CA2032627C - Process for producing sodium carbonate and ammonium sulphate from sodium sulphate - Google Patents
Process for producing sodium carbonate and ammonium sulphate from sodium sulphateInfo
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- CA2032627C CA2032627C CA 2032627 CA2032627A CA2032627C CA 2032627 C CA2032627 C CA 2032627C CA 2032627 CA2032627 CA 2032627 CA 2032627 A CA2032627 A CA 2032627A CA 2032627 C CA2032627 C CA 2032627C
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- sodium
- sulphate
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- C—CHEMISTRY; METALLURGY
- C01—INORGANIC CHEMISTRY
- C01D—COMPOUNDS OF ALKALI METALS, i.e. LITHIUM, SODIUM, POTASSIUM, RUBIDIUM, CAESIUM, OR FRANCIUM
- C01D7/00—Carbonates of sodium, potassium or alkali metals in general
- C01D7/18—Preparation by the ammonia-soda process
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- C—CHEMISTRY; METALLURGY
- C01—INORGANIC CHEMISTRY
- C01C—AMMONIA; CYANOGEN; COMPOUNDS THEREOF
- C01C1/00—Ammonia; Compounds thereof
- C01C1/24—Sulfates of ammonium
- C01C1/244—Preparation by double decomposition of ammonium salts with sulfates
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- C—CHEMISTRY; METALLURGY
- C01—INORGANIC CHEMISTRY
- C01D—COMPOUNDS OF ALKALI METALS, i.e. LITHIUM, SODIUM, POTASSIUM, RUBIDIUM, CAESIUM, OR FRANCIUM
- C01D7/00—Carbonates of sodium, potassium or alkali metals in general
- C01D7/02—Preparation by double decomposition
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- Engineering & Computer Science (AREA)
- Materials Engineering (AREA)
- Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)
- Fertilizers (AREA)
Abstract
A new continuous process for producing sodium carbonate and ammonium sulphate from naturally occurring sodium sulphate is disclosed. The process comprises reacting sodium sulphate in aqueous solution with ammonia and carbon dioxide to precipitate sodium bicarbonate which is separated by filtration and converted by calcining to sodium carbonate. The mother liquor from the precipitation of sodium bicarbonate is concentrated by evaporation to precipitate unreacted sodium sulphate, cooled to precipitated ammonium sulphate and further cooled to precipitate a double salt of sodium and ammonium sulphate. The double salt is added to the mother liquor from the precipitation of sodium bicarbonate, prior to the evaporation, whereas the mother liquor from the precipitation of the double salt is concentrated by evaporation and added to the mother liquor from the separation of sodium sulphate. The process, in addition to producing sodium carbonate, also produces ammonium sulphate in a purity such that it can be immediately used as a fertilizer. The process does not produce any unwanted byproducts and avoids the use of complicated equipment, such as absorption towers.
Description
PROCESS FOR PRODUCING SODIUM CARBONATE
AND AMMONIUM SULPHATE FROM SODIUM SULPHATE
This inventlon relates to manufacturing of sodium car-bonate (soda ash). More particularly, this invention relates to manufacturing of sodium carbonate and ammonium sulphate.
Sodium carbonate (Na2CO3) is a white crystalline hygroscopic powder known in the chemical trade as ash, soda ash, soda and calcined soda. Most of the world's soda ash is produced by the ammonia-soda process, also known as the Solvay process. In this process soda ash is produced from common salt (sodium chloride), ammonia, carbon dioxide and limestone by a sequence of reactions involving recovery and reuse of practically all the ammonia and part of the carbon dioxide. In what is conventionally considered as the first step, limestone is calcined in a kiln to produce lime and carbon dioxide. Carbon dioxide is dissolved in a purified brine containing additionally the ammonia, with resultant precipitation of sodium bicarbonate. The latter is filtered off, dried and thermally decomposed, at about 200C, to sodium carbon-ate. The comparatively expensive ammonia is recovered from ammon-ium chloride contained in the filtrate by adding to the heatedfiltrate a slurry of lime. Calcium chloride produced in this step is a major byproduct.
The Solvay process suffers from several disadvantages.
Firstly, the process consumes large amounts of fuel (to burn limestone, calcine sodium carbonate and produce steam for recovery of ammonia from ammonium chloride). Secondly, the chlorine from common salt is not recovered but discarded in the form of calcium chloride effluent, which pollutes the environment. Largely because of high energy costs and strict pollution controls, most of the synthetic production of soda ash by the Solvay process in North America has been or is being abandoned in favour of the product obtained from natural deposits.
The most abundant natural form of sodium carbonate is trona (Na2CO3.NaHCO3.2H2O). Two processes simpler than the Solvay process are used to refine trona ore. In the first, so called monohydrate process, the trona ore is calcined to impure soda ash which is then purified. In the second, a sodium sesquicarbonate process, the soda ash is produced by calcination of sodium sesquicarbonate which had previously been purified. Both pro-cesses are characterized by relatively favourable production costs and reduced environmental hazards.
Soda ash is also produced in the United States from the natural brine at Searles Lake, California. More complicated processes are required in this case than for processing of trona, because of the complex nature of the brines. In one of the pro-cesses, brine is evaporated to give sodium carbonate in the form of burkeite (Na2CO3.2Na2SO4), which is then separated in a complex sequence of fractional crystallizations into sodium carbonate and sodium sulphate. In another process brine is first treated with carbon dioxide in carbonation towers, the precipitated sodium bicarbonate is filtered off, washed and converted by heating to soda ash.
Several processes described in the prior art, but not commercialized, propose to make use of sodium sulphate for sodium carbonate. Canadian Patent No. 1,099,892 discloses a process for dry conversion of alkali metal sulphate to alkali metal carbonate.
In the process alkali metal sulphate is reduced with carbon to alkali metal sulphide which is subsequently converted with carbon dioxide and steam to alkali metal carbonate. Similar processes comprising the step of reducing alkali metal sulphate with carbon are disclosed in U.S. Patents No. 1,979,151 and 3,127,237. A
process for producing an alkali metal carbonate by reacting solid alkali metal sulphate with a gas containing hydrogen and carbon monoxide is described in U.S. Patent No. 3,401,010.
U.S. Patent No. 3,134,639 discloses a wet process for converting alkali metal sulphates into alkali metal carbonates using lime, hydrogen sulphide and carbon dioxide. No ammonia is used in this process.
U.S. Patent No. 3,493,329 discloses a wet process for producing soda ash from sodium chloride which process eliminates the use of calcareous material. The process comprises a liquid state cycle and a solid state cycle. A solution of sodium sul-phate is used in the liquid state cycle to precipitate sodium bicarbonate under conditions similar to those of ammonia-soda 20 process. Ammonium sulphate produced in this step is then precipi-tated in the following dissolving-crystallizing step by adding solid sodium sulphate to the solution from which the precipitated sodium bicarbonate has been separated. The ammonium sulphate precipitated by addition of the solid sodium sulphate, either pure or in the form of a double salt of ammonium sulphate and sodium sulphate, is reacted in its solid form in the solid state cycle by heating it in the presence of sodium sulphate to produce ammonia and sodium bisulphate which is subsequently reacted in the form of fused salt with sodium chloride to produce regenerated sodium sulphate and hydrochloric acid. It should be noted that in this process sodium chlorlde is the actual process feed material, whereas sodium sulphate is in fact used only for recovery of ammonia and is itself finally recovered and reused.
Canadian Patent No. 543,107 discloses a process for the recovery of ammonium sulphate from the filtrate obtained by fil-tering off sodium bicarbonate precipitated in the ammonia-soda process using sodium sulphate as a starting material. Ammonium sulphate is precipitated from the filtrate by dissolving in it ammonia under superatmospheric pressure and filtering off the precipitated salts under the ammonia pressure. However, the product so obtained contains less than 75% of ammonium sulphate and over 25% of sodium sulphate. Such a product has a very little practical use because it is a mixture of salts. It cannot be used, for example, as a fertilizer because of the high percentage of sodium sulphate. The patent does not provide any indication whether and how this mixture of salts can be separated to produce essentially pure and therefore marketable ammonium sulphate.
A process for the separation of sodium sulphate and ammonium sulphate from aqueous solutions of a mixture of them to obtain substantially pure sulphates is disclosed in Canadian Patent No. 821,457. The process uses an evaporator-crystallizer system which produces a crystal mixture in which crystals of each salt have different granulometric and density characteristics.
This mixture is separated, e.g. by dry or wet screening, into fractions consisting mainly of sodium sulphate and ammonium sul-phate crystals, respectively, which fractions are subsequently washed to dissolve away the sulphate present in the least quantity and leave the other sulphate substantially pure. The process, even though finally producing substantially pure sodium and ammonium sulphate from a mixture thereof, is lengthy and requires a rela-tively complicated equipment to separate the mixture of crystals obtained in the evaporation-crystallization step.
