Coke Oven By-Product Plant

Coke Oven By-Product Plant

The coke oven by-product plant is an integral part of the by-product coke making process. During the production of coke by coal carburization in a coke oven battery a large quantity of gas is generated because of the vapourization of volatile matter in the coal. The gas is generated over majority of the coking period, the composition and the rate of evolution changing during this period and being normally complete by the time coal charge temperature reaches 700 deg C. This gas is known as raw or crude coke oven gas and is processed in the by-product plant.

The functions of the by-product plant are to process the raw coke oven gas to recover valuable coal chemicals and to condition the gas so that it can be used as a clean, environmentally friendly fuel gas. The by-product plant is required to  carry out the functions of (i) to cool the coke oven gas for condensing of water vapour and contaminants, (ii) to remove tar aerosols to prevent gas line / equipment fouling, (iii) to remove ammonia to prevent gas line corrosion, (iv) to remove naphthalene to prevent gas line fouling by condensation, (v) to remove light oil for recovery and sale of benzene, toluene, and xylene (BTX), and (vi) to remove hydrogen sulphide to meet local emissions regulations governing the combustion of coke oven gas. In addition to treating the coke oven gas, the by-product plant is also required to condition the flushing liquor which is returned to the coke oven battery, and treat the waste water which is generated by the coke making process. The final yield of clean gas after treatment in the by-product plant is around 300 normal cubic meter per ton (N cum/t) of dry coal, the yield depending upon a number of factors, including coal volatile material available and the carbonization conditions.

The raw coke oven gas is sprayed with flushing liquor as it leaves the individual oven chambers, the objective being to reduce the temperature to a reasonably low value and to condense the most easily condensable (high boiling point) components. The gas is hence cooled by adiabatic evaporation of some the spray liquor and the mixed gases from the battery are reduced to around 80 deg C to 84 deg C water saturated, this temperature being sufficiently low for handling the gas in the gas collecting mains etc. The gas, together with flushing liquor and condensed tar, pass along the collecting main and through a butterfly control valve before leaving the battery area. This control valve is set to maintain a slight positive pressure, around 10 kilo-pascal (kPa), in the collecting main to provide safe working conditions in the coke oven chambers.

The coke oven gas is a complex mixture and contains several components, including typically around by volume hydrogen (H2) – 48 % to 55 %, methane (CH4) – 28 % to 30 %, carbon monoxide (CO) – 5 % to 7.5 %, carbon di-oxide (CO2) – 1.5 % to 2.5 %, nitrogen (N2) 1 % to 3 %, high paraffins and unsaturated hydrocarbons – 2.5 % to 4 %, and oxygen – less than 0.5 %. Minor components include ammonia (NH3), hydrogen sulphide (H2S), hydrogen cyanide (HCN), ammonium chloride (NH4Cl), benzene (C6H6), toluene (C7H8), xylene (C8H10), and naphthalene (C10H8) and other aromatics, tar components, tar acid gases (phenolic gases), tar base gases (pyridine bases), and carbon di-sulphide (CS2). Of these, the hydrogen, methane, carbon monoxide, paraffinic and unsaturated gases are useful components of the final clean coke oven gas. The small quantities of carbon di-oxide, oxygen, and nitrogen are allowed to remain in the final gas as inert but harmless components. The remainder of the components are removed in the by-product plant as far as is practical.

Spraying of the gas leaving the oven chamber with ammonia liquor condenses from the gas the high-boiling tar vapour compounds and ammonium chloride. The tar forms a separate liquid phase with the hot spray liquor and the ammonium chloride dissolves in the liquor. The tar and liquor are separated from the gas stream at the downcomer and pass to the tar decanting plant. The flow of coke oven gas is withdrawn from the battery by a centrifugal blower, the exhauster, and delivered through the gas vessels of the by-product plant at a sufficient pressure for distribution as a fuel. Before passing to the exhauster, the gas is cooled in the primary gas coolers to around 25 deg C. By this means a large proportion of the condensable material is removed, e.g., tar and naphthalene. Most of the water vapour is also condensed from the gas and the gas contracts in cooling from around 80 deg C to 25 deg C. The volume of gas which the exhauster has to handle is hence very much reduced.

For ensuring reliable and efficient coke production from the coke oven battery, it is absolutely necessary to have continuous exhauster operation to remove the gas as it is generated, and also to have a reliable continuous supply of flushing liquor to quench this coke oven gas to a reasonably low temperature. On the majority of the coke oven and by-product plants, the exhausters are located in the gas stream immediately following the primary gas coolers, however, some plants have the exhausters located at the end of the gas stream. Whilst this latter position provides a cleaner gas for exhauster operation, location of the exhauster after the primary gas coolers provides for majority of the gas stream to operate at a slight positive pressure.

Following initial cooling and passage through the exhausters, the coke oven gas flows in sequence through a number of vessels which incorporate means for removing undesired or saleable components. Throughout the by-product plant and associated equipment, it is always necessary to assume the presence of both tar and naphthalene, even when these materials are not to be present in the system. Although these materials can be present in only trace concentrations, they can accumulate over a period of time to create blockages, typically in plant drains and vents. The gas components which are normally removed are described below.

Tar – It is condensable vapour forming an essentially chemically neutral liquid, containing however the tar acids (phenols) and tar bases (pyridines etc.). The tar consists mainly of a mixture of various aromatic hydrocarbons which range from the volatile benzene to pitches which are solids at room temperature. Tar is removed as a saleable material. Since it is a rich source of mainly aromatic chemicals, it can be processed in a tar distillation plant for producing these chemicals. Around 70 % of the tars, including the least volatile materials, are condensed in the gas quenching operation using sprayed liquor. A further around 20 % is removed when the gases are cooled in the primary cooler, and the remaining around 10 % is removed in the electrostatic tar precipitators. The electrostatic tar precipitators are designed to remove the fine tar droplets remaining in the gas after the exhausters and are normally located in the gas stream immediately following the exhausters.

Naphthalene – It is present as a vapour which condenses directly to a solid, chemically neutral aromatic material. The raw coke oven gas normally has a naphthalene concentration corresponding to a dew point of around 50 deg C to 55 deg C, depending on the carbonization conditions. Hence, there is almost no naphthalene removal when the gas is quenched to 80 deg C to 84 deg C, but there is substantial removal when the gas is cooled to 25 deg C in the primary coolers. Naphthalene dissolves freely in tar. Majority of the naphthalene remaining in the coke oven gas as a vapour after the primary gas cooling stage is removed subsequently by absorption in either mineral or creosote oil in the benzol washing stage or a special unit. As a general principle, it is necessary to reduce the naphthalene in gas concentration corresponding to a dew point temperature appreciably below a temperature which can be experienced by the walls of the gas distribution pipework.

