 1 * 10

v% P20 Ns/m kg/m3

m kPa 2

m

Table 1: Data for a commercial filter.

The filterability of the HH-phosphoric acid slurry obtained from chemically pure reagents is good, because the values obtained for the permeability of the HH cakes are substantially higher than those obtained for the DH cake from a commercial plant.

96

6.5. Nomenclature A filter area h cake height K permeability Re cake resistance Rm filter medium resistance t time u. liquid velocity u particle velocity V filtrate volume V„ volume of solids w mass of solids deposited per unit area a specific cake resistance AP pressure drop e cake porosity e volume fraction of liquid in feed n viscosity p solids density

6.6. Literature

1. Becker, P., Phosphates and Phosphoric Acid, Fertiliser Science and Technology Series, Vol. 3-, 369-^OiJ, Marcel Dekker Inc., New York (1983).

2. Coulson, J.M., Richardson, J.F., Backhurst, J.R. and Harker, J.H., Chemical Engineering, Vol. 2, 3rd ed, pg. 126, 321, Pergamon Press Inc. New York (1978).

3. Dahlgren, S.E. and Hakanson, R., Acta Polytech. Scand. Chem. Met. Ser., 52 (1966) 3.

4. Slack, A.V., Ed., Phosphoric Acid, Fertiliser Science and Technology Series, Marcel Dekker Inc., New York (1968), Vol. 1, pg. 107.

97

m m

m m s m * s m s

1 -1

kg * m m * kg Pa

N * s * m kg -3

7. FLUORIDE DISTRIBUTION COEFFICIENTS (G/L) IN WET PHOSPHORIC ACID PROCESSES

7.1. Summary

Phosphoric acid for use in fertiliser applications is mainly produced by digestion of phosphate ore with sulphuric acid. The phosphate ore contains about 2-4 w$ F. During the digestion of the ore, the fluoride is released as hydrogen fluoride, which reacts with silica to fluorosilicic acid. Part of it precipitates with alkaline ions in the early stages of the phosphoric acid production process and part of it decomposes in SiFj. and HF, which on their turn partly evade into the air. The residual fluoride in the solution is distributed between the phosphoric acid and the byproduct, a calcium sulphate modification.

The disposal of this byproduct as well as the application of both the phosphoric acid and the byproduct can be hampered by their fluoride content. Removal of the fluoride from the phosphoric acid is therefore a necessity and has to be carefully controlled, because the fluoride concentration in the air is limited too. So accurate fluoride distribution coefficients between the phosphoric acid and the ambient air have to be known.

These fluoride distribution coefficients were measured by saturation of a nitrogen gas stream passing through several absorption bottles in series. The absorption bottles were filled with fluorosilicic acid dissolved in mixtures of distilled water, chemically pure phosphoric acid and chemically pure sulphuric acid at various temperatures.

A theoretical model, based on the description of the acid mixture with Elmore's model of phosphoric acid and the ion specific interaction model of Pitzer, is used to describe the results. An expression is derived, which allows calculation of the fluoride distribution coefficients between the acid and the ambient air at temperatures between 70 and 95 °C, for solutions containing 30-50 w$ PJD,., up to 6 w? H„S0„, maximal 4 w$ H„SiF, and a molar F/Si ratio of six.

Comparison of the fluoride distribution coefficients predicted by the derived expression with those measured for commercially "wet process" phosphoric acid with molar F/Si ratios below 10, showed no significant difference.

7.2. Introduction

The production of phosphoric acid is almost directly related to the world phosphate fertiliser consumption, which still tends to increase [28]. Phosphoric acid for use in fertiliser applications is mainly produced by wet processes. In

98

these processes the phosphate ore is digested by a mixture of sulphuric and phosphoric acid and huge amounts of hydrated calcium sulphate are precipitated as a byproduct. The phosphate ore is mainly fluoroapatite (Ca (P0jJ,F ) with some additional compounds such as calcite, quarts, clay etc.

During digestion of the ore the fluoride is released as hydrogen fluoride. The hydrogen fluoride reacts with the silica, present in the ore or added as a clay, to form fluorosilicic acid. Some of the fluorosilicic acid precipitates during the production process with sodium or potassium ions as Na SiF>, K SiF,, NaKSiF, or as more complex compounds [14]. The remaining fluorosilicic acid in o the phosphoric acid partly decomposes and evades as SiF,. and HF. The residual fluoride in solution is distributed between the phosphoric acid and the byproduct during its precipitation [2,28]. In this byproduct the fluoride is

2- 2-assumed to be present as A1F,. ions, which can substitute for the SO,. ions in b 4 the calcium sulphate lattice [19].

The fluoride content of the byproduct hampers its application as e.g. building material or as a settling retarder in cement [42], when the fluoride content is higher than about 0.15 wj F [15]. The disposal of the byproduct in natural water is restricted by an environmental limitation of 6 ppm F in seawater [2],

The fluoride remaining in the phosphoric acid is eventually deposited on the meadows through the application of phosphate fertilisers. This can give an enlarged fluoride content of the grass, which can cause fluorosis of e.g. cows [1.35].

An accepted limit for the fluoride content of phosphoric acid, used in fertiliser applications, is about 1 wj F [40]. If the phosphoric acid were used in cattle feed production, this limit is even lower (0.2 w$ F) [35].

The evaded fluoride has to be recovered, because the fluoride concentration in air is limited to about 2 ppm for environmental reasons [2,40].

The fluoride, evaded as SiF^ and HF, can be absorbed in water to give a fluorosilicic acid solution [2,28]. These H SiF, solutions are applied for fluoriding potable water in the U.S.A. and for sterilisation of wood [28].

In addition, processes to produce SiO. and HF from these solutions [12] are gaining field. The Si0o produced from H„SiF, solutions is a suitable basic

£ 2 b material for the production of ceramics [12,40]. This market for application of Si02 is growing, due to the increasing interest in ceramic materials [24],

Hydrogen fluoride is mainly used for the production of chloro-fluoro-carbons [28], but is also applied for the production of highly resistant

99

engineering plastics like PTFE and PVDF, which use is increasing [6] and for the production of cryolite (Na^AlF,) for use in the aluminum industry [28].

The fluorosilicic acid from phosphoric acid production processes is a welcome fluoride source, because the fluoride ore used nowadays for the production of hydrogen fluoride, fluorite (CaF„) is getting scarse [28].

For all reasons mentioned above the fluoride content of phosphoric acid has to be reduced and almost all fluoride has to be recovered. Possible ways to remove fluoride from phosphoric acid are extraction [10,14], precipitation [6,7,16,25], stripping and flashing [20,23,31].

If extraction methods are used, a second liquid phase is generated, which separation from the phosphoric acid phase can cause problems. Moreover, a relatively huge stream has to be purified, so the extraction equipment is expected to be large.

Another way to remove fluoride from the phosphoric acid is precipitation by addition of alkaline cations. The disadvantages of this method are the increased scaling in the process equipment caused by the permanent state of supersaturation with respect to the alkaline silicates in the phosphoric acid solution and the difficulties connected with the fluoride and SiO„ recovery from the mixture of solids.

So only stripping and flashing seem to be commercially feasible unit operations to remove fluoride from the phosphoric acid and to recover the fluoride. An advantage of these methods is the possibility to obtain the fluoride in a directly salable form.

In order to calculate the amount of fluoride evaded into the gas phase during stripping or flashing under various process conditions, the fluoride distribution coefficients between the phosphoric acid solution and the gas phase are indispensable data. In the next section a survey of previous measurements of these coefficients will be given and the objective of this study will be defined.

7.3. Literature survey

The fluoride distribution coefficients between the the gas phase and the phosphoric acid, which contains several additional compounds, such as sulphuric acid, hydrofluoric acid, fluorosilicic acid, iron, magnesium, aluminum and calcium ions, are essential data for an accurate design of the fluoride stripping unit of a phosphoric acid plant.

100

Most of the research on fluoride distribution coefficients presented in literature is done on H„SiFg-H 0 systems. Illarionov [18], Baur [1] and Shrambam [36] studied the H„SiF,-H„0 system. At low H„SiF,-concentrations they found the 2 o d. d o vapour pressure of HF to be higher than the vapour pressure of SiF.. The molar HF over SiF. ratio in the gas phase lessened with increasing H?SiF, concentration. According to Lehr [22] the hydrolysis of SiF,, is reduced and thus its evasion stimulated, due to the decrease in the water content of the solution with increasing H„SiF, concentration. Teslenko e.a. [39] even suggested that the fluoride is mainly present as SiF.(H-O),, in strongly acidic media.

Borisov and Mel'nikova [3] and Kopylov e.a [20] performed measurements with the chemically pure H.PO.-H SiF,-H„0 system. These investigations were carried out by passing an inert gas stream through the acid until saturation was reached. They report that the predominance of SiF. over HF in the gas phase increases with increasing phosphoric acid concentration in the system, but an increase of the temperature raises the molar HF over SiF. ratio in the gas phase. They measured the fluoride evasion, however, only at temperatures between 60 and 90 °C in phosphoric acid solutions with P_CL percentages between 10 and

2 D

55 w%, in which range they could only detect SiF. evasion. In a later article, however, Kopylov e.a. [21] were able to detect HF in

the gas phase as well, in the same concentration range during concentration of "wet process" phosphoric acid (W.P.A.). They measured an increase of the molar F/Si ratio in the liquid phase from 6 to approximately 100. So the vapour pressure of SiF. above phosphoric acid must be higher than of HF [21], probably due to the stronger interaction of HF with phosphoric acid. This results in an increasing HF concentration in the liquid phase during concentration of W.P.A.

Also Senetova e.a. [33,3*0 and Spijker [38] studied the evasion of fluoride into the gas phase during concentration of W.P.A.. Spijker investigated the evasion of fluoride not only during concentration of the H P0.-H2SiF,-H?0 system, but also during concentration of the H PO.-HF-HpO system. He presents the percentage of evaded fluoride as a function of the Po0c content of the

2 D phosphoric acid during concentration. From these experiments a sharp raise in the amount of evaded fluoride was noticed above approximately 10 wj P_0C, if the

2 5 fluoride is initially present as H SiF,. This must be assigned to the reduced hydrolysis of SiF^ by water at the decreasing water content of the W.P.A. during concentration, resulting in an enhanced SiF. evasion. Only the fluoride content and not the silicon content in the gas phase was measured, so the molar F/Si ratio remained unknown.

101

The amount of evaded fluoride will thus decrease near the end of the concentration of W.P.A., partly due to the lower Henry coefficient of HF compared to SiF^ and partly due to the formation of mono-fluoro-phosphoric acid, H.PO.F, above approximately 10 w$ P..CL. The amount of mono-f luoro-phosphoric acid increases with an increasing molar F/Si ratio in the solution [21].

Senetova e.a. [33] also measured the influence of the addition of sulphuric acid and Si0? to the W.P.A. on the evasion of fluoride during concentration of W.P.A.. Both variables are found to enhance the fluoride evasion.

Odintsova e.a. [27] measured the influence of Mg ions on the fluoride evasion. An increase of the Mg concentration in the solution diminishes the partial vapour pressure of fluoride owing to the precipitation of MgF?, which even predominates the formation of MgSiF,-.

To remove almost all fluoride from phosphoric acid and to obtain the highest fluoride distribution coefficients it is therefore necessary to maintain a molar F/Si ratio of approximately 6 in the solution by continuous addition of silicon as for instance a clay.

Furthermore the corrosion of the equipment used for the removal and recovery of the fluoride as well as for the production of phosphoric acid is found to be less if the fluoride is present as fluorosilicic acid instead of hydrogen fluoride [13,26].

Although a lot of results and general trends are presented, fluoride distribution coefficients between the liquid and the gas phase for phosphoric acid solutions between 70 and 90 °C containing 140-55 wj Pp05, 0-5 w$ H SO, and 0-4 wj F and a molar F/Si ratio of approximately six are lacking.

The objective of this study is therefore to determine fluoride distribution coefficients between phosphoric acid solutions and the gas phase, while maintaining a molar F/Si ratio of about six in the phosphoric acid solution. The influence of the phosphoric acid concentration, temperature, sulphuric acid and fluoride content of the phosphoric acid on the fluoride distribution coefficients is studied.

7.4. Model development

7.4.1. General remarks

In order to predict the fluoride distribution coefficient between the liquid and the gas phase as a function of the temperature and process

102

parameters, such as the concentration of H„SiF,, H„SO„ and H PO. in the 2 O 2 4 3 1

solution, an expression has to be developed. The composition and the temperature of the solution are believed to be the

most essential parameters, which determine the amount of evading fluoride. In the description of the composition the activities of the species have to be taken into account, instead of their concentrations since we are dealing with highly concentrated electrolyte solutions. The activity coefficients of the species in such solutions can be calculated with the ion specific interaction model of Pitzer [30].

7.H.2. Calculation of the activity coefficients

Pitzer [29D showed that the ion activity coefficients of an electrolyte MX can be calculated from the following equations:

I n YM = z**F + E m *(2*B M + Z*CM =) + E m * ( 2 * 0 M + E m *ï- ) M M a Ma Ma c Mc a Mca a c a

+ I * E E m *m ,** ,M + | z J * E E m *m *C (1 ) 2 , a a ' a a ' M ' M ' c a c a a a ' c a

in Tx = z£*F ♦ E V ( 2 * B X c + Z"CXc> + l m a * ( 2 * Q X a + Z V W c a c

+ I * I E m *m *<? , + | z v | *E E m *m *C (2) 2 , c c ' cc'X ' X' c a ca c c ' c a

where the subscripts M, c and c' refer to cations and X, a and a' to anions and where:

'/ F = _«*(___! + I * ln(1+b*I /a)) + E E m *m *B' * ,.h,T 'z t> c a ca

1+b*I z c a + 5 * E E m *m *0' , + \ * Z E m *m *6' (3) 2 c c' cc' 2 , a a' aa' J

c c' a a'

Z = E m *|z. I = 2.* E m *z = 2 * E m *|z I (H) i ' l ' c c a ' a' l c a

For 1-1 and 1-2 electrolytes the coefficients B.„, which are functions of the MX

ionic strength have the following form:

103

BMX " *m + 6MX) * *(x> (5)

B ,MX = B M X } * « , ( X ) / I ( 6 )

g(x) = 2 * (1 - (1 + x) * e"x)/x2 (7)

g'(x) = - 2 » (1 - (1 + x + | * x2) * e"X)/x2 (8)

x = 2 * /I and b = 1.2. (9)

In these calculations the third virial coefficients can be neglected, since these coefficients do not have a great influence on the activity coefficients [30]. This implies that the C and i|i terras can be omitted. Because we are dealing with concentrated electrolyte solutions the Debye-Huckel term A can also be omitted. Without these terms the equations (1) and (2) can be rewritten as:

In YM = E m *2*B„ + Z m *2*0M + M a Ma c Mc a c

+ z2 * ( Z Z m *m *B' + \* Z Z m *m *0' , + \* Z Z m *m ,*Q' ,) (10) M c a c a ca 2 c c, c c' cc* 2 & &, a a' aa"

in Yx = I mc*2*BXc + Z ma*2*0Xa + c a

+ z2 * ( Z Z m *m *B' + I* Z Z m *m ,*0' , + \* Z Z m *m »0' ,) (11) X c a ca 2 . c c' cc' 2 , a a' aa' c a c c' a a'

Due to the high ionic strength of the solutions, g(x) in equation (7) and g'(x) in equation (8) will tend to zero and the derivatives of the second virial coefficients can be considered to equal zero. These approximations are simplifying the equations (10) and (11) considerably:

In Y = Z 2*m *BM + Z 2*m «0.. (12) M „ a Ma c Mc a c

in Yx - Z 2*mo*BXc ♦ Z 2*ma*0Xa (13) c a

These expressions for the activity coefficients of an electrolyte MX in a concentrated electrolyte solution can be used for a description of the liquid

■\0k

phase in terms of activities of the species, if the composition of this phase in concentrations of the species is known.

The hydrogen ion concentration of the phosphoric acid solutions can be calculated by using Elmore's model [8], as will be shown in the next section.

In section T.H.H. it will be demonstrated that only the hydrogen ion concentration has to be calculated, because the concentrations of the other species can directly be obtained from the amount of H-PO., H.SCK and H SiF, applied for preparing the solutions.

7.4.3• Calculation of the hydrogen ion concentration in phosphoric acid

In order to calculate the hydrogen ion concentration in phosphoric acid, Elmore e.a. [8] assumed the following three equilibria to take place mainly:

HgPO^ ■* H+ + H,,PO~ K1 (11))

H6 P2°8 t H+ + H5P2°8 % (15)

H3 P°4 + H2P0~H t H5P2°8 K0 (16)

where the K-values are given in molarities as: K1 = 7.1 * 10" 3, KQ = 1.263 and K 4 = 0.3.

The existence in concentrated phosphoric acid of the dimeric (H P O J or HgPpOg was recently confirmed by Wertz and Cook [41],

The hydrogen ion concentration at 25 °C can be calculated for a 1-1 electrolyte with equation (25) from an article of Elmore e.a. [8] as:

1 O 8 ( K I - H 3 P ° ' . ) + l 0 8 ( 1 + K°* aH3pV = iStislisfe+ 2*log ° c +

+ 0.3387298 * ac + 0.030961 * (ac) 2 (17)

where a = activity of the undissociated part of the phosphoric acid monomer

expressed in mole per litre, c = the total phosphate concentration expressed as H.PO^ monomer

per litre and a = cH

+/c. Substitution of cR+ for ac in equation (17) leads to:

105

-1.0182/c + ^V^PO^ + ^V^PO^ m ÜÜ396557Ö-; + 2*10g °H+ +

+ 0.3387298 * cH+ + 0.030961 * c + (18)

The activity of the undissociated part of the phosphoric acid monomer, au on , can be calculated, according to Van Rysselberghe [32] with the equation: H3POt

d log a = -§5i5_(_d log a„ .) (19) H3P°4 BH P04

H2°

in which 55.5 is the number of moles H O per kg HO. The hydrogen ion molarities can be converted into hydrogen ion molalites

with the following formula:

pl CH PO,, H3P01( Elmore e.a. [9] calculated the values of a„ _ as a function of the phosphoric

acid concentration from vapour pressure data above phosphoric acid solutions at 25 °C with:

log a - log jj- (21) H20 PQ

where P/Pn = vapour pressure of the phosphoric acid solution divided by the vapour pressure of pure water.

Then they graphically evaluated the values of a at 25 °C from a plot of M3 4

-log a„ - versus 1/IIL, on for phosphoric acid concentrations between 7.5 and 80 "2° 3 H

molal. Equation (17) has also been used for phosphoric acid solutions at temperatures from 70 - 95 °C, because no information is available about the temperature dependence of the K-values and the constants in equation (17). Hereby it is assumed that the values of a„ for phosphoric acid solutions containing 30 - 55 w$ P_0,_ between 70 and 95 °C are the same as those obtained at 25 °C. This is 2 5 not unrealistic, if the ratio P/Pn in equation (21) is assumed to be hardly temperature dependent. From the equations (18) and (20) and the values of a„ p n given by Elmore e.a.

3 H [9] the hydrogen molality can be obtained as a function of the phosphoric acid

106

concen t ra t ion . Between 30 and 55 w$ P 0 a l i n e a r increase of the hydrogen mola l i ty with the phosphoric ac id mola l i ty i s found, which could be f i t t e d with the equation:

n^+ = 0.551 * n^ p o - 1.7179 (22)

7.H.H. Determination of an expression for the fluoride distribution coefficients

For the derivation of an expression for the fluoride distribution coefficients at the various phosphoric acid concentrations and temperatures, a model of the H„P0„-H„SiF,-H„0 system is needed. The following simplified model

3 1 <= t> <= is taken, by lack of any other adequate model of the system:

SiFg" + 2 H+ + SiF4(l) + 2 HF(1) (23)

SiF4(l) + SiFl4(g) (24)

HF(1) ; HF(g) (25)

The fluorosilicic acid is a strong acid and is assumed to be completely 2- + dissociated into SiF, and H in pure water as well as in phosphoric acid [2].