Demonstrated worldwide demand for sodium-based chemicals, particularly sodium carbonate (soda ash) has been on the rise in recent years. This strong demand, which is forecast to continue, keeps soda ash in a tight supply position thereby holding the price at a high level. Because of environmental concerns and the vast reserves of natural sodium carbonate, the production of soda ash in the United States is primarily from natural sources. In Canada, the known natural reserves of sodium carbonate are not as vast as those in the United States. However, Canada is endowed with vast deposits of sodium sulphate, located mostly in Southern Saskatchewan, that could potentially become a source of soda ash.
Moreover, statistics indicate that the demand for sodium sulphate is dwindling and its prices declining. This declining trend in the demand for and prices of sodium sulphate together with strong demand for and relatively high prices of soda ash created a need for an economical process for producing sodium carbonate (soda ash) from naturally-occurring sodium sulphate as feedstock.
Thus, the present invention provides a continuous pro-cess for preparing sodium carbonate, which process comprises reacting sodium sulphate in aqueous solution with ammonia and carbon dioxide to precipitate sodium bicarbonate, separating the sodium bicarbonate and calcining it to convert it to sodium car-bonate, sujecting the mother liquor from the precipitation of the sodium bicarbonate to evaporation to precipitate unreacted sodium sulphate, cooling the mother liquor to precipitate ammonium sul-phate, further cooling the mother liquor to precipitate a double salt of sodium sulphate and ammonium sulphate and adding the double salt to the mother liquor from the precipitation of sodium bicarbonate, prior to the evaporation.
The present invention has the considerable advantage that, in addition to producing sodium carbonate, it also produces ammonium sulphate in a purity such that it can immediately be used as a fertilizer. In contrast to the Solvay process, it is not necessary to seek to recycle all the expensive ammonia used in the process; a substantial amount of that ammonia appears in a pro-duct, ammonium sulphate, of high commercial value. Also, the process does not produce any unwanted byproducts, such as calcium chloride produced in the Solvay process. The products of the process are the required sodium carbonate, ammonium sulphate that can be used as fertilizer, a double salt of sodium and ammonium sulphate that is recycled and therefore does not accumulate, and sodium sulphate that can be recycled. As the sodium sulphate recovered in the process is of enhanced purity when compared with the raw salt normally used as starting material, it may be a valuable commercial product. Another advantage over the Solvay process is the avoidance of absorption towers, which have a tendency to become clogged with precipitated bicarbonate.
To carry out the precipitation of sodium bicarbonate according to the invention, any continuously fed reactor capable of operating under superatmospheric pressure and assuring an efficient gas-in-liquid dispersion can be used. In a preferred embodiment a closed, cylindrical reactor equipped with two gas dispersion impellers mounted on a coaxial shaft is used. The precepitation is normally conducted under superatmospheric pressure of from 0 kPa to about 700 kPa. A pressure of about 70 kPa is preferred.
Since the reaction generates heat, the reactor is equipped with a cooling system to maintain a temperature in an optimum range preferably from about 20C to about 60C. A tem-perature of about 40C is particularly preferred.
A sodium sulphate solution having a concentration of from about 400 g/L to about 500 g/L is continuously fed into the reactor. A low temperature of the brine fed into the reactor helps in maintaining the content of the reactor in the optimum temperature range. However, since the solubility of sodium sul-phate in water decreases with decreasing temperature, this tem-perature should not be excessively low, to keep the solubility of sodium sulphate in the indicated range of concentrations. A
temperature for the brine solution in the range of from 20C to 60C is suitable. In a preferred embodiment a saturated solution of sodium sulphate (concentration of about 488 g/L) having tem-perature of about 40C is used.
2032627 608l8-l7 Carbon dioxide and ammonia are fed into the reactor in the liquid and/or gaseous form. According to a preferred embodi-ment, a major part of the two gases is injected into the brine in the liquid form which has the advantage of producing a cooling effect compensating for the exothermic effect of the processes taing place in the reactor.
It is also preferred that a gaseous mixture of ammonia and carbon dioxide recovered from the brine after the precipita-tion of sodium bicarbonate is recycled into the reactor. It is therefore preferred to heat the mother liquor remaining after the precipitation of the sodium bicarbonate, to recover dissolved or unreacted carbon dioxide and ammonia before subjecting the liquor to evaporation.
The pH of the content of the reactor is maintained in a range of from about 7.0 to about 9Ø The value of pH of about 8.0 is preferred. The value of pH can be maintained by control-ling the supply of liquid carbon dioxide and/or ammonia.
In a preferrred embodiment the mother liquor from the precipitation of the double salt is added to the mother liquor from the precipitation of sodium sulphate. The mother liquor from the precipitation of the double salt can be subjected to concen-tration, for example by evaporation, before it is added to the mother liquor from the precipitation of sodium sulphate.
The mother liquor from the precipitation of sodium sulphate may be alternatively added to the liquor from the preci-pitation of sodium bicarbonate, before this liquor is evaporated.
However, it is preferred to reduce the amount of the recycled salts by cooling the mother liquor from the precipitation of sodium sulphate to precipitate ammonium sulphate, further cooling the mother liquor to precipitate a double salt of sodium sulphate and ammonium sulphate and adding this double salt to the mother liquor from the precipitation of sodium bicarbonate, prior to the evaporation of this liquor.
It is preferred that the mother liquor in the first evaporation stage shall be at a temperature not lower than about 80C, preferably not lower than about 95C, so that the precipi-tated sodium sulphate is not contaminated with any significantamount of sodium sulphate. It is preferred that the precipitation of ammonium sulphate shall take place at a temperature above about 40C, preferably above about 60C, so that is is not contaminated with any significant amount of the double salt. It is possible to include a concentration step between the precipitation of the sodium sulphate and the ammonium sulphate. It is also possible to include a concentration step between the precipitation of the ammonium sulphate and the double salt.
The various evaporation, precipitation and concentration stages can be carried out under reduced pressure, in which case temperatures different from those mentioned above will be appro-priate. Determing the appropriate temperatures for a particular chosen reduced pressure will be within the competence of a person skilled in the art.
The invention will be further described by way of a preferred embodiment and with references to the accompanying drawings in which:
Figure 1 represents schematically a process for purifi-cation of raw sodium sulphate and preparing sodium sulphate brine, Figure 2 represents schematically a process for produc-ing soda ash according to one preferred embodiment of the inven-tion, Figure 3 represents schematically a process for recover-ing ammonium sulphate and sodium sulphate according to one pre-ferred embodiment of the invention, Figure 4 represents schematically an experimental bench-scale reactor for conducting the process according to the inven-tion, and Figure 5 represents a flow diagram of the experimental system using the reactor of Figure 4.
The process according to the invention requires as a feed a solution of sodium sulphate. According to a preferred embodiment of the invention, this solution is prepared from raw mined sodium sulphate, in a process shown schematically in Figure 1. Raw sodium sulphate (Na2SO4.10H2O) is continuously fed into a classifier 1 and mixed with hot brine (an overflow from cyclone 5). The heat from the brine is sufficient to dissolve the crystals of sodium sulphate. Any stones and dirt settle to the bottom of the classifier 1 and are discarded as waste. The salt brine from the classifier 1 is put through a 30 mesh screen 6 to remove any fine particles suspended in the brine.
Once screened, the brine is transferred by a pump 7 to two submerged combustion evaporators No. 1 and 2 to evaporate some water, thus causing some sodium sulphate to crystallize out. An air and natural gas mixture which is burned under the brine in a burner tank 2 causes violent bubbling action. This causes some brine to splash over a weir and run down into a settling tank 3.
Here the crystals of sodium sulphate settle towards the bottom of the tank and are drawn off by a pump 10 and put through cyclones 5 to spin off more of the brine. This brine is the cyclone overflow which is used to dissolve the raw salt in the classifier 1. The solid sodium sulphate coming out of the cyclones 5 contains about 20% of water and is pumped by a pump 11 into a rotary dryer 55 shown in Figure 3 to finish drying the salt and produce sodium sulphate (salt cake).
The brine that comes off the top of the settling tank 3 is normally recycled to the burner tank 2 for further evaporation when only sodium sulphate is to be produced. In the process according to the invention, some of this brine is transferred by a pump 8 into a thickener 4, where a flocculating agent, for example Percol 156, is added to remove fine clay particles still suspended in the brine. The flocculating agent causes the clay particles to stick together and settle out faster at the bottom of the thick-ener 4, to be removed as a mud. The resulting brine is a cleansaturated solution of sodium sulphate (concentration of about 30%) in water having temperature of about 40C. This brine is fed into the reactor 21 shown in Figure 2 to be mixed with carbon dioxide and ammonia.
A cylindrical reactor 21 is equipped with gas dispersion impellers 22 and 23 and cooling coils 24. For better agitation cooling coils in the form of single vertical tubes are placed radially along the reactor wall. These tubes also act as baffles when the content of the reactor is agitated.