Ammonia – It is present as a vapour in the coke oven gas, readily dissolving in water to form a weakly alkaline solution. Since the ammonia is associated with the problems of corrosion when hydrogen sulphide and ammonia are present, it is normally removed at the earliest possible stage, immediately following the tar precipitators. There are few economically favourable outlets for the ammonia or its associated compounds. Processes for ammonia removal make use of its properties of a weakly alkaline compound to react with acidic reagents of one type or another.

Benzol – It consists of volatile aromatic hydrocarbons, including mainly benzene, toluene, xylene. Benzol is removed presently as a valuable saleable product, and also to reduce their concentrations of benzene, toluene, and xylene in the coke oven gas used as a fuel in the coke oven battery and in the steel plant. The benzol vapour components are normally removed from the gas by absorption in an oil fraction, derived either from coal tar or mineral oil. The benzol absorption process also removes a substantial proportion of the naphthalene vapour still present in the gas in addition to some of the carbon di-sulphide. Some of the benzol compounds initially present in the gas leaving the coke ovens are absorbed in the early tar condensing and primary gas cooling sections of the plant.

Hydrogen sulphide – It is a weakly acidic gas, readily absorbed in alkaline solutions and also readily desorbed as a gas from them. Hydrogen sulphide has to be removed from the gas for three main reasons namely (i) it is a poisonous component, (ii) in association with other components it causes plant corrosion, and (iii) its combustion products (sulphur di-oxide / sulphur tri oxide, SO2 / SO3) are environmentally very undesirable if released to the atmosphere when the coke oven gas is burnt. There are several hydrogen sulphide removal processes which basically rely upon contacting the gas with a circulating absorbent, from which sulphur is recovered by a controlled oxidation reaction. The general principle observed in deciding the sequence of operations in the by-product plant is to follow the tar / liquor separation from the gas and the primary gas cooling stages by sequential removal of components in such a way as to minimize downstream disturbances. The operations on the gas stream are normally carried out at a slightly positive pressure (7 kPa to 15 kPa). This avoids the possibilities of undetected air in-leakage into the plant at sub-atmospheric pressure without the expense and complexities because of the high-pressure plant operation.

A normal sequence for a plant hence is (i) tar and liquor separation, (ii) primary gas cooling, (iii) compression in exhausters, (iv) electrostatic tar droplet removal, (v) secondary / final gas cooling (frequently combined with ammonia removal), (vi) ammonia removal, (vii) benzol removal, (viii) naphthalene removal (if necessary), and (ix) hydrogen sulphide removal. The sequence can be varied if, for example, two processes are inter-dependent. As an example, in some process streams, the ammonia and hydrogen sulphide removal processes are inter-dependent. Following the primary gas cooling stage majority of the gas treatment processes are designed to operate at slightly above atmospheric temperature, and at 25 deg C to 30 deg C. Since several of the absorption processes are very temperature dependent, for example, ammonia, benzol and hydrogen sulphide removal, it is advantageous to maintain the gas stream as cool as practical.

It is very important to be aware of relevant dew point temperatures in the gas scrubbing system in order to prevent undesirable condensation. Both the water and naphthalene dew point temperatures are relevant, for example to avoid condensing water into oil circulation systems and to avoid naphthalene deposition into gas mains or water circulation systems. Both water and naphthalene dew point temperatures correspond to the primary cooler exit temperature and pressure, gas passage through the exhausters raises these temperatures typically by 2 deg C to 3 deg C.

Some by-product plants have an entirely different operating basis, by chilling the main gas stream to much lower temperatures than normal, it is possible to produce a clean stream of hydrogen for a synthesis process. Other materials removed from the gas are separated from the condensed liquids. The cleaned coke oven gas is normally available at a pressure of around 7 kPa and 25 deg C, it is then available for distribution to different consumers as a fuel gas. Frequently, for consumers at the same steel plant, no further compression is needed. However, when necessary for more distant consumption, the gas pressure can be increased by use of booster compressors. In this case, it is also necessary to reassess the relevant dew points to avoid downstream condensation issues.

It is normal to provide some means of cleaned gas storage in a gas holder to provide a reliable source in case of gas treatment plant delays or hold-up. The capacity, type, and location of this facility is decided from a study of operating plant and gas consumer characteristics. The design of several elements of modern coke oven plant by-product units has been affected quite profoundly by modern personnel safety and environmental pollution considerations. These features can be typified by brief mention of the main characteristics. Acknowledgement of the carcinogenic properties of aromatic hydrocarbons (benzol, tar compounds) means that equipments are to be provided to prevent release of these materials to the surroundings or exposure of plant operators to them. In practice, this means providing efficient vapour collection and treatment facilities in addition to providing secured holding for any such liquids.

Gas and vapours containing sulphur or sulphur oxides are to be treated to substantially reduce their release to the atmosphere. In general, all the process liquids are to be contained by drainage and recovery to prevent leakage to the ground. Combustion units are to be designed to minimize release of oxides of nitrogen gases to the atmosphere. In general terms, the modern coke oven gas by-product plant consists of a sequence of operations carried out on the main gas flow stream, together with associated side-stream processing equipment. For almost every operation several options of plant designs are available.

Tar and liquor decanting and storage 

Tar and ammonia liquor are produced in the coal carbonization process from materials contained in the coal feedstock and water added to the process. The main tar and liquor combined flows in the by-product plant arise as a result of the gas cooling as it leaves the oven chamber and the subsequent primary gas cooling. Minor flows of these materials arise at different other locations such as plant drains or where other plant condensates form. The functions of the plant sections dealing with tar and liquor can be summarized as (i) continuous rapid separation of a suitable flushing liquor stream, (ii) separation of a clean and tar free excess ammonia liquor for further processing, and (iii) separation of a clean tar essentially free from water and solids for sale.

The continuous separation of the flushing liquor stream is particularly an important feature, because this flow being needed to cool the hot oven exit gases down to a temperature which can be handled in the gas collecting system. Since the flushing liquor supply is so important for good coke oven battery operation, it is normal for plant and equipment dealing with flushing liquor decanting and recirculation to be sized and duplicated as necessary to provide adequate standby resources when some units are to be taken out of service for maintenance. Further, because of the presence of hazardous and carcinogenic aromatic compounds, it is a necessary feature of modern coke oven and by-product plant design and operation that all vessels and equipment containing ammonia liquor or tar are furnished with effective methods of preventing leakage or escape of vapours, liquids, or solids to the atmosphere or ground. Hence, all the equipment is these sections are required to have efficient vent and drain collection systems, with methods of treating materials collected. The main sources for tar and ammonia liquor on a by-product plant are indicated schematically in Fig 1, which also shows the flushing liquor circulation.

