Recovery of Ammonia during Production of Coke from Coking Coal
Recovery of Ammonia during Production of Coke from Coking Coal
Ammonia (NH3) is a by-product produced during the production of coke from coking coal in the by-product coke ovens. It is a constituent of the coke oven gas (COG) leaving the coke ovens, with a typical concentration in raw COG of 6 grams per normal cubic meters (g/N cum). The solubility of NH3 in water leads to its presence in the flushing liquor of coke oven battery (COB) with a typical concentration of 5 grams per litre (g/l) to 6 g/l of total NH3. Therefore, due to the net production of flushing liquor in the COB, also sometimes being referred to as excess flushing liquor, there arises a liquid stream as well as a gas stream from which NH3 is required to be removed. The quantity of excess liquor is around 12 % of the dry coal throughput, which depends on the coal moisture content.
Removal of NH3 from the gas stream is a universal feature of a coke oven and by-product plant. This is because NH3, in the presence of the other COG contaminants hydrogen cyanide (HCN), hydrogen sulphide (H2S), oxygen (O2), and water, is extremely corrosive to pipelines made of carbon steel. Also, when ammonia is uncontrollably burnt in any combustion chamber, it forms nitrogen oxides (NOx) which causes air pollution. Hence, removal of NH3 from COG and liquid stream is required to be also done due to environmental reasons.
The primary NH3 handling process in the coke oven and by-product plant deals with the removal and disposal of the NH3 present in the COG. However, NH3 recovery systems often include facilities to handle the NH3 arising in the excess flushing liquor. For proper understanding of how these facilities are incorporated into the overall NH3 handling system, the treatment of NH3 in the excess flushing liquor is described first and then the main processes for removal of NH3 from COG is described.
Treatment of excess flushing liquor
In some places, excess flushing liquor can be disposed of without prior treatment, using deep well injection. A once common practice is to use the excess flushing liquor for quenching the hot coke, although for environmental reasons this practice is no longer acceptable. In the absence of such simple disposal methods, the remaining alternatives are the removal of the majority of NH3 from the liquor by distillation, normally followed by final treatment in a biological effluent treatment (BET) plant. It is possible to use BET alone to remove NH3 from excess flushing liquor, however, the size and operating cost of a BET plant is considerably reduced when preliminary removal of NH3 is done by distillation.
Distillation of excess flushing liquor involves feeding the liquor to the top of a distillation column with trays, usually called an NH3 still, and feeding a counter current flow of stripping steam at the bottom. The stripping steam distils off the NH3 which leaves with the overhead vapours and passes on for further treatment. The stripped liquor is pumped from the bottom of the still and cooled before discharge to the local sewer or to the BET plant. Typical levels of total NH3 in the stripped liquor range from less than 50 ppm (parts per million) to 150 ppm.
It is not readily possible to strip all the dissolved NH3 present in excess flushing liquor with steam. The many chemical species present in flushing liquor lead to the formation of various ammonium salts in solution. These include ammonium carbonate, chloride, and sulphate among others. Salts such as ammonium carbonate are easily decomposed by heat in the still to yield free molecules of NH3. However, other salts such as ammonium chloride and sulphate are not decomposed and retain the NH3 in a fixed form. The fraction of fixed NH3 to total NH3 in excess flushing liquor is typically in the range of 20 % to 50 %. To allow the distillation of the fixed NH3, the excess flushing liquor is needed to be made alkaline. The following typical reaction then takes place, liberating free molecules of NH3.
NH4+ + OH- = NH3 + H2O
The addition of alkali is determined by mass balance, based on chemical analysis of the excess liquor to give the concentration of the fixed NH3 present.
The form of alkali used in the distillation of excess flushing liquor has changed over the years. For many years, a suspension of calcium hydroxide (lime) was used. This material had the advantage of low cost and ready availability, but the formation of insoluble calcium salts such as calcium carbonate created a major problem with fouling. The NH3 stills required a considerable overdesign to allow continued operation while partially fouled, but even so they had to be taken out of service regularly for cleaning. To avoid the problem of fouling, sodium carbonate (soda ash) has also been used. This has the advantage that insoluble salts are not formed, but impurities can be present to cause fouling and on site storage and mixing equipment is required. An environmental factor, when using reagents such as lime and soda ash, is that any insoluble deposit formed can create a solid waste disposal problem.
