fmc P rocess

FMC was a major player in phosphates in the United States and through its Spanish operation, FMC Foret in Huelva, in Europe. In the United States, its business was founded on phosphorus operations in Idaho. Foret, however, relied on wet acid for its processes and installed a small purification plant based on IMI technology in the early 1970s. FMC had undertaken a process development program in the 1960s in the United States and was granted a patent [45] utilizing a 3:1 by volume TBP/kero-sene mixture as solvent, columns for extraction and stripping, and solvent regenera-tion. The process described was capable of producing a water white acid to meet food grade standards. In the 1970s, FMC evaluated different solvent extraction

TABLE 2.13

Budenheim Process Stream Composition

Premixer–Settler Stream Postmixer–Settler Stream Final Product

% P₂O₅ 9.0 8.9 54.0

% isopropanol 72.3 72.3

ppm SO₄ 20 50 30

ppm F 10 50 20

ppm Na 400 10

ppm Fe 10 1 5

ppm Al 90 1 5

ppm Ca 10 1 5

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equipment including Karr, Scheibel, and rotating disk contractor (RDC) columns and Podbielniak centrifugal contactors. In the early 1980s, further extensive labora-tory and pilot-scale study was carried out to facilitate the design of the new purifica-tion plant at Huelva. During the 1980s, FMC was granted a number of patents. Of note were improvements to the scrubbing (washing) stage [46] by the use of metal phosphates (one example was the addition of sodium ions, thus creating an aqueous stream with a Na/P ratio of 0.05–0.14) and solvent regeneration [19]. Further patents were filed, extending the use of sodium ion addition creating a sodium phosphate stream suitable for conversion to STPP [47] and modifying the solvent for improved efficacy by replacing kerosene with diisobutyl ketone (DIBK). None of these later patents were implemented on an industrial scale.

In 1983, a 12,000 tpa P₂O₅ plant was installed at Foret, principally to supply the STPP plant. Figure 2.41 shows a block diagram of the process; Figure 2.42 is a photo-graph of the plant itself. Other views are currently available via a well-known search engine. Feed acid came from the existing wet acid plant on-site and was pretreated to

Storage

NaOH

NaOH TBP/kerosene

RO water

Condensate Steam

Vacuum concentration Defluorination

Decolorization Filtration Raffinate

treatment

Solvent storage

Solvent regeneration Feed acid

16,000 tpa P₂O₅ 47%–52% P₂O₅ acid

Fertilizer grade

acid 4000 tpa P₂O₅ Product acid

12,000 tpa P2O5

28%–37% P₂O₅ purified acid

storage 3

2 1

H₂O₂

FIGURE 2.41 FMC Foret Huelva plant block diagram.

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reduce sulfate, arsenic and other heavy metals, and fluoride levels to those compat-ible with the final product specifications. Finally, the feed acid was concentrated to 47%–52% P₂O₅.

Feed acid at 50°C entered the top of the extraction column, a Kühni column with 30 compartments equivalent to 5 mixer–settler stages, and descended the column contacting the TBP/kerosene solvent. The solvent to feed acid ratio was approxi-mately 4:1 by weight. The organic stream leaving the top of the column was 12%

P₂O₅ with Fe and SO₄ levels less than 50 ppm. The aqueous raffinate stream leaving the column base was 20% P₂O₅ with Fe and SO₄ levels close to 1%.

The organic stream was pumped via intermediate storage to a cascade of three coun-tercurrent mixer–settlers and scrubbed with a sodium-enriched aqueous stream. The aqueous stream flowed back to the extraction column via the feed storage, the organic stream forward to stripping in a second Kühni column entering at the column base.

The scrubbed organic stream contacted water countercurrently in the column and phosphoric acid was transferred from the organic to aqueous phase. The ratio of water and organic flow was controlled to give an acid concentration in the range 28%–37% P₂O₅, depending on the desired purity of the final product acid.

During production, a side stream of solvent passed continually to the solvent regen-eration unit as the TBP became steadily contaminated with silica, organic species, FIGURE 2.42 FMC Foret Huelva PWA 2004 (author’s photograph).