The equilibrium constant, K, of the dissociation of H?SiF, in the liquid phase (equation 23) is represented by:

a2 * a 3 HF aSiFü

K = (26) a2+ * a 2-H SiFf 0

The activities of the volatile SiFa and HF compounds in the liquid phase 2-are both assumed to be negligible compared to the activity of SiF, , at a molar

F/Si ratio of six, which has to be maintained in the liquid phase if a maximal fluoride evasion is pursued. The fluoride is therefore assumed only to be

2-present as SiF, in the liquid phase. In Pitzer's model as well as in Elmore's model molalities instead of

concentrations are used for the liquid phase. Since furthermore the gasphase can be assumed to behave as an ideal gas at about 100 kPa, the fluoride distribution coefficient can conveniently be defined as:

107

total_F_concentration in_the_gas_ghase SiF, molality in the liquid phase

Here H is defined in such a way that its value can easily be calculated from the measured total fluoride concentration in the gas phase and from the measured corresponding amount of F in solution, since the SiF. and HF concentrations in the solution are neglected. Equation (27) can be rewritten as:

(4 * CSiFl) + eHF>g H = i- (28)

The individual SiFü and HF concentrations in the gas phase are unknown, but these concentrations are related to the SiFj. and HF activities in the liquid phase by Henry's law:

CHP ( g ) h(HF) ■ a"(i) (29)

CSiFlJ(g)

*(SiV - I-U)

The activities of SiF,. and HF in the liquid phase are assumed to be 2-negligible compared to the activity of SiFfi , but it will be assumed that they

always remain in the same ratio q:

aHF ( 1 ) q = a--"(l) (31)

Substitution of equation (31) into eq. (26) and combination with eq. (29) and eq. (30) leads to:

CHF K*q*a*+ , m "2= = h(HF) * ( } * YSiF 2_ (32) mSiF2 (HF) a» 2- SlF6 6 SiF6

and:

108

CS1F, K*a*+ , - — « = h,„.„ , * ( ) /a * Y„ 2- (33) aSiF2 ( S l IV q 2*a*2- SiF6 6 SiFg

Substitution of eq. (32) and eq. (33) in eq. (28) gives:

H = 1»h, . * (— -tt ) ' 3 * Y 2- ♦ h * ( S-) '3 * Y 2- (34) < S l V q2*a%.p 2- SlF6 (HF) a%._2- SlF6 SiFg blt6

Since K, h,,. , , h. . and q are constant only the activity coefficients (HF) (MFjJ 2-and the hydrogen ion molality have to be known. The SiF, concentration follows

from the total amount of F added as H?SiF, (see 7.6). Taking the logaritm of equation (32) and rearranging the terms leads to:

lnH = V -f"1" V + -f"1" V " -f"ln mSiFf + -3"ln ^SiF2' (35) If the equations (12) and (13) for the activity coefficients of an

electrolyte in a concentrated electrolyte solution are applied on the H-PCL-HpSiFg-H-O system, containing as most important species H , H_P_0„ and SiF, , this delivers: 6

l n Y + = 2 * m 2 - * B + 2- + 2 * m - * B + - (36) 'H SiFf H -SiF^ HcPo0Q H -HcPo00 ^ J o o 5 2 8 b ^ o

ln ^SiF2" " 2 * V * V-SiF2" + 2 * mH5P2°8 ^ 0SiF6"-H5P2°8 ^

ln V208" ■ 2 # V * V-H^o" * 2 " ™SiF2- * VA-SiF2- (38)

In equation (35) only t he values of the a c t i v i t y c o e f f i c i e n t s given by equat ions (36) and (37) have t o be i n s e r t e d .

If a l so su lphur ic acid i s present in the l i q u i d phase, i t i s assumed to be + -

d i s soc ia t ed in H and HSCL, and t he equat ions (36) and (37) become:

I n Y + = 2*171 2 - * B + 2 - + 2*m - * B + - + ?*m - * R + - f W l i n TH m s . F £ BH _ s i F £ t mH5P2o8 uH - H 5 P 2 O 8

+ 2 m ^ V-HSO^ (39)

109

In Y 2- = 2*m. + * B + 2- + 2*m - * e 2-YSiF6

2 % H -SiF6 H5P20g 0SiF^ -H5P2Og

2*m - * 0 2- (MO) HSOj, SiFg -HSO^ K u'

The hydrogen ion molality is difficult to calculate, since no literature data for H PO.-H SO.-H SiFfi mixtures are available. In calculating the hydrogen ion molality the H_SiF, is assumed to be completely dissociated and one mole of

d 0 it therefore gives two moles of hydrogen ions. One mole of H SO. is assumed to deliver one mole of hydrogen ions. The hydrogen ions resulting from the dissociation of H.POj. can be calculated as indicated in the former section with equation (22). Because of the presence of H SO,, and H?SiF, the degree of dissociation of H,POü will lessen. The degree of dissociation of phosphoric acid was calculated to lessen about 5-7 % for phosphoric acid solutions containing 38 wj P„0C and up to 17 w? HoS0... This relatively small decrease is neglected, do d " because such high H-SO. or H?SiFfi concentrations will never be present in the phosphoric acid solutions prevailing in the CTPA process and the following equation is used:

m + = a * m - 1.7179 + m + 2 * m. H H,PO„ '•'"* 1 " H S O „ K siF, 3 4 d 1 do (Ml)

Equation (35) can be rearranged by inserting the equations (39), (40), (41) and (12) (with a = 0.551):

mH 5P 20- " -■»H3P0J, " K 7 1 7 9 {"2)

into:

In H = C* , | » in V - I * in m ♦ d 0

2 H 2 m x (i * a * a + 2 - + - * a * B + - + - * a * 0 2 - - ) m H 3 P0 4

( 3 BH - S i F ^ 3 BH -H 5 P 2 0g 3 ° S i F ^ - H ^ O g '

Q

+ m * ( - * B + 2 - ) + H 2 S i F 6 ^3 H - S i F g ;

U 2 ? m * (_ * g + _ + _ * e 2- - + - * n + „.„2-) -H2S0^ ^3 H -HS0^ 3 SiFg -HS0y 3 H -SiFg

110

U P 2 1 717Q * (- * B + - + - * 0 2- - + - * B + 2-) (H?1 '•<''* 3 DH -HSO,, 3 SiF^ -HcP„0o 3 H -SiF^ ; ^i>

1 0 D £. 0 D

Replacing the terms with the constant C., a, B and Q factors in equation (43) by the constants C to C^ leads to:

in H - C, ♦ f • in V - f • in m ^ ^ ♦ O, * m ^ + C3 * m ^ . ^

By this equation the dependence of the fluoride distribution coefficient, H upon known or calculable species in the phosphoric acid solutions, is given. The validity of this equation can now be verified by comparing the H-values calculated from this equation with those calculated from measured fluoride concentrations in the liquid and in the gas phase by using equation (27).

7.5. Experimental

7.5.1. Chemicals

Chemically pure phosphoric acid (85 wj H.PCO, chemically pure sulphuric acid (98 wj HjSOjj), technical grade fluorosilicic acid (36 w$ H2SiF6) with a molar F/Si ratio of 6, and distilled water are used to prepare the selected acid mixtures. The technical grade fluorosilicic acid is obtained from the fluoride recovery unit of a single filter HH/DH (Nissan-H) process. A molar F/Si ratio of six is maintained in the liquid phase, since the fingered filters of boro-silicate glass present in the solution acted as a continuous SiO source.

In all experiments the gas phase consists of nitrogen gas saturated with water vapour.

7.5.2. Equipment

In figure 1 the apparatus for the determination of the fluoride distribution coefficients is shown. It consists of ten wash bottles made of boro-silicate glass with an inside diameter of 3 cm each. The end of each long tube situated in the liquid was a fingered filter needed to generate small nitrogen gas bubbles in order to create a large gas-liquid interface for masstransfer.

111

Wash bottle 1 is needed to protect the nitrogen supply pipe against acid disposure. The wash bottles 2 till 4 contain the selected acid mixture without fluoride. In these wash bottles the nitrogen gas is heated to the selected temperature and saturated with water vapour. The water vapour pressure in the wash bottles 5 till 8 is assumed to be the same as in the wash bottles 2 till 4, since the composition of the liquid phase is only slightly different. The wash bottles 5 till 8 contain the same acid mixture as the bottles 2 till 4, but this time with the selected amount of fluoride added as H„SiF,.

The HF and SiF^ vapours are removed from the nitrogen gas stream by absorption in distilled water in the wash bottles 9 and 10. All bottles are filled with about 50 ml of liquid.

WASTE 6AS

WASTE 6AS

Figure 1: Apparatus for the determination of the fluoride distribution coefficients.

7.5.3. Procedure

The nitrogen gas is supplied at a temperature of about 20 °C by a nitrogen bottle. The flowrate of the nitrogen gas is measured by a dry-gas flowmeter and set with a flow controller. The maximum flowrate of the gas at a temperature of 20 °C is found to be 2.M 10 mVhr. At higher values the fluoride content of the

112

nitrogen gas stream will not be in equilibrium with the fluoride content of the liquid phase.

The real gas flowrate through the wash bottles at the selected temperatures of the experiments can be calculated with:

6 = * (20°C) * (-1-) * ( 1 2 ! — - ) (D5) " in5 - p

H20

After the nitrogen gas stream is set, the gas is led through the wash bottles 2 till 8 and led to the waste-pipe via three-way valve 2. After checking for leakages, the wash bottles 2 till 10 are placed in a thermostated water bath of the selected temperature.

When the content of the wash bottles has reached the required temperature after about 20 min., three-way valve 2 is reset and the gas is led through the absorption bottles 9 and 10. After a fixed time of about 30 to 60 min. the gas is led again to the waste-pipe by resetting three-way valve 2. Thereafter the complete apparatus is taken out of the thermostated bath.

The contents of the absorption bottles 9 and 10 is assembled in a 250 ml flask and diluted with distilled water to 250 ml.

The fluoride content in the absorption liquid from the wash bottles 9 and 10 as well as in the wash bottles 5 and 8 is measured. The contents of the bottles 5 and 8 are analysed for their fluoride concentrations to check if the fluoride content of wash bottle 8 remained constant and equal to the added amount of fluoride and if the molar F/Si ratios in the bottles 5 to 8 remained six. Constancy of both quantities ensured saturation of the gas stream with fluoride under the selected conditions.

7.5.4. Analyses

The fluoride content of the absorption bottles is measured with a fluoride selective electrode (F1052F of Radiometer). A TISAB-IV buffer [37] is used to keep the ionic strength of the solution at a constant value. Hereto 25 ml of the absorption liquid and 25 ml of the TISAB-IV buffer are well mixed. These acid probes are diluted with water by a factor 50-100 and analysed for their fluoride content with a calibration procedure.

113

The silicon concentration in the absorption liquid from the wash bottles 9 and 10 as well as in the wash bottles 5 and 8 is measured with the ICP-technique.

7.6. Results and discussion

In all experiments a molar F/Si ratio of six was found in the wash bottles 5 and 8 as well as in the absorption liquid from the wash bottles 9 and 10. This molar F/Si ratio in the liquid was maintained during the experiments, due to reaction in the liquid phase of the HF formed by partial dissociation of the fluorosilicic acid with the fingered filters of glass at the long tube ends in the wash bottles. The molar F/Si ratio in the gas phase was six or more, because the molar F/Si ratio in the absorption liquid was found to be also six. The surface of the bottles showed no signs of reaction with HF contrary to the fingered filters in the solution.

The fluoride distribution coefficients are calculated from the total amount of fluoride present in the wash bottles 9 and 10, the known quantity of gas led through the absorption liquid in the given fixed period of time and the fluoride content of the acid mixture in wash bottle 8 by applying equation (27). The procedure given in the appendix (7.10) was used to convert the concentrations given in wj P„0_, w? H_S0,, and w$ F into molalities.

The fluoride distribution coefficients are given in table 1 as a function of the temperature and the H_P0Ü, H.S0Ü and H.SiF, concentrations in the liquid phase. Table 1 shows that the fluoride distribution coefficients vary over a wide range, due to the influence of both the temperature and the acid concentration.

The influence of the phosphoric acid concentration on the fluoride 3 distribution coefficient (expressed in mg F/ m vapour/ w$F in the solution, i.e.

units normally used in the industry) is illustrated in figure 2 (curve 1). In this figure also the fluoride distribution coefficients obtained by Spijker [38] in a simulation of an industrial evaporator (curve 3) as well as the values of the fluoride distribution coefficients calculated from the results obtained by Kopylov e.a. [20] from gas saturation experiments are given (curve 2). This figure clearly shows that the maximal values of the fluoride distibution coefficients are not reached in an industrial evaporator. This is not surprising, because the molar F/Si ratio in the liquid is not equal to six and industrial evaporators are designed to evaporate water from W.P.A. and not to remove fluoride from W.P.A.. The values of the fluoride distribution coefficient calculated from the results of Kopylov e.a. [20] are significantly lower than

114 •

those obtained here. They could only detect SiF. in the gas phase and therefore the molar F/Si ratio in the solution probably increased with time in their experiments, although they started with a molar F/Si ratio of six in the solution also. Another explanation can be that not all fluoride present in the gas phase was detected in their experiments, due to the precipitation of part of the evaded fluoride in the alkaline solutions used for the absorption of the fluoride from the gas phase.

The influence of the temperature on the fluoride distribution coefficients is illustrated in figure 3. This figure shows that the fluoride distribution coefficient is proportional to the reciprokal value ofthe absolute temperature. The temperature dependence of the fluoride distribution coefficient can thus be expressed by a Clausius-Clapeyron type relationship [3.18].

If the temperature dependence is accounted for by adding a C./T term to 5

equation (14), this equation can be rewritten as:

in H = | * in m H + - | * In m ^ g ^ ♦ C, ♦ O, * m ^ ♦

+ C3 * mH2SiF6 + S * "HJJSO, + S/T W)

The parameters C., derived from our experiments and optimised with a computer program are:

°1 = 32,i|' °2 = °-156' C3 = 1 - ° ' Cn = °'H57 and c5 = " 9 5 5°

The above equation with the given C-values was found to describe the experimental results in the following area:

- wï P 0 in the phosphoric acid solution between 30 and 50. - w$ H SO between 0 and 6. - w* H2SiF, below 1. - temperatures between 70 and 95 °C. - a molar F/Si ratio of 6.

115

exp. no. '

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 21 25 26 27 28 29 30 31 32 33 31 35 36

CPo°c 2 5

(w*)

35 38 10 10 13 17 52 55 17 17 40 10 38 38 47 17 47 17 17 10 10 10 10 10 10 17 17 17 17 47 47 47 47 40 40 40

CH„S0„ 2 4

(w$)

0 0 0 0 0 0 0 0 2 4 2 4 4 8 0 0 0 0 0 0 0 2 2 4 4 2 2 2 2 2 2 4 4 4 4 4

CF

(wj)

1.81 1;86 1.90 1 .92 1 .85 1.85 1.77 1.77 1.83 1.83 1.83 1.93 1.91 1.83 0.83 3.3 1.7 1.88 1.85 1.86 1.83 2.07 2.07 2.06 2.05 1 .90 2.03 2.00 3.82 4.22 4.22 2.10 2.10 4.06 1.06 1.11

T

(°C)

91 90 90 94 90 90 90 86 90 90 94 94 90 90 90 90 90 73 83 70 81 70 80 70 80 70 80 90 70 80 90 70 80 70 80 90

V P O , 3 4

, mole . lkg"H~5 J

9.99 11.84 13-291 13.30 15.81 20.19 28.18 35.46 21.48 22.97 13.91 11.69 13.02 11.37 19.125 21.383 23.678 20.212 20.188 13.275 13.263 11.026 11.026 14.753 14.749 21.543 21.657 21.633 23.390 23.815 23.815 23.248 23.248 15.799 15.799 15.841

"HOSO,, 2 4

. mole .

0 0 0 0 0 0 0 0 0.662 1.416 0.504 1.064 0.990 2.192 0 0 0 0 0 0 0 0.508 0.508 1.069 1 .069 0.664 0.668 0.667 0.721 0.734 0.734 1.434 1 .434 1.145 1 .145 1.148

VsiR 2 6

, mole .

0.327 0.361 0.393 0.398 0.424 0.495 0.597 0.711 0.521 0.557 0.387 0.441 0.111 0.131 0.211 0.935 1.412 0.504 0.195 0.385 0.378 0.453 0.153 0.173 0.171 0.512 0.582 0.573 1.184 1.332 1.332 0.647 0.647 0.999 0.999 1.020

-4 H.10

. mg F/m3 vap_our ^inoïe H^SÏFg/kg H^>'

2.51 2.77 2.46 3.02 7.31

16.16 43.55

134.34 31.7 57.5

6.70 9.80 7.20

12.0 14.0 . 19.0 28.0

5.23 9.72 0.726 1.31 1.01 2.10 1.29 3.14 6.29

12.98 27.20

9.37 18.94 47.74

9.20 16.15

2.78 7.43

12.96

Table 1: Fluoride distribution coefficients as a function of various concentrations and temperatures.

116

}Q°-\

io5H

10"

io3H

10'

[mg F/m 3 vapour"!

%~F -I

Temperature 90°C

10 20 —f— 40 30 50

- * - w%P205

60

Figure 2: The fluoride distr ibution coefficient as a function of the phosphoric acid concentration.

12H

1H

10H

• 40w%P205,Ow%H2SO((

X40w%P205.2w%H2S0i>

a it0w%P205.4w% H2S0i(

I I Ü 2.8

1 2.9

- ^ ! x10 3

Figure 3: The fluoride dis t r ibut ion coefficient as a function of the temperature.

117

ggp. T m H H-raeas. * 10

J ^kg~H~0'' lmöIe~H SiF 6 /kg -H 2Ö ;

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36

94 90 90 94 90 90 90 86 90 90 94 94 90 90 90 90 90 73 83 70 81 70 80 70 80 70 80 90 70 80 90 70 80 70 80 90

4.44 5.528 6.391 6.406 7.841

10.40 15.00 19.24 11;82 13-47 7.224 8.322 7.274 9.254 9.413

11.93 14.15 10.43 10.40 6.365 6.364 7.425 7.425 8.426 8.418 11.90 12.05 12.01 14.26 14.80 14.80 13.82 13-82 10.13 10.13 10.20

2.510 2.770 2.460 3.020 7.310

16.16 43.55

134.3 31.70 57.50

6.700 9.800 7.200

12.00 14.00 19.00 28.00

5.227 9.719 0.726 1.307 1 .015 2.103 1.291 3.137 6.293

12.98 27.20

9.367 18.94 47.74

9.198 16.15

2.778 7.425

12.96

Table 2: Experimental and ca l cu la t ed

c o e f f i c i e n t .

mg F/m^ vap_our_ , möTe~H2SiF"7kg H Ö'

2.246 2.521 3.396 4.522 5.650

13-06 56.53 54.1 23-34 44.92

6.834 10.60 5.504

13.67 14.32 17.53 34.40

3-596 7.784 0.734 1.744 1 .107 2.434 1 .727 3.797 5.087

11.42 23.94

8.876 23.07 48.59 10.28 22.62

2.447 5.383

11.56

- 1 0 . 5 - 9 . 0 38.1 49.8

-22 .7 - 1 9 . 2

29.8 14.7

-26 .4 -21 .9

2.0 8.1

-23 .6 13.9 2.3

-7 .7 22.9

-31 .2 - 1 9 . 9

1.1 33.4

9.0 15.7 33.8 21 .0

-19 .2 -12 .0 - 1 2 . 0

- 5 . 2 21.8

1.8 11.8 40.0

-11.9 -27 .5 -10 .8

values for the fluoride distr ibution

118

In Table 2 the values of the fluoride distribution coefficient calculated from the measured data and indicated by H-meas. are compared with the H-cal values predicted by equation (46). The derived expression fits the H-meas. values reasonably well. Considering the high sensitivity of this kind of experiments for stoppages and leakages an error of about 10-20? is unavoidable.

It turned out to be important to check whether leakages occur and whether sufficiently small bubbles are formed in the liquid at the bottom of the wash bottles. After a while the fingered filters are severely attacked and large bubbles are formed. Then the fluoride content of the nitrogen gas might not be in equilibrium with the fluoride content of the liquid phase.