Liquid carbon dioxide and liquid ammonia are injected into the brine through injection nozzles 24 and 25, respectively.
A part of carbon dioxide and ammonia used in the process is sup-plied to the reactor in the gaseous form through a tube 26, as a mixture recovered from the gas recovery boiler 31, as explained below.
As carbon dioxide and ammonia dissolve in the brine, the following reactions take place:
NH3 + H2O = NH40H
Na2SO4 + 2NH4HCO3 = 2NaHCO3 + (NH4)2SO4 These reactions produce considerable heat which is to some extent compensated by the cooling effect produced by injecting liquid carbon dioxide and ammonia into the brine. Any excess heat is removed by the cooling coils 24 supplied with cooling water having temperature of about 7C to maintain the temperature of the reac-tor at an optimum level of about 40C. The pressure maintained in the reactor 21 is about 70 kPa. The pH of the content of the reactor is maintained at about 8 by controlling the supply of carbon dioxide.
Because of its limited solubility in water, sodium bicarbonate (NaHCO3) produced in the above shown series of reac-tions precipitates out. It is continuously removed from the reactor 21 in the form of a slurry by a pump 35 and transferred to a brine cooler 27, where the slurry is cooled to about 20C to 20~2627 60818-17 further reduce the solubility of sodium bicarbonate and to cry-stallize out as much of the product as possible. The sodium bicarbonate is then screened out of the brine on a 21 micron screen 28 and subsequently washed on a screen 29 with water to remove any ammonium and sodium sulphate present in the entrained brine. The washed sodium bicarbonate is then fed into a rotary dryer 30 to be dried and calcined to sodium carbonate (soda ash).
The brine remaining after screening off the solid sodium bicarbonate contains a mixture of unreacted sodium sulphate, ammonium sulphate, ammonium bicarbonate and minor amounts of sodium bicarbonate. This brine is transferred by a pump 36 into a gas recovery boiler 31 where it is heated to a temperature of 95 to 100C. Under these conditions the ammonium bicarbonate breaks down and sodium bicarbonate dissolved in the brine reacts with ammonium sulphate to produce sodium sulphate, carbon dioxide and ammonia. Carbon dioxide and ammonia dissolved in the brine boil off, leaving in the solution a mixture composed mostly of sodium and ammonium sulphate. The carbon dioxide and ammonia so regener-ated are cooled in a gas cooler 32 and returned to the reactor 21 by a blower 33 after being further cooled in a gas cooler 34.
This regeneration step minimizes the amount of carbon dioxide and ammonia used in the process.
The brine remaining after removing ammonia and carbon dioxide has a temperature of 95 to 100C and contains approxi-mately 15% of sodium sulphate and 15% of ammonium sulphate. This brine is added by a pump 37 to double salt crystals separated on the screen 61 shown in Figure 3, as explained below. The 1:1 ratio of sodium sulphate to ammonium sulphate in the double salt is the same as the ratio of these two salts in the brine, so that the mixing of the double salt with the brine do not affect the proportion of the two salts in the brine.
The mixture of the double salt from the screen 61 and the brine from the gas recovery boiler 31 is then fed into the submerged combustion evaporator No. 3 where the brine is concen-trated up to the point where both salts reach their saturation point. Since ammonium sulphate has a much higher situation point than sodium sulphate, most of the latter crystallizes out before the brine is saturated with ammonium sulphate. When the brine approaches saturation also with respect to ammonium sulphate, the brine is drawn off from the settling tank 52 by a pump 65 and transferred onto a screen 53, where sodium sulphate crystals are screened off. These crystals are then washed on a screen 54 to remove any ammonium sulphate present in the entrained brine.
Although sodium sulphate so recovered could be redissolved in water and returned to the reactor 21, this product is so pure that it is more economical to dry it to produce sodium sulphate. This is normally done by combining the solids from the screen 54 with the underflow of the cyclone 5 shown in Figure 1 and feeding the mixture into a dryer 55 for final drying.
The brine remaining after screening out sodium sulphate has a temperature of about 95C and is saturated with respect to both sodium and ammonium sulphate. This brine via a pump 67 is combined with brine from a submerged combustion evaporator No. 4 used to crystallize out ammonium sulphate. The brine from this evaporator is saturated in both salts and contains some crystals of ammonium sulphate. By cooling the mixture of both brines in a crystallizer 56 down from 95C to 60C, the solubility of ammonium sulphate is decreased while the solubility of sodium sulphate increases. This forces more ammonium sulphate to crystallize out while keeping sodium sulphate in the solution. The brine is then drawn off from the crystallizer 56 by a pump 68 and transferred onto a screen 57, where ammonium sulphate crystals are screened off. The crystals are then washed on a screen 58 to remove any sodium sulphate from the entrained brine and finally dried in a dryer 59.
After removing the ammonium sulphate crystals, the brine is transferred by a pump 69 to a double salt crystallizer 60 where it is further cooled to about 15C. At this temperature sodium sulphate and ammonium sulphate crystallize out in the form of a double salt (Na2SO4.(NH4)2SO4.2H20). The brine with suspended crystals of the double salt is transferred from the crystallizer 60 by a pump 70 onto a screen 61 where the solids are screened out and added to the brine from the gas recovery boiler 31 shown in Figure 2.
By removing the double salt from the brine the ratio of sodium sulphate to ammonium sulphate becomes 1:3 in the remaining brine. This brine is fed into submerged combustion evaporator No.
4 where it is concentrated until sodium sulphate reaches its saturation point, by which time some of the ammonium sulphate crystallizes out. At this point the brine passing from a burner tank 62 into a settling tank 3 is transferred from the settling tank 63 by a pump 66 into the crystallizer 56 where it is combined with the brine left after screening out sodium sulphate on the screen 53.
EXPERIMENTAL
The operating parameters of the process of the precipi-tation of sodium bicarbonate were studied and optimized in a bench-scale reactor unit.
Desiqn and Construction of the Laboratory Bench Unit The bench unit comprises three main sections: reaction system, feed system, and product recovery systems.
The reaction system consists of three main sections:
reactor, agitator, and cooling jacket.
A schematic of the reactor is shown in Figure 4. The reactor 1 is fabricated of a 0.076 m (3 in) diameter x 0.457 m (18 in) long mild steel pipe and has a volume of 2 litres. The reac-tor head 2 carries an agitator 3, a 0.0063 m (1/4 in) NPT feed inlet port 4, 0.0063 m NPT reaction gaseous vapour outlet port 5, and 0.0032 m (1/8 in) and 0.003 m NPT ports for thermocouple 6 and back-pressure gauge 7, respectively. The bottom of the reactor is a 0.051 m (1 in) long, 45-degree cone 8 to which is attached a 0.0095 m (3/8 in) diameter product drain connector 9. The reactor is designed for pressures up to 1034 kPa (150 psi) and tempera-tures up to 30C.
The feed mixer is a variable speed (640-930 rpm) motor-driven agitator 3 welded onto the reactor head 2 and centrally located in the reactor. Attached to the part of the agitator inside the reactor is a 0.069 m (2-3/4 in) wide by 0.28 m (11 in) long 16 x 16 mesh adjustable stainless-steel screen (not shown in the drawing). The screen enables good liquid/gas mixing by allow-ing the gas to pass freely through the screen pores to contact the feed.
The cooling jacket 10 is constructed of 0.102 m (4 in) diameter by 0.406 m (16 in) mild steel pipe, which is welded to the reactor in such a way that it is concentric with the reactor.
The coolant (glycol/water mixture) circulates through the system by the inlet and outlet ports (11 and 12, respectively).
The feed system consists of the feedtank and transfer line, metering pump and gas delivery manifold. The freshly pre-pared feed is siphoned into the feedtank 20 shown in Figure 5 from where it is fed to the reactor. To prevent crystallization of the feedstock, the tank is maintained at a constant temperature of about 40C. The feed is delivered to the reactor by means of a metering pump 21. The feed transfer line is 0.0063 m (1/4 in) diameter vinyl tubing heat traced at a constant temperature of about 40C and connected to the reactor by the metering pump.
Heat tracing the feed lines is necessary to prevent plugging problems caused by crystallization of sodium sulphate in the feed solution. The gaseous feed line is also heat traced to provide a constant feed rate.
Gas distribution within the reactor is accomplished by means of a manifolds 13 and 14. This system has eighteen gas inlet nozzles: nine for carbon dioxide and nine for ammonia. The manifolds are located near the bottom of the reactor so that the gas can be fed counter currently with the sodium sulphate solution, which is fed into the reactor from the top. The gas inlet nozzles 15 and 16 are arranged in a circle inside the reac-tor such that each gas is fed by an alternate nozzle. This design enables good dispersion of the gas mixture within the reactor.