The generally ready availability and ease of handling of sodium hydroxide (NaOH) (caustic soda) solution is presently the alkali of choice for present day design of NH3 stills. NaOH is the most expensive of the alkalis traditionally used, but its consumption can be closely controlled which is of great advantage where limitations are imposed upon still effluent pH. The non-fouling characteristics of NaOH allows the use of more economical still designs, incorporating valve trays in place of the traditional bubble cap trays. In practice, NaOH is injected into the NH3 still near to, but not at, the top tray. This allows dissolved acid gases such as HCN and H2S to be stripped from the liquor first, before they can react with NaOH to form fixed salts.
Materials of construction for NH3 stills are chosen for their corrosion resistance. Cast iron has traditionally been used, with generous allowances for corrosion, although it is now often more economical to use materials such as Hastelloy (a nickel based alloy), titanium and stainless steel (grade 316). The upper sections of NH3 stills where both NH3 and acid gases are present generally require the use of highly corrosion resistant materials.
A common operating problem in the distillation of excess flushing liquor is the presence of tar carryover which can lead to serious fouling in the still. The usual solution to this problem is to install sand or gravel bed filters in the excess flushing liquor supply line. These are manually or automatically operated. The removed tar is back-flushed to the tar and liquor decanters.
Recovery of ammonia from coke oven gas
Presently there are three methods which are being used for the recovery of NH3 from the COG. These are known as (i) ammonium sulphate process, (ii) Phosam process, and (iii) water wash process.
Ammonium sulphate process – This process removes NH3 from the COG by absorption in a solution of ammonium sulphate [(NH4)2SO4] and sulphuric acid (H2SO4). The absorption reaction is 2NH3 + H2SO4 = (NH4)2SO4. The ammonium sulphate produced by the reaction of NH3 with H2SO4 is recovered by crystallization. The crystals are then centrifuged, washed and dried. The various ammonium sulphate processes in operation differ in the type of gas/liquor contacting device and the type of crystallization equipment used. The process equipments of the ammonium sulphate process are described below.
An early method and still very commonly used process employs a dip tube extending below the surface of the acid/ammonium sulphate solution in a vessel referred to as a saturator. Saturator is a cylindrical vessel with conical bottom. It is provided with a bubbler hood, which is duct prolonged to middle of the saturator. The duct has a hood at the bottom equipped with vanes like arrangement. Another ring like structure with small openings is at the conical portion, which is used for nitrogen (N2) feeding. Hot water rings are provided at the top of saturator. Saturator is always maintained with acid bath called liquor, which contains 4 % to 5 % of H2SO4. COG is fed through the dip tube and gas/liquid contact is caused as the gas bubbles moves up through the solution in the saturator. Acid is continuously added to the saturator. During this period, the NH3 present in the gas reacts with the H2SO4 in the liquor to form ammonium sulphate. The heat of reaction between NH3 and H2SO4 causes the evaporation of water into the COG. The concentration of ammonium sulphate reaches saturation, causing crystals to form directly in the saturator where they are allowed to grow until they are removed from the process. By means of agitation and circulation of the solution, the fine crystals are retained within the process solution of ‘mother liquor’ as it is known. Ammonium sulphate thus formed settles at the bottom of saturator.
Nitrogen is fed to the saturator for the purpose of agitation. The flow of N2 to the saturator is normally maintained at 450 cubic meters per hour (cum/h) to 550 cum/h. Pure N2 is purged into the saturator through N2 rings at a pressure of around 6 kg/sq cm. N2 purging increases the crystal growth.