Fig 1 Schematic diagram for flushing liquor circulation

Source and composition of the ammonia liquor

The water in the excess flushing liquor originates from (i) free moisture in the coal charge, (ii) bound, or combined water in the coal, and (iii) additional water introduced to the plant, for example by steaming during oven charging or for hosing down plant. On a typical coke oven and by-product plant, the liquor production rate is around 120 litres/ton (l/t) to 130 l/t of coal carbonized. The flushing liquor supply to the battery is around 4,000 l/t of coal carbonized, depending upon the battery and collecting main arrangement. The liquor contains dissolved ammonia and a wide variety of compounds. These are normally grouped as ‘fixed’ or ‘free’ compounds, Fixed compounds are relatively stable while the ‘free’ compounds can be decomposed solely by heating to around 100 deg C. The ‘free’ compounds include the hydroxide, carbonate, bicarbonate, sulphide and cyanide while the ‘fixed’ compounds include the chloride, thiocyanate, thiosulphate and sulphate.

In addition, the ammoniacal liquor contains smaller but important concentrations of dissolved neutral oils, phenols, and tar bases (pyridine etc.). The ammonium chloride content of the ammonia liquor is important, since this compound represents almost the entire chlorine present in the coal charge and the ammonium chloride is contained in the condensates arising in the gas collecting main and the primary gas cooling. In view of the particularly corrosive nature of chlorides, and the fact that they do not vapourize readily at normal process temperatures, their concentration has to be avoided in any system for ammonia liquor treatment or disposal. For example, use of this liquor or derivative as a make-up source to coke quenching systems has frequently aggravated the corrosion problems in these systems. Because of this, particularly when new plant processes are being considered, it is necessary to make a rough chloride balance to confirm that high concentrations are prevented.

Source and composition of the tar

The compounds present in the coal tar are formed ultimately from complex organic materials in the coal charge. The heating process during coal carbonization breaks down some of these materials into simpler compounds, and further reforming reactions also take place as these vapours pass through the hot coke mass forming in the oven chamber and through the hot space above the charge. The tar is an extremely complex mixture of organic compounds which condense to medium to dark liquids or solids at room temperature, mostly aromatic compounds. Typically, they range from light oils such as benzene, in a fraction boiling up to 200 deg C, through to the pitches, the residue after distillation up to 350 deg C.

The tar normally has a specific gravity higher than 1 and is somewhat viscous, both values depending on the source of the fraction within the plant and the features of the coal carbonization. The first fraction which condenses when the flushing liquor cools the oven exit gases to the temperature in the range 80 deg C to 84 deg C, corresponds to the heavier and more viscous fraction consisting of the higher boiling point compounds. The second fraction, which condenses in the primary gas cooler, has an appreciably lower specific gravity and lower viscosity. This fraction is removed as droplets in the electrostatic tar precipitator and has even lower specific gravity and viscosity. Typically, the collected bulk tar has a specific gravity of around 1.2 and viscosity in the range 300 centistokes (cS) to 3,000 cS at 35 deg C, with the viscosity reducing rapidly as the tar temperature increases.

The types and properties of the coal tars vary considerably depending on the nature of the coal charge and the coal carbonization conditions. High temperature carbonization, such as in the modern coke oven batteries, subjects the tar vapours to more rigorous thermal cracking than lower temperature processes. Hence, the tars from the higher temperature batteries contain higher proportions of the aromatic types of hydrocarbons. The yield of coal tar is around 30 l/t to 50 l/t of dry coal.


Each of the tar fractions collected in the by-product plant contain quantities of solid materials. These range in size from near-colloidal particles to appreciably large pieces. While almost all the solids are generated in the coking process, some of the larger pieces are foreign in nature, e.g., refractory materials or substances inadvertently introduced during the by plant operation into the system. The smaller particles are normally carbonaceous and can form agglomerates, while almost all of the larger particles can be separated from the liquid tar by sedimentation procedures, which are normally carried out in stages.

Tar / liquor separation

The initial separation of tar and liquor is carried out in a flushing liquor decanter, a unit based on gravity separation. Since the ammonia liquor has specific gravity of around 1 and that of the tar is around in the range of 1.1 to 1.2, hence, in principle such a separation is not difficult. There are, however, several features in an operating plant which make the process more complex. These features include the very high flushing liquor circulation rate which does not allow sufficient time to make a thorough tar / liquor separation. The presence of solids in the system is to be taken into account.

Because of the low liquor viscosity and reasonable difference of tar / liquor densities, it is expected to be easy to achieve a tar-free liquor. However, some tar has an apparent density less than that of water and is hence float. The high tar viscosity retards separation of small liquor droplets and fine solids from the tar. The operating temperature (around 75 deg C to 80 deg C) of the decanting unit means that the vapours above the ammonia liquor and tar contain corrosive, condensable and toxic components such as hydrogen sulphide and hydrogen cyanide, water, benzene, and naphthalene.

In the case of the flushing liquor, the requirement of the coke ovens and by-product operating plants are met by making a satisfactory compromise by using the decanting equipment to remove larger solids and almost all of the tar by a high-rate gravity settling process. Since the flushing liquor has to pass through spray nozzles at the point of use, any solids large enough to cause blockage problems, are removed by a strainer located in the flushing liquor pump delivery line. While the earlier coke ovens and by-product plants used basket type filters for this purpose which needed manual cleaning, the modern equivalent motor-driven filters are continuously automatically back-flushed and solids removed from the flushing liquor are returned to the decanting plant.

In addition to the main use for cooling the gases leaving the individual coke ovens, flushing liquor is used in by-product plants as a convenient liquid to clean or maintain pipework and equipment free from blockage. There are several advantages in using flushing liquor in this way. Since the ammoniacal liquor contains only components dissolved from the coke oven gas, it is chemically compatible with the majority of the other condensates and process liquors in the by-product plant, and does not introduce additional corrosion or precipitation problems. For a similar reason, it is advantageous if the flushing liquor contains some free tar, since several of the plant piping and equipment deposits contain such materials as naphthalene, these materials are normally readily soluble in hot tar.

A most important advantage to using flushing liquor to keep plant free, rather than use fresh water, is that it helps to maintain a tight plant water balance. In modern by-product plant practice, it is necessary to minimize plant effluent discharges, any additional water added to the process appears ultimately as extra aqueous effluent flow which has to be treated before discharge.

As described above, all the tar streams generated on the plant contain some types of solids. The larger particles and pieces are normally readily separated from the tar in the initial flushing liquor decanting equipment. Removal and handling of solids removed from tar is never an easy operation, it is complicated by the ever-present tar. Although a substantial proportion of the larger solids can be removed when the tar is held for a period of 6 hours to 12 hours, the finer particles are difficult to separate, even after days or weeks of storage. Over the years, several attempts have been made to use centrifuges to eliminate this issue. While in some cases, it has been possible to achieve useful separations, it has never been really acceptable for long-term adoption as a reliable process. Ultimately the near colloidal sized particles are present in the tar shipped out from the plant and the downstream user has to be aware of the situation and develop the process or plant accordingly.