Feedstock Preparation The solution of sodium sulphate used in all the experi-ments was prepared as follows:
Approximately 12.4 kg of fresh natural Glauber salt crystals (Na2SO4.10H2O) and 12 L of hot tap water were added into a metal pail and stirred continuously until the aqueous liquid was saturated. The temperature of the stirred mixture was maintained at about 40C by means of a drum heater wrapped around the pail.
After about two hours, the stirring was stopped and the mixture left to settle for approximately four hours at 40C to remove silt, salt crystals and other fine clays in the mixture The supernatant (clear sodium sulphate solutions) was withdrawn, without disturbing the sediments, and transferred to the feed storage tank 20. This decanting of the solution was accomplished by siphoning. A sample of the feed solution was analyzed to determine the sodium sulphate content. It was found to contain 2.13 g mol/L or 24.6 wt % of sodium sulphate. The density of the solution at 40C was determined to be 1229.7 kg/m3.
Experimental set-uP and Procedure A flow diagram describing the experimental system is illustrated in Figure 5. A typical procedure involves feeding the sodium sulphate solution to the reactor 22 by means of a metering pump 21 calibrated to maintain the desired flow rate of the solution. The feed solution is fed from the top of the reactor The carbon dioxide and ammonia gases at desired flow rates and pressures are metered with capillary flowmeters 23 and 24, respectively, into the reactor 22 through the gas manifold inlets.
The gas manifold is located near the bottom of the reactor en-abling the gases to flow countercurrently with the downflowing sodium sulphate liquid solution. The desired speed of the agita-tor (640-930 rpm) is set so that the reactor contents are ade-quately mixed.
The heat generated by the exothermic reactions of the liquid feed/gas mixture is removed by the circulating glycol/water mixture maintained at a constant temperature of about 20C. The liquid and gaseous reaction vapours are routed through the liquid trap 25 and back-pressure regulator (BPF) 26 into the vent. The sodium bicarbonate/ammonium sulphate product mixture is manually withdrawn from the coned-bottom of the reactor through the drain connector at intervals determined by a stopwatch (=51-53 mL/min).
Ideally the product recovery is accomplished at a uniform rate by means of a metering pump 27.
Process OPeratinq Conditions The experimental runs were conducted based on the fol-lowing reactor unit operating conditions:
Feed flow rate, mL/min 55 Ammonia flow rate, L/min 0.96-1.0 Carbon dioxide flow rate, L/min 1.1-1.2 Pressure, kPa 550-620 Temperature, C 23 Agitator speed, rpm 640-930 Product withdrawal, mL/min 51-53 The carbon dioxide flow rate is an estimated flow rate based on the ammonia rate because the CO2 flow exceeded the maxi-mum flow capacity of the flowmeter.Results The aqueous sodium sulphate feedstock was evaluated primarily for its effectiveness to convert to sodium bicarbonate when reacted with carbon dioxide and ammonia.
Crystallization of sodium sulphate in the feed solution occurred at relatively low temperatures (<40 C) necessitating heat tracing the feed lines (both liquid and gaseous lines) to prevent plugging problems.
Ammonia flow rate is an important parameter because it directly affected the quantity and quality of the product. For instance, excess amount of ammonia inflow resulted in the product being a mixture of sodium bicarbonate and sesquicarbonate instead of predominantly the less soluble sodium bicarbonate.
The process reactions were exothermic, requiring con-tinuous circulation of the glycol/water cooling mixture. An efficient cooling system had to be installed to provide the required cooling because of an excessive amount of ammonia gas and C2 had to be fed to the reactor under insufficient cooling condi-tions. This resulted in the formation of excessive quantities of intermediate product, ammonia bicarbonate/carbonate.
A vacuum filtration system and a medium speed (3000 rpm) centrifuge were used as a means of recovering the bicarbonates and carbonates from the product stream. The vacuum filtration system produced a purer product (0.86 wt % sulphate after only one wash-ing of 35 mL washing solution/70 g wet precipitates) than the centrifugation system after two washings (5.0 wt % sulphate). The washing solution was a saturated ammonium bicarbonate aqueous solution.
After the carbonate/bicarbonate products were removed (at 23C), the supernatant solution contained mainly 281 g/L
ammonium sulphate, and 65.8 g/L of sodium compounds (carbonates, sulphates, and bicarbonates). This solution was treated by two different processes to recover the ammonium sulphate. The first process used one litre of methanol per litre of solution, which recovered 157 g of precipitates per litre of fluid. The precipi-tates were found to contain 18 wt % of sodium. In the second process, ammonia gas was added to the solution. This process recovered 115 g of precipitates per litre of fluid. The precipi-tates were found to contain 28 wt % of sodium.
The experimental data are presented in Tables 1 and 2.
The percent recoverable carbonate/bicarbonate products were esti-mated from the product composition data as follows:
2032627 608l8-l7 A = 33.5 wt % NaHCO,/100 x wt % recoverable (Table 1) = 28.14 wt %
B = 65.2 wt. % Na2C03 . NaHC03 . 2H20/100 x wt %
recoverable (Table 2) = 43.0 wt %
C = 1.3 wt % Na2S04 Thus, the calculated % recoverable of products -A=B=C = 72.5 wt %.
The actual percentage of recovered products (Table 1) was 73.6 wt %.
The amount of sodium remaining in solution was 25.8 g/L.
When this solution is heated, the sodium bicarbonate/carbonate compounds react with the ammonium sulphate to produce Na2S04, C02 and NH3.
Table 1 Percentage Recoverable at 0C, Theoretical vs. Experimental % Recovery (Calculated Based Experimental Results O
on Solubility Data) (% Recovery Basis Na @ O C) Wt % (Basis Na) Compound at OC 1 st 2 hr of 2nd 2 hr of Trial Run Trial Run 74.0 Sodium Carbonate 70.0 Sodium Carbonate, decahydrate 73.6 73.5 66.0 Sodium Sesquica,l,on~le (Na2C03-NaHC03.2H20) 84.0 Sodium Bica,l,on~le 2 0 3 2 ~ 27 60818-17 Table 2 Percentage Recovery of Sodium at 23C and OC
SulphateSodium Content in Recovery pH ContentSu~,e",ala"l g/L (basis Na) @ 22C(wt %) @ 23C @ 0C @ 23C @ 0C
Supernatant from 7.9 - 65.9 25.8 32.6 73.6 product stream recovered in 1st two hours of Trial run Supernatant from 8.6 - 65.7 25.9 32.8 73.8 product stream recovered in 2nd two hours of Trial Run
AND AMMONIUM SULPHATE FROM SODIUM SULPHATE
This inventlon relates to manufacturing of sodium car-bonate (soda ash). More particularly, this invention relates to manufacturing of sodium carbonate and ammonium sulphate.
Sodium carbonate (Na2CO3) is a white crystalline hygroscopic powder known in the chemical trade as ash, soda ash, soda and calcined soda. Most of the world's soda ash is produced by the ammonia-soda process, also known as the Solvay process. In this process soda ash is produced from common salt (sodium chloride), ammonia, carbon dioxide and limestone by a sequence of reactions involving recovery and reuse of practically all the ammonia and part of the carbon dioxide. In what is conventionally considered as the first step, limestone is calcined in a kiln to produce lime and carbon dioxide. Carbon dioxide is dissolved in a purified brine containing additionally the ammonia, with resultant precipitation of sodium bicarbonate. The latter is filtered off, dried and thermally decomposed, at about 200C, to sodium carbon-ate. The comparatively expensive ammonia is recovered from ammon-ium chloride contained in the filtrate by adding to the heatedfiltrate a slurry of lime. Calcium chloride produced in this step is a major byproduct.
The Solvay process suffers from several disadvantages.
Firstly, the process consumes large amounts of fuel (to burn limestone, calcine sodium carbonate and produce steam for recovery of ammonia from ammonium chloride). Secondly, the chlorine from common salt is not recovered but discarded in the form of calcium chloride effluent, which pollutes the environment. Largely because of high energy costs and strict pollution controls, most of the synthetic production of soda ash by the Solvay process in North America has been or is being abandoned in favour of the product obtained from natural deposits.
The most abundant natural form of sodium carbonate is trona (Na2CO3.NaHCO3.2H2O). Two processes simpler than the Solvay process are used to refine trona ore. In the first, so called monohydrate process, the trona ore is calcined to impure soda ash which is then purified. In the second, a sodium sesquicarbonate process, the soda ash is produced by calcination of sodium sesquicarbonate which had previously been purified. Both pro-cesses are characterized by relatively favourable production costs and reduced environmental hazards.
Soda ash is also produced in the United States from the natural brine at Searles Lake, California. More complicated processes are required in this case than for processing of trona, because of the complex nature of the brines. In one of the pro-cesses, brine is evaporated to give sodium carbonate in the form of burkeite (Na2CO3.2Na2SO4), which is then separated in a complex sequence of fractional crystallizations into sodium carbonate and sodium sulphate. In another process brine is first treated with carbon dioxide in carbonation towers, the precipitated sodium bicarbonate is filtered off, washed and converted by heating to soda ash.