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sulfides, etc., and itself started to hydrate to DBP. The solvent regeneration unit com-prised three mixer–settlers where firstly, the solvent was contacted with water to remove phosphoric acid and soluble impurities; secondly, the solvent was contacted with sodium hydroxide to remove any remaining contaminants into the aqueous phase—this phase then passed to effluent treatment; and thirdly, the treated solvent contacted the aqueous phase from the first mixer–settler neutralizing any sodium hydroxide from the second mixer–settler. The regenerated solvent is pumped to the solvent storage.

The weak product acid passed to a double-effect evaporator for concentration to 54% then to a single-effect evaporator to achieve 62% P₂O₅.

To achieve fluoride and color specification, food grade acid passed first to a deflu-orination column. The defludeflu-orination column was PTFE lined and fitted with graph-ite plates operating under vacuum. The acid was steam stripped to reduce fluoride levels to <10 ppm.

Finally, the defluorinated acid was mixed in a static mixer with hydrogen perox-ide to remove color bodies leaving a water white food grade acid.

The Foret plant operated satisfactorily until its closure in 2010 due largely to the decline of the STPP market in Europe.

Toward the end of the 1990s, FMC’s phosphorus operation in Idaho was faced with expensive environmental upgrades. Consequently, given most of the phosphorus was subsequently burned to make phosphoric acid for STPP and other phosphates, a purified acid plant project was evaluated. FMC chose to apply its own technology, developed at its technical center in Princeton and implemented in Huelva, in Idaho at the site of a wet acid and fertilizer producer, Agrium.

The original plan was to build an 80,000 tpa P₂O₅ single-train plant that would take pretreated acid at 52% P₂O₅, purify it to food grade, and return the raffinate to site for ammonia fertilizer (MAP/DAP) production. During discussions with the site owners, FMC agreed to design and build the pretreatment stages and utilities. The feed acid was now an untreated 28% P₂O₅ made from a Western States ore. Broadly speaking, the standout characteristics of this rock are high organics and vanadium content. Prior to the date of implementation, the rock was calcined to destroy organ-ics as part of the rock treatment processes prior to acidulation to wet acid; however, a combination of environmental and cost drivers led to the decision to use uncalcined rock and manage the organics in the wet process chemically.

The block diagram of FMC’s process at Idaho is shown in Figure 2.43.

The overall pretreatment process goals were to reduce sulfate levels to 0.2%, fluoride to 0.3%, arsenic to less than 1 ppm, and cadmium to less than 15 ppm.

120,000 tpa 28% P₂O₅ was pumped via acid coolers to the desulfation tank where it was mixed with 4000 tpa P₂O₅ phosphate rock slurry. Other sulfate-reducing agents were considered including soda ash but rock was chosen on price and availability criteria. The site received the rock as slurry via pipeline; consequently, it was only available in this form. The quantity required on the purified acid plant for desul-fation was a small percentage of the total site demand, and in practice the control of the rock addition, and therefore sulfate and other impurity levels, proved difficult.

Following desulfation, the acid was clarified and the slurry from the clarifier filtered on a belt filter with condensate washing. The solid waste from belt filtration was col-lected with other filter waste for disposal; in total P₂O₅ losses through filtration were

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significant; the diagram simplifies total P₂O₅ loss, ascribing it to filtration alone. The clarified acid was then pumped to the 52% P₂O₅ vacuum-pumped circulation con-centrator with double-effect steam ejectors creating the system vacuum. Following concentration, the acid was cooled and held in intermediate storage.

The cooling of 52% P₂O₅ wet acid always causes the precipitation of complex phosphates. Therefore, following storage, the acid passed to a large vacuum drum filter that was precoated with diatomaceous earth. Filter cake, loaded with com-plex phosphates and some free acid, was sent for disposal with cake from the belt filter. The filtered acid was pumped to the sulfiding unit where it contacted sodium

Phosphate

FIGURE 2.43 FMC Idaho plant block diagram.

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hydrogen sulfide in a pressurized reactor to precipitate heavy metals, principally arsenic and cadmium, passed through a hydrogen sulfide stripper, and was subse-quently filtered under pressure. In the filtration step, both filter aid and activated carbon were added to the acid.