It is a normal objective to remove the tar moisture content to 2 % to 3 % in the initial separation stage. As indicated above, the separation of a relatively tar-free flow of ammoniacal liquor is helped by the low liquor viscosity and normally heavier tar density. The large flow needed for return to the coke oven battery as flushing liquor is separated at an early stage, the remaining liquor flow represents the surplus, or production, quantity and passes forward to a secondary decanting / settling stage which takes place over a longer period.

Types of flushing liquor decanters

Rectangular – It is by far the most widely used type of decanter which is shown diagrammatically in Fig 2. This is basically a flat-bottomed rectangular container. At one end the base slopes upwards at an angle to house the solids discharge mechanism. The mixture of tar, liquor, and solids to be separated flow into one end of the decanter and pass along the unit, bulk separation of the two liquid phases taking place rapidly into an upper liquor layer and low tar layer. Individual units can be up to around 350 cum capacity, normally the capacity allows up to 15 minutes to 20 minutes liquor residence time and several hours for the tar. Separated liquor flows over an outlet weir and out to an intermediate storage tank, from which the flushing liquor is pumped to the coke oven battery and the surplus liquor is transferred to main storage before processing.

Fig 2 Rectangular tar decanter

Solids and heavy tar, a mixture of solids and viscous tar, fall to the bottom of the decanter. This material is moved along the decanter base and over the sloping outlet riser by a very slow-moving scraper conveyor to be discharged into a receiving hopper. The settled tar is withdrawn through an adjustable overflow. This can be used to raise or lower the interface between the tar and liquor in the decanter. Raising this interface increases the volume, and hence the settling time available, of the tar. This feature allows some control to be exercised over the tar moisture content as it passes to storage.

Tar decanter designs include low pressure steam heating probes or coils to maintain tar temperature to reduce tar viscosity. The operating temperature of these units is in the range of 75 deg C to 80 deg C, depending mainly on the temperature of the tar and liquor stream entering. Some units have been built to operate under a slight positive pressure in order to allow the tar temperature to be higher so that solids and water separation are easier. These units are built as horizontal cylindrical vessels, rather than rectangular, but otherwise retained the same general configuration for separation and decanting.

Double-cone – It consists of an inner conical tank mounted within an outer vessel, the primary rapid separation of tar and liquor to produce the flushing liquor return flow taking place in the annular space between the cones. The flushing liquor flow is withdrawn from a point diametrically opposite to the introduction of the combined tar and liquor input. The main rapid flow is hence around the conical annulus. Solids fall to the bottom of this annular space and are discharged periodically through a double-valved hopper arrangement, the upper valve of which remains open in the solids collection phase.

Surplus liquor and tar flow up through a central riser and enter the inner conical vessel. This arrangement provides good separation conditions for the tar and liquor, with the positive features being that there is no moving mechanism present and that in particular the tar section temperature is maintained by the hot liquor in the annular space. Settled excess liquor flows continuously out from the upper section of the inner cone, and settled tar is withdrawn continuously from the bottom of this cone through an external line. The level of the nominal tar / liquor interface is controlled by the variable height adjustable overflow vessel on the tar outlet line. This type of unit is shown schematically in Fig 3.

Fig 3 Double cone decanter

Vertical – In this decanter, the decanting and separation of tar and liquor takes place in a vertical cylindrical vessel. This vessel is of simple construction, it is divided into two separate compartments by an inner cone in the lower part. Before entering the decanter, the combined battery tar and liquor flows pass through a solids’ pre-separator. This small vessel incorporates a scraper conveyor to remove solids and any heavy tar from the system to prevent the blockages in the decanter.

The tar and liquor flow enters the decanter chamber through an inlet pipe which feeds it into the unit somewhat above the nominal tar / liquor interface. Tar settles in the lower section and is withdrawn continuously from the base through an adjustable overflow for transfer into intermediate and final storage. Liquor overflows from the upper level of the chamber and passes into the lower compartment, where it helps to maintain the tar temperature. Flushing and excess liquors are withdrawn from this lower compartment continuously for recirculation or storage respectively. The vertical type of decanter is shown in Fig 4.

Fig 4 Vertical decanter


Liquor – It is normal to provide a common flushing liquor tank to collect liquor from all the decanters, this tank only has a capacity corresponding to a few minutes of the recirculation rate but provides a valuable buffer storage volume to allow for example pump change over. The main ammonia liquor storage system on a coke oven and by-product plant has to be such as to further prepare this material for the downstream processing. As indicated above, the main impurities in the ammonia liquor at this stage are small quantities of tar droplets which sink naturally frequently together with some tar material having an apparent density less than the liquor and hence a floating fraction. As a result, the ammonia liquor storage tanks have to be provided with a bottom tar withdrawal means, the liquor outlet being somewhat above the tank bottom.

Similarly, a top-level overflow can be provided to remove any floating fraction. The tank vents have to be maintained free from solid naphthalene deposits, e.g., by clean gas purging, or heating, or both. The number and capacity of ammonia liquor storage tanks are determined by predicted upstream or downstream plant maintenance outage times as well as the need to provide an extended period for settling to provide a clean and tar-free liquor. It is also to be remembered that this liquor contains very low but important concentrations of such dissolved organic materials as benzol, naphthalene, phenols, and tar bases.

Tar – The treatment and storage equipment for tar discharged from the primary decanting system is to be designed to produce a tar of low moisture and solids content to suit market conditions. The first stage of storage / treatment for the tar is frequently carried out in horizontal cylindrical tanks, sized to contain individually 24 hours of plant tar production. These units are heated by internal or external low-pressure steam coils to bring the tar temperature up to 90 deg C to 95 deg C, and are operated batch-wise. The high temperature, and hence reduced tar viscosity, assist in allowing the small water droplets in the tar to be separated. These units are frequently referred to as tar dehydrators.

Tar from the dehydrators, if used, is transferred to larger intermediate and final storage tanks. These tanks are normally provided with a form of low-pressure steam heating to maintain tar temperature at around 60 deg C to 80 deg C. Tar storage is frequently operated stage-wide to allow finer solids particles to be deposited before final discharge. The finest particles can be almost impossible to remove by gravity settling. At this stage the storage tanks can be fitted with slow moving side-entry mixers to prevent further solids settlement. Since majority of the tar storage tanks are to be cleaned out by physical removal of settled solids deposits, the tanks are to be fitted with large side doorways for equipment access, and the steam heaters can take the form of layer coils in the base or top-entry heater units. The latter type can be removed and cleaned without taking the tank out of service.

Ammonia removal processes

Ammonia (NH3) is a by-product of the coking of coal. It is a constituent of the coke oven gas leaving the coke ovens, with a typical concentration in raw coke oven gas of 6 grams per normal cubic meter (g/N cum). The solubility of ammonia in water leads to its presence in the coke oven battery flushing liquor with a typical concentration of 5 g/l to 6 g/l total ammonia. As a result, because of the net production of flushing liquor in the coke oven and by-product plant, referred to as excess flushing liquor but also known as coal water, virgin liquor, or weak ammonia liquor, there arises a liquid stream as well as a gas stream from which ammonia is to be removed. The quantity of excess liquor is around 12 % by weight of the dry coal throughput, depending on the coal moisture content.