Several processes described in the prior art, but not commercialized, propose to make use of sodium sulphate for sodium carbonate. Canadian Patent No. 1,099,892 discloses a process for dry conversion of alkali metal sulphate to alkali metal carbonate.
In the process alkali metal sulphate is reduced with carbon to alkali metal sulphide which is subsequently converted with carbon dioxide and steam to alkali metal carbonate. Similar processes comprising the step of reducing alkali metal sulphate with carbon are disclosed in U.S. Patents No. 1,979,151 and 3,127,237. A
process for producing an alkali metal carbonate by reacting solid alkali metal sulphate with a gas containing hydrogen and carbon monoxide is described in U.S. Patent No. 3,401,010.
U.S. Patent No. 3,134,639 discloses a wet process for converting alkali metal sulphates into alkali metal carbonates using lime, hydrogen sulphide and carbon dioxide. No ammonia is used in this process.
U.S. Patent No. 3,493,329 discloses a wet process for producing soda ash from sodium chloride which process eliminates the use of calcareous material. The process comprises a liquid state cycle and a solid state cycle. A solution of sodium sul-phate is used in the liquid state cycle to precipitate sodium bicarbonate under conditions similar to those of ammonia-soda 20 process. Ammonium sulphate produced in this step is then precipi-tated in the following dissolving-crystallizing step by adding solid sodium sulphate to the solution from which the precipitated sodium bicarbonate has been separated. The ammonium sulphate precipitated by addition of the solid sodium sulphate, either pure or in the form of a double salt of ammonium sulphate and sodium sulphate, is reacted in its solid form in the solid state cycle by heating it in the presence of sodium sulphate to produce ammonia and sodium bisulphate which is subsequently reacted in the form of fused salt with sodium chloride to produce regenerated sodium sulphate and hydrochloric acid. It should be noted that in this process sodium chlorlde is the actual process feed material, whereas sodium sulphate is in fact used only for recovery of ammonia and is itself finally recovered and reused.
Canadian Patent No. 543,107 discloses a process for the recovery of ammonium sulphate from the filtrate obtained by fil-tering off sodium bicarbonate precipitated in the ammonia-soda process using sodium sulphate as a starting material. Ammonium sulphate is precipitated from the filtrate by dissolving in it ammonia under superatmospheric pressure and filtering off the precipitated salts under the ammonia pressure. However, the product so obtained contains less than 75% of ammonium sulphate and over 25% of sodium sulphate. Such a product has a very little practical use because it is a mixture of salts. It cannot be used, for example, as a fertilizer because of the high percentage of sodium sulphate. The patent does not provide any indication whether and how this mixture of salts can be separated to produce essentially pure and therefore marketable ammonium sulphate.
A process for the separation of sodium sulphate and ammonium sulphate from aqueous solutions of a mixture of them to obtain substantially pure sulphates is disclosed in Canadian Patent No. 821,457. The process uses an evaporator-crystallizer system which produces a crystal mixture in which crystals of each salt have different granulometric and density characteristics.
This mixture is separated, e.g. by dry or wet screening, into fractions consisting mainly of sodium sulphate and ammonium sul-phate crystals, respectively, which fractions are subsequently washed to dissolve away the sulphate present in the least quantity and leave the other sulphate substantially pure. The process, even though finally producing substantially pure sodium and ammonium sulphate from a mixture thereof, is lengthy and requires a rela-tively complicated equipment to separate the mixture of crystals obtained in the evaporation-crystallization step.
Demonstrated worldwide demand for sodium-based chemicals, particularly sodium carbonate (soda ash) has been on the rise in recent years. This strong demand, which is forecast to continue, keeps soda ash in a tight supply position thereby holding the price at a high level. Because of environmental concerns and the vast reserves of natural sodium carbonate, the production of soda ash in the United States is primarily from natural sources. In Canada, the known natural reserves of sodium carbonate are not as vast as those in the United States. However, Canada is endowed with vast deposits of sodium sulphate, located mostly in Southern Saskatchewan, that could potentially become a source of soda ash.
Moreover, statistics indicate that the demand for sodium sulphate is dwindling and its prices declining. This declining trend in the demand for and prices of sodium sulphate together with strong demand for and relatively high prices of soda ash created a need for an economical process for producing sodium carbonate (soda ash) from naturally-occurring sodium sulphate as feedstock.
Thus, the present invention provides a continuous pro-cess for preparing sodium carbonate, which process comprises reacting sodium sulphate in aqueous solution with ammonia and carbon dioxide to precipitate sodium bicarbonate, separating the sodium bicarbonate and calcining it to convert it to sodium car-bonate, sujecting the mother liquor from the precipitation of the sodium bicarbonate to evaporation to precipitate unreacted sodium sulphate, cooling the mother liquor to precipitate ammonium sul-phate, further cooling the mother liquor to precipitate a double salt of sodium sulphate and ammonium sulphate and adding the double salt to the mother liquor from the precipitation of sodium bicarbonate, prior to the evaporation.
The present invention has the considerable advantage that, in addition to producing sodium carbonate, it also produces ammonium sulphate in a purity such that it can immediately be used as a fertilizer. In contrast to the Solvay process, it is not necessary to seek to recycle all the expensive ammonia used in the process; a substantial amount of that ammonia appears in a pro-duct, ammonium sulphate, of high commercial value. Also, the process does not produce any unwanted byproducts, such as calcium chloride produced in the Solvay process. The products of the process are the required sodium carbonate, ammonium sulphate that can be used as fertilizer, a double salt of sodium and ammonium sulphate that is recycled and therefore does not accumulate, and sodium sulphate that can be recycled. As the sodium sulphate recovered in the process is of enhanced purity when compared with the raw salt normally used as starting material, it may be a valuable commercial product. Another advantage over the Solvay process is the avoidance of absorption towers, which have a tendency to become clogged with precipitated bicarbonate.
To carry out the precipitation of sodium bicarbonate according to the invention, any continuously fed reactor capable of operating under superatmospheric pressure and assuring an efficient gas-in-liquid dispersion can be used. In a preferred embodiment a closed, cylindrical reactor equipped with two gas dispersion impellers mounted on a coaxial shaft is used. The precepitation is normally conducted under superatmospheric pressure of from 0 kPa to about 700 kPa. A pressure of about 70 kPa is preferred.
Since the reaction generates heat, the reactor is equipped with a cooling system to maintain a temperature in an optimum range preferably from about 20C to about 60C. A tem-perature of about 40C is particularly preferred.
A sodium sulphate solution having a concentration of from about 400 g/L to about 500 g/L is continuously fed into the reactor. A low temperature of the brine fed into the reactor helps in maintaining the content of the reactor in the optimum temperature range. However, since the solubility of sodium sul-phate in water decreases with decreasing temperature, this tem-perature should not be excessively low, to keep the solubility of sodium sulphate in the indicated range of concentrations. A
temperature for the brine solution in the range of from 20C to 60C is suitable. In a preferred embodiment a saturated solution of sodium sulphate (concentration of about 488 g/L) having tem-perature of about 40C is used.
2032627 608l8-l7 Carbon dioxide and ammonia are fed into the reactor in the liquid and/or gaseous form. According to a preferred embodi-ment, a major part of the two gases is injected into the brine in the liquid form which has the advantage of producing a cooling effect compensating for the exothermic effect of the processes taing place in the reactor.
It is also preferred that a gaseous mixture of ammonia and carbon dioxide recovered from the brine after the precipita-tion of sodium bicarbonate is recycled into the reactor. It is therefore preferred to heat the mother liquor remaining after the precipitation of the sodium bicarbonate, to recover dissolved or unreacted carbon dioxide and ammonia before subjecting the liquor to evaporation.
The pH of the content of the reactor is maintained in a range of from about 7.0 to about 9Ø The value of pH of about 8.0 is preferred. The value of pH can be maintained by control-ling the supply of liquid carbon dioxide and/or ammonia.
In a preferrred embodiment the mother liquor from the precipitation of the double salt is added to the mother liquor from the precipitation of sodium sulphate. The mother liquor from the precipitation of the double salt can be subjected to concen-tration, for example by evaporation, before it is added to the mother liquor from the precipitation of sodium sulphate.
The mother liquor from the precipitation of sodium sulphate may be alternatively added to the liquor from the preci-pitation of sodium bicarbonate, before this liquor is evaporated.
However, it is preferred to reduce the amount of the recycled salts by cooling the mother liquor from the precipitation of sodium sulphate to precipitate ammonium sulphate, further cooling the mother liquor to precipitate a double salt of sodium sulphate and ammonium sulphate and adding this double salt to the mother liquor from the precipitation of sodium bicarbonate, prior to the evaporation of this liquor.