H2SO4 (98 %) is fed to the saturator to maintain the acidity in the saturator. The gas collected at top of the saturator is fed to the acid trap. As the gas rises up, some of the crystal can be carried with the gas and they get stacked to walls of the saturator at top. Then the hot water is sprayed through the ring provided. The crystal attached to the walls of saturator is washed away. When hot water is sprayed them the concentration of the liquor decreases. So inlet acid concentration is increased to around 6 % to 7 % during the period. After the reaction, mother liquor is continuously drawn to the circulating tank provided at the side of the saturator. This also acts as a seal for the saturator. From circulating tank, mother liquor is fed to the mother liquor tank. The crystals collected at the bottom are fed to the crystal receiver tank by using pump
The outlet of the saturator carries some acid mist. In order to remove the acid mist, the gas is sent to the acid trap. It is hollow cylindrical vessel. The COG from saturator enters tangential to the trap. Due to the centrifugal motion, the acid mist gets separated. The acid collected at the bottom is fed to the circulating tank. After the removal of the acid mist, the COG is fed to the benzol recovery section of the by-product plant.
Ammonium slurry from the bottom of the saturator is pumped to crystal receiver tank having conical bottom where the ammonium sulphate crystals settle at the conical portion of the tank. The mother liquor from the top of the receiver is fed to the saturator. The slurry from the bottom is fed to the centrifuge. The crystal from centrifuge contains some amount of moisture. To remove this moisture, crystals are to be dried. Drier is usually a fluidized bed drier. The principle is based on making the particles loose so that they acquire fluidity due to the action of air flow with a definite air velocity. The drier is provided with a screen at the bottom, ceramic rings are arranged at the bottom of the screen. The drier is provided with air with forced draft air fan, heated in the duct. A spreader at the feed chute of the drier spreads the feed in all directions.
Forced draft fan sucks the atmospheric air and feeds to the drier. The discharge chute of the fan is divided into two sections. The air is heated to 120 deg C to 150 deg C by using the steam and the hot air is fed form the bottom of the screen. The ceramic rings distribute the air in all directions and making the crystals to fluidize. The temperature of the air is to be maintained sufficiently high to remove the moisture from the crystals. At the discharge end of the drier, the ammonium sulphate crystals are cooled.
When the pressure level of the fluidized bed reaches the set point which is around 300 mm water column (WC) to 400 mm WC, an automatic discharge feeder discharges the dry ammonium sulphate to the bucket elevator. The elevator discharges the dry product into the bunker, which in turn feeds the product to the bagging machine. The zone above the fluidized bed is kept in the range of 5 mm WC to 10 mm WC in order to avoid the carry-over of the ammonium sulphate particles out of the drying unit to the dust catcher.
The air from the drier is sucked by the suction fan and fed to the cyclone separators. Cyclone separator separates fine ammonium sulphate crystals in the air and feed to the bunker. The crystal dust laden air from the cyclones is fed to the bottom of the dust collecting tank which contains flushing liquor up to certain level. Here the crystal dust is dissolved in the water and the air is vented into the atmosphere.
The excess liquor from the saturator enters the mother liquor tank. Each saturator is provided with two mother liquor tanks. One is vertical and is horizontal. First the liquor enters the horizontal tank. As the liquor has less density than the tar, it floats. Then the clear mother liquor is fed to the vertical tank. From the bottom of the vertical tank mother liquor is fed to the saturator through the pumps provided. The concentration of the liquor is maintained at around 10 % to 12 %.
The typical flowsheet of ammonium sulphate plant incorporating excess flushing liquor treatment is shown in Fig 1.
Fig 1 Typical flowsheet of ammonium sulphate plant
Ammonium sulphate is also produced by the direct reaction of concentrated H2SO4 and gaseous NH3 in an evaporative crystallizer. The process is described here. Liquid NH3 is evaporated in an evaporator using steam at a pressure of 16 kg/sq cm and preheated using low pressure steam. The stoichiometric quantities of preheated gaseous ammonia and concentrated H2SO4 (98.5 %) are introduced to the evaporator- crystallizer which is operating under vacuum. These quantities are maintained by a flow recorder controller and properly mixed by a circulating pump (from upper part of the crystallizer to the evaporator).