Stoppages are mainly caused by the unwanted formation of SiO_, when evaded SiF. reacts with condensed water [5,38,40] according to:

3 SlFy + 2 H20 + Si02 + 2 H^iFg (47)

5 SiF„ + 2 H„0 -> Si0o + 2 H.SiF,.SiF„ (48)

In addition to the experiments with chemically pure acids, also experiments were performed with W.P.A. obtained from a single stage HH/DH (Nissan H) process to determine the influence of impurities on the fluoride distribution coefficients. The measured fluoride distribution coefficients as well as the values predicted by eq. (16) are given in table 3.

P 20 5 H ^ Si F m m m m H F F/Si H-exp H-cal 3 1 2 1 2 O ,

(wï) (wt) (wï) f„ï)(-m2i§-w_m2iË.w_!!!°ie_w_!nole_.mole , __mg_F/m__vapour W ; KW W ; l W % M kg H 2 0 M kg H 2 0 H kg H 2 0 K k g H 2 0 , r a " ° (rnole H2SiF6 /kg H20 )

28.0 3 6 . t 45.2 50.6 54.2

2.0 2.5 3.0 3.4 3.7

0.236 0.249 0.1 24 0.013 0.017

1.29 1.24 0.72 0.49 0.44

6.824 11.209 18.864 27.170 36.333

0.353 0.558 0.907 1 .167 1 .797

0.1458 0.1945 0.1313 0.0177 0.029

0.300 0.260 0.335 0.877 0.931

8 7 9

55 38

2.297 3.285

21.515 293-592 212.705

1.705 4.442

29.180 513.635

2575.171

Table 3: F luor ide d i s t r i b u t i o n c o e f f i c i e n t s for W.P.A. a t 90 °C.

119

The values of the fluoride distribution coefficient predicted by equation (16) for concentrated phosphoric acid (> 50 wj P2°c:)> having high molar F/Si ratios, differ largely from the experimental values. In this case the fluoride is not only present as H„SiF, but also as HF and therefore the equations (3) and (9) are no longer valid.

For the prediction of the fluoride distribution coefficient of W.P.A. during concentration from 30 to above 50 w$ P-,0,., when the molar F/Si ratio changes from about 6 to approximately 60, the model presented here is too simple and leads to erroneous results for phosphoric acid solutions containing a molar F/Si ratio above ten. This again stresses the importance of the addition of sufficient SiO if a high fluoride evasion were pursued.

7.7. Conclusions

- Stripping and flashing seem to be the best methodes for the removal of fluoride from phosphoric acid production processes if the fluoride has to be recovered and is used as a new fluoride source, which can substitute for fluorite.

- If almost all fluoride must be removed from the phosphoric acid pure as well as W.P.A. a molar F/Si ratio of about six should be maintained in the solution to obtain the highest fluoride distribution coefficients.

- From a theoretical model an equation is derived describing the fluoride distribution coefficients between phosphoric acid solutions and the gas phase as a function of various process parameters. This equation was found to describe the measured values with reasonable accuracy between 70 and 90 °C, for chemically pure as well as for commercial phosphoric acid solutions containing 30-50 w$ P 0 , up to 6 w$ H SO., .maximal 4 w% H„SiF, and a molar F/Si ratio below 10.

- The model used for the derivation of an expression for the fluoride distribution coefficient is too simple to predict fluoride distribution coefficients for W.P.A. during concentration, when the molar F/Si ratio in the liquid phase changes from about 10 to approximately 60.

- The simplified model cannot predict the molar F/Si ratio in the gas phase.

120

7 .8 . Nomenclature

A, Debye-Huckel c o e f f i c i e n t 9

b thermodynamic parameter B binary interaction parameter c concentration C ternary interaction parameter h Henry coefficient H fluoride distribution coefficient I K m M P T z a B Y P e * V

ion strength equilibrium constant molality molar mass vapour pressure absolute temperature ion charge number degree of dissociation thermodynamic parameter activity coefficient density binary interaction parameter ternary interaction parameter volume flow

subscripts:

a,a' c,c' g 1 1

supers

t

anions cations gas phase component liquid phase

cript:

derivative

mol/1

mg F/m^ vap_our möl"H~SiF77kg"H~Ö 2 0 2

mole/kg HO g/mole

Pa K

kg/mJ

m3/hr

121

7.9. Literature

1. Baur, E., Zeitschrift fur Phys. Chera., 48 (1984) 483. 2. Becker, P., Phosphates and Phosphoric Acid, Fertiliser Science and

Technology Series, Vol. 3 (1983), Marcel Dekker Inc., New York. 3. Borisov, V.M. and Mel'nikova, S.V., Zh. Prikl. Khira., 57, 3 (1984) 705. 4. Buskus, H., Chemisch Magazine, Nov. (1985) 735. 5. Charleston, A.G., New Zealand Journal of Science, 27, 3 (1984) 279. 6. Chemtob, E.M. and Beer, G.L., US Patent 4.505.884 (1982). 7. Drechsel, E.K., Chem. Eng. News, 24 (1979) 37. 8. Elmore, K.L., Hatfield, J.D., Dunn, R.L. and Jones, A.D.

J. Phys. Chera., 69, 10 (1968) 3520. 9. Elmore, K.L., Mason, CM. and Christensen, J.H.,

J. Am. Chem. Soc, 68 (1946) 2528. 10. Eyal, A. and Baniel, A., Solvent Extraction and Ion Exchange, 2 (1984) 677. 11. Eyal, A., Hajdu, K. and Baniel, A.,

Solvent Extraction and Ion Exchange, 2 (1984) 659. 12. Flemmert, G.L., Proceedings of the Fertiliser Society of London,

no 163 (1977). 13. Fokin, M.N., Petrovskaya, V.A., Dobrolyubov, V.V., Kurteva, 0.1.,

Nikolaeva, G.N. and Grabyl'nikova, V.I., Khim. Prom., 9 (1984) 543. 14. Frazier, A.W., Lehr, J.R. and Dillard, E.F.,

Environ. Sci. Technol., 11, 10 (1977) 1007. 15. Getting rid of Phosphogypsum 2,

Phosphorus and Potassium, 89 (1977) 36. 16. Hodges, W.A., US Patent 2.917.367 (1959). 17. Homann, J., Werkstoffe und Korrosion, 37 (.1986) 532. 18. Illarionov, V.V., Smirnova, Z.G. and Knyazeva, K.P.,

Zh. Prikl. Khim., 36, 2 (1963) 237. 19. Kitchen, D. and Skinner, W.J., J. Appl. Chem. Biotechnol., 21 (1971) 53, 65. 20. Kopylov, V.A., Senetova, G.I. and Pozin, M.E.,

Zh. Prikl. Khim., 48, 11 (1975) 2455. 21. Kopylov, V.A., Senetova, G.I. and Pozin, M.E.,

Zh. Prikl. Khim., 49, 5 (1976) 1149. 22. Lehr, J.R., Proceedings of the Environmental Symposium of the

Fertiliser Institute, New Orleans, March 6-8, (1978) 277.

122

23. Miller, F.D. and Biggers, E.D., US Patent 3.193-351 (1965). 24. Ministerie van Economische Zaken, Directie R. en 0.,

Technieuws Washington W-84-09, Technische Keramiek, 's Gravenhage (1984). 25. Nabiev, M.N., Gafurov, K. and Korosteleva, V.I.,

Khim. Prom., 6 (1983) 318. 26. Nassif, N., Surf. Technol., 26, 3 (1985) 189. 27. Odintsova, G.S., Guller, B.D., Zhdanova, M.V. and Zinyuk, R.Yu.,

Zn. Prikl. Khim., 54, 1 (1981) 37. 28. Phosphoric Acid, Outline of the Industry,

British Sulphur Corporation Ltd., London, (1984). 29. Pitzer, K.S., Am. Chem. Soc. Symp. Ser., 123,23 (1980) 451. 30. Pitzer, K.S., J. Phys. Chem., 77, 2 (1973) 268. 31. Rushton, W.E., Che. Eng. Progr., 74, 11 (1978) 52. 32. Rysselberghe, P. von, J. Phys. Chem., 39 (1935) 403. 33. Senetova, G.I., Kopylov, V.A. and Pozin, M.E.,

Zh. Prikl. Khim., 49, 6 (1976) 1371. 34. Senetova, G.I., Novikov, A.A., Kopylov, V.A., Khoipunov, N.F.,

Repenkova, T.G. and Pozin, M.E., Zh. Prikl. Khim., 49, 6 (1976) 1374. 35. Shape, J.L., Miner, M.L., Harris, L.E. and Greenwood, D.A.,

Am. J. Vet. Res., 23, 95 (1962) 777. 36. Shr.amban, B.I., Pavlukhina, L.D., Afonina, N.D. and Gorskaya, A.P.,

Zh. Prikl. Khim., 52, 10 (1979) 2344. 37. Spijker, R., UKF Internal report no. YB0-RP-81-208 (1981). 38. Spijker, R., UKF Internal report no. YB0-RP-83-248 (1983). 39. Teslenko, V.V. and Rakov, E.G., Khim. Prom., 12 (1981) 744. 40. UKF, Private communications. 41. Wertz, D.L. and Cook, G.A., J. Sol. Chem., 14, 1 (1985) 41. 42. Weterings, K., The Utilisation of Phosphogypsum,

Proc. of the Fertiliser Society of London no. 208 (1982).

123

7.10 Appendix: Conversion of w% to molalities

For conversion of w$ to molalities for a phosphoric acid solution containing only chemically pure phosphoric acid, sulphuric acid and fluorosilicic acid the following procedure is used:

2 * x x wj P„0_ = Ü mole H-PO.. per 100 grams of solution. (1) 25 . y ^ 3 *

y w$ H SO. = --^ mole H SO. per 100 grams of solution d H

2S0n

(2)

1 z z w£ F = r * r.- mole H„SiF, per 100 grams of solution (3) b M„ d 0

The water content of 100 grams of solution, w is then given by:

luu d M^ ^ H^O^ y 6 MF H2SiF6 (H>

2 5

and the molalities are calculated with the following equations:

x w$ P 0 = 2 * M"--- / w mole H PO^ per kg H20 P2°5

(5)

y w? H SO. = - - x / w mole H SO. per kg H?0 (6)

z w * F = j r * - ; - / w mole H.SiF, per kg H.0 (7) D M d O d r

12H

8. MASS AND HEAT BALANCES OF THE CTPA PROCESS

8.1. Introduction

Phosphoric acid, used for the manufacturing of fertilisers is mainly produced by digestion of phosphate ore with sulphuric acid, the so-called "wet-process". The production of phosphoric acid is expected to rise from about 20 million tons P„0._ in 1982 to about 28 million tons in 1990 [26].

c. D

In all wet processes, impurities such as cadmium, radium and fluoride ions, originating from the phosphate ore (Ca Q(P0.)gF + a few w? CaCO and quartz), are distributed between the phosphoric acid and the byproduct, a calcium sulphate modification. The use of the byproduct is limited by its phosphate, radium and fluoride content, while the disposal is hampered by its cadmium content [1]. The phosphate reduces the strength of the byproduct, while the fluoride reduces its hardening time [26]. Fluorides can enter the atmosphere via the production process too. They are regarded as serious pollutants [26]. The decay of radium yields the radio-active radon gas. The most problematic contaminant is, however, cadmium. It is very toxic and can enter the food chain, both via the byproduct and the phosphoric acid. Too high cadmium contents might even prohibit the use of phosphate fertilisers in the future.

In most processes the digestion of phosphate ore and the crystallisation of CaSO^HgO (DH) or CaSO^-rH-O (HH) take place more or less simultaneously in the same reactors [1], The easiest way to remove cadmium therefore seems to be from the ore, before it is introduced in the process. One method to achieve this purpose is calcination of the ore at about 700 °C, the cadmium then evades as CdO. This process, however, requires about 20 MW for a 1000 tons P_0_ per day producing plant. Moreover, the solubility of phosphate ore is reduced by calcination, while the fluoride and radium ions remain in the ore. Another method is selective leaching of cadmium from the carbonate fraction of the phosphate ore, which sometimes contains up to 50 \*% of the cadmium, by means of digestion of the carbonate containing part of the ore with diluted sulphuric acid with a high chloride content [9,31]. The chloride ions keep the cadmium as cadmium-chloride complexes In the solution and thus prevent the precipitation of cadmium either as cadmium sulphate or phosphate. Incorporation of cadmium in the calcium sulphate modification produced can therefore also be reduced by feeding complexing agents, such as halogen!des to the process liquid. These methods, however, only appear to be satisfactory if the cadmium content of the ore is not

125

too high. Furthermore the cadmium still has to be removed from the phosphoric acid or the leaching solution. In all cases the radium remains in the "phosphogypsum".

A completely different method is based on the separation of the digestion and the crystallisation stages, in order to remove the cadmium ions in between. In this process the possibility is created to optimise the digestion and crystallisation stages independently from each other. Moreover, the radium can be removed before the majority of the calcium ions from the phosphate ore are precipitated as HH. A disadvantage, however, is the large recycle stream of phosphoric acid needed for complete digestion of the phosphate ore.

The aim of this new process, referred to as a Clean Technology Phosphoric Acid (CTPA) process is the production of commercially competitive concentrated phosphoric acid (40 w$ P-0_) with a low cadmium and fluoride content as well as the production of calcium sulphate hemihydrate (HH) with a low cadmium, phosphate and radium content.

This study considers the mass and heat balance of the process. It is the first step towards the engineering flow diagram of a production plant, which can be used to compare this process with existing processes.

8.2. Process description

8.2.1. General approach

To meet the foregoing requirements the digestion of the phosphate ore and the crystallisation of the HH are separated in two almost independent stages. A simplified flowsheet of the process is shown in figure 1.

First the digestion of phosphate ore proceeds in recycled phosphoric acid. This acid contains sulphate ions in a concentration dictated by the operating conditions in the crystalliser. Part of the calcium ions from the ore will therefore precipitate as HH. Since more then 2 w? P„0_ is incorporated in this HH, it cannot be disposed, without lowering the overall phosphate efficiency of the process too much. This HH is thus recrystallised into DH to reduce the phosphate content of the crystals to an acceptable low level (< 0.5 w% P 2

0c)-Before the cadmium ions can be removed, the precipitated HH from the

digestion stage, together with the insoluble residue of the ore, have to be separated from the slurry to avoid problems in the cadmium removal stage.

126

The cadmium can be removed from the almost clear calcium-di-hydrogen-phosphate (CDHP) solution by anion exchange of cadmium halogenide complexes formed by addition of either bromide or iodide ions [33].

1 digestion

H 2 S<H

HH crystallisation

recrystallisahon H H - — D H DH

recycle add product

Figure 1: Simplified process flowsheet.

The CDHP solution with a strongly reduced cadmium content is then fed to the crystalliser, together with the recrystallisation liquid and concentrated sulphuric acid. In this crystalliser the calcium ions from the CDHP solution are precipitated as HH. . Thereafter the HH is separated by filtration, yielding concentrated phosphoric acid. About 90 w% of this phosphoric acid is recycled to the digestion stage.

The product acid is fed to the iodide or bromide removal stage and finally to the fluoride removal stage. After the removal of the iodide or bromide and the reduction of the fluoride level, concentrated phosphoric acid (10 w$ PjCv) with a low cadmium and fluoride content is obtained.

8.2.2. Digestion of phosphate ore

The digestion stage is proposed to consist of three equally sized stirred tanks in series. The phosphate ore is introduced in the first tank, together with the recycled phosphoric acid and some clay. The clay is used to stimulate the evasion of fluoride. It is the easiest way to add it here, because it can be mixed with the phosphate ore in advance. The quantity of clay (15 w? SiCv and 38 »% Al 0.) is determined by the additional amount of active silica needed, above the amount already present in the ore, to obtain a molar F/Si ratio of six in the acid.

The composition of the phophate ore used (10 w$ Khouribga, 60 w$ Zin) is given in table 1.

The phosphoric acid must have a concentration of at least 10 w% P 0. for direct use in the production of mono- or dl-ammonium-phosphate (MAP or DAP) [1].

127

The temperature in the d iges t ion i s l i m i t e d t o 95 °C to avoid format ion of .0H2O (AH).

The main r e a c t i o n occuring under these cond i t ions i s :

CaSO4.0H2O (AH)

C a 1 0 ( P V 6 F 2 + U H3

P °4 F l u o r o a p a t i t e

component

10 Ca(H2P04)2

CDHP 2 HF

concentration | component concentration

moisture CaO P2°5 so4 F A1203

Fe2°3 MgO co2

1 50.8 31.55 2.4 4.2 0.36 0.17 0.33 5.9

w% wj wjt vf w? w% wj w* w*

Cd Cl Na K20 Si02 (total) Si02 (active) Hg Pb Organic C

21 380 0.56 0.08 1.35 0.35 5 2 0.3

mg/kg mg/kg wjt wj wj wj ug/kg mg/kg w%

Table 1: Composition of the phosphate ore (Khouribga/Zin 40/60 w$),

The rate of the digestion is found to be determined by diffusion of calcium ions from the surface of the ore into the bulk of the solution. The maximum solubility of CDHP in phosphoric acid can be expressed as 5.5 wj CaO in 40 w? P20 at about 95 °C [1]. To maintain a safety margin 90 °C, 40 wj P20? and only 5 w$ CaO were chosen as the operating conditions.

The amount of recycled acid needed, is determined by the w% CaO in the CDHP solution after the digestion stage and the sulphate content of the recycled phosphoric acid. For a sulphate content of the recycle stream below 2 wj H SO., as required to avoid blinding of the ore particles, more than 99 w$ of the soluble part of the ore (maximum particle diameter 2 mm) is digested within 90 minutes.

The carbonate present in the ore evades as carbon dioxide, saturated with fluoride and water vapour. Due to the presence of organic carbon in the ore, the evasion of carbon dioxide causes excessive foaming. The fluoride in the gas phase is a mixture of hydrogen fluoride and silicon-tetra-fluoride with a mean molar F/Si ratio of six.

128

The precipitated HH contains cocrystallised PpCL, Al, Cd and fluoride. 1 1

These species are assumed to be incorporated as CaHPO^.^H 0, CaAlF .-H_0 and cadmium ions.

It is furthermore assumed that 50 w$ of the fluoride originating from the ore is either cocrystallised with the HH or precipitated as sodium-fluorosilicate (Na.SiF,-). Although this fluorosilicate precipitate will partly be deposited as scale, it is assumed, that it leaves the digestion stage with the HH and the insoluble residue of the phosphate ore, which is mainly quartz. These solids are separated from the CDHP solution by a gravity settler, followed by a decanter centrifuge. Because of the low permeability of the cake obtained after pressure filtration of this slurry, filtration is not a realistic option.

The almost clear CDHP solution is introduced in the cadmium removal stage, while the slurry, containing 50 w$ of solids, is fed to the recrystallisation stage.

The heat of digestion of the phosphate ore, the heat of crystallisation of HH, the heat losses through the vessel walls and the energy dissipation of the turbine stirrers all contribute to the overall heat balance of the digestion stage. The enthalpies of the incoming and outgoing streams are calculated relative to their values at 0 °C with the specific heats given in the appendices. The heat of evaporation of water is taken into account in the enthalpy of the water vapour stream.

8.2.3- Cadmium removal

The cadmium removal is assumed to be done in a tubular co-current adsorption. The feed consists of the CDHP solution, a slurry of the strongly basic anion exchange resin in a CDHP solution and iodide ions as complexing agents for cadmium. After passing the adsorber, the anion exchange resin is separated from the solution and almost completely recycled to the entrance of the adsorber. Part of the anion exchange resin is regenerated with water.

The excess iodide is circulating through the whole process and is removed from the product acid, by passing this acid through packed columns of a weakly basic anion exchange resin. This last anion exchange resin is regenerated with concentrated ammonia, after washing of the resin with water to remove the phosphoric acid.

129

To prevent oxidation of iodide during its passage through the whole process, an adapted amount of iron is added in the digestion stage.