Ammonia removal from gas streams remains a universal feature of coke oven and by-product plants. The reason for this is that ammonia, in the presence of the other coke oven gas contaminants such as hydrogen cyanide, hydrogen sulphide, oxygen, and water, is extremely corrosive to carbon steel. Earlier, an added incentive was the profitable sale of by-products such as ammonium sulphate, anhydrous ammonia, and concentrated ammonia solutions. With the possible exception of anhydrous ammonia, the reduced market value of these by-products, especially in the industrialized nations, no longer makes their production profitable.

Ammonia removal from liquid streams is performed mainly for environmental reasons. The primary ammonia handling equipment in the by-product plant deals with the removal and disposal of the ammonia present in the coke oven gas. However, these systems frequently include facilities to handle the ammonia arising in the excess flushing liquor. To help understand how such facilities are incorporated into the overall ammonia handling system, First the ammonia removal processes are described with the treatment of ammonia in the excess flushing liquor. The main processes for removal of ammonia from coke oven gas are then described. As these processes generate in several cases a concentrated ammonia vapour stream, and finally the alternatives available for treatment of this stream are described.

Treatment of excess flushing liquor

Available options – In some of the locations, 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 ammonia from the liquor by distillation, normally followed by final treatment in a biological effluent treatment plant (BETP). The biological effluent treatment plant can be on site or it can be operated by the local authority in which the coke oven and by-product plant is located. It is possible to use biological effluent treatment alone to remove ammonia from excess flushing liquor, however, the size and operating cost of a biological effluent treatment plant is considerably reduced when preliminary removal of ammonia by distillation is performed.

Distillation of excess flushing liquor – The distillation of excess flushing liquor involves feeding the liquor to the top of a trayed distillation column, normally called an ammonia still, and feeding a counter current flow of stripping steam at the bottom. The stripping steam distils off the ammonia 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 on-site biological effluent treatment plant. Typical levels of total ammonia in stripped liquor range from less than 50 parts per million (ppm) to 150 ppm.

Not all the dissolved ammonia present in excess flushing liquor is readily steam strippable. Several chemical species present in the flushing liquor, lead to the formation of different ammonium salts in solution. These include ammonium carbonate, ammonium chloride, and ammonium sulphate among others. Salts such as ammonium carbonate are easily decomposed by heat in the still to yield free molecules of ammonia. However, other salts such as ammonium chloride and ammonium sulphate are not decomposed and retain the ammonia in a ‘fixed’ form. The fraction of fixed ammonia to total ammonia in excess flushing liquor is typically 20 % to 50 %. To allow the distillation of the fixed ammonia, the excess flushing liquor is to be made alkaline. The typical reaction which then takes place, liberating free molecules of ammonia is NH4+ + OH = ΝΗ3 (g) + Η2O.

In addition, alkali is determined by mass balance, based on chemical analysis of the excess liquor to give the concentration of the fixed ammonia present. The form of alkali used in the distillation of excess flushing liquor has changed over the years. For several 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 issue with fouling. The ammonia stills needs a considerable marking of overdesign to allow continued operation while partially fouled, but even so they had to be taken out of service regularly for cleaning.

For avoiding the issue 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 needed. 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 (caustic soda) solution has made this the alkali of choice for present day design of ammonia stills. Caustic soda is the most expensive of the alkalis traditionally used, but its consumption can be closely controlled which is of great benefit where limitations are imposed upon still effluent pH.

The non-fouling characteristics of caustic soda allows the use of more economical still designs, incorporating valve trays in place of the traditional bubble cap trays. In practice, the caustic soda is injected into the ammonia still near to, but not at, the top tray. This allows dissolved acid gases such as hydrogen cyanide and hydrogen sulphide to be stripped from the liquor first, before they can react with the caustic soda to form fixed salts.

Materials of construction for ammonia stills are chosen for their corrosion resistance. Cast iron has traditionally been used, with generous allowances for corrosion, although it is now frequently more economical to use materials such as Hastelloy, a nickel-based alloy, titanium and 316 grade stainless steel. The upper sections of ammonia stills where both ammonia and acid gases are present normally need the use of highly corrosion resistant materials.

A common operating issue in the distillation of excess flushing liquor is the presence of tar carryover which can lead to serious fouling in the still. The normal 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.

Ammonia removal from coke oven gas

The presently used methods for removal of ammonia from coke oven gas are variations of three basic processes namely (i) the ammonium sulphate process, (ii) the Phosam process, and (iii) the water wash process.

The ammonium sulphate process – The ammonium sulphate process removes ammonia from the coke oven gas by absorption in a solution of ammonium sulphate and sulphuric acid. The ammonium sulphate produced by the reaction of ammonia with sulphuric acid is recovered by crystallization. The crystals are then centrifuged, washed, and dried. The different ammonium sulphate systems in operation differ in the type of gas / liquor contacting device and the type of crystallization equipment used. An early and still very commonly used system employs a dip tube extending below the surface of the acid / ammonium sulphate solution in a vessel referred to as a saturator (Fig 5).

Fig 5 Ammonium sulphate process

The solution strength is maintained around 4 % acid. Coke oven gas is fed through the dip tube and gas / liquid contact is affected as the gas bubbles move up through the solution in the saturator. Acid is continuously added to the saturator. The heat of reaction between ammonia and sulphuric acid causes the evaporation of water into the coke oven gas. 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 system. 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. The traditional material of construction for the saturator and all wetted surfaces is lead lined carbon steel. Alloys such as Monel and 316 grade stainless steel 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 316 grade stainless steel has allowed the development of the modern ammonia absorber systems. In these systems, a circulating stream of ammonium sulphate / sulphuric acid solution is sprayed counter currently to the coke oven gas flow in an absorber vessel. Absorption of ammonia 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 systems results in an increase in the heat content of the coke oven gas leaving the absorber or saturator. The reason for this is that to maintain the water balance in the system, especially in the case of saturators, water is required to be evaporated into the gas stream. In addition to the water added to the system 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 ammonia absorbers / saturators. The evaporation of water into the coke oven gas 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 hydrogen sulphide removal to be operated effectively, the gas is to be cooled in ‘final’ or ‘secondary’ gas coolers. The ammonium sulphate processes accommodate ammonia from excess flushing liquor to be feeding the overhead vapours from flushing liquor distillation into the coke oven gas main upstream of the ammonia saturator / absorber. The major economic disadvantage with ammonium sulphate processes is the price relationship between sulphuric acid and ammonium sulphate. The sulphuric acid needed to make ammonium sulphate can cost up to two times the value of the ammonium sulphate product.