It is preferred that the mother liquor in the first evaporation stage shall be at a temperature not lower than about 80C, preferably not lower than about 95C, so that the precipi-tated sodium sulphate is not contaminated with any significantamount of sodium sulphate. It is preferred that the precipitation of ammonium sulphate shall take place at a temperature above about 40C, preferably above about 60C, so that is is not contaminated with any significant amount of the double salt. It is possible to include a concentration step between the precipitation of the sodium sulphate and the ammonium sulphate. It is also possible to include a concentration step between the precipitation of the ammonium sulphate and the double salt.
The various evaporation, precipitation and concentration stages can be carried out under reduced pressure, in which case temperatures different from those mentioned above will be appro-priate. Determing the appropriate temperatures for a particular chosen reduced pressure will be within the competence of a person skilled in the art.
The invention will be further described by way of a preferred embodiment and with references to the accompanying drawings in which:
Figure 1 represents schematically a process for purifi-cation of raw sodium sulphate and preparing sodium sulphate brine, Figure 2 represents schematically a process for produc-ing soda ash according to one preferred embodiment of the inven-tion, Figure 3 represents schematically a process for recover-ing ammonium sulphate and sodium sulphate according to one pre-ferred embodiment of the invention, Figure 4 represents schematically an experimental bench-scale reactor for conducting the process according to the inven-tion, and Figure 5 represents a flow diagram of the experimental system using the reactor of Figure 4.
The process according to the invention requires as a feed a solution of sodium sulphate. According to a preferred embodiment of the invention, this solution is prepared from raw mined sodium sulphate, in a process shown schematically in Figure 1. Raw sodium sulphate (Na2SO4.10H2O) is continuously fed into a classifier 1 and mixed with hot brine (an overflow from cyclone 5). The heat from the brine is sufficient to dissolve the crystals of sodium sulphate. Any stones and dirt settle to the bottom of the classifier 1 and are discarded as waste. The salt brine from the classifier 1 is put through a 30 mesh screen 6 to remove any fine particles suspended in the brine.
Once screened, the brine is transferred by a pump 7 to two submerged combustion evaporators No. 1 and 2 to evaporate some water, thus causing some sodium sulphate to crystallize out. An air and natural gas mixture which is burned under the brine in a burner tank 2 causes violent bubbling action. This causes some brine to splash over a weir and run down into a settling tank 3.
Here the crystals of sodium sulphate settle towards the bottom of the tank and are drawn off by a pump 10 and put through cyclones 5 to spin off more of the brine. This brine is the cyclone overflow which is used to dissolve the raw salt in the classifier 1. The solid sodium sulphate coming out of the cyclones 5 contains about 20% of water and is pumped by a pump 11 into a rotary dryer 55 shown in Figure 3 to finish drying the salt and produce sodium sulphate (salt cake).
The brine that comes off the top of the settling tank 3 is normally recycled to the burner tank 2 for further evaporation when only sodium sulphate is to be produced. In the process according to the invention, some of this brine is transferred by a pump 8 into a thickener 4, where a flocculating agent, for example Percol 156, is added to remove fine clay particles still suspended in the brine. The flocculating agent causes the clay particles to stick together and settle out faster at the bottom of the thick-ener 4, to be removed as a mud. The resulting brine is a cleansaturated solution of sodium sulphate (concentration of about 30%) in water having temperature of about 40C. This brine is fed into the reactor 21 shown in Figure 2 to be mixed with carbon dioxide and ammonia.
A cylindrical reactor 21 is equipped with gas dispersion impellers 22 and 23 and cooling coils 24. For better agitation cooling coils in the form of single vertical tubes are placed radially along the reactor wall. These tubes also act as baffles when the content of the reactor is agitated.
Liquid carbon dioxide and liquid ammonia are injected into the brine through injection nozzles 24 and 25, respectively.
A part of carbon dioxide and ammonia used in the process is sup-plied to the reactor in the gaseous form through a tube 26, as a mixture recovered from the gas recovery boiler 31, as explained below.
As carbon dioxide and ammonia dissolve in the brine, the following reactions take place:
NH3 + H2O = NH40H
Na2SO4 + 2NH4HCO3 = 2NaHCO3 + (NH4)2SO4 These reactions produce considerable heat which is to some extent compensated by the cooling effect produced by injecting liquid carbon dioxide and ammonia into the brine. Any excess heat is removed by the cooling coils 24 supplied with cooling water having temperature of about 7C to maintain the temperature of the reac-tor at an optimum level of about 40C. The pressure maintained in the reactor 21 is about 70 kPa. The pH of the content of the reactor is maintained at about 8 by controlling the supply of carbon dioxide.
Because of its limited solubility in water, sodium bicarbonate (NaHCO3) produced in the above shown series of reac-tions precipitates out. It is continuously removed from the reactor 21 in the form of a slurry by a pump 35 and transferred to a brine cooler 27, where the slurry is cooled to about 20C to 20~2627 60818-17 further reduce the solubility of sodium bicarbonate and to cry-stallize out as much of the product as possible. The sodium bicarbonate is then screened out of the brine on a 21 micron screen 28 and subsequently washed on a screen 29 with water to remove any ammonium and sodium sulphate present in the entrained brine. The washed sodium bicarbonate is then fed into a rotary dryer 30 to be dried and calcined to sodium carbonate (soda ash).
The brine remaining after screening off the solid sodium bicarbonate contains a mixture of unreacted sodium sulphate, ammonium sulphate, ammonium bicarbonate and minor amounts of sodium bicarbonate. This brine is transferred by a pump 36 into a gas recovery boiler 31 where it is heated to a temperature of 95 to 100C. Under these conditions the ammonium bicarbonate breaks down and sodium bicarbonate dissolved in the brine reacts with ammonium sulphate to produce sodium sulphate, carbon dioxide and ammonia. Carbon dioxide and ammonia dissolved in the brine boil off, leaving in the solution a mixture composed mostly of sodium and ammonium sulphate. The carbon dioxide and ammonia so regener-ated are cooled in a gas cooler 32 and returned to the reactor 21 by a blower 33 after being further cooled in a gas cooler 34.
This regeneration step minimizes the amount of carbon dioxide and ammonia used in the process.
The brine remaining after removing ammonia and carbon dioxide has a temperature of 95 to 100C and contains approxi-mately 15% of sodium sulphate and 15% of ammonium sulphate. This brine is added by a pump 37 to double salt crystals separated on the screen 61 shown in Figure 3, as explained below. The 1:1 ratio of sodium sulphate to ammonium sulphate in the double salt is the same as the ratio of these two salts in the brine, so that the mixing of the double salt with the brine do not affect the proportion of the two salts in the brine.
The mixture of the double salt from the screen 61 and the brine from the gas recovery boiler 31 is then fed into the submerged combustion evaporator No. 3 where the brine is concen-trated up to the point where both salts reach their saturation point. Since ammonium sulphate has a much higher situation point than sodium sulphate, most of the latter crystallizes out before the brine is saturated with ammonium sulphate. When the brine approaches saturation also with respect to ammonium sulphate, the brine is drawn off from the settling tank 52 by a pump 65 and transferred onto a screen 53, where sodium sulphate crystals are screened off. These crystals are then washed on a screen 54 to remove any ammonium sulphate present in the entrained brine.
Although sodium sulphate so recovered could be redissolved in water and returned to the reactor 21, this product is so pure that it is more economical to dry it to produce sodium sulphate. This is normally done by combining the solids from the screen 54 with the underflow of the cyclone 5 shown in Figure 1 and feeding the mixture into a dryer 55 for final drying.
The brine remaining after screening out sodium sulphate has a temperature of about 95C and is saturated with respect to both sodium and ammonium sulphate. This brine via a pump 67 is combined with brine from a submerged combustion evaporator No. 4 used to crystallize out ammonium sulphate. The brine from this evaporator is saturated in both salts and contains some crystals of ammonium sulphate. By cooling the mixture of both brines in a crystallizer 56 down from 95C to 60C, the solubility of ammonium sulphate is decreased while the solubility of sodium sulphate increases. This forces more ammonium sulphate to crystallize out while keeping sodium sulphate in the solution. The brine is then drawn off from the crystallizer 56 by a pump 68 and transferred onto a screen 57, where ammonium sulphate crystals are screened off. The crystals are then washed on a screen 58 to remove any sodium sulphate from the entrained brine and finally dried in a dryer 59.
After removing the ammonium sulphate crystals, the brine is transferred by a pump 69 to a double salt crystallizer 60 where it is further cooled to about 15C. At this temperature sodium sulphate and ammonium sulphate crystallize out in the form of a double salt (Na2SO4.(NH4)2SO4.2H20). The brine with suspended crystals of the double salt is transferred from the crystallizer 60 by a pump 70 onto a screen 61 where the solids are screened out and added to the brine from the gas recovery boiler 31 shown in Figure 2.
By removing the double salt from the brine the ratio of sodium sulphate to ammonium sulphate becomes 1:3 in the remaining brine. This brine is fed into submerged combustion evaporator No.