The cadmium and iodide removal stages as well as the amount of iron needed, to prevent oxidation of iodide will be reported elsewhere [34]. These stages are not yet included in the mass and heat balances in detail. They can be treated, however, as black boxes, with very little exchange of mass and heat.

Only the amount of water needed for the regeneration of the strongly basic anion exchange resin and for the washing of the weakly basic anion exchange resin is roughly estimated. This water comes from the fluoride removal stage and will be used in the second step of the HH filter cake washing stage, after removal of cadmium from this stream, by either electrolysis or precipitation as a sulphide.

8.2.4. Crystallisation of calcium sulphate hemihydrate (HH)

The main reaction in this stage is:

2 Ca(H2P01))2 + 2 H2S04 + H20 — > 2 CaSO^.^O + 4 H PO,,

This reaction requires a single stirred tank in which the CDHP solution, the concentrated sulphuric acid and the liquid from the recrystallisatlon stage are introduced simultaneously.

To obtain HH crystals with a good filterability the sulphuric acid concentration in the crystalliser must be kept constant at about 1.7 wj H.SOj., while a residence time of about 30 minutes is recommended [26].

The precipitated HH contains cocrystallised P20c> A1i c d a n d fluoride in an amount which can be calculated form the correlations and data given in the appendices for all these ions.

The most important heat effect in this stage is the heat of crystallisation of HH. The heat of mixing of the acids is quite low, because the major part of the sulphuric acid is almost directly precipitated as HH. The temperature in the crystalliser is maintained at about 90 °C, to prevent formation of AH. The excess heat is removed by evaporation of water in a vacuum slurry cooler. The enthalpies of the streams are calculated, relative to their values at 0 °C, with the heat capacities, given in the appendices.

130

8.2.5. Recrystallisation of HH into gypsum

The HH formed during digestion of the ore is recrystallised into DH under conditions, similar to those prevailing in commercially known HH to DH recrystallisation processes, i.e. at about 60 °C in a solution containing 27 -30 w* P20 and about 5 w$ H^O^.

The recrystallisation stage is assumed to consist of three equally sized stirred tanks in series. The slurry from the digestion stage, containing about 50 wj CDHP solution, is fed to the first tank, together with the wash water from the HH and the DH filters, concentrated sulphuric acid and half of the DH slurry stream from the third recrystalliser. This slurry stream is recycled to supply enough DH seed material to ensure a high recrystallisation rate of HH into DH. A residence time of about five hours should be sufficient for complete conversion of HH into DH [26].

The produced DH also contains cocrystallised Po0_, Cd, A1 and fluoride in an amount, which can be obtained from the data and correlations given in the appendices. The phosphate, Al, and fluoride are assumed to be incorporated as CaHP0^.2H20 and CaALF .2H20 respectively.

The main heat effects are caused by the heat of recrystallisation of HH into DH, the heat of mixing of sulphuric acid with water and phosphoric acid. To maintain a temperature of about 60 °C in the first recrystalliser, water must be evaporated by vacuum slurry cooling. The enthalpies of the streams are calculated, relative to their values at 0 °C, with the heat capacities given in the appendices.

8.2.6. Solid-liquid separations

A belt filter is used for the filtration of both the HH slurry and the DH slurry. The HH as well as the DH crystals are washed three times in counterflow with water. The wash water temperature for the HH washing must be at least 60 °C to avoid recrystallisation into DH on the filter. This would reduce the permeability of the filter cake. The wash water of the DH crystals is assumed to be about 15 °C. The HH as well as the DH crystal cake contain aproximately 30 w$ liquid. Neither HH nor DH is assumed to dissolve in the wash water. The HH and DH crystal cake contain 30 w? pure water after the third wash step. This implies negligible P„0_ losses due to washing. A small amount of water evaporates during filtration of the DH and HH slurry.

131

The total amount of water, which can be used as wash water in this process is not only determined by the required concentration of the product acid (to w$ P 0 ) and the sulphuric acid concentration, but also by the amount of water needed for the regeneration and washing of the anion exchange resins. The washing procedures of the HH and DH cakes are assumed to proceed isothermal.

8.2.7. Fluoride removal

After the iodide removal, the fluoride level in the product acid is reduced to about 0.5 w? F in a packed stripping column. Stripping is achieved with a mixture of carbon dioxide and water vapour of about 90 °C. The stripping gas is, together with the off gasses from the digestion stage, fed to a crossflow scrubber, in which part of the fluoride is removed from the gas mixture through countercurrent washing with water. A 25 w$ H_SiF, solution is produced in this scrubber. The largest part of the gas mixture is recycled to the stripping column. The residual gas mixture is disposed, after passing a second scrubber to meet environmental requirements.

The water, to be applied for the HH cake washing, is used in the second scrubber as a scrubbing medium. This scrubber operates at room temperature. The temperature of this wash water is further raised by using it in the direct contact condensor, in which the water vapour from the recrystallisation stage is condensed. Thereafter the temperature is raised to about 60 °C, by applying this wash water for direct contact condensation of the water vapour from the crystalliser.

Since the heat capacities of the streams in the fluoride removal stage are not well known, only a rough estimation of this stage is included in the mass and heat balances.

8.3. Combined mass and heat balances for each stage

The balance equations needed to calculate the overall mass and heat balances of the complete CTPA process are for each stage: - the overall mass balance - the mass balances for the four components CaO, SO,, P20_ and water - the overall heat balance

The fixed operating conditions were: - 40 w$ P„0C in the product acid - 5 wj CaO in the CDHP solution

132

- 1 J 4 H2SOn in tne cystalliser - 5 w$ H SO. in the recrystallisers

If the Incoming and outgoing streams of a process stage are known, the overall balance equations of each stage can be written. The mass balances for the components CaO and SO. for each stage, can only be applied to solve the overall mass and heat balances of the CTPA process, if the solubility products of HH and DH (appendix 8.6.1) and the compositions of HH and DH are given.

The P„0_ mass balance of each stage can be used to solve the total mass and heat balances, if accurate data on the incorporation of phosphate in the HH and DH are known. These data were either measured (see chapter 5) or obtained from literature and are given in appendix 8.6.2.

The composition of the sulphuric acid as well as of the filter cakes and accurate data on the vapour pressure of water above the process streams are needed to be able to use the water balance of each stage. These vapour pressure data are obtained from literature and given in appendix 8.6.3.

The heat balance of each stage requires the enthalpies of reaction, which are given in appendix 8.6.7, and the heat capacities or heat contents of the process streams, which are given in appendix 8.6.6. Information about the heats of mixing of the acids with each other and with water are also needed and are therefore given in appendix 8.6.8.

For the less important components, from the viewpoint of process control, distribution coefficients are used, which for cadmium, aluminum and fluoride are given in appendix 8.6.2 and appendix 8.6.4.

The complete set of mass and heat balance equations can be solved simultaneously with the software of the flowsheeting simulation package TISFLO-2, developed by DSM. The number of equations must of course be equal to the number of unknown parameters. In this way the overall mass and heat balances of the complete CTPA process can be obtained as well as the composition of the process streams as shown in appendix 8.6.9. From the results given in appendix 8.6.9 a mass and a heat balance of each stage can be deduced.

In the following part of this section the overall mass and heat balances of each stage are given and discussed. First of all the digestion stage will be dealt with. The incoming and outgoing process streams of the digestion stage are given in figure 2. In this figure the most important process conditions are given also.

133

The mass and heat balances of this stage are given in table 2a and 2b respectivily.

It is important to notice that the solids content of the slurry, to be fed to the recrystallisation stage, is a critical parameter in this stage. The phosphate concentration in the recrystallisers will raise above 30 w% P20c> if the solid content of this slurry is less than 50 w$. This leads to a reduced conversion rate of HH into DH. As a consequence the recrystallisation of HH into DH will not be completed before the filtration. This can cause a considerable reduction in the overall phosphate efficiency of the whole CTPA process.

15°C ore c t ay ,

L: vapour (PH Q =0,5 bar)

5 w % C a 0 40.6 w% P 20 5

92.3 °C

recycle phosphoric acid 90.8 °C

Figure 2: The digestion stage.

re-

\ \

\

HH ore ' scale

CDHP inert

CDHP solution

component in

ore i n e r t clay r ecyc le a c i d

36.7 0.4 0.8

288.0

componen

vapour

t out

CDHP s o l u t i o n HH i n e r t ore s c a l e CDHP

3.1 300.9

10.1 0 . 1

0 .3 0 . 2

10.9

t o t a l 325.9 t o t a l 325.9

Table 2a: Mass balance of the d ige s t i on s tage (mass s t reams in k g / s ) .

134

The heat of digestion of the phosphate ore is the second parameter, which has to be discussed. The temperature in this stage can be higher in reality, because the heat of digestion is not accurately known. A twice as high value as used now will raise the temperature of the CDHP solution, however, at most 2 °C, due to the large heat content of the recycle acid stream.

component in

ore, inert and clay recycle acid heat of digestion heat of crystallisation

0.1* 67.2 3.6

2.0

component out

vapour CDHP solution HH, inert, ore and scale CDHP heat losses

2.8 66.7

0.9 2.it

0.»l

total 73.2 total 73.2

Table 2b: Heat balance of the digestion stage (enthalpies in MW).

After the digestion stage, the almost clear CDHP solution is fed to the cadmium removal stage, which will be reported elsewhere [34], as was already mentioned. This stage is regarded as a black box with very little exchange of mass and heat. The most important incoming and outgoing streams are shown in figure 3, while an estimation of the mass and heat balances is given in table 3a and 3b respectively.

w a t e r

CDHP

92.3°C

60°C

Cd removal

646 mg Cd/s

60°C " water

CDHP 92.3°C

Figure 3: The cadmium removal stage.

135

The water needed for the regeneration of the strongly basic anion exchange resin and the washing of the weakly basic anion exchange resin, was roughly estimated to be about 2.5 kg per 300 kg of CDHP solution. This number is given only, to be able to deal with the complete process. The amount of water needed is not yet known accurately.

component in component out

CDHP solution 300.9 CDHP solution 300.6 water 2.5 water 2.8

total 303.1 total 303.4

Table 3a: Mass balance of the cadmium removal stage (mass streams in kg/s).

The water needed for the regeneration and washing of the resins is part of the wash water for the HH filter cake. After its use, the cadmium is removed from the water by electrolysis or by precipitation as a sulphide. Thereafter it is introduced into the process again in the second wash step of the HH filer cake washing stage, because it will contain an appreciable amount of phosphate.

component in

CDHP solution 66.7 water 0.6

component out

CDHP solution water

66.7 0.6

total 67.3 total 67.3

Table 3b: Heat balance of the cadmium removal stage (enthalpies in MW).

After the cadmium removal stage, the CDHP solution is introduced into the crystalliser, together with the solution from the recrystalliser and the concentrated sulphuric acid. Figure 4 shows the incoming and outgoing process streams of the crystalliser and the most important operating conditions. The mass and heat balances are given in the tables 4a and 4b.

The most important process parameter in this stage is the sulphate content of the solution in the crystalliser, because this parameter determines the

136

permeability of the HH cake, obtained after filtration of the slurry. The temperature of the CDHP solution, together with the heat of

crystallisation of HH mainly determines the temperature in the crystalliser, which should not exceed about 92 °C, to avoid AH formation.

92.3°C CDHP

15°C 98% H2S04

59.9°C recr.acid

40w% P205

1.7w% H2S04

90.8°C

vapour (87°C.PH Q =0, ; bar)

M phosphoric acid

Figure H: The crystallisation stage.

The excess heat can be removed by evaporation of water in a vacuum slurry cooler. As was already mentioned, the heat of digestion can be higher than the value used in this mass and heat balance. As a consequence, the temperature of the CDHP solution will be higher and more heat has to be removed in this stage.

component in

CDHP so lu t ion 98 wj H^Ojj r e c r y s t a l l i s a t i o n acid

300 28

Hi

.6 8

2

t o t a l

component out

phosphoric acid 333.0 vapour 0. H HH 37.2

370.6 total 370.6

Table Ha: Mass balance of the c r y s t a l l i s a t i o n s tage (mass s treams in k g / s ) .

This heat , however, can adequately be used for a fu r the r r a i s e of the temperature of the wash water for t he HH f i l t e r cake. The hea t removed here i s

137

already used for that purpose. This heating of the wash water will be discussed in the fluoride removal stage.

To avoid local primary nucleation of HH or even of AH at the entrance of the CDHP or the sulphuric acid solution in the crystalliser, the mixing of the contents of the crystalliser has to be thoroughly.

component in

CDHP solution 98 w* H2S0^ recrystallisation acid heat of HH crystallisation

66.7 0.7

7.2

7.5

component out

phosphoric acid vapour HH heat losses heat of acid mixing

77.8 1.0 2.9 0.3

0.1

total 82.1 total 82.1

Table Mb: Heat balance of the crystallisation stage (enthalpies in MW).

The recrystallisation acid, mentioned in the tables Ha and 4b, comes from the recrystalliser, in which the HH from the digestion stage is recrystallised into DH.

The incoming and outgoing process streams of the recrystallisation stage are shown in figure 5. The mass and heat balances of this stage are given in the tables 5a and 5b.

92.3°C CDHP

ore + inert HH ♦ scale

15°C 98% H2S04

IM

5 w % H2S0;

28 w% P205

60°C

vapour (55.6°C. Pu^rr 0 . 1 4 b a r >

DH inert + scale

recrystallisation acid

wash water HH 76.3°C

wash water DH 36.5°C

Figure 5: The recrystallisation stage.

138

In this stage there are two important parameters, which have to be controlled. First of all, the solids content of the slurry coming from the digestion stage must remain equal to or must be higher than 50 w$. If the solids content is lower, more CDHP solution will be fed to the recrystallisatlon stage. This raises the phosphate content of the solution, and therefore reduces the recrystallisatlon rate of HH into DH and unacceptable phosphate losses can occur.

component in

CDHP solution 98 v% H2S04 wash water HH wash water DH HH, ore, scale inert

10.9 3.7 27.6 10.3

11.0

component out

recryst. acid vapour DH inert scale

illisation 17.5 0.4 11.0 0.4 0.2

total 63.5 total 63.5

Table 5a: Mass balance of the recrystallisatlon stage (mass streams in kg/s).

In this mass and heat balances, it is already necessary to feed the wash water of the DH as well as of the HH filter cake to the first recrystalliser, to obtain a phosphate content in the recrystalliser, which is low enough to allow recrystallisatlon of HH into DH in a reasonable time. A recycle stream of 50 w$ of the DH slurry from the third recrystalliser, is fed again into the first recrystalliser to achieve a fast conversion of HH into DH.

The second parameter is the heat removal, which is necessary to limit the temperature in the recrystallissation stage to a maximum of about 60 °C. Unless the phosphate content in the recrystallisatlon stage is chosen to be much lower than 30 v% Pp°c> tne temperature should not be higher, because then conversion of HH into DH will not take place, -as follows from the phase diagram of calcium sulphate ([27], chapter 5, figure 1).

The excess of heat has to be removed by evaporation of water in a vacuum slurry cooler, because the vapour pressure of water is very low under these conditions. This vapour pressure can be calculated from the correlations given in appendix 8.6.3.

139

Table 5b shows that the enthalpy of the wash water of the HH filter cake and the enthalpy of the slurry from the digestion stage, mainly determine the temperature in the recrystallisation stage. If the heat of digestion were higher in reality even more heat has to be removed in this stage.

component in

CDHP solution 98 wj H^Oy wash water HH wash water DH HH, ore, scale inert heat of mixing of acids and water heat of recrystallisation

2 0 6 1

0

1

1

4 1 4 1

9

0

.0

component out

recrystallisation acid vapour DH, scale, inert heat losses

8.3 3.5 0.9 0.2

total 12.9 total 12.9

Table 5b: Heat balance of the recrystallisation stage (enthalpies in MW).

req. water60°C . . , , o r * wash water 60°C

f wash water HH 76.3 °C recycle and product

phosphoric acid

Figure 6: The HH filtration and washing stage.

110

After the recrystallisation of HH into DH as well as after the crystallisation of HH, the crystal slurries have to be filtered. First the HH filtration and washing stage will be dealt with and thereafter the DH filtration and washing will be discussed. In figure 6 the incoming and outgoing streams of the HH filtration and washing stage are shown. The mass and heat balances are given in the tables 6a and 6b respectively.

component in

phosphoric acid HH wash water

333 37 24

0 2 8

component out

wash water from the cadmium removal stage 2.8

phosphoric acid HH water wash water HH

317.0 37.2 16.0 27.6

total 397.8 total 397.8

Table 6a: Mass balance of the HH filtration and washing stage (mass streams in kg/s).

component in component out

phosphoric acid 77.9 HH 2.9 wash water 6.2 wash water from the cadmium removal stage 0.6 heat of mixing of acid and water 0.5

total 88.0

phosphoric acid HH water wash water HH

7^.0 2.4 5.2 6.4

total 3.0

Table 6b: Heat balance of the HH filtration and washing stage (enthalpies in MW).

141

The heat losses on the filter are neglected. So the temperature of the wash water going to the recrystallisation stage is assumed to be equal to the temperature of the HH leaving the process. If this stage cannot be considered to behave isothermal, the amount of heat to be removed in the recrystallisation stage can change considerably.

The HH leaving the process contains 30 w? water. In practice this water might still contain an amount of phosphate, which cannot be neglected. The liquid content of the HH cake can be influenced by process disturbances. It might be necessary to adjust the amount of wash water used.

wash water 15°C

60 °C solids

recr. acid F

' 1

"

W

\ wash wat er DH 36.

solids water

recrystallisation acid

Figure 7: The DH filtration and washing stage.

36.5 °C

component in component out

recrystallisation acid 47.5 DH, inert, scale 14.6 wash water 10.3

total 72.4

r e c r y s t a l l i s a t i o n acid DH, i n e r t , s c a l e water

wash water DH

41 .2

14.6

6.3 10.3

total 72.4

Table 7a: Mass balance of the DH filtration and washing stage (mass streams in kg/s).

142

In figure 7 the incoming and outgoing process streams of the DH filtration and washing stage are shown. The mass and heat balances are given in the tables 7a and 7b.

component in component out

recrystallisation acid 8.3 DH, inert, scale 0.9 wash water 0.6 heat of mixing of acid and water 0.1

recrystallisation acid 7.2 DH, inert, scale 0.6 water 1.0 wash water DH 1.1

total 9.9 total 9.9

Table 7b: Heat balance of the DH filtration and washing stage (enthalpies in MW).

This stage is also considered to behave isothermal. The phosphate content of the water leaving the process, together with the DH, is neglected too. After washing of the DH filter cake, the wash water is introduced in the first recrystalliser.

The temperature of this wash water is set equal to the temperature of the DH leaving the process. This wash water stream, however, will hardly affect the temperature in the recrystallisation stage.

The last step of the CTPA process is the fluoride removal. This stage Is discussed in two parts. The first part deals with the reduction of the fluoride level in the product acid, while in the second part the heating of the wash water of the HH filter cake with the fluoride containing off gasses, Is discussed.

Figure 8 shows the incoming and outgoing process streams of the first part, which consists of a stripping column and the first scrubber. The rough values of the mass and heat balances are given in table 8a and 8b.

The stripping column as well as the first scrubber are assumed to operate isothermal. The heat capacity of the fluorosilicic acid solution is set equal to the heat capacity of water and quite rough values for the fluoride distribution coefficient are used. So the data given in table 8a and 8b are only

113

approximations of what will happen in practice.

phosphoric acid

digestion stage gas

water 15°C

SCRUBBER

STRIPPING COLUMN

-»~ stripping gas

-»► phosphoric acid

- * - 25w%H2SiF6

Figure 8: The stripping column and first scrubber

component in component out

phosphoric acid 28.9 water 3.0 off gasses from the digestion stage 3.1

phosphoric acid 28.9 25 u% H 2SiF 6 3.1 stripping gas 3.0

total 35.0 total 35.0

Table 8a: Mass balance of the stripping column and first scrubber (mass streams in kg/s).

The stripping gas is mainly carbon dioxide. Before this gas is recycled to the stripping column, the fluoride content of the stipping gas is reduced in the first scrubber, which operates at about 90 °C. A part of the stripping gas, as large as the quantity off gasses from the digestion stage, is disposed.