The now pretreated acid was then held in another intermediate storage with integral heater to ensure the acid was at 55°C when introduced to the top of the 2.6  m diameter extraction column, supplied by Kühni with 42 compartments.

TBP/kerosene solvent was pumped to the base of the extraction column; flow ratio control ensured that the acid/solvent feed ratio was approximately 5 by volume.

Approximately 80% of the acid migrated to the solvent; the aqueous phase raf-finate containing 21% P₂O₅ passed to raffinate treatment and then to the fertilizer plant. The loaded solvent was pumped to a series of six mixer–settlers and was scrubbed with an aqueous stream of phosphoric acid from the stripping column with a solvent-to-aqueous ratio of 26. Additionally, liquid ammonia was added to the mixer–settlers to increase scrubbing efficiency. Ammonia was chosen to aid  the scrubbing process as it would contribute to the ammonia content of the MAP/DAP fertilizers, whereas sodium hydroxide scrubbing would not and would raise the sodium level in the fertilizers to unacceptable levels. The scrubbed loaded solvent was then pumped to the stripping column, also supplied by Kühni. Reverse osmosis–treated water entered the upper section of the column and contacted the loaded solvent passing up the column, stripping it  of the acid. Purified acid passed out of the column base at 36% P₂O₅ concentration,  of which about 10%

was pumped back to scrubbing and 90% pumped forward to concentration. The acid was concentrated up to about 63% P₂O₅ in a double-effect vacuum concentra-tor in two stages, the first to 50% P₂O₅, at 75°C, and the second at 105°C, both at 14.3 kPa vacuum. Hot acid was then pumped forward, via an economizer into the defluorination column where it was contacted with live steam. The column was held at an operating temperature to 160°C and reduced fluoride levels to less than 10 ppm, required for food grade acid. Acid from the defluorination column was cooled passing through the economizer and mixed in line with hydrogen peroxide to decolorize the acid. Finally, the acid was pumped through a polishing filter to final product storage ready for dispatch.

Although it started up on time in the spring of 2000, the plant underwent an extended and troubled commissioning and never achieved nameplate rates reliably.

This is surprising from a technical standpoint, but there were a number of reasons for the lack of progress. The basic FMC process, as demonstrated in Spain, was good;

the Idaho flow sheet was perfectly satisfactory if somewhat pinched in places, and much of the equipment was identical to that used on the other purified acid plants at Aurora and Geismar and operated well. Nevertheless, the commissioning team faced challenges on all fronts:

1. Firstly, the project scope included the provision of most utilities (steam, air, RO water, nitrogen, cooling water), all of which had to be commissioned and brought their own challenges.

2. Secondly, the pretreatment section was difficult to operate because the phosphate rock slurry system was difficult to control; in turn, this led to

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poor control of the clarification unit that in part led to more frequent block-ages in the concentration system than would normally be the case; further-more, the drum filter proved problematic which may have been due both to aspects of its mechanical design and/or the duty required of it. These problems built on each other as well as hampering the sulfiding operation so that in the end the treated acid ready to feed the solvent extraction section was often of poor quality; indeed, at times so poor, it was remarkable that any purified acid was made.

3. The third challenge was that the host company was expanding and commis-sioning its own wet acid production at the same time and making changes to its rock feed. Consequently, at least in the early days of commissioning, the purified acid plant did not receive a stable wet acid feed of consistent composition.

The net effects of these challenges were very high P₂O₅ and solvent losses and low output. Despite these challenges, the plant operation did steadily improve;

however, during this period, the whole industrial phosphate industry was undergo-ing major change. FMC and Solutia (itself spun out of Monsanto) formed a joint venture, Astaris, which comprised most of the phosphate operations of the two companies. The Idaho plant was included in the joint venture. At the same time, overcapacity in the STPP market led to a loss in sales for Astaris and consequent drop in internal demand for acid to make STPP. The final nail in the coffin for the Idaho plant was the expansion by PCS of its purified acid plant, with the lowest costs in the United States. Taking all these factors into account, Astaris decided to close the plant in 2003.