The Phosam process – The Phosam process was developed by United States Steel as a means of producing a saleable, commercially pure anhydrous ammonia product from the ammonia present in raw coke oven gas. This high value product makes the process much more economically viable than the ammonium sulphate processes. In the Phosam process, ammonia is selectively absorbed from the coke oven gas by direct contact with an aqueous solution of ammonium phosphate in a two-stage spray absorption vessel.

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 (i) H3PO4 + NH3 = NH4H2PO4, (ii) NH4H2PO4 + NH3 = (ΝΗ4)2HPO4, and (iii) (NH4)2HPO4 + NH3 = (ΝΗ4)3PO4. The ammonia 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 of around 1.3 MPA. The reason for this is that the reversible reactions which liberate the ammonia from solution are favoured by higher temperatures. Hence, 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 ammonia. These vapours are condensed and then fed to a fractionating column where anhydrous ammonia is recovered as the condensed overhead product. The fractionator bottoms product, mainly water, leaves the system as effluent.

The Phosam process can become contaminated by tar and by absorption of acid gases (hydrogen cyanide, hydrogen sulphide, and carbon di-oxide) 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 ammonia rich solution and feeding it into a vessel referred to as contractor. In this vessel, the preheat causes vapourization of water and acid gases from the solution. These vapours are vented back to the coke oven gas 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 ammonia, is to add sodium hydroxide to the fractionator feed. The sodium hydroxide fixes the acid gas compounds as non-volatile sodium salts which remain in the fractionator bottoms effluent stream. Fig 6 shows the schematic diagram of Phosam process.

Fig 6 Schematic diagram of Phosam process

An important operational feature is the control of the water balance in the system. Substantial quantities of steam are condensed in the solution stripper, and this condensate is to be re-evaporated from the circulating solution into the coke oven gas stream. The temperature of the solution returning to the absorber is around 60 deg C, and hence the coke oven gas becomes heated as it flows through the absorber. The increased gas temperature normally makes it necessary to install a final gas cooler after the Phosam absorber. The addition of phosphoric acid to the absorption solution is needed only to account for operating losses such as spillage. It is added at weekly intervals at a rate equivalent to 7.5 grams phosphoric acid (H3PO4) per kilogram ammonia produced.

The Phosam process is very efficient, capable of achieving higher than 99 % recovery of ammonia from coke oven gas. Other plant configurations are possible in which, for example, aqueous ammonia solution is produced instead of the anhydrous ammonia. Materials of construction are stainless steel for all areas in contact with phosphate solution of aqueous ammonia, and carbon steel for other areas. As in the sulphate process, ammonia present in excess flushing liquor is handled first by distillation, with the vapours being fed to the coke oven gas upstream of the Phosam absorber.

Water wash process – One of the simplest and most frequently used methods of removing ammonia from coke oven gas is to absorb it in water. Aqueous absorption liquor is fed counter currently to the flow of coke oven gas in an ammonia washer vessel. The vessel can be designed as a spray type absorber with several liquor respray stages, or as a packed tower as is common in several German designed plants. The type of packing normally used is vertically arranged expended metal sheets which promote gas / liquor contact but resist playing and fouling.

The rich ammonia solution formed, with a typical concentration of 5 g/l to 8 g/l, is then fed to a distillation column where the ammonia is stripped from the aqueous liquor using steam. The ammonia and water vapours leaving the top of the stripping column are passed on for subsequent treatment in a variety of ways which are described earlier. After stripping, the absorption liquor is cooled and returned to the washer. There is a continuous blowdown of stripped liquor from the circuit which is equivalent to the volume of steam condensed in the stripper column. This blowdown is plant effluent and needs biological effluent treatment to fully remove the residual ammonia. As no chemical reactions are involved, other than the dissolving of ammonia in water, the water wash process is temperature dependent and is most efficient at low coke oven gas temperatures (20 deg C to 30 deg C).

The ammonia washer vessel is normally 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 is operated at positive pressure. To promote ammonia removal efficiency, gas cooling is needed to remove this superheat and to cool the gas to the optimum temperature range. The gas cooling stage is frequently incorporated into the ammonia washer vessel itself. The ammonia 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 considerable quantities of acid gases (hydrogen cyanide, hydrogen sulphide, and carbon di-oxide) from the coke oven gas. Hence, the ammonia stripping column is nowadays being frequently constructed of corrosion resistant materials such as titanium and 316 grade stainless steel, although several plants continue to operate cast iron stills. The stripping columns is equipped with bubble cap trays or with more economical valve trays. Because of the lower liquor temperatures in the washer and hence the reduced rate of corrosion, this vessel can be constructed entirely in carbon steel. Fig 7 shows the schematic diagram of the water wash process.

 Fig 7 Schematic diagram of water wash process

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 ammonia in the washer. The advantage of doing this is that as the excess liquor is going 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 ammonia. For other plant effluent streams, it frequently makes sense to perform steam stripping as a preliminary effluent treatment step. Combining this with the ammonia absorption process minimizes overall steam consumption for the by-product plant. The excess flushing liquor is added at a point in the washer where its free ammonia concentration most closely matches the free ammonia concentration of the absorption liquor.

The presence of fixed ammonia does not influence absorption of free ammonia from the coke oven gas. 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 ammonia can be removed in the stripping column by the addition of caustic soda. Alternatively, the blowdown stream can be fed to a separate fixed ammonia still. The use of a separate fixed ammonia still avoids the presence of alkali in the recirculating absorption liquor, which can be responsible for forming fixed compounds with acid gases such as hydrogen cyanide and leading to the presence of these compounds in the by-plant plant effluent stream.

Treatment alternatives for ammonia vapour streams

After the removal of the ammonia from the coke oven gas, a vapour stream containing ammonia, water, and acid gases(hydrogen cyanide, hydrogen sulphide, and carbon di-oxide) gets generated. The processes which have been used or proposed for the further treatment of these vapours are (i) incineration of the vapours, (ii) production of concentrated ammonia liquor, (iii) catalytic ammonia destruction, (iv) production of ammonium sulphate, and (v) production of anhydrous ammonia. The latter two are modification of the ammonium sulphate process and the Phosam process. In each case, the ammonia vapour stream replaces the coke oven gas feed stream. Naturally the size of the equipment needed for absorption of the ammonia is reduced. In the case of production of ammonium sulphate, this plant arrangement is referred to as the ‘indirect’ process. The remaining three alternatives are described below.

Incineration – Incineration of ammonia vapours has been widely practiced from the 1960s. The process simply reacts the ammonia with air in specially designed burners. The process is exothermic and no support fuel is needed other than for a pilot flame to initiate the combustion. The incinerator is refractory lined for heat resistance. The products of combustion, mainly nitrogen and water vapour, are emitted to atmosphere through a high stack. Depending on the source of the vapours to be incinerated, they can contain considerable quantities of the acid gases (hydrogen sulphide and hydrogen cyanide). These components are also incinerated, with any sulphur compounds present producing sulphur di-oxide.