4 where it is concentrated until sodium sulphate reaches its saturation point, by which time some of the ammonium sulphate crystallizes out. At this point the brine passing from a burner tank 62 into a settling tank 3 is transferred from the settling tank 63 by a pump 66 into the crystallizer 56 where it is combined with the brine left after screening out sodium sulphate on the screen 53.
EXPERIMENTAL
The operating parameters of the process of the precipi-tation of sodium bicarbonate were studied and optimized in a bench-scale reactor unit.
Desiqn and Construction of the Laboratory Bench Unit The bench unit comprises three main sections: reaction system, feed system, and product recovery systems.
The reaction system consists of three main sections:
reactor, agitator, and cooling jacket.
A schematic of the reactor is shown in Figure 4. The reactor 1 is fabricated of a 0.076 m (3 in) diameter x 0.457 m (18 in) long mild steel pipe and has a volume of 2 litres. The reac-tor head 2 carries an agitator 3, a 0.0063 m (1/4 in) NPT feed inlet port 4, 0.0063 m NPT reaction gaseous vapour outlet port 5, and 0.0032 m (1/8 in) and 0.003 m NPT ports for thermocouple 6 and back-pressure gauge 7, respectively. The bottom of the reactor is a 0.051 m (1 in) long, 45-degree cone 8 to which is attached a 0.0095 m (3/8 in) diameter product drain connector 9. The reactor is designed for pressures up to 1034 kPa (150 psi) and tempera-tures up to 30C.
The feed mixer is a variable speed (640-930 rpm) motor-driven agitator 3 welded onto the reactor head 2 and centrally located in the reactor. Attached to the part of the agitator inside the reactor is a 0.069 m (2-3/4 in) wide by 0.28 m (11 in) long 16 x 16 mesh adjustable stainless-steel screen (not shown in the drawing). The screen enables good liquid/gas mixing by allow-ing the gas to pass freely through the screen pores to contact the feed.
The cooling jacket 10 is constructed of 0.102 m (4 in) diameter by 0.406 m (16 in) mild steel pipe, which is welded to the reactor in such a way that it is concentric with the reactor.
The coolant (glycol/water mixture) circulates through the system by the inlet and outlet ports (11 and 12, respectively).
The feed system consists of the feedtank and transfer line, metering pump and gas delivery manifold. The freshly pre-pared feed is siphoned into the feedtank 20 shown in Figure 5 from where it is fed to the reactor. To prevent crystallization of the feedstock, the tank is maintained at a constant temperature of about 40C. The feed is delivered to the reactor by means of a metering pump 21. The feed transfer line is 0.0063 m (1/4 in) diameter vinyl tubing heat traced at a constant temperature of about 40C and connected to the reactor by the metering pump.
Heat tracing the feed lines is necessary to prevent plugging problems caused by crystallization of sodium sulphate in the feed solution. The gaseous feed line is also heat traced to provide a constant feed rate.
Gas distribution within the reactor is accomplished by means of a manifolds 13 and 14. This system has eighteen gas inlet nozzles: nine for carbon dioxide and nine for ammonia. The manifolds are located near the bottom of the reactor so that the gas can be fed counter currently with the sodium sulphate solution, which is fed into the reactor from the top. The gas inlet nozzles 15 and 16 are arranged in a circle inside the reac-tor such that each gas is fed by an alternate nozzle. This design enables good dispersion of the gas mixture within the reactor.
Feedstock Preparation The solution of sodium sulphate used in all the experi-ments was prepared as follows:
Approximately 12.4 kg of fresh natural Glauber salt crystals (Na2SO4.10H2O) and 12 L of hot tap water were added into a metal pail and stirred continuously until the aqueous liquid was saturated. The temperature of the stirred mixture was maintained at about 40C by means of a drum heater wrapped around the pail.
After about two hours, the stirring was stopped and the mixture left to settle for approximately four hours at 40C to remove silt, salt crystals and other fine clays in the mixture The supernatant (clear sodium sulphate solutions) was withdrawn, without disturbing the sediments, and transferred to the feed storage tank 20. This decanting of the solution was accomplished by siphoning. A sample of the feed solution was analyzed to determine the sodium sulphate content. It was found to contain 2.13 g mol/L or 24.6 wt % of sodium sulphate. The density of the solution at 40C was determined to be 1229.7 kg/m3.
Experimental set-uP and Procedure A flow diagram describing the experimental system is illustrated in Figure 5. A typical procedure involves feeding the sodium sulphate solution to the reactor 22 by means of a metering pump 21 calibrated to maintain the desired flow rate of the solution. The feed solution is fed from the top of the reactor The carbon dioxide and ammonia gases at desired flow rates and pressures are metered with capillary flowmeters 23 and 24, respectively, into the reactor 22 through the gas manifold inlets.
The gas manifold is located near the bottom of the reactor en-abling the gases to flow countercurrently with the downflowing sodium sulphate liquid solution. The desired speed of the agita-tor (640-930 rpm) is set so that the reactor contents are ade-quately mixed.
The heat generated by the exothermic reactions of the liquid feed/gas mixture is removed by the circulating glycol/water mixture maintained at a constant temperature of about 20C. The liquid and gaseous reaction vapours are routed through the liquid trap 25 and back-pressure regulator (BPF) 26 into the vent. The sodium bicarbonate/ammonium sulphate product mixture is manually withdrawn from the coned-bottom of the reactor through the drain connector at intervals determined by a stopwatch (=51-53 mL/min).
Ideally the product recovery is accomplished at a uniform rate by means of a metering pump 27.
Process OPeratinq Conditions The experimental runs were conducted based on the fol-lowing reactor unit operating conditions:
Feed flow rate, mL/min 55 Ammonia flow rate, L/min 0.96-1.0 Carbon dioxide flow rate, L/min 1.1-1.2 Pressure, kPa 550-620 Temperature, C 23 Agitator speed, rpm 640-930 Product withdrawal, mL/min 51-53 The carbon dioxide flow rate is an estimated flow rate based on the ammonia rate because the CO2 flow exceeded the maxi-mum flow capacity of the flowmeter.Results The aqueous sodium sulphate feedstock was evaluated primarily for its effectiveness to convert to sodium bicarbonate when reacted with carbon dioxide and ammonia.
Crystallization of sodium sulphate in the feed solution occurred at relatively low temperatures (<40 C) necessitating heat tracing the feed lines (both liquid and gaseous lines) to prevent plugging problems.
Ammonia flow rate is an important parameter because it directly affected the quantity and quality of the product. For instance, excess amount of ammonia inflow resulted in the product being a mixture of sodium bicarbonate and sesquicarbonate instead of predominantly the less soluble sodium bicarbonate.
The process reactions were exothermic, requiring con-tinuous circulation of the glycol/water cooling mixture. An efficient cooling system had to be installed to provide the required cooling because of an excessive amount of ammonia gas and C2 had to be fed to the reactor under insufficient cooling condi-tions. This resulted in the formation of excessive quantities of intermediate product, ammonia bicarbonate/carbonate.
A vacuum filtration system and a medium speed (3000 rpm) centrifuge were used as a means of recovering the bicarbonates and carbonates from the product stream. The vacuum filtration system produced a purer product (0.86 wt % sulphate after only one wash-ing of 35 mL washing solution/70 g wet precipitates) than the centrifugation system after two washings (5.0 wt % sulphate). The washing solution was a saturated ammonium bicarbonate aqueous solution.
After the carbonate/bicarbonate products were removed (at 23C), the supernatant solution contained mainly 281 g/L
ammonium sulphate, and 65.8 g/L of sodium compounds (carbonates, sulphates, and bicarbonates). This solution was treated by two different processes to recover the ammonium sulphate. The first process used one litre of methanol per litre of solution, which recovered 157 g of precipitates per litre of fluid. The precipi-tates were found to contain 18 wt % of sodium. In the second process, ammonia gas was added to the solution. This process recovered 115 g of precipitates per litre of fluid. The precipi-tates were found to contain 28 wt % of sodium.
The experimental data are presented in Tables 1 and 2.
The percent recoverable carbonate/bicarbonate products were esti-mated from the product composition data as follows:
2032627 608l8-l7 A = 33.5 wt % NaHCO,/100 x wt % recoverable (Table 1) = 28.14 wt %
B = 65.2 wt. % Na2C03 . NaHC03 . 2H20/100 x wt %
recoverable (Table 2) = 43.0 wt %
C = 1.3 wt % Na2S04 Thus, the calculated % recoverable of products -A=B=C = 72.5 wt %.
The actual percentage of recovered products (Table 1) was 73.6 wt %.
The amount of sodium remaining in solution was 25.8 g/L.
When this solution is heated, the sodium bicarbonate/carbonate compounds react with the ammonium sulphate to produce Na2S04, C02 and NH3.