The fluoride content of the gas mixture to be disposed is lowered in a second scrubber, operating at room temperature, to meet environmental requirements.

1H1(

component in component out

phosphoric acid 6.8 water 0.2 off gasses from the digestion stage 2.8

phosphoric acid 6.8 25 w$ H2SiF& 0.5 stripping gas 2.5

total 9.8 total 9.8

Table 8b: Heat balance of the stripping column and first scrubber (enthalpies in MW).

In figure 9 the incoming and outgoing process streams of the second part of the fluoride removal stage are shown. The mass and heat balances of the second scrubber and the two direct contact condensor are given in the tables 9a and 9b.

stripping gas

vapour from HH/DH flasher

vapour from HH flasher

second scrubber and

condensors

j

15°C wafer

i

-■ '/ ^ / /

wast

wash

heat

Figure 9: The second scrubber and direct contact condensors.

Table 9b shows that no external heat is required to raise the temperature of the wash water for the HH filter cake and for the regeneration and washing of the ion exchange resin, to 60 °C. If, however, a higher temperature is required, external heat must be used to raise the temperature unless the heat of digestion of the phosphate ore is much higher than used in these calculations. The calculations for the mass and heat balances of the fluoride removal stage are only roughly, and therefore some external heat might be necessary in practice.

115

component in component out

stripping gas 3.0 water 25.4 vapour from the recrystallisation 1.4 vapour from the crystallisation 0.4

total

waste gas wash water

30.2 total

2.9 27.3

30.2

Table 9a: Mass balance of the second scrubber and direct contact condensors (mass streams in kg/s).

component in

stripping gas water vapour from the recrystallisation vapour from the crystallisation

2.5 1.6

3.5

1.0

component out

waste gas wash water

1.8 6.8

total 8.6 total 8.6

Table 9b: Heat balance of the second scrubber and direct contact condensors (enthalpies in MW).

The unwanted deposition of Si0_ on the walls of the equipment, which is frequently observed in phosphoric acid plants, can have a large influence on the operating efficiency of the scrubbers, of the direct contact condensors as well as of the stripper. It is, however, possible by very adequate temperature control to avoid formation of SiO. in the equipment. Therefore, the method proposed for removal of fluoride from the acid and the off gasses should be possible.

146

8.4. Discussion and conclusions

The results given in appendix 8.6.9 and presented in the former sections can be put together to give a mass and a heat balance of the complete CTPA process. In table 10 and 11 the mass and heat balances are given of a plant, producing 1000 tons P?0,- per day and operating according to the CTPA process. The composition of the products is given in table 12.

The overall phosphate efficiency of the process is calculated to be about 99 % from the amount of phosphate ore used, the amount of phosphoric acid produced and their phosphate content. The phosphate losses, due to incomplete washing of the HH and DH filter cakes are, however, neglected. In practice this assumption will not be fulfilled.

Most of the data and correlations, given in the appendices and used for the calculations, are obtained for chemically pure systems. The influence of impurities is not accurately known, but hardly any influence on the phosphate and cadmium incorporation in the HH crystals is expected ([27], chapter 9).

In practice, the overall phosphate efficiency of the process will therefore, be less than 99 %, but the specification of the products (a maximum cadmium concentration in the phosphoric acid of 5 ppm and in the HH and the DH of 1 ppm) can still be reached.

component in

phosphate ore c lay water 98 v% H^OJJ

37.1 0.8

38.7 32.4

component out

phosphoric acid HH-slurry DH-slurry 25 w$ H2SiF6

waste gas

28.9

53.2 20.9

3.1 2.9

total 109.0 total 109.0

Table 10: Overall mass balance of the CTPA process (mass streams in kg/s).

Table 11 shows that no external heat is required in the CTPA process. The results obtained so far, show that the production of concentrated and "clean" phosphoric acid and "clean phosphogypsum" should be possible with the CTPA process.

147

component in component out

ore and clay 98 M% H2S0^

water heat of digestion heats of mixing heats of crystallisation heat of recrystallisation

0.4 0.8 2.4 3.6 1.5

9.5

1.0

phosphoric acid HH-slurry DH-slurry 25 w% H2SiF5

waste gas heat losses

6.8 7.6 1.6 0.5 1.8 0.9

total 19.2 total 19.2

Table 11: Overall heat balance of the CTPA process (enthalpies in MW).

phosphoric acid

40 wj P205

1.7 w* H2S04

0.5 w? F 4 ppm Cd

HH

0.4 ppm Cd 0.2«J P205

DH

0.2 ppm Cd 0.3 w$ P205

Table 12: Composition of the products of the CTPA process.

Additional information about the heat of digestion, about the cadmium removal stage as well as about the fluoride removal stage is, however, required to make a complete engineering flow diagram of the CTPA process.

A reasonably sized pilot plant, in which all process steps are incorporated is needed to confirm the results given here. Only in a process in which all impurity levels are build up, accurate data for the description of the cadmium as well as of the fluoride removal stage can be obtained.

A preliminary study of the investment needed to build a plant operating according to the CTPA process and to build a plant operating according to a single filter HH/DH process, with the same production capacity, showed hardly any difference in investment cost.

148

8.5. Literature

1. Becker, P., Phosphates and Phosphoric Acid, Fertiliser Science and Technology Series, Vol. 3, Marcel Dekker Inc., New York, (1983).

2. Bos, W.S., Private Communications, Windmill B.V., Vlaardingen, (1985). 3. Edukos, A.T. and Ikonomu, E., Zh. Prikl. Khim., 58, 9, (1985) 1937. 4. Egan, E.P.jr. and Luff, B.B., Ind. Eng. Chem., 47 (1955) 1280. 5. Egan, E.P.jr. and Luff, B.B., J. Chem. Eng. Data, 11 (1966) 509. 6. Egan, E.P.jr., Luff, B.B. and Wakefield, Z.T.,

J. Phys. Chem., 62 (1958) 1091. Elmore, K.L. and Farr, T.D., Ind. Eng. Chem., 32 (1940) 580. Farr, T.D., Tennesee Valley Authority, Chem. Eng. Report 8 (1950) 44. Frankenfield, K., Verfahren zur Entfernung von Cadmium aus Phosphaterzen, Chemische Fabrik Budenheim, DE-3332698 A1, 21 maart 1985.

10. Frochen, J., Laine, J., Allyot, H. and Gosse, G., Concentrated Phosphoric Acid Manufacturing Process by Double Crystallisation, Bulletin d*Information HEURTEY, 57 (1972).

11. Gaubert, P., C.r., 197 (1933) 72. 12. Giauque, W.F., Hornung, E.W., Kunzler, J.E. and Rubin, T.R.,

J. Am. Chem. Soc, 82 (1960) 62. 13. Glazyrlna, L.N., Savinkova, E.I. and Grinevich, A.V.,

Zh. Prikl. Khim., 53, 11 (1980) 2524. 14. Ivanchenko, L.G., Guller, B.D., Zinyuk, R.Yu. and Peresvetova, S.A.,

Issled. V. Obi. Tekhnol. Mineral Udobr., L (1983) 58. 15. Kelley, K.K., Southard, J.C. and Anderson, T.C.,

Techn. Pap. Bur. Mines, nr. 625 (1941) 1. 16. Knobeloch, J.B. and Schwartz, C.E., J. Chem. Eng. Data, 7 (1962) 386. 17. Klocko, M.A. and Kurbanov, M.S.,

Ivzestija Sektora Fiz.-Chim. Anal., 24 (1954) 252. 18. Laptev, V.M., Kopylov, B.A., Varshavskii, V.L. and Ovsyannikova, L.G.,

Sbornik Trudov/Leningradskii Tekhol. Inst. Im Lensovita, 4 (1973) 11. 19. Linck, G. and Jung, H., Z. Anorg. Chem., 137 (1924) 407. 20. Perry, R.H. and Green, D.W., Ed., Perry's Chemical Engineers' Handbook,

6th. ed., McGraw-Hill Book Co., New York, 1984.

149

21. Rakshit, J.N., Z. Electrochemie, 32 (1926) 276. 22. Riddell, W.C., Rock. Prod., 53 (1950) 68, 102. 23. Rosincky, V. and Kockta, J., Spisy Priradovedeckou Fak. Musarykovy Univ.,

nr. 210 (1935) 3. 24. Skauge, A., Fuller, N. andHepler, L.G.,

Thermochimica Acta, 61 (1983) 139. 25. Slack, A.V., Ed., Phosphoric Acid, Fertiliser Science and Technology Series,

Vol. 1, Marcel Dekker Inc., New York, 1968. 26. Sluis, S. van der, Meszaros, Y., Wesselingh, J.A. and Rosmalen, G>M> van,

Proc. Fertiliser Society of London, nr. 249, (1986). 27. Sluis, S. van der, PhD Thesis, Technical University of Delft, Delft, 1987. 28. Smith, A. and Huffman, E.O.,

Ind. Eng. Chem. Chem. Eng. Data Ser., 1, 1 (1956) 99. 29. Sokolik, S., Z. Physik. Chem., A158 (1932) 305. 30. Spencer, L.J., Mineralog. Mag., 21 (1926/1928) 337. 31. Spijker, R. and Hovenkamp, H.,

UKF-rapport YBO-ME-85-3974, 12 december 1985. 32. Taperova, A.A., J. Appl. Chem. USSR., 18 (1945) 521. 33. Tjioe, T.T., Weij, P. and Rosmalen, G.M. van,

Proc. World Congress 111 Chem. Eng., Tokyo, Japan, Vol 2, (1986) 925. 34. Tjioe, T.T., PhD Thesis, To be published, Technical University of Delft. 35. Tutundzic, P.S., Liler, M. and Kosanovic, D.,

Glasnik. Hem. Drustva. Beograd., 20 (1955) 1. 36. Wakefield, Z.T., Luff, B.B. and Reed, R.B.,

J. Chem. Eng. Data, 17, 4 (1972) 420. 37. Weast, R.C. and Astle, M.J., Handbook of Chemistry and Physics,

CRC Press, Cleveland, Ohio, 59th. ed. (1979). 38. Weterings, K., The Development of a Clean Technology Phosphoric Acid

Process, DSM Research B.V., Report of 5 april 1985. 39. Wulff, P. and Schaller, A., Z. Krist., 87 (1934) 43. 40. Yudenkova, I.N. and Ishchenko, N.A.,

Zh. Prikl. Khim., 55, 5 (1982) 1162. 41. Zinyuk, R. Yu., Rogova, G.I. and Pozin, M.E.,

J. Appl. Chem. USSR., 44, 10 (1971) 2316.

150

8.6. Appendices

8.6.1. The solubility of HH and DH

The solubility product KHH for HH, defined as: ie = wj CaO * v% SO., is calculated from data given in literature [12, 13, 14, 18, 25, 32] as:

y « „ » .((-1277 + 36 * w* P 0 )/T) HH = 3 2 5 ,

where the weight percentages , a r e those present in s o l u t i o n and T i s the abso lu te temperature . The s o l u b i l i t y product KD„ for DH, defined a s : w? Ca * (w$ SO. ) n , i s given by Becker [1 ] as :

w* P O , KDH = 0.160 * " Q Q - - 2 * ( (2 .40 - l og (T-273) ) /0 .912 ) ,

in which n = 1.25 - 0.01 * (T-273).

The weight percentages are those in so lu t i on and T i s the abso lu te tempera ture .

8.6.2. Incorporation in HH and DH

8.6.2.1. Incorporation in calcium sulphate hemihydrate

The phosphate incorporation in HH can be expressed as [26]: wj P„0C in HH ., vt% Po0,. in solution 2_5 11 » 10 » w* SO in HH ' w* SO in solution

The cadmium incorporation in HH was estimated by [26]:

52i_Q5_BêE_!$B_öü „ i * 1 0 -3 * 52i_Sd_per_kg_solution moï Ca per kg HH = . mol Ca per kg solution

In the presence of 240 ppm I-, the value of 1 * 10 must be replaced by -4 5 * 10 , because the given iodide content reduces the cadmium incorporation by

about a factor two ([27], chapter 5).

151

Data obtained in the bench-scale plant ([27], chapter 9) revealed that:

wJ_Al_in_HH = k wjj Al in solution

2-Since the fluoride in the HH crystals is incorporated as AlF^ , which replaces ?-SO^ in the HH lattice ([27],chapter 5), the fluoride content of the HH

crystals also follows from the former equation.

8.6.2.2. Incorporation in calcium sulphate dihydrate

According to Frochen e.a. [10], the phosphate incorporation in DH is given by the equation:

w$ P o i„ DH - 0.01 * e(-°-055it ' °' 3 2 7 6 * f s c * °' 0 2 1 8 * fsc2), d o

in which fsc « percentage of free sulphuric acid in the solution. The free sulphuric acid content of a solution equals the total sulphate

concentration in the solution, expressed as mol per kg, minus the calcium concentration in the solution, also expressed as mol per kg.

The cadmium incorporation in the DH is given by:

p_Dm_Cd_in_HH . ppm Cd in solution ~ ' '

which value has been found in a commercial plant, operating under similar conditions [38]. In the presence of 240 ppm iodide the value of 0.1 must be replaced by 0.05, due to the reduction of the cadmium incorporation by formation of cadmium-iodide complexes. The incorporation of Al and F is assumed to be the same as for HH.

8.6.3. Vapour pressure of the P.0 -H„0 system . . do d

The correlations, used for predicting the vapour pressure of phosphoric acid solutions in the CTPA process are based on data measured by Farr [8].

152

log P - - (2935.61 - 0.122287 * n + 0.022738 * n 2) /T - 4.9373 * log T + 1.6939 * 10~5 * T + 4.874 * 10 " 8 * T2 + 25.5525 - 1.1237 * 10~3 * n +

1.1216 * 10"5 * n 2

for 0 £ w* P„0C 2 23-5 and

log P = - (2944.79 - 1.6128 * n + 1.6939 * 10 - 5 * T + 4.874 2.3186 * 10~5 * n2

for 23.5 £ w$ P.0C i 38.2 and

log P = - (3667.44 - 33.802 * n + 1.6939 * 10~5 * T + 4.874 4.4405 * 10~5 * n2

for 38.2 £ wj P.0,. £ 63 .7 ,

in which P = water vapour pressure in Pa T = temperature in K n = w$ p_0,. in the solution ^ D

These correlations are used for all the streams in the CTPA process. Sulphuric acid or other species, present in the solution, were treated as water.

In the digestion stage, the vapour pressure of water above the phosphoric acid is equal to the partial vapour pressure of water. The absolute pressure of 1 bar is the sum of the partial vapour pressure of water and the vapour pressures of carbon dioxide and the various fluoride compounds.

8.6.4. Fluoride distribution coefficients

8.6.4.1. Liquid-gas

If the composition of the solution is known, the concentration of fluoride in the gas phase can be estimated by the equation ([27], chapter 7):

0.02113 * nc)/T - 4.9373 * log T + * 10 "8* T 2 + 25.4825 2.7695 * 10

0.42947 * n2)/T - 4.9373 * log T + * 10 " * T 2 + 25.7501 - 5.0125 * 1 0~3

153

in H = 32.4 + -f- * ln(m(H )) - -\- * ln(m(H„SiF,)) + 0.156 * m(HoP0„) + 3 3 Q^n 2 6 3 * m(H2SiF5) + 0.457 * mCHgSO,,) - ^ 2 - ,

in which H - 2S FJ.1 ^.Y§B2ur mol H SiF, / kg H O

m(x) = mol of species x per kg of water in the solution T = temperature in K

For the digestion, the crystallisation and the fluoride removal stages a value 4 3 of 10 mg F/m vapour/w% F in the solution was used, while a value of 100 mg F/

m vapour/w$ F in the solution was used for the recrystallisation stage.

8.6.4.2. Liquid-solid

The total amount of fluoride deposited as Na SiF,, K2SiF, and NaKSiF,, is difficult to estimate, because a supersaturated solution of these compounds can easily be maintained. It was therefore quite arbitrary assumed, that in the CTPA process, half of the fluoride present in the ore, leaves the process as solids, as was also found in a commercial single stage HH process [2]. The incorporation of fluoride in the HH and the DH has already been defined (8.6.2.1.). The amount of fluoride precipitating with the HH and DH is less than 50 w% of the fluoride present in the phosphate ore. So the remaining part of the fluoride, to be precipitated, will be deposited as a fluorosilicate precipitate, also called scale. Furthermore, it is assumed that, this scaling will be formed only in the digestion stage, where the fluoride is added and in the recrystallisation stage, where the temperature is lowered from 90 to 60 °C. The fluoride deposited as scale, is assumed to be sodium fluorosilicate only. The fictitious solubility products used in the separate stages are:

K(digestion) = [Na+] * [SiFg2"] = 1.766 * 10_2 mol2/kg2

K(recrystallisation) = 3.295 * 10~3 mol2/ kg2

The clay, introduced in the digestion stage, was assumed to contain no sodium ions. Other values of the solubility products have to be used, if the clay contains sodium ions. The fluoride deposited as scale as well as the fluoride incorporated in the HH and DH is assumed to leave the process with the HH and DH

154

respectively. In practice the scale will be deposited on the walls of the reactors and cleaning of the reactors is needed frequently.

8.6.5. Density correlations

8.6.5.1. Phosphoric acid

The density of chemically pure phosphoric acid was derived from literature data [4,17] a n d fitted between 0 - 95 °C and 0 - 70 wj Po0,_ with the following

c- D correlation:

p = ( ( -6 .135 * w - 1.2450) * 10 -11 * (T - 273) + 1.0073 + 0.7487 * w +

0.4258 * w2 + 0.3237 * w3) * 103

in which p = dens i ty of phosphoric acid in kg/m w = w$ P.0C in s o l u t i o n T temperature in K

8.6.5.2. Sulphuric acid

The density of sulphuric acid was derived from literature data [20] and fitted between 0 - 100 °C and 0 - 100 w$ H-SO^ with the following correlation:

(0.83222 * w - 0.00046944 * (T - 273) + 6.9128) P = e . (-14.9985 + 0.069634 * (T - 273) + (24.484 - 0.086606 » (T - 273)) * w)

in which p = density of sulphuric acid in kg/m w = w$ SO. in solution T = temperature in K

8.6.5.3. CDHP solutions

The following correlations were obtained by fitting the literature data [7,28] between 0 - 70 w$ P.O., 0 - 95 °C and between 0 w? CaO and the saturation concentration of CaO in the solution:

p(T,w,c) = p(T,w) + c * fc * 103

155

in which p(T,w,c) = density of the CDHP solution, containing w w$ Po0c, c w$ CaO 2 D

at a temperature of T kelvin. p(T,w) = density of chemically pure phosphoric acid at the given

temperature and phosphate content c = w% CaO fc = 1.05 + 0.00225 * (T - 273) + 1.9 * w T = temperature in K

8.6.5.4. H PO^-H SOjj-HgO mixtures

According to Tutundzic e.a. [35] and Yudenkova and Ishchenko [40], the sum of the fractional contributions of the pure components can be used.

8.6.5.5. Calcium sulphate hemihydrate (HH)

HH can occur in two different modifications, the so-called a and B forms. The (5 form is dehydrated gypsum, while the a form develops from solutions. So in the CTPA process the a form of HH is precipitated. The density of the a form is taken to be the mean value of the data given in literature [11,15,19,23]:

p = 2746 kg/m3, st. dev. 11 kg/m3

8.6.5.6. Calcium sulphate dihydrate (DH)

The mean value of the density, obtained from literature data [15,20,21,30,39] is:

3 3 p =2312 kg/m , st. dev. 6 kg/m

8.6.5.7. The phosphate ore

The density of the Zin phosphate ore was determined to be approximately 3

2700 kg/m , as measured by the pyknometer method ([27], chapter 4).