The single stage combustion of ammonia occurs at high flame temperatures in excess air. These circumstances promote the formation of oxides of nitrogen (NOx) which leave with the stack gases. To combat this effect, ‘low NOx’ incinerator design uses a two-stage combustion with intermediate cooling (by water sprays) between the stages. The first stage is used to incinerate majority of the ammonia, with carefully controlled air flow rates so that excess air is avoided. In the second stage, excess air is allowed to complete the incineration but at relatively low flame temperatures. By these means the NOx content in the stack gases can be held to less than 100 ppm. The presence of oxides of both nitrogen and sulphur in the stack gases has led to a close review of the suitability of this process with regard to atmospheric pollution regulations. New installations of this process are frequently nowadays limited to standby or emergency applications as a short-term gap measure, for example, during a maintenance outage of a more environmentally friendly process.

Production of concentrated ammonia liquor – Concentrated ammonia liquor in different grades of purity can be prepared from the ammonia vapour stream. Normally, a preliminary condensation is performed to reduce the quantity of water vapour present. The condensate is returned to the ammonia distillation column. The remaining vapours are then condensed to form the product ammonia liquor. The vapour temperature after the preliminary condensation determines the remaining water content and hence determines the concentration of the ammonia liquor produced. With the presence of acid gases in the ammonia vapours, particularly carbon dioxide, the concentration of ammonia liquor produced by condensation of the vapours is limited to 15 % to 20 % ammonia. This is since ammonium carbonate is also formed and crystallizes on the condenser surfaces at higher concentrations.

The problems caused by acid gases can be eased to an extent by a prior removal of these compounds from the enriched wash liquor before it is fed to the ammonia still. This is done in a ‘de-acidification’ column, which is a steam stripping column. Relatively small flow rates of steam are used to strip the more volatile acid gases from the wash liquor, leaving the ammonia largely in solution. The acid gases can be returned to the coke oven gas main, and the de-acidified solution then passes to the ammonia distillation column. A common method of cooling and condensing ammonia vapours is by direct contact with recirculated cooled ammonia solution in a packed column. This method avoids localized sub-cooling of the condensate which can lead to crystallization and fouling.

Carbon steel can be used to handle concentrated ammonia liquor below 50 deg C. At higher temperatures more corrosion resistant materials such as 316 grade stainless steel, Hastelloy, and titanium can be necessary. The production of concentrated ammonia liquor is also used to provide a low cost standby in plants where the normal ammonia handling facilities are out of service for maintenance.

Catalytic ammonia destruction – The process was first developed and installed by the company ‘Firma Carl Still’ in 1968, this process has gained popularity because of its ability to dispose of the ammonia removed from coke oven gas without the raw material cost of the sulphate process and without the emission of pollutants. The process begins with the ammonia vapours coming from the ammonia distillation column. Normally, a partial condenser is used to reduce the content of water in the vapours. The condensate is returned to the column. The remaining vapours, including any acid gases, flow into a top mounted burner installed on a refractory lined ammonia destruction reactor. Coke oven gas support fuel and a stoichiometric amount of combustion air are also fed to the burner, at closely controlled flow rates. Within the reactor is a nickel catalyst bed. Fig 8 shows schematic diagram of the catalytic ammonia destruction.

Fig 8 Schematic diagram of catalytic ammonia destruction

The reaction kinetics favour the combustion of the coke oven gas with the air. This results in a reducing atmosphere consisting of the ammonia vapours and coke oven gas combustion products at a temperature of 1,150 deg C. Over the nickel catalyst, ammonia is cracked to nitrogen and hydrogen. Hydrogen cyanide reacts with the water vapour present to form nitrogen, hydrogen, and carbon monoxide. Hydrogen sulphide passes through the catalyst bed unreacted. The reactor gases then flow through a waste heat boiler where they are cooled to around 300 deg C, by raising high pressure steam.

The gases can then be passed directly to the raw coke oven gas main, closing the process loop, or they can first be further cooled by water quenching in a tail gas cooler vessel. With careful control of the coke oven gas and air flow rates, no oxides of nitrogen or sulphur are formed, and no excess air is returned to the coke oven gas main. The addition of the tail gas results in a small reduction in the calorific value of the coke oven gas. This is because of the volume of nitrogen which enters the process as combustion air. Materials of construction in the ammonia destruction plant are primarily carbon steel, with Hastelloy or titanium for the ammonia vapour feed lines. Refractory lining is used where necessary for heat protection.

General overview of a modern by-product plant

The scheme of a modern by-product plant is described here. The raw coke oven gas of the coke oven batteries including the loaded flushing ammonia liquor, is taken over from the crude gas mains at defined points. Leaving the downcomer, the gas is directed to the primary gas coolers. In the primary gas cooler, the coke oven gas is cooled down within indirect cooling stages using both cooling water and chilled water. Using a continuously operated flushing system inside the primary gas cooler for cleaning purposes, the operation time is optimized. The tar water / ammonia water mixture leaving the downcomer is directed to the static tar separation plant.

From the tar separation plant the flushing liquor is returned to the collecting main for flushing the coke oven gas coming from coke oven batteries, while coal water is discharged through a tank to the gravel filter plant for final tar removal. Leaving the tar separation plant the crude tar, as product, is led through the storage and loading station for further disposal. After collecting and preparing within the tar separation plant, by tar solid decanters, solid particles are finally discharged into the battery feed coal.

Coming from the primary gas coolers, the cooled coke oven gas flows through the electrostatic tar precipitators for reducing the tar content upstream the gas exhausters. The cooled and tar reduced coke oven gas passes the gas exhausters. The needed performance and suction pressure of the gas exhausters are to be considered for the pressure difference of the entire gas treatment plant, beginning from coke oven batteries down to the inlet of the coke oven gas network. After leaving the gas exhausters, the coke oven gas passes a scrubber system with integrated final and internal cooling stages for reducing the hydrogen sulphide / ammonia components. By these cooling stages, operated indirectly with chilled water, the reached temperature level secures the needed performance values of the scrubber system. The scrubbing liquor for the hydrogen sulphide /ammonia removal consists of stripped and de-acidified water supplied by the distillation plant.

As final gas cleaning stage a benzene-toluene-xylene (BTX) scrubber/ naphthalene scrubber is applied. In order to reduce the BTX (benzene, toluene, and xylene) and naphthalene content in the coke oven gas coming from the hydrogen sulphide / ammonia scrubbers, the gas is scrubbed by special tar-based wash oil. For the regeneration of the wash oil and the production of crude BTX as product, a stripping system is provided. In order to strip the hydrogen sulphide and free ammonia compounds from the enriched scrubbing liquor coming from the hydrogen sulphide / ammonia scrubbers, a distillation plant consisting of de-acidifiers and ammonia strippers is provided with the capability to handle the entire scrubbing liquor. The excess coal water, leaving the gravel filter plant, is stripped under the presence of caustic soda in a fixed ammonia still column section to remove the fixed ammonia compounds. The surplus outlet stream is led to the biological effluent treatment plant.