The reaction takes place in the crystallizer where the generated heat of reaction causes evaporation of water making the solution supersaturated. The supersaturated solution settles down to the bottom of crystallizer where it is pumped to vacuum metallic filter where the ammonium sulphate crystals are separated, while the mother liquor is recycled back to the crystallizer.
In the ammonium sulphate process of NH3 recovery, the generation of fines which act as the seeds for new crystals and the removal of fines are the main processes which determine the size of the crystals that can be produced. There are three basically different types of crystallizers namely (i) the forced circulation (FC) type, (ii) the turbulence or draft tube baffled (DTB) type, and (ii) the fluidized bed type. In the FC type crystallizer, the entire contents of the crystallizer are completely mixed and pass the circulation pump and heat exchanger hundred times per hour. This gives most mechanical stress on the crystals and leads to the formation of a large amount of fragments that act as seed crystals. These fine particles grow into many crystals and hence smaller product crystals are produced by this crystallizer. In the DTB crystallizer, the crystal suspension does not pass the heat exchanger and as a result the power input of the circulation pump is a factor ten times lowers than for the FC crystallizer leading to much lower fines generation. In addition, fines which are present are taken out of the crystallizer through a clarifying zone and are dissolved as a result of the heating in the heat exchanger. These effects lead to a crystal size which is 3 to 4 times larger than what can be achieved in the FC type. In the fluidized bed type crystallizer, there is no pump which is in contact with the crystal suspension but the crystals are kept in a fluidized bed. As a result the mechanical stress on the crystals is even lower. Because of this, the largest possible crystals are produced by this crystallizer, but for ammonium sulphate the crystal size which can be reached is only slightly larger than with the DTB crystallizer and if the plant is overloaded the crystal size collapses because then the crystallizer starts acting similar to the FC crystallizer.
In ammonium sulphate crystallization, it is necessary to distinguish between reaction crystallization and evaporative crystallization. In addition the performance of the different types of crystallizers needs to be considered. In reaction crystallization, such as that of ammonium sulphate from H2SO4 and NH3, both the reaction which creates the super-saturation of the solute, and the subsequent crystallization of the solute, occur inside the crystallizer vessel. The heat of dissolution and reaction for ammonium sulphate, when using reasonably concentrated reactants, is sufficient to operate a reaction in the ammonium sulphate crystallizer without any external energy source for evaporation. The super-saturation profile and the ammonium sulphate crystallization kinetics, as well as the method of operation of such a crystallizer are different from those of a classic evaporative crystallizer. In an evaporative unit, the feedstock is brought in under-saturated and a heater needs to be operated in conjunction with crystallizer, to evaporate the water in the feedstock. The majority of the ammonium sulphate crystallizers, estimated at around 80 % to 90 %, are operated in the evaporative mode. The general trend is that reactive crystallization produces smaller crystals but has an energetic advantage over evaporative crystallization.
Since ammonium sulphate is produced as a by-product during the recovery of NH3 from COG, there are normally a lot of impurities from different sources. These are organic impurities as well as the inorganic impurities. The combined effects of these impurities are complex and influence the purity and the crystal shape and size.
The traditional material of construction for the saturator and all wetted surfaces is lead lined carbon steel. Alloys such as Monel and stainless steel (grade 316) are also used. Brick lining is used to protect the lead lining, which suffers from “creep” and damage by erosion.
The availability of acid resistant materials such as stainless steel (grade 316) has allowed the development of the modern NH3 absorber systems. In these systems, a circulating stream of ammonium sulphate / H2SO4 solution is sprayed counter currently to the COG flow in an absorber vessel. Absorption of NH3 from the gas takes place on the spray droplet surfaces. A portion of the circulating liquor is continuously withdrawn and fed to a separate continuous crystallizer. Here, the liquor is concentrated using heat and negative pressure to evaporate the water and so promote crystallization. The crystals are removed and a stream of mother liquor is continuously fed back to the absorber circuit.