8.6.5.8. Water

From literature data [37] the following expression for the density of water is

156

obtained:

p = 1006,287 - 0.12876 * (T - 273)

3 in which p = density of water in kg/m and T = temperature in K

8.6.6. Heat capacities and heat contents

8.6.6.1. Phosphoric acid

From literature data [6,36] the following correlation was obtained for chemically pure phosphoric acid:

cp = (3725.80 + 2.23326 * (T - 273) " 32.5439 * w - 0.00310846 * (T - 273)2 + 0.0236263 * w2 + 0.00887146 * (T - 273) * w) * 10~3

in which cp = heat capacity in kJ/kg/K, w = w$ P?0_ in sollution and T = temperature in K

8 . 6 . 6 . 2 . Sulphuric acid

According to l i t e r a t u r e data [ 2 9 ] , the heat capaci ty of chemically pure su lphur ic ac id i s nea r ly temperature independent and can be approximated with the following c o r r e l a t i o n :

cp = 1.19 * (1 - (9.8177 * 1 0 - 3 * w - 4.9327 * 10~5 * w2 + 1.5081 * 10~7 * w 3 ) )

in which cp = hea t capaciry in kJ/kg/K w = wj SO In so lu t i on

8 . 6 . 6 . 3 . CDHP s o l u t i o n s

From l i t e r a t u r e data [ 5 ] , the following equation was obtained for chemically pure CDHP s o l u t i o n s :

cp = (4175.09 - 49.1067 * c - 45.1056 * w + 1.59092 * c * w + 0.0974336 * w2 - 0.0270606 * c * w2 + 2 * (T - 298)) * 10~3

157

c = M% CaO in solution, T » temperature in K

8.6.6.1. H POjj-H^O^-HgO mixtures

For lack of data on t h i s system, the c o r r e l a t i o n given in 8 .6 .6 .1 was used . The su lphur ic ac id conten t of the s o l u t i o n was neglec ted and considered t o be water . After a l l c a l c u l a t i o n s were performed, a r e i n v e s t i g a t i o n of the data g iven in l i t e r a t u r e [ 1 6 ] , made i t poss ib le to de r ive the fol lowing express ion for the heat capaci ty of t h i s mix tures :

cp = (x * (18.0726 - 4.4635 * 10~3 * T + 6.7299 * 10~5 * T2 -1.4873 * 10 - 7 * T3) / 18.0152 + x2 * (32.07 + 3.923 * 10~2 * T) / 98.0774 + x3 * (0.1568 + 8.748 * 10 - " * T - 2.316 » 10~7 * T2 ) ) * (x3 / 0 . 2 ) 0 ' 0 5

in which cp = hea t capac i ty in c a l / g / ° C , T = temperature in °C x1 = w* H20, x 2 = wjt H2S04 and x = wit PgO

The dev ia t ions between the experimental v a l u e s , between 30 -90 °C, and those predic ted by the c o r r e l a t i o n given above i s maximal 3 %.

8 . 6 . 6 . 5 . Calcium sulphate hemihydrate (HH)

The heat capac i ty of the a form of HH i s given by [15,22] and equa l s :

cp = (488.4 + 1.13 * T) * 10~3

in which cp = heat capacity in kJ/kg/K and T = temperature in K

8.6.6.6. Calcium sulphate dlhydrate (DH)

According to Riddell [22], the following correlation, based on data of Kelley [15], can be used to estimate the heat capacity of DH:

cp = (530.7 + 1 .84 * T) * 10~3

in which, cp is given in kJ/kg/K and the temperature in K.

in which cp = heat capacity in kJ/kg/K, w = w$ P O in solution and

158

8.6.6.7. The phosphate ore

According to literature data [6], the following expression can be used to estimate the heat capacity of phosphate ore:

cp = (937.6 + 0.1189 * T - 20.25 * 106 * T~2) * 10~3

in which, cp is given in kJ/kg/K and the temperature in K.

8.6.6.8. Water

The heat content data, given by Weast and Astle [37], are used to derive the following expression for the heat content of water, relative to 0 °C:

H = 4.80296 * 10~2 + 4.2003 * (T - 273) " 5.05431 * 10_6 * (T - 273)2 + 4.10118 * 10-6 * (T -273)3

in which H = heat content in kj/kg and T = temperature in K.

8.6.6.9. Water vapour

The heat content of water vapour can be c a l c u l a t e d from l i t e r a t u r e da ta [20,37] with the following equa t ion , obtained from TISDATA, the databank for phys ica l parameters of DSM:

H = (44861.1 + 36.8545 * (T - 273) - 506.193 * P + 2.88044 * (T - 273) * P "

1.351176 * 10~2 * (T - 273) 2 - 14.1243 * P2) / 18.0153

in which H = hea t conten t in k j / k g , T = temperature in K and P = vapour p re s su re of water in bar

8.6.6.10. Carbon dioxide

From the databank of DSM, TISDATA, the following expression for the heat capacity of carbon dioxide was obtained [20]:

159

cp = (21.3548 + 6.425188 * 10 2 * T - 4.103002 * 10~5 * T 2 ) / 14.01

in which cp = heat capac i ty in kJ/kg/K and T = temperature in K

8 . 6 . 6 . 1 1 . Other compounds

The con t r ibu t ion of SiFü and HF to the hea t capac i ty of the gas s treams i s neg lec t ed , while the heat capaci ty of the Na?SiF, s c a l e i s a r b i t r a r y s e t equal to 0.75 kJ/kg/K. After a l l c a l c u l a t i o n s were performed, i t was found t h a t a value of 1 kJ/kg/K i s more r e a l i s t i c [ 2 0 ] . The heat capac i ty of clay i s 0.94 kJ/kg/K [ 1 7 ] . This value i s used, because the type of c l ay to be used i s not yet known. If the type of clay t o be appl ied i s known, then t h e heat capaci ty of the c l a y , as a funct ion of the temperature , can probably be obtained from the data and c o r r e l a t i o n s given by Skauge e . a . [ 2 4 ] . The h e a t capac i ty of the HpSiFg s o l u t i o n i s se t equal t o the heat capac i ty of wa te r , because t he re a r e no l i t e r a t u r e data found for t h i s s o l u t i o n .

8 . 6 . 7 . Entha lp ies of r e a c t i o n

8 . 6 . 7 . 1 . Digest ion of phosphate ore

C a 1 0 ( P 0 4 ) 6 F 2 + 1 l t H3P°4 " _ > 1 ° C a ( H 2 P V 2 + 2 H F

AH = - 100 kJ/kg ore [3]

8 . 6 . 7 . 2 . C r y s t a l l i s a t i o n of calcium su lpha te hemihydrate (HH)

2 Ca(H2P01))2 + 2 H^Ojj + H20 > 2 C a S O ^ . ^ O + 4 H^O^

AH = -200 kJ/kg HH [41]

8.6.7.3. Recrystallisation of HH into DH

2 CaSO^.^O + 3 H20 — > 2 CaS0r2H20

AH = - 100 kJ/kg HH [22]

160

All values are rough approximat ions , because they a r e mainly determined in water and no accura te data a re a v a i l a b l e a t the condi t ions p r e v a i l i n g in the CTPA process . These data , however, w i l l be measured in the near fu ture and for the time being, the values of the hea t s of r e a c t i o n , given above a r e used.

8 . 6 . 8 . Heat of mixing in the H-P0^-H2S0^-H2O system

8 . 6 . 8 . 1 . General procedure

Knobeloch and Schwartz [ 1 6 ] , have given a method to c a l c u l a t e the heat of mixing in a te rnary system, from the hea t s of mixing in the th ree binary systems (see f igure 1 ) .

H70 H70

H2S04 H3PO4 H2S04 H3P04 H2S04 H3POif

Figure 1: The t h r e e binary systems, in which the t e rna ry system can be s p l i t up

The binary h e a t s of mixing have t o be c a l c u l a t e d in each b inary system a t t h e following composit ions:

A1H12 = A H12 ( x1 ' 1 " V A,H 1 3 = AH 1 3 (x 1 t 1 - X l )

A1H23 = AH 2 3(x 2 / (x 2 + x 3 ) , x 2 / ( x 2 + x ))

A2H12 = ^ Hl 2

( 1 " x 2 , x 2 )

A2H1 3 = AH 1 3 (x 1 / ( x 1 + x 3 ) , x l / ( x 1 + x 3 )) A2H2 3 = AH 2 3 ( 1 -x 2 , x 2 )

161

A3H12 = AH 1 2 (x 1 / (x l + x 2 ) , x l / ( x 1 + x 2 ) ) A3H13 = A H ^ O - X - . x - ) A3H23 = AH 2 3 (1-x 3 ,x 3 )

in which x = mole f r a c t i o n water , A. = h e a t s in binary system 1 x = mole f r a c t i o n H„SO., A. = h e a t s in binary system 2 x = mole f r a c t i o n H PCL, A, = h e a t s in binary system 3 AH = heat of mixing component 1 (water) and 2 (H SO.) , e t c .

For each binary system a t e rna ry hea t of mixing can be obta ined from the t r i p l e t

of binary hea t s of mixing:

* 1 H 1 2 3 - x-%--x3 * V l 2 + x 2 - h - * *1H13 + / ( 1 ' X 1 ) * A1H23

A2H123 " x j T x - * * 2 H1 2

+ x--%2-x-- * A2H23 + / ( 1 " X 2 ) * *2H1 3

A3H123 " X-4-X-- * V l 3 + x-;-!-x-- * A3H23 + ^ " V * A3H12

The heat of mixing in the t e rna ry system i s then s e t equal to t h e mean va lue of t he three t e rna ry hea ts of mixing for the binary systems:

123 " 3

AH is given in c a l o r i e s per mol mix tu re .

The binary hea t s of mixing a r e given in the s e c t i o n s 8 .6 .8 .1 to 8 .6 .8 .3 Although a l l c a l c u l a t i o n s are based on a temperature of 25 °C, these c o r r e l a t i o n s may a l s o be used at higher t empera tu res , because only a s l i g h t inf luence of the temperature was found [ 1 6 ] . The d i f fe rences between t h e experimental va lues and the c a l c u l a t e d values of the t e rna ry heat of mixing a r e found t o be smal ler than 10 %, if t h e mole f r a c t i o n of water in the system exceeds 0 . 5 .

8 . 6 . 8 . 2 . Binary heat of mixing in the H PO.-H 0 system

From l i t e r a t u r e data [ 8 ] , the fol lowing express ion was obtained for the hea t of

162

mixing of chemically pure phosphoric acid and water:

AH13 = (F - 5.87) * x3

in which F = 56.968 * w - 2.9512 * w2 + 0.09298 * w3 -1.1159 « 10-3 * uH + H.8388 * lo"6 » w5

AH. _ = heat of mixing in cal. per mol mixture w = w$ H^PO^ in solution and x- = mole fraction H PO,. in solution

8.6,8.3. Binary heat of mixing in the H^SO^-H-O system

The data given by Giauque e.a. [12], were used to derive the following expression for the heat of mixing of chemically pure sulphuric acid and water:

AH = w * (100 - w) * (0.35801 - 5.639 * )0~H * w + H.3483 * 10~k * w2 -8.5131 * 10~6 * w3 + 8.2961 * 10~8 * w4)

with w = wj H.SOy in solution and AH.. = heat of mixing in cal. per mol mixture.

8.6.8.1. Binary heat of mixing in the H PO^-H-SO^ system

The expression for calculating the binary heat of mixing of chemically pure phosphoric acid and sulphuric acid is [16]:

AH23 = x2 * x3 * (1935 " 26i)1 * ( x2 " x 3 ) + 510,5 * (X2 + x 3 ) 2 )

in which x? = mole fraction sulphuric acid in solution x- = mole fraction phosphoric acid in solution

AH_, = heat of mixing in cal per mol mixture

2 The last term in the correlation, given above, i.e. 510.5 * (x + x ) , is probably a printing error in the original article, because x? + -x_ equals 1 in the binary system phosphoric-sulphuric acid. No significant error was found, however, in the ternary heat of mixing, if the above correlation is assumed to be correct.

163

8.6.9. Total mass and heat balances of the CTPA process

The total mass and heat balances of a 1000 ton P„Cv per day producing plant, operating according to the CTPA process was obtained by simultaneously solving 1268 linear or linearized equations. This large number of equations was necessary to obtain the composition of all process stream. In figure 1 the largest part of the flowsheet of the CTPA process is given, while figure 2 shows the fluoride removal stage. The numbers given on the flowsheets, represent either an amount of heat, a process stream or a piece of equipment and are explained below. The composition of all process streams as well as the mass and heat streams are given in table 1.

List of equipment

A1 First fluoride scrubber and stripping column A2 Second fluoride scrubber C1 Direct contact condensor for the vapour from the flasher in the

recrystallisation stage C2 Direct contact condensor for the vapour from the flasher in the

crystallisation stage El Heat exchanger for heating the wash water of the HH washing stage F1 Solid-liquid separation by gravity and a centrifuge after the digestion F2a,2b DH filtration and washing respectively F3a,3b HH filtration and washing respectively R1 First digester R2 Second digester R3 Third digester R4 First recrystalliser R5 Second recrystalliser R6 Third recrystalliser R7 Crystalliser 51 Product acid splitter 52 Sulphuric acid splitter 53 DH slurry recycle splitter SH Wash water splitter T1 Filter feed tank for HH slurry U1 Cadmium removal unit

164

Figure 1: Flowsheet of the CTPA process (except the fluoride removal stage),

165

G02 G01 G03

^ i 7 G 5 1 > — * A 1

G32L - r -1

L51

A2

L59

* 611 .

G52 r

L52

C1

G31>

L53 "

C2

i

L32 '

Figure 2: Flowsheet of the fluoride removal stage of the CTPA process.

List of process streams (kg/s)

CD = Cadmium D, DR = Calcium sulphate dihydrate (DH) E = Phosphate ore (soluble part) G = Off gasses H = Calcium sulphate hemihydrate (HH) I = Inert (insoluble part of the ore) K = liquid L = liquid S = c l a y

SC, SCR = sca le

T01 = E01 + 101

T02 = E02 + H02 + 102 + K02 + SC02

T03 = E03 + H03 + 103 + K03 + SC03 TOU = E04 + H0>i + 1014 + K01 + SC04

T11 = E11 + H11 + 111 + K11 + SC11

T12 = D12 + H12 + 112 + K12 + SC12

T13 = D13 + H13 + H 3 + K13 + SC13

T14 = DIM + 111 + K i t + SC1H

T15 = D15 + 115 + K15 + SC15

T16 =' D16 + 116 + K16 + SC16

166

^

T17 «■ D17 + 117 + K17 + SC17 T18 = D18 + 118 + K18 + SC18 T31 = H31 + K31 T32 = H32 + K32 T33 = H33 + K33

T3H = H34 + K34

List of heat streams (kJ/s)

QE1 -Q1E1 QF2 =Q1F2 QF3 =Q1F3 QR1 =Q1R1 + Q2R1 + Q3R1 + Q1R1 QR2 =Q1R2 + Q2R2 + Q3R2 + Q^R2 QR3 =Q1R3 + Q2R3 + Q3R3 + Q4R3 QR1 =Q1R4 + Q2R1» + Q3R4 + Q4R4 + Q5R4 QR5 =Q2R5 + Q3R5 + Q4R5 QR6 =Q2R6 + Q3R6 + Q1R6 QR7 =Q2R7 + Q3R7 + Q1R7 + Q5R7 QT1 =Q1T1 + Q2T1 Q1R1-Q1R4 = hea t of d i g e s t i o n of phosphate ore Q2R1-Q2R3.Q2R7 = hea t of c r y s t a l l i s a t i o n of HH Q2R4-Q2R6 = hea t of r e c r y s t a l l l s a t i o n of HH in to DH Q3R1-Q3R7.Q1T1 = Heat l o s se s through the wal ls Q4R1-Q4R7.Q2T1 = Power input by the s t i r r e r s Q5Rt,Q5R7,Q1F2,Q1F3 = heat of mixing Q1E1 = HH wash water heat ing

Table 1: The CTPA process : Composition of the process streams as well as the mass and heat s t r eams .

167

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174

9. THE BENCH-SCALE PLANT

9.1. Introduction

In the early stages of the development of a new process for e.g. the production of bulk chemicals, mainly batch experiments are performed on a laboratory scale. In these experiments small parts of the total process are studied. Such experiments are relatively cheap and simple, so the exploration of different process conditions and optimalisation of parts of the process is possible.

However, especially in processes where the role of impurities is predominant, as in phosphoric acid production processes, these experiments have a major drawback, because the results might not be representative for a continuous process. In a continuous process Impurity levels build up with time.

Many impurities are known for their striking influence on the size and shape of the growing crystals, and therefore on the filterability and washability of these crystals. This also counts for the calcium sulphate hemihydrate (HH) crystals precipitated during phosphoric acid production.

Neither in batch experiments nor in continuous experiments with chemically pure reagents these impurity levels are completely imitated. Even by purposeful addition of impurities, which are known to be released from the phosphate ore some minor impurities might be overlooked. Trace amounts of specific impurities like e.g. cerium [4] can already have a strong impact upon the growth kinetics of the developing crystals.

For this reason the performance of at least a few experiments, in which the real process is closely simulated on a bench-scale, is strongly recommended. The most relevant parts of the new Clean Technology Phosphoric Acid (CTPA) process were therefore tested in a bench-scale plant. The flowsheet of the bench-scale plant is shown in figure 1. This plant has been operated for sufficiently long periods (> 100 hours) to obtain an idea of the steady state behaviour of the CTPA process.

The aim of this study was not only to determine the influence of impurities on the permeability of the calcium sulphate hemihydrate (HH) cake obtained by filtration of the HH-phosphoric acid slurry, but also to determine their effect on the cadmium and phosphate uptake by the HH crystals at the selected optimal operating conditions.

175

AIR 98w%H2SO/+

ORE FEE OER

. D1

1 ,

D2

ANTI-FOAMING AGENT

»- n,

-* RETURN ACI 3

m O . T

PRESSURE FILTER

" PRESSURE

FILTER 2

'

i

CDHP BUFFER VESSEL

\ CRYSTALLISER

H2O

' ♦ FILTER

' ACID

BUFFER VESSEL

PRC )0UCT ACID

Figure I: Flowsheet of the bench-scale plant.

9.2. Experimental

9.2.1. Process conditions

The process conditions to be maintained during the operation of the bench-scale plant are given in table 1 and can be obtained from the chapters 4, 5 and 8 of this thesis: - at least 10 wj P.O.- has to be produced for direct use of the phosphoric acid

2 D for the production of mono- or di-ammonium-phosphate (MAP and DAP).

- the maximum solubility of CaO in phosphoric acid containing HO u% P.O,. at d. D

90 °C is about 5 w$ CaO. This solubility, together with the sulphate content of the return acid determines the quantity of recycle acid needed for complete digestion of the phosphate ore.

- in separate digestion studies no reduction in the digestion rate due to blinding of the phosphate ore was observed if the sulphate content of the return acid was less than 2 w% H_S0Ü.

- in separate crystallisation studies no AH formation occurred if concentrated sulphuric acid was fed directly into the crystalliser.

176

I crystalliser / return acid: HO v% P205> 2 u% H^O^, 90 °C | | digesters / CDHP solution : HO w% ?20 , H.5 w$ CaO, 90 °C |

Table 1: Process conditions to be maintained in the bench-scale plant.

The CaO concentration in the CDHP solution was chosen to be 1.5 w$ CaO instead of 5 w$ CaO to maintain a reasonable safety margin.

9.2.2. Chemicals

The phosphate ore used was a mixture of Khouribgha and Zin HO/60 w%. After thorough mixing the phosphate ore was sieved and all particles larger than 2 mm were removed. The composition of the phosphate ore is given in chapter 8 of this thesis, table 1 .

The anti-foaming solution to be added during the digestion to prevent excessive foaming was a 50/50 vol$ mixture of distilled water and anti-foaming agent (defoamer A-50 of Chemy B.V., Emmerich).

The sulphuric acid used was chemically pure 96 w% H.SCv. By using a special acid mixture, which already contains a basic impurity level, for the startup of the plant, the operating time of the plant needed to approach a steady state can be reduced. The acid mixture used for the startup of the plant was prepared by mixing commercially wet process phosphoric acid (53 u% P 0 and H w$ H SO., produced from the same mixture of phosphate ores), water and technical grade phosphoric acid (62 w$ P ? 0 ^ ' I n t n i s w a v t n e startup solution was obtained, containing 10 w% P.O., about 2 w% H_S0n and a basic

2 5 c. A impurity level.