Table 1 Percentage Recoverable at 0C, Theoretical vs. Experimental % Recovery (Calculated Based Experimental Results O
on Solubility Data) (% Recovery Basis Na @ O C) Wt % (Basis Na) Compound at OC 1 st 2 hr of 2nd 2 hr of Trial Run Trial Run 74.0 Sodium Carbonate 70.0 Sodium Carbonate, decahydrate 73.6 73.5 66.0 Sodium Sesquica,l,on~le (Na2C03-NaHC03.2H20) 84.0 Sodium Bica,l,on~le 2 0 3 2 ~ 27 60818-17 Table 2 Percentage Recovery of Sodium at 23C and OC
SulphateSodium Content in Recovery pH ContentSu~,e",ala"l g/L (basis Na) @ 22C(wt %) @ 23C @ 0C @ 23C @ 0C
Supernatant from 7.9 - 65.9 25.8 32.6 73.6 product stream recovered in 1st two hours of Trial run Supernatant from 8.6 - 65.7 25.9 32.8 73.8 product stream recovered in 2nd two hours of Trial Run
Claims (15)
1. A process for producing sodium carbonate, which process comprises:
reacting within a reactor sodium sulphate in aqueous solution with ammonia and carbon dioxide to precipitate sodium bicarbonate and form a first mother liquor;
separating the sodium bicarbonate and calcining it to convert it to sodium carbonate;
subjecting the first mother liquor from the precipitation of the sodium bicarbonate to evaporation to precipitate unreacted sodium sulphate, forming a second mother liquor;
cooling the second mother liquor from the precipitation of the unreacted sodium sulphate to precipitate substantially independently ammonium sulphate in a purity of greater than approximately 75 wt. percent, forming a third mother liquor;
further cooling the third mother liquor from the precipitation of the ammonium sulphate to precipitate a double salt of sodium sulphate and ammonium sulphate, forming a fourth mother liquor; and adding the double salt to the first mother liquor from the precipitation of sodium bicarbonate prior to the evaporation.
reacting within a reactor sodium sulphate in aqueous solution with ammonia and carbon dioxide to precipitate sodium bicarbonate and form a first mother liquor;
separating the sodium bicarbonate and calcining it to convert it to sodium carbonate;
subjecting the first mother liquor from the precipitation of the sodium bicarbonate to evaporation to precipitate unreacted sodium sulphate, forming a second mother liquor;
cooling the second mother liquor from the precipitation of the unreacted sodium sulphate to precipitate substantially independently ammonium sulphate in a purity of greater than approximately 75 wt. percent, forming a third mother liquor;
further cooling the third mother liquor from the precipitation of the ammonium sulphate to precipitate a double salt of sodium sulphate and ammonium sulphate, forming a fourth mother liquor; and adding the double salt to the first mother liquor from the precipitation of sodium bicarbonate prior to the evaporation.
2. A process according to claim 1, wherein the mother liquor from the precipitation of the double salt is added to the mother liquor from the precipitation of the sodium sulphate.
3. A process according to claim 2, wherein the mother liquor from the precipitation of the double salt is subjected to concentration before being added to the mother liquor from the precipitation of the sodium sulphate.
4. A process according to claim 3, wherein the concentration is carried out by evaporation.
5. A process according to any one of claims 1 to 4, which process further comprises recycling to the reactor carbon dioxide and ammonia from the mother liquor from the precipitation of the sodium bicarbonate prior to evaporation.
6. A process according to claim 1, 2, 3 or 4, wherein the content of the reactor is maintained at a temperature of from about 20°C to about 60°C.
7. A process according to claim 6, wherein the content of the reactor is maintained at a temperature of about 40°C.
8. A process according to claim 1, 2, 3 or 4, wherein the reactor operates under a pressure of from 0 kPa to about 700 kPa.
9. A process according to claim 8, wherein the reactor operates at a pressure of about 70 kPa.
10. A process according to claim 1, 2, 3 or 4, wherein at least a part of carbon dioxide and ammonia are fed into the reactor in the liquid form.
11. A process according to claim 1, 2, 3 or 4, wherein at least a part of carbon dioxide and ammonia is fed into the reactor in the gaseous form.
12. A process according to claim 1, 2, 3 or 4, wherein the content of the reactor is maintained at a pH of from about 7 to about 9.
13. A process according to claim 12, wherein the content of the reactor is maintained at a pH of about 8.
14. A process according to claim 13, wherein the pH is maintained by regulating the supply of carbon dioxide and/or ammonia.
15. A process according to claim 1, 2, 3 or 4, wherein the sodium sulphate solution is a saturated solution having a temperature of about 40°C.
Priority Applications (2)
| Application Number | Priority Date | Filing Date | Title |
|---|---|---|---|
| CA 2032627 CA2032627C (en) | 1990-12-18 | 1990-12-18 | Process for producing sodium carbonate and ammonium sulphate from sodium sulphate |
| US08/494,073 US5654351A (en) | 1990-12-18 | 1995-06-23 | Method for sodium carbonate compound recovery and formation of ammonium sulfate |
Applications Claiming Priority (1)
| Application Number | Priority Date | Filing Date | Title |
|---|---|---|---|
| CA 2032627 CA2032627C (en) | 1990-12-18 | 1990-12-18 | Process for producing sodium carbonate and ammonium sulphate from sodium sulphate |
Publications (2)
| Publication Number | Publication Date |
|---|---|
| CA2032627A1 CA2032627A1 (en) | 1992-06-19 |
| CA2032627C true CA2032627C (en) | 1997-01-14 |
Family
ID=4146686
Family Applications (1)
| Application Number | Title | Priority Date | Filing Date |
|---|---|---|---|
| CA 2032627 Expired - Fee Related CA2032627C (en) | 1990-12-18 | 1990-12-18 | Process for producing sodium carbonate and ammonium sulphate from sodium sulphate |
Country Status (1)
| Country | Link |
|---|---|
| CA (1) | CA2032627C (en) |
Families Citing this family (15)
| Publication number | Priority date | Publication date | Assignee | Title |
|---|---|---|---|---|
| US5654351A (en) * | 1990-12-18 | 1997-08-05 | Ormiston Mining And Smelting Co. Ltd. | Method for sodium carbonate compound recovery and formation of ammonium sulfate |
| US5830422A (en) * | 1995-06-23 | 1998-11-03 | Ormiston Mining And Smelting Co. Ltd. | Method for production of sodium bicarbonate, sodium carbonate and ammonium sulfate from sodium sulfate |
| US6692716B1 (en) | 1998-10-13 | 2004-02-17 | Airborne Industrial Minerals, Inc. | Method of formulating alkali earth salts |
| US6106796A (en) * | 1998-10-13 | 2000-08-22 | Airborne Technologies Inc. | Method of ammonium sulfate purification |
| CA2284967A1 (en) * | 1998-10-13 | 2000-04-13 | Aristos Capital Corporation | Method of formulating alkali earth salts |
| CN107572562A (en) * | 2017-10-27 | 2018-01-12 | 四川省洪雅青衣江元明粉有限公司 | A kind of glauber salt prepares the device and method of soda ash and ammonium sulfate |
| CN114291827B (en) * | 2021-12-09 | 2022-06-28 | 中国科学院过程工程研究所 | Preparation method of large-particle sodium bicarbonate |
| CN114538471B (en) * | 2022-01-13 | 2023-08-01 | 宁波弗镁瑞环保科技有限公司 | Comprehensive utilization method of sodium sulfate-sodium chloride mixed salt |
| CN114436297B (en) * | 2022-02-21 | 2023-12-08 | 恒信润丰科技开发（北京）有限公司 | Method for preparing sodium carbonate from mirabilite |
| CN114751430B (en) * | 2022-04-07 | 2023-12-01 | 湖南化工设计院有限公司 | Method for producing sodium bicarbonate and co-producing ammonium chloride by low-temperature bidirectional salting-out circulation method |
| CN114702047B (en) * | 2022-06-07 | 2022-08-23 | 中国科学院过程工程研究所 | Method for preparing sodium carbonate and co-producing ammonium sulfate by using sodium sulfate |
| CN114715920B (en) * | 2022-06-07 | 2022-09-06 | 中国科学院过程工程研究所 | Method for preparing sodium carbonate and co-producing ammonium sulfate and ammonium chloride by using mixed sodium salt |
| CN115420696B (en) * | 2022-09-19 | 2024-08-02 | 安徽鲁控环保有限公司 | Method for detecting component content of quaternary water salt system |
| CN116282081A (en) * | 2023-05-17 | 2023-06-23 | 中国科学院过程工程研究所 | Method for preparing sodium bicarbonate from baking soda desulfurization ash |
| CN116917242A (en) * | 2023-05-26 | 2023-10-20 | 广东邦普循环科技有限公司 | Treatment method of high-salt high-organic wastewater |
-
1990
- 1990-12-18 CA CA 2032627 patent/CA2032627C/en not_active Expired - Fee Related
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| Publication number | Publication date |
|---|---|
| CA2032627A1 (en) | 1992-06-19 |
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