The operation of ammonium sulphate processes results in an increase in the heat content of the COG leaving the absorber or saturator. The reason for this is that to maintain the water balance in the process, especially in the case of saturators, water is needed to be evaporated into the gas stream. In addition to the water getting added to the process with the acid, regular desaturations are necessary in which water is added to the mother liquor to dissolve crystal deposits and reduce fouling. In some installations, gas heaters are provided upstream of the NH3 absorbers/saturators. The evaporation of water into the COG results in an outlet gas with a higher dew point than at the inlet. In order for downstream gas cleaning processes such as naphthalene, benzol and H2S removal to be operated effectively, the gas is required to be cooled in final gas coolers.
The ammonium sulphate processes accommodate NH3 from excess flushing liquor be feeding the overhead vapours from flushing liquor distillation into the COG main upstream of the NH3 saturator / absorber.
The major economic disadvantage with ammonium sulphate processes is the price relationship between H2SO4 and ammonium sulphate. The H2SO4 required to make ammonium sulphate can cost upto two times the value of the ammonium sulphate product.
The Phosam process – This process was developed by United States Steel as a means of producing a saleable, commercially pure anhydrous NH3 product from the NH3 present in raw COG. The anhydrous NH3 produced by this process is a high value product compared to the ammonium sulphate produced by the other processes described earlier.
In the Phosam process, NH3 is selectively absorbed from the COG by direct contact with an aqueous solution of ammonium phosphate in a two stage spray absorption vessel (Fig 2). The absorption solution actually contains a mixture of (i) phosphoric acid (H3PO4), (ii) mono ammonium phosphate (NH4H2PO4), (iii) di-ammonium phosphate [(NH4)2HPO4], and(iv) tri-ammonium phosphate ](NH4)3PO4]. The reversible absorption reactions which take place are as given below.
H3PO4 + NH3 = NH4H2PO4
NH4H2PO4 + NH3 = (NH4)2HPO4
(NH4)2HPO4 + NH3 = (NH4)3PO4
The NH3 absorbed is recovered by steam stripping. This regenerates the absorption solution which is returned to the spray absorber. The steam stripping is performed at high pressure, around 13 kg/sq cm. The reason for this is that the reversible reactions which liberate the NH3 from solution are favoured by higher temperatures. Therefore by operating at high pressure (and hence higher temperature) the consumption of stripping steam is minimized.
The overhead vapours from the stripper are virtually only water vapour and NH3. These vapours are condensed and then fed to a fractionating column where anhydrous NH3 is recovered as the condensed overhead product. The fractionator bottoms product, mainly water, leaves the process as effluent. The schematic flowsheet of Phosam process is shown in Fig 2.
Fig 2 Schematic flowsheet of Phosam process
The Phosam process can become contaminated by tar and by absorption of acid gases (HCN, H2S and CO2) in the recirculated solution. To remove the tar, a froth flotation device is installed in the solution circuit between the absorber and the stripper. Acid gases are removed by preheating the NH3 rich solution and feeding it into a vessel referred to as a contactor. In this vessel, the preheat causes vapourization of water and acid gases from the solution. These vapours are vented back to the COG main and the remaining rich solution is fed to the stripper. A subsequent step to deal with any remaining acid gases and prevent them from contaminating the anhydrous NH3 is to add sodium hydroxide (NaOH) to the fractionator feed. The NaOH fixes the acid gas compounds as non-volatile sodium salts which remain in the fractionator bottoms effluent stream.
An important operational feature is the control of the water balance in the process. Substantial amounts of steam are condensed in the solution stripper, and this condensate is required to be re-evaporated from the circulating solution into the COG stream. The temperature of the solution returning to the absorber is around 60 deg C, and therefore the COG becomes heated as it flows through the absorber. The increased gas temperature usually makes it necessary to install a final gas cooler after the Phosam absorber.