9.2.3. Equipment

The flowsheet of the bench-scale plant is shown in figure 1 . Separate digestion studies showed that three digesters were needed for digestion of more than 99 % of the phosphate ore within one and a half hour. The phosphate ore was fed into the first digester by an "Accu-rate" feeder of TARCO BV, Vleuten. The weight of the feeder and the phosphate ore together was registered with time to control the phosphate ore feed.

177

The air stream needed to prevent condensation of water in the ore feed pipe was supplied by a separate air compressor.

The anti-foaming solution supply vessel was a magnetically stirred stop erlenmeyer with a volume of 500 ml. A 10 rev/min Watson Marlowe peristaltic tube pump equiped with a Viton B tube (inside diameter 0.8 mm) was used to feed the anti-foaming solution into the first digester.

The return acid was fed to the first digester with a 100 rev/min Watson Marlowe peristaltic tube pump equiped with a Marprene tube (inside diameter 3.2 mm).

All three digesters were double walled, thermostated vessels made of glass with a polypropene lid and a volume of about 2.5 litre.

All stirrers were six bladed turbine stirrers made of a Hastelloy B rod and covered with PVDF. The diameter of the stirrers in the digesters was 1.5 cm and the stirring speed varied from 1500 rpm in the first to 1200 rpm in the third digester.

The slurry from the first digester was led through a glass pipe (inside diameter 3 cm) to the second digester and in the same way from the second to the third digester. The third digester was equiped with an outlet, which could be closed by a glass valve (inside diameter 1 cm). The pump between the third digester and the pressure filters was a 100 rev/min Watson Marlowe peristaltic tube pump equiped with a Marprene tube (inside diameter 3-2 mm).

Pressure filters were needed for separation of the HH/ore residue from the CDHP solution, because vacuum filtration caused severe boiling of the CDHP solution after a few minutes filtration. The pressure filters were double walled, thermostated and made of PVDF. Two valves were installed at the top of each pressure filter, one for the CDHP slurry feed and one for airation.

The filter medium consisted of a perforated Hastelloy C support plate (diameter 11.1 cm) with 121 holes of 3 mm diameter, a terylene 23-12 filter cloth and a precoat layer. This precoat layer was necessary to avoid the passage of small particles through the filter cloth at the start of the filter procedure and consisted of 10 grams of Dicalite 1.15.8, put on the filter as a 10 w? slurry in water. The precoat layer was protected against the CDHP slurry stream fed to the pressure filter by a Hastelloy C break plate.

. The nitrogen gas used to squeeze the largest part of the CDHP solution out of the pressure filter cake was supplied from a nitrogen bottle through a back­wash vessel made of perspex and connected with a manometer.

The CDHP solution from the pressure filters was led through a Viton B tube (inside diameter 8 mm) to the double walled, thermostated CDHP solution buffer

178

vessel made of glass and equiped with a polypropene lid. The volume of this buffer vessel was about 3 litre. The stirrer diameter was 1.5 cm and the stirrer speed was about 600 rpm.

A 100 rev/min Watson Marlowe peristaltic tube pump equiped with a Marprene tube (inside diameter 3.2 mm) was used to pump the CDHP solution from this buffer vessel through a glass pipe (inside diameter 1 cm) into the crystalliser. This glass pipe was covered with "Heat by the Yard" heating tape, and isolated. The current through the heating tape was set at 1.2 A by adjustment of the voltage with a variac.

A 50 rev/min Watson Marlowe peristaltic tube pump equiped with a Viton B tube (inside diameter 0.8 mm) was used to feed the sulphuric acid solution from the supply vessel at room temperature into the crystalliser. The weight of the sulphuric acid solution supply vessel with a volume of about 3 litre, was registered as a function of time to control the sulphuric acid solution feed.

The crystalliser was a double walled, thermostated vessel made of glass, with a volume of about 3 litre and equiped with a polypropene lid. The stirrer diameter was 5 cm and the stirring speed was 1200 rpm.

The slurry from the crystalliser was led through a PVC tube (inside diameter 2.5 cm) into the HH filter baskets made of terylene 23_1I2, which were standing on P. porous glass plates at the bottom of double walled, thermostated filters made of glass. Thereafter the liquid was assembled in double walled, thermostated round flasks, made of glass and placed underneath the filters. Two of these filtration units were used to be able to operate them alternately. The round flasks, also called filter holders, were connected with each other by a Viton B tube (inside diameter 8 mm) and a three-way valve (inside diameter 1 cm). The filter baskets were used to move the HH quickly from the filters to the equipment used for the HH washing procedure (appendix 9.6.3.). The liquid from the filter holders was periodically fed through the three-way valve and a Viton B tube (inside diameter 8 mm) to the return acid buffer vessel.

The return acid buffer vessel was a double walled, thermostated vessel with a volume of about 2.5 litre, made of glass and equiped with a polypropene lid. At the bottom a valve was installed to be able to withdraw the excess return acid as product phosphoric acid. The stirrer diameter was M.5 cm and the stirring speed was about 600 rpm. From this buffer vessel the return acid was fed into the first digester through the same type of glass pipe as used for feeding the CDHP solution into the crystalliser.

To be able to heat all thermostats, equipment, slurries, liquids, heating tape and other materials to 90 °C, to stir each vessel and to weight the

179

phosphate ore feeder and sulphuric acid solution supply vessel a new 20 kW electricity supply-cable was installed.

In table 2 the residence times in each vessel are given. These times are only average values, since owing to adjustments in the process streams and the variation in the liquid levels in the buffer vessels, fluctuations in the residence times were unavoidable.

apparatus residence time in minutes

first, second digester third digester CDHP solution buffer vessel crystalliser HH filters and filter holders return acid buffer vessel

40 10 60 30 10 40

Table 2: Estimated residence times in the vessels in the bench-scale plant.

The use of relatively expensive materials, such as PVDF, Hastelloy B and C, Marprene and Viton B, was necessary to avoid excessive corrosion during the operation of the bench-scale plant.

9.2.1. Procedures

9.2.4.1. Safety

Before the start of a run with the bench-scale plant, each operator recieved an instruction map. This map contained not only the working procedures, but also names and phone numbers of persons qualified to give medical assistence and additional safety instructions.

For the sake of safety, the whole bench-scale plant was covered with PVC shields to protect the operators against potential acid disposure in case of malfunctioning.

The off gasses from the plant were removed by a ventilator, installed in the upper part of the plant.

Furthermore protective clothing for all operators was bought, like face shields, acid resistant overalls and handshoes.

180

The cadmium containing phosphoric acid and phosphogypsum, produced in the bench-scale plant were removed by the "DAAS", who take care for disposal of waste products of the TU Delft in such a way that environmental pollution remains minimal.

9.2.1.2. Startup procedure

First a mass balance was calculated for the bench-scale plant. The most relevant values of the process streams are given in table 3.

phosphate ore 150 g/hr sulphuric acid 100 g/hr return acid 3690 g/hr CDHP solution 3910 g/hr anti-foaming solution 20 ml/hr

Table 3: Process streams in the bench-scale plant.

A rough estimation was made of the operating time needed to build up the impurity levels in the bench-scale plant. It was found that the plant had to operate at least 100 hours to approach a steady state, unless the acid used for the startup of the plant already contained a basic impurity level. Therefore the so-called startup solution was made from commercially wet process phosphoric acid produced from the same ore mixture, water and technical grade phosphoric acid.

Reserve materials and equipment were ordered or made for the equipment used in the plant, in case replacement was expected to be necessary during the run.

In addition the ICP apparatus was made available for analysing the samples day and night at regular intervals of 3 to 1 hours.

About 150 glass sample bottles for the HH samples from the crystalliser and the HH/ore residue samples from the pressure filter were consecutively numbered.

Also 200 plastic sample bottles for the liquid samples from the crystalliser, CDHP solution and return acid buffer vessel were consecutively numbered and weighted. Thereafter an accurately weighted amount of about 3 ml distilled water was introduced in the plastic sample bottles. This amount of

181

water was needed to prevent precipitation of calcium sulphate or calcium phosphate from the liquid samples before their analyses.

A work-schedule was made for the operators needed to perform a continuous run with the bench-scale plant of about 100 hours. Each 24 hours three shifts were operating with in each shift four operators.

To initiate the process, all thermostats were filled with distilled water and heated to about 90-95 °C. Approximately 20 kg of the startup solution was heated to about 90 °C also. The thermostats had to be put on at least four hours before the real start of the run to heat the startup solution and the equipment to the desired temperature of 90 °C.

In the meantime the phosphate ore feeder was filled with about 8 kg of phosphate ore, the anti-foaming supply vessel was filled with about 400 ml anti-foaming solution and the sulphuric acid solution supply vessel was filled with approximately 5 kg of 96 w? H-SOj.. The ventilator used for removal of the off gasses was put on.

The presence of sufficient sulphuric acid (> 50 kg), Dicalite 4.15.8 (> 2.5 kg), phosphate ore (> 50 kg), anti-foaming agent (> 3 kg) and sodium-hydrogen-carbonate (> 25 kg) was checked.

Now the return acid buffer vessel was filled with startup solution. When the temperature of this solution was 90 °C-, the return acid pump was started, together with the phosphate ore, air and anti-foaming solution supply. The current through the heating tape around the return acid supply pipe was set at ■1.2 A.

Then the CDHP solution buffer vessel was filled with startup solution. From this vessel it was pumped into the crystalliser, HH filters and filter holders.

One of the pressure filters was made ready for operation by the precoat procedure (appendix 9.6.2.1.)

The thermostat for the HH washing equipment was put on and two stop erlenmeyera were filled each with about 65 ml of distilled water, closed and allowed to heat to 90-95 °C in the thermostat.

A stop erlenmeyer was filled with about 100 grams of water, closed and put in the thermostat next to the pressure filters to have wash water for the pressure filter cake at the right temperature available.

After about one and a half hour the third digester was half filled and the pump to the pressure filter was started. Then the second pressure filter was made'ready for filtration by the precoat procedure (appendix 9.6.2.1.). The CDHP solution buffer vessel was almost completely filled after about one and a half hour, whereafter the CDHP solution and sulphuric acid pumps were started. The

182

current through the heating tape around the CDHP solution feed pipe was set at 1.2 A. In about half an hour the HH crystals from the crystalliser had filled the first filter basket and the run proceeded further according to the working procedure.

9.2.1.3. Working procedure

During the run with the bench-scale plant each three or four hours solid and liquid samples were taken for analyses. After analyses of the liquid samples the sulphate, phosphate and calcium ion concentrations in the CDHP solution and return acid were known. If these concentrations differed from the process conditions to be maintained, these differences were used to adjust the process streams in the plant.

Problems caused by faillure of the equipment in the plant had to be recognized if possible in advance to avoid disturbances of the process. Therefore a so-called service round was made each half an hour (appendix 9.6.1.).

The plant was further operated according to the pressure filter procedures (appendix 9.6.2.) and the HH washing procedure (appendix 9.6.3.).

During the run one operator was responsable for the HH washing procedure, one operator for the pressure filter procedures and one operator for the analyses. The fourth operator made the service rounds, repaired damaged equipment, corrected process disturbances and took over from the other operators, if they had to leave for a while.

The last step of a run with the bench-scale plant was the stop procedure.

9.2.1).4. Stop procedure

First samples of the solids and the liquids were taken for the last time. - the stop procedure started by stopping the supply of phosphate ore, anti-

foaming solution and air to the first digester. The return acid pump was put in reverse to empty the return acid feed pipe.

- the current through the heating tape around this pipe was turned off as well as the return acid pump, after the return acid pipe was empty.

- the contents of the return acid buffer vessel was siphoned over into the product acid storage vessel.

183

- the contents of the first digester was siphoned over into the third digester, after the stirrer was stopped.

- if the third digester was almost empty, the contents of the second digester was siphoned over into the third digester, after the stirrer was stopped.

- when the third digester was empty again, the pump between this digester and the pressure filter was stopped as well as the stirrer.

- thereafter the phosphate ore feeder and anti-foaming solution supply vessel were emptied in the appropriate vessels.

- the digestion stage as well as the pressure filters were cleaned with a sodium-hydrogen-carbonate solution in water and thereafter with potable water.

- In the meantime the CDHP solution buffer vessel was empty and the CDHP solution and sulphuric acid pump were stopped and the current through the heating tape around the CDHP solution feed pipe was turned off.

- the contents of the crystalliser was siphoned over into the HH filter baskets. - the sulphuric acid supply vessel was emptied in the appropriate vessel. - all return acid was recovered as product phosphoric acid. - the tubes were removed from the pumps. - thereafter all equipment, inclusive the HH wash equipment and the materials

used for taking samples were cleaned with a sodium-hydrogen-carbonate solution in water and then with potable water.

- after everything was cleaned, the ventilator was turned off and a list was made of the materials and equipment, which were damaged and had to be repared or replaced before the next run.

9.2.5. Analyses

The liquid samples taken from the CDHP solution buffer vessel and the crystalliser or return acid buffer vessel were analysed with the ICP technique during the run with the bench-scale plant. The CDHP solution was analysed for its P20,-, CaO, Si and Cd content, while the return acid or the liquid from the crystalliser were analysed for its P-Cv, H„SO„, Si and Cd content.

The solid samples were analysed with the ICP technique for their P20j- and Cd content after the run.

Sometimes a few liquid and solid samples were analysed with the ICP technique for their Al and Fe content.

The INAA technique was used to determine the radium content of some solid and liquid samples.

181

9.3. Results and discussion

9.3.1. The performance of the bench-scale run

The Po°5' C a 0 a n d Cd concentrations during the last run in the CDHP solution buffer vessel are plotted in the figures 2, 3 and 4 as a function of time.

The P205> H2SOi( a n d C d concentrations in the crystalliser and the P 0 and

Cd concentration in the HH samples from the crystalliser are plotted versus time in the figures 5 to 9.

The figures 2 to 7 clearly demonstrate that the operating conditions were oscillating around the preset process conditions in the CDHP solution and in the solution in the crystalliser and that their average values were reasonably constant. Due to the oscillation of the operating conditions the phosphate and cadmium incorporated in the HH samples as shown in the figures 8 and 9, also show an oscillating behaviour.

w % 50-1

40-

30-

20-

10-

(

P2°5

^^Ww/W

1 1 ) 50 100

time (hours)

w 6 -

5-

3-

2-

1 -

0

C

%CaO

f\l

^

A/lyV/v/^W V y\

1 1 50 100

time (hours)

Figure 2: P_0_ concentration in the CDHP solution in the buffer vessel.

Figure 3= CaO concentration in the CDHP solution in the buffer vessel.

The consequences of the operating conditions in combination with the impurity levels reached in the bench-scale plant on the phosphate and cadmium

185

incorporation in the HH as well as on the permeability of the f i l t e r cakes will be discussed in the following sect ions.

ppm Cd 30-

2 0 -

10-

vs/V^A/vWV

50 100 time (hours)

w %P70c 50-1

40-

30-

20-

10-

^^A/vTW/VY

n 1 -

50 100 time (hours)

Figure 4: Cd concentration in the CDHP solution in the buffer vessel.

Figure 5: P„0_ concentration in the solution in the crystalliser.

w %H 7S0, 10-,

8-

4-

2-

_ 1 j _ 50 100

time (hours)

ppm Cd 30-1

25-

20-

15-

10-

5

—I T" 50 100 time (hours)

Figure 6: H SO. concentration in the solution in the crystalliser.

Figure 7: Cd concentration in the solution in the crystalliser.

186

w % P 2 0 5 0.4-,

0.3-

0.2-

0.1-

0.0- -I T -

50 100 time (hours)

Figure 8: P_0C concentration in the 2 5 HH crystals from the crystalliser.

ppm Cd 10-,

8-

6-

4

2H

1 I 0 50 100

h'me (hours) Figure 9: Cd concentration in the

HH crystals from the crystalliser.

run | w% CaO in the | conditions in the crystalliser | cadmium incorporation | CDHP solution | wj P 20 w$ H ^ O ^ ppm Cd | ppm Cd KCd

« I 5 I * . 6 I 7 I

4.9

4

4

4 .2

| 37

I 36 | 40

I 40

2.3

2.5

2

2.1

18

25

19

23

4.5

6

6

6

0.25

0.24

0.32

0.26

Table 4: Cadmium incorporation in HH during the bench-scale runs. * addition of Dicalite 4.15.8 as body feed. .

9.3.2. The cadmium incorporation in HH

The results of the cadmium incorporation in the HH precipitated from the crystalliser during the last run, as well as those obtained from former runs are given in table 4. The w? CaO in the CDHP solution is only given as a measure for the solids content in the crystalliser, while also the KCd values, representing the cadmium concentrations in the HH divided by the corresponding cadmium concentrations in the solution, are presented.

187

The KCd value obtained in the 6th run might be influenced by the addition of Dicalite 4.15.8. This was added in the digestion stage as body feed to improve the separation of the HH/ore residue mixture from the CDHP solution by pressure filtration.

The cadmium incorporated in the HH from the digestion stage (< 0.07 w$ HpSOj.) varied between 2 and 3 ppm at a cadmium concentration in the CDHP solution of 20-25 ppm. This implies a KCd value of about 0.1 for the HH from the digestion stage.

9.3.3. The phosphate incorporation in HH

The results of the P.0,. uptake in the HH, precipitated in the crystalliser during the last four runs with the bench-scale plant are given in table 5.

run I w? CaO in the I conditions in the crystalliser | w% P_0C incorporated 1 1 ' d o

I CDHP solution | w$ P20 wj H2S04 | in the HH

37 2.3 I 0.13 36 2.5 I 0.2 HO 2 I 0.6 40 2.1 | 0.2

Table 4: Phosphate incorporation in the HH formed in the crystalliser during the bench-scale runs. * addition of Dicalite 4.15.8 as body feed.

From this table it can be concluded that the addition of SiO in the form of Dicalite 4.15.8 strongly influences the phosphate incorporation in the HH. A possible explanation can be that the addition of silica influences the Al-F-Si equilibrium in the solution. The molar F/Al ratio will change and it is known that this ratio influences the growth kinetics of the HH crystals [1]. This change of the molar F/Al ratio may cause a higher supersaturation and thus an increase in the phosphate incorporation (chapter 5).

The value of the phosphate incorporation to be expected in practice, without Dicalite 4.15.8 addition, is about 0.2 w$ P 0 in the HH from the crystalliser.

4.9 4 4 4.2

188

The amount of phosphate incorporated in the HH obtained from the digestion stage appeared to be about 3"3-5 w$ P 0 in all four runs.

9.3.4. The permeabilities of the filter cakes

The permeability of the HH cake obtained by filtration of the HH from the crystalliser is determined according to the procedure given in chapter 6. The slurries were obtained directly from the overflow of the crystalliser. The values of the permeability are given in the tables 6, 7 and 8 for the last three runs with the bench-scale plant. The oscillation of the operating conditions around the preset process conditions, as already mentioned, is clearly visible in these tables. The influence of the CaO content of the CDHP solution on the w% solids in the solution in the crystalliser is, for example shown in table 7.

conditions in the crystalliser | permeability of the w% P205 wj H2S04 w? solids | HH cake * 10 1 2 m 2

39 3.7 7.3 | 16.5 39 3.5 7.3 | 22.7

Table 6: Permeability of the HH cake obtained by filtration of the HH slurry from the crystalliser in the 5th run with the bench-scale plant.

cond i t ions in the w? P205

39 40

37 140 •

39

w$ H 2 S0 4

H.2 2.0 2.2 3.0 1.6

c r y s t a l l i s e r J w% s o l i d s |

5 I 9.1 |

10 |

6.3 | 10.2 |

permeabi l i ty of the HH cake * 10 1 2 m2

4.5 2.0 3-0

5.1 0.9

Table 7: Permeability of the HH cake obtained by filtration of the HH slurry from the crystalliser in the 6th run with the bench-scale plant. In this run Dicalite 4.15.8 was added as body feed.