The vapours leaving the distillation plant, mainly consisting of hydrogen sulphide, ammonia, hydrogen cyanide, carbon di-oxide, and hydrocarbons, are treated in a combined ammonia cracking / elementary sulphur unit (Claus unit). In this process the ammonia / hydrogen cyanide components are cracked while the hydrogen sulphide content is converted to liquid sulphur. The tail gas flow of the Claus plant is returned to the crude gas mains upstream the primary gas coolers. Fig 9 shows the overview flow diagram of the by-product plant.

Fig 9 Overview flow diagram of the by-product plant

The technical descriptions of the three main units are given as examples for given blow.

Chilled water plant – A high performance scrubbing system needs for the absorption process a temperature decrease. Hence, a chilled water system is needed. This chilled water system allows keeping optimum coke oven gas and process water temperatures even during hot periods. The technical standard and the structure of the chilled water plant, comprising of absorption type (e.g., lithium bromide system) or compressor type refrigerating units. The type which is to be considered depends on the availability of steam and electrical power as well as the required technical specifications. The chilled water plant is to be designed with respect to naphthalene precipitation risks in the primary cooling area and regarding the minimization of heat exchanger areas. Feed temperature of the chilled water is around 14 deg C and return temperature amounts to 24 deg C.

BTX scrubbing and recovery units – After leaving the ammonia scrubber the coke oven gas is directed to the BTX scrubbing system for removal the BTX compounds. Fig 10 shows a schematic flow diagram. The coke oven gas enters the scrubber at the bottom and moves upwards through the scrubbing sections before it leaves the scrubber at the top. The coke oven gas, reduced by BTX components, is directed to the coke oven gas distribution network. The scrubbing procedure is carried out by wash oil, coming from the stripped oil buffer tank. The wash oil enters at the top of the scrubber flowing downwards in counter flow to the rising coke oven gas. The scrubbing sections consist of expanded metal or structured packings and distribution trays. The enriched wash oil leaves the scrubber at the bottom through a sealing pot by gravity and is collected in a buffer tank before it is fed to the wash oil regeneration plant.

Fig 10 Schematic flow diagram of BTX scrubbing and recovery unit

Wash oil losses are to be compensated by adding fresh wash oil into the enriched wash oil buffer tank. The fresh wash oil is supplied from fresh wash oil tank. The enriched wash oil, leaving the BTX scrubbing system, is fed to the wash oil regeneration plant for reusing it for further wash processes in the BTX scrubber. Hence, a saleable BTX fraction is produced. Enriched wash oil from the BTX-scrubber, intermediately stored in a buffer tank, is pumped to the BTX stripper, passing two heat exchanger groups for indirect heating with hot stripped oil and steam, respectively. The BTX fraction is stripped by adding steam in the bottom area of the distillation column. The BTX fraction loaded with water vapour, leaves at the top of the BTX stripper and passes a condenser / cooler, sequentially, before entering a separation tank. The separation tank is equipped with special packings to reach an excellent separation between water / crude benzene. The crude benzene is fed to the storage tank by means of pumps while the separation water is discharged to the distillation unit. As a site fraction of the BTX stripper, the naphthalene oil fraction is drawn, collected in a drain tank and mixed to crude tar coming from tar storage tanks. Fig 11 shows flow diagram of wash oil regeneration plant.

Fig 11 Flow diagram of wash oil regeneration plant

The stripped wash oil leaves the bottom of the stripper and is pumped through three groups of heat exchangers. In the first heat exchanger group, the heat of the stripped wash oil is used for indirect heating up the cold enriched wash oil, coming from the BTX-scrubber plant. By applying the following second heat exchanger groups, the stripped wash oil is indirectly cooled with cooling and chilled water. Then, the stripped wash oil is directed to the buffer tank of the scrubbing system for further use in the BTX scrubber. Before feeding the hot stripped wash oil to the above-mentioned oil / oil heat exchanger, a small part of wash oil is led to the pitch column, supported by injected steam. By removing high viscose pitch components, the wash oil is kept in a feasible range of viscosity. The used wash oil is generated by a distillation fraction of crude tar from coke plants. The discharged residuals from the pitch column are mixed to the naphthalene oil / crude tar mixture, which is fed to the tar storage tanks. No additional treatment of waste oil / muck is necessary.

Combined ammonia cracking / elementary sulphur unit (Claus plant) – For processing the hydrogen sulphide / ammonia, vapours coming from the distillation plant, a combined ammonia cracking / elementary sulphur plant unit is provided. It consists of a high efficiency process with a two-stage Claus-reactors. Under the top pressure of the de-acidifier and ammonia stripping system, the hydrogen sulphide / ammonia vapours are led to the burner system of the cracking reactor. Operating at sub-stoichiometrical combustion conditions, at a proper temperature for reaction, a certain ratio of hydrogen sulphide is burned to sulphur di-oxide. The combustion air is supplied by an air blower. After preheating in a steam operated heater, the air is directed into the burner system. Especially, for start-up operation and after process interruptions partly cleaned coke oven gas, or natural gas is used for heating up and stabilizing the combustion.

The coke oven gas / natural gas is supplied by a gas blower. Inside the catalyst bed of the crack reactor the ammonia and hydrogen cyanide compounds of the vapours are cracked. Downstream of the catalytic bed, secondary air is supplied to adjust the stoichiometrical ratio of hydrogen sulphide / sulphur di-oxide for the reaction in the following Claus reactors. The hot process gas leaves the crack reactor and passes the waste heat boiler system. In this boiler system the process gas is indirectly cooled by generating steam through boiler feed water, added by pumps. The parameters of the produced steam can be selected in a certain range. It is possible to generate high pressure steam in combination with low pressure steam. During cooling down of the process gas within the low-pressure boiler, the first sulphur is condensed. After separation it is directed to the sulphur sealing pot. Fig 12 shows the schematic flow diagram of combined ammonia cracking / elementary sulphur unit.

Fig 12 Combined ammonia cracking/ elemental sulphur unit

By mixing the outlet gas of the high-pressure boiler with the outlet of the low-pressure boiler the needed inlet temperature to the first Claus reactor stage is adjusted. Within the first and second stage of Claus reactors, the process gas passes a catalyst bed. By that, the reaction between sulphur di-oxide and hydrogen sulphide takes place to produce liquid elementary sulphur plus water. Leaving the first stage of the Claus reactor, the process gas is directed to the first stage of a sulphur condenser. Similar to the low-pressure boiler, the process gas is indirectly cooled by boiler feed water generating low pressure steam. Then the process gas passes a separator for precipitation of the condensed sulphur. While the liquid sulphur is led to