The addition of phosphoric acid to the absorption solution is required only to account for operating losses such as spillage. It is added at weekly intervals at a rate equivalent to 0.0075 kg H3PO4 per kilogram NH3 produced. The Phosam process is very efficient, capable of achieving greater than 99 % recovery of NH3 from COG. Other plant configurations are possible in which, for example, aqueous NH3 solution is produced instead of the anhydrous NH3. Materials of construction are stainless steel for all areas in contact with phosphate solution of aqueous NH3, and carbon steel for other areas.
As in the sulphate process, NH3 present in excess flushing liquor is handled first by distillation, with the vapours being fed to the COG upstream of the Phosam absorber.
The water wash process
One of the simplest and most frequently used methods of removing NH3 from COG is to absorb it in water. Aqueous absorption liquor is fed counter currently to the flow of COG in an NH3 washer vessel (Fig 3). The vessel can be designed as a spray type absorber with several liquor respray stages, or as a packed tower as is common in many plants. The type of packing generally used is vertically arranged expended metal sheets which promote gas/liquor contact but resist playing and fouling. The rich NH3 solution formed, with a typical concentration of 5 g/l to 8 g/l, is then fed to distillation columns where the NH3 is stripped from the aqueous liquor using steam. The NH3 and water vapours leaving the top of the stripping column are passed on for subsequent treatment in a variety of ways. The schematic of the wash process is shown in Fig 3.
Fig 3 Schematic of the water wash process
After stripping, the absorption liquor is cooled and returned to the washer. There is a continuous blow down of stripped liquor from the circuit which is equivalent to the volume of steam condensed in the stripper column. This blow down is the plant effluent and requires biological effluent treatment to fully remove the residual NH3.
As no chemical reactions are involved, other than the dissolving of NH3 in water, the water wash process is temperature dependent and is most efficient at low COG temperatures (20 deg C to 30 deg C). The NH3 washer vessel is usually placed immediately after the tar precipitator in a typical coke oven by-product plant. At this point, the gas retains some superheat from the gas exhauster, if the by-product plant is operated at positive pressure. To promote NH3 removal efficiency, gas cooling is required to remove this superheat and to cool the gas to the optimum temperature range.
The gas cooling stage is often incorporated into the NH3 washer vessel itself. The NH3 washer is not to be operated at a lower temperature than the outlet temperature of the gas cooling stage, otherwise fouling by naphthalene can result.
Use of aqueous absorption liquor results in the simultaneous absorption of significant quantities of acid gases (H2S, CO2, and HCN) from the COG.
Therefore, the NH3 column is nowadays often constructed of corrosion resistant materials such as titanium and stainless steel (grade 316), although many plants continue to operate cast iron stills. The stripping column is equipped with bubble cap trays or with more economical valve trays. Due to the lower liquor temperatures in the washer and hence the reduced rate of corrosion, this vessel can be constructed entirely in carbon steel.
An advantage of the water wash process is that excess flushing liquor and other aqueous plant streams (such as benzol plant effluent) can be used to absorb NH3 in the washer. The advantage of doing this is that as the excess liquor is to be steam stripped in any case, there is a net saving of stripping steam if the excess flushing liquor is also used to absorb NH3.
For other plant effluent streams, it often makes sense to perform steam stripping as a preliminary effluent treatment step. Combining this with the NH3 absorption process minimizes overall steam consumption for the byproduct plant. The excess flushing liquor is added at a point in the washer where its free NH3 concentration most closely matches the free NH3 concentration of the absorption liquor. The presence of fixed NH3 does not influence absorption of free NH3 from the COG. If excess flushing liquor is used in the water wash process, the flow rate of the blow down effluent stream is increased to maintain the circulating liquor inventory. The fixed NH3 can be removed in the stripping column by the addition of caustic soda. Alternatively, the stream of blow down can be fed to a separate fixed NH3 still. The use of a separate fixed NH3 still avoids the presence of alkali in the recirculating absorption liquor, which can be responsible for forming fixed compounds with acid gases such as HCN and leading to the presence of these compounds in the plant effluent stream.