189

conditions in the crystalliser | permeability of the w$ Pp05 u% HgSOjj w? solids | HH cake * 10 1 2 m2

40

40

39

140

40

2

3

3

2.5

2

6.7

6.6

7.1

6.7

7.1

Table 8: Permeability of the HH cake obtained by filtration of the HH slurry from the crystalliser in the 7th run with the bench-scale plant.

In table 9 the permeability is given for the DH cake obtained by filtration of a DH-phosphoric acid slurry from a commercial single filter HH/DH (Nissan H) process.

conditions in the recrystalliser | permeability of the v% P_0C wj H„S0„ w$ solids I DH cake * 10 1 2 m 2

2 5 2 4

26.4 3.2 34.5 | 1.0

Table 9: Permeability of the DH cake obtained by filtration of the DH slurry (at 60 °C) formed after recrystallisation of HH into DH in a commercial single filter HH/DH (Nissan H) process

These results indicate that addition of Dicalite 4.15.8 lowers the permeability of the HH cake significantly to about 20 % of the permeability obtained under normal operating conditions. By comparing the results of table 6

_i 2 p and 8 a mean permeability of about 15 * 10 m is obtained for the HH cake produced by the CTPA process, where the sulphuric acid concentration in the crystalliser is about 2 wj H SO.. This mean value of the permeability of the HH cake produced by the CTPA process will approach the value to be obtained in a commercial CTPA process, because the influence of impurities on the permeability is taken into account in these bench-scale runs. Comparison with the value in table 9 shows that the permeability of the DH cake is less than 10 $ of the

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value for the HH cake under the operating conditions prevailing in the two different processes.

The permeability of the HH/ore residue cake obtained by pressure filtration of the slurry from the digestion stage, consisting of a HH/ore residue mixture in the CDHP solution, could not be measured directly, due to its extreme low value. If, however, some basic data given in table 10 are used for an estimation

~11 2 of the permeability of the HH/ore residue cake, a value of about 5 * 1 0 m is obtained, which is even less than 10 times the permeability of the DH cake.

viscosity of the CDHP solution density of the CDHP solution filter area cake height pressure drop over the cake mass stream of CDHP solution filtered

Table 10: Data used for the calculation of the permeability of the HH/ore residue cake, obtained by pressure filtration of the CDHP slurry from the digestion stage in the runs with the bench-scale plant.

9.3.5. Additional results

Some additional information could be deduced from the continuous runs with the bench-scale plant. For the HH precipitated in the crystalliser from a solution containing 37 w% P-,0,. and 2.3 w$ H.SO.., distribution coefficients for Al and Fe ions could be derived. For Al a K-value of approximately 0.H and for Fe a K-value of about 0.05 was obtained.

Furthermore the radium content of some solid and liquid samples taken during a run with the bench-scale plant, was determined. According to a DSM patent (NL-Patent no. 8006946 dd. 16-6-1982) all radium should precipitate in the digestion stage, if 10 $ of the total amount of sulphuric acid used in the process were fed into the digestion stage. In our process about 20 % of the required sulphuric acid is indirectly fed into the digestion stage with the return acid.

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3 * 10 J

1500 9.5 * 10 2.5 100 3.7

Pa * s kg * m 2

m m kPa kg * hr

The measured radium concentrations in the various phases are given in table 11 for the 5th run with the bench-scale plant. The process conditions maintained during this run are given in table 12.

Compound radium content (Beq/g)

phosphate ore 1.3 HH from the crystalliser 0.7 HH/ore residue from the digestion stage 1.2 product phosphoric acid 0.2

Table 11: Radium concentrations in the feed material and products of the 5th run with the bench-scale plant.

process stream w? P 0. wj CaO w? H-SCL

CDHP solution 36 4 return acid 36 - 2.5

Table 12: Process conditions during the 5th run with the bench-scale plant.

These results do not confirm the findings reported in the patent, probably because the sulphuric acid was fed into the first digester after dilution with return acid and the results given in table 11 are obtained in a continuous process, while the results reported in the patent are obtained in batch experiments.

The radium concentration in the HH precipitated in the crystalliser in this process and given in table 11 is comparable with the radium content of the DH obtained by recrystallisation of HH in a commercial single filter HH/DH (Nissan H) process [3].

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9.4. Conclusions

From the continuous runs with the bench-scale plant the following conclusions can be drawn: - the cadmium incorporation in the HH precipitated in the crystalliser in the

presence of impurities during the runs with the bench-scale plant is only slightly higher than in HH produced from chemically pure reagents (KCd » 0.25 and 0.2 respectively).

- the phosphate incorporation in the HH precipitated in the crystalliser in the presence of impurities during the runs with the bench-scale plant is also hardly different from the incorporation found in the HH produced from chemically pure reagents (w$ P?0,- = 0.2 and 0.15 respectively).

- The permeability of the HH cake obtained by filtration of the HH slurry from the crystalliser in the bench-scale plant, i.e. by filtration of crystals formed in the presence of impurities, is a factor five less than the values obtained for the HH cake produced in experiments with chemically pure reagents. This clearly shows the striking influence of impurities on the permeability of the HH cake. For this reason pilot plant data normally are used to obtain a reasonable estimation of the permeability of the filter cake to be formed in a commercial process.

-12 -12 2 (the permeability was 15 * 10 and 65 * 10 m respectively)

- under the conditions prevailing in the CTPA process (90 °C, 40 w? P201'

2 w$ H So or 4.5 w$ CaO) corrosion is a serious problem. - running a bench-scale plant is a tough job, due to the consequences of

even small variations in feed, process streams or concentrations. - the search for a method to remove radium from the HH or to prevent the

coprecipitation with the HH has to be continued.

9.5. Literature

1. Becker, P., Phosphates and Phosphoric Acid, Fertiliser Science adn Technology Series, Vol. 3, Marcel Dekker Inc., New York (1983).

2. Slack, A.V., Phosphoric Acid, Fertiliser Science and Technology Series, Vol. 1, Marcel Dekker Inc., New York (1968).

3. Weterings, K., The Utilisation of Phosphogypsum, Proceedings of the Fertiliser Society of London, nr. 208 (1982).

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4. Witkamp, G.J., Schuit, S.P.J. and Rosmalen, G.M. van, Recrystallisation of Calcium Sulphate Modifications in Phosphoric Acid, Condensed Papers of the Second International Symposium on Phosphogypsum, Miami, Florida (1986) 106.

9.6. Appendices

9.6.1. Service round

Use a wake-up signal to perform the following actions each half an hour. 1. Control the air supply 2. Control the amount of phosphate ore in the ore feeder and refill

if necessary. 3. Control the ore supply to the first digester. 4. Control the return acid and anti-foaming solution supply to the

first digester. 5. Report caking of phosphate ore in the first digester and suspend it

in the solution. 6. Control the amount of anti-foaming solution in the supply vessel and

refill if necessary. 7. Control the stirring speed in the three digesters. 8. Control the slurry overflow from the first and the second digester. 9. Control the pressure on the operating pressure filter. 10. Control the flow of the CDHP solution into the buffer vessel. 11. Control and report the height of the CDHP solution level and the stirring

speed in the CDHP solution buffer vessel. 12. Control the temperatures in the orystalliser, third digester, return acid

and CDHP solution buffer vessel and report deviations from 90 °C. 13. Control the supply of CDHP solution and sulphuric acid to the orystalliser. 14. Control the slurry overflow from and the stirring speed

in the orystalliser. 15. Control the amount of sulphuric acid in the supply vessel and refill

if necessary. 16. At the end of each shift the water levels in the thermostats are

controlled and adjusted, if necessary, by feeding distilled water.

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9.6.2. Pressure filter procedures

9.6.2.1. The precoat procedure

- 10 grams of Dicalite 't. 15.8 are suspended in 90 grams of distilled water. This slurry is thoroughly mixed and poured into the pressure filter.

- The break plate for the CDHP slurry stream is put in the filter. - The CDHP slurry supply pipe through the lid of the filter is cleaned with

a rod. - The filter lid is put on the lower part of the filter and the whole

is screwed up. - The CDHP slurry supply tube is cleaned with water and fitted to the

pressure filter valve. - The pressure filter valve is opened. - The airation valve is closed. - The pressure filter is ready for the filter procedure.

9.6.2.2. The filter procedure

- The pump between the third digester and the pressure filter is started. - After about one hour the pressure in the pressure filter is 1.8 to 2.0 bar

excess. - Now the filter change procedure is started.

9.6.2.3. The filter change procedure

1. The pressure filter valve is closed. 2. The pump is put in reverse. 3. If the tube between the pressure filter and the third digester is empty

the pump is stopped. H. The valve near the third digester is closed. 5. The tube is removed from the pump and the third digester. 6. The tube connected to the second pressure filter is put into the pump

and fitted to the third digester. 7. The pump is started, pumping in normal direction. 8. The valve near the third digester is opened.

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9. The CDHP solution in the loaded pressure filter is almost completely squeezed out by supplying nitrogen gas to maximal H bar excess.

10. The nitrogen supply is stopped. 11. The pressure filter is airated by opening the airation valve. 12. The pressure filter is opened. 13. About 100 grams of distilled water per hour operating time of that

pressure filter is put into the filter. 14. The stop erlenmeyer is filled with 100 grams of distilled water,

closed and put in the thermostat again. 15. After removing the break plate, the cake and the water are stirred gently. 16. The pressure filter and the airation valve are closed. 17. Repeat point 9 to 12. 18. A sample of the upper part of the cake is taken. 19. The remaining solids are removed from the filter with a spoon in the

residue storage box. 20. The residue storage box is weighted and the weight, date and time are noted

down 21. Now the precoat procedure starts.

9.6.3. The HH washing procedure

Each quarter of an hour a filter basket filled with HH slurry is removed from the filters located after the crystalliser in the bench-scale plant. All filters are made of glass, so be carefulll! - the contents of the filter basket is spread out over the filter. - during one minute the slurry is filtered by suction and in the meanwhile firmly pressed by a stamper to squeeze almost all acid out of the cake.

- the suction is turned off and the wash water erlenmeyer, containing the wash water of the second wash step of the former washing procedure is emptied on the filter.

- the cake and the solution are stirred gently. - during one minute the slurry is filtered by suction, while the cake is firmly pressed by a stamper to squeeze out the liquid.

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- the suction is turned off and the erlenmeyer underneath the filter is emptied in the filters placed after the crystalliser In the bench-scale plant.

- the wash water erlenmeyer is put underneath the filter. - the stop erlenmeyer containing about 65 grams of distilled water is removed

from the thermostat and emptied on the filter. - the stop erlenmeyer is filled with 65 grams of distilled water and put

into the thermostat again. - The cake and the water are stirred gently. - the slurry is filtered one minute by suction, meanwhile firmly pressing the

cake with a stamper to squeeze out the liquid. - the suction is turned off and the wash water erlenmeyer is removed, closed and put into the thermostat.

- a sample is taken from the upper part of the HH cake. - the HH cake is removed with a spoon into the HH storage box. - the HH storage box is weighted and the weight,date and time are noted. - the first erlenmeyer is put underneath the filter again. - the filter basket is cleaned with water and set apart for later use on

the filters placed after the crystalliser in the bench-scale plant. - The HH washing procedure can start again.

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SUMMARY

The fertiliser industry is the main user of phosphoric acid produced by digestion of phosphate ore with sulphuric acid. Impurities, such as the heavy metal ion cadmium, originating from the phosphate ore are distributed between the phosphoric acid and the byproduct, a calcium sulphate modification.

In the Netherlands, this byproduct, the so-called phosphogypsum, is disposed in the river Rhine. Due to environmental restrictions the. use of phosphoric acid with a high cadmium content as well as the disposal of phosphogypsum with a high cadmium content will be prohibited in the near future. The Dutch Government therefore started a project in cooperation with DSM and the TU-Delft to develop a Clean Technology Phosphoric Acid (CTPA) process, in which phosphoric acid is produced with less than 5 ppm cadmium and phosphogypsum with less than 1 ppm cadmium.

In phosphoric acid processes operating nowadays the digestion of the phosphate ore and the crystallisation of calcium sulphate takes place in the same reactors. The process route selected for the CTPA process has separate digestion and crystallisation stages, thus enabling removal of the cadmium in between.

In this thesis the results are presented of an investigation of the technological aspects of this new process.

First of all the mechanism which controls the digestion rate of phosphate ore in recycled phosphoric acid was determined to be the diffusion of calcium ions from the surface of the ore into the bulk of the solution. The masstransfer coefficients were determined as a function of the phosphoric acid concentration and the temperature.

Thereafter the cadmium and phosphate incorporation in calcium sulphate hemihydrate (HH) formed in the crystallisation stage was measured in dependence of the process parameters. The cadmium incorporation increased with an increase in the sulphate content of the solution in the crystalliser, while the phosphate content of the HH crystals decreased.

Furthermore, the filtration of this phosphoric acid-HH slurry was studied. The permeability of the HH cake obtained in HO vi% P.Cv is maximal, if the sulphate content of the solution in the crystalliser is about 1.8 w$ H„SCv at a residence time of about 20 minutes.

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An aditional study was performed to determine the fluoride distribution coefficients between phosphoric acid solutions and the ambient air. An expression was determined, which predicts the fluoride distribution coefficients as a function of the temperature, the fluoride, phosphate and sulphate content of the phosphoric acid solutions.

Finally all the results obtained in the former studies as well as some additional results were used to solve the mass and heat balances of a plant operating according to the CTPA process and producing 1000 tons P?0_ per day. This plant was found to be nearly self supporting with respect to energy consumption.

Furthermore a bench-scale plant was build in which the influence of impurities on the cadmium and phosphate uptake in the HH crystals as well as on the permeability of the HH cake obtained by filtration of the phosphoric acid-HH slurry were studied. The influence of impurities on the incorporation of cadmium and phosphate ions in the HH crystals was hardly detectable, but the permeability of the HH cake was about five times less than the permeability of the HH cake obtained from separate crystallisation experiments, in which chemically pure reagents were used.

The results point out that a CTPA process is in principle feasible. Validation on a pilot plant scale is, however, still required.

The feasibility of the cadmium removal section will be reported elsewhere (PhD Thesis T.T. Tjioe, Technical University of Delft, 1987).

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SAMENVATTING

Fosforzuur, dat gemaakt is door het ontsluiten van fosfaaterts in zwavelzuur, wordt voornamelijk gebruikt door de kunstmest industrie.

In het fosfaaterts bevinden zich veel verontreinigingen, waarvan het zware metaal ion cadmium,, er een is. Tijdens het ontsluiten van het erts verdeelt het cadmium zich over het fosforzuur en het bijproduct, een gehydrateerd calciumsulfaat. Dit bijproduct, fosfogips genaamd wordt in Nederland geloosd in de Nieuwe Waterweg. Tengevolge van milieutechnische bezwaren zal het gebruik van fosforzuur met een hoog cadmium gehalte evenals het lozen van fosfogips met een hoog cadmium gehalte in de nabije toekomst verboden worden.

De Nederlandse Overheid heeft daarom samen met DSM en de TU-Delft een project opgestart, om een mogelijkheid te vinden om schoon fosforzuur met minder dan 5 ppm cadmium en schoon fosfogips met minder dan 1 ppm cadmium te produceren.

In de hedendaagse fosforzuur processen vindt de ontsluiting van het fosfaaterts en de kristallisatie van het calciumsulfaat plaats in dezelfde reactoren. De proces route, die geselecteerd is voor het Schone Technologie Fosforzuur Proces (STFP) berust op een scheiding van deze twee stappen, waardoor het mogelijk is om het cadmium tussen deze twee stappen in te verwijderen.

In dit proefschrift worden de resultaten gepresenteerd van het onderzoek naar de technologische aspecten, die voor de ontwikkeling van dit nieuwe proces onderzocht moesten worden.

Als eerste is de ontsluitsnelheid van fosfaaterts in fosforzuur onderzocht. Er is gevonden dat diffusie van calcium ionen van het erts oppervlak naar de bulk van de oplossing de snelheidsbepalende stap is. Ook zijn de stofoverdrachtscoefficienten als functie van de fosforzuur concentratie en de temperatuur bepaald.

Daarna is de cadmium en fosfaat inbouw in het calciumsulfaat hemihydraat (HH), dat geproduceerd wordt in de kristallisatie stap van het proces, onderzocht. De invloed van verschillende proces parameters op de mate van inbouw is bekeken. Het meest opvallende resultaat is dat de cadmium inbouw in het HH toeneemt bij een stijgende zwavelzuur concentratie in de oplossing in de kristallisator, terwijl de fosfaat inbouw juist afneemt.

Verder is de filtreerbaarheid van deze HH kristallen onderzocht. In HO w? P_0_ wordt de maximale permeabiliteit van de HH koek bereikt, als de HH gevormd wordt in een fosforzure oplossing (MO wï P20c)> d i e ongeveer 1,8 w$ H-SOn bevat bij een gemiddelde verblijftijd in de kristallisator van circa 20 minuten.

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Omdat fluor een waardevol byprodukt is, is nog een apart onderzoek opgezet om de verdelingscoefficient van fluor tussen het fosforzuur en de lucht te bepalen. Een uitdrukking voor deze fluorverdelingscoefficient is afgeleid, die als functie van de temperatuur en van het fluoride, het fosfaat en het zwavelzuur gehalte van het fosforzuur, voorspelt wat de concentratie in de lucht zal zijn.

Tenslotte is een massa en warmte balans opgesteld voor een fabriek, die volgens het STFP werkt en 1000 ton P_0C per dag produceert. Hierbij zijn gegevens uit de hierboven beschreven onderzoeken gebruikt en aangevuld met resultaten van additionele proeven en gegevens uit de literatuur. Deze fabriek blijkt ongeveer in zijn eigen energie behoefte te kunnen voorzien.

Als laatste onderwerp worden de resultaten gepresenteerd van een onderzoek met een fosforzuur fabriekje op laboratorium schaal. In dit fabriekje is gekeken naar de invloed van verontreinigingen op de cadmium en fosfaat inbouw in het HH. Tevens is gekeken naar de filtreerbaarheid van dit HH. De invloed van verontreinigingen op de cadmium en fosfaat inbouw was erg klein, maar de perméabiliteit van de HH koek was ongeveer een factor vijf lager dan van de HH koek, die geproduceerd werd in de afzonderlijke kristallisatie proeven, waarbij chemisch zuivere reagenten gebruikt werden.

Alhoewel een pilot plant stadium zeker niet overgeslagen kan worden, is wel vast komen te staan dat het STFP in principe een haalbaar proces is.

De haalbaarheid van de cadmium verwijderingsstap zal op een andere plaats gepresenteerd worden (Proefschrift T.T. Tjioe, Technische Universiteit Delft, 1987).

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PUBLICATIONS

Various chapters of this Thesis have been published or are submitted for publication:

A Clean Technology Phosphoric Acid Process S. van der Sluis, Y. Meszaros, J.A. Wesselingh and G.M. van Rosmalen. Proceedings of the Fertiliser Society of London, nr 219, 1986.

The Predigestion Stage of a New Phosphoric Acid Process S. van der Sluis, Z. Murach, J.A. Wesselingh and G.M. van Rosmalen. Proceedings of World Congress III of Chemical Engineering, Tokyo, Japan, Volume U (1986) 96.

The Digestion of Phosphate Ore in Phosphoric Acid S. van der Sluis, Y. Meszaros, W.G.J. Marchee, J.A. Wesselingh and G.M. van Rosmalen. Submitted to Industrial and Engineering Chemistry Research.

Crystallisation of Calcium Sulfate in Concentrated Phosphoric Acid S. van der Sluis, G.J. Witkamp and G.M. van Rosmalen. Journal of Crystal Growth 79 (1986) 620.

Fluoride Distribution Coefficients in Wet-Phosphoric Acid Processes S. van der Sluis, J.H.M. Schrijver, F.P.C. Baak and G.M. van Rosmalen. Submitted to Industrial and Engineering Chemistry Research.

Mass and Heat Balances of a Clean Technology Phosphoric Acid Process S. van der Sluis, J.M.P. Oomens, Y. Meszaros, J.A. Wesselingh and G.M. van Rosmalen. The Second International Symposium on Phosphogypsum, Miami, Florida, December 10-12 (1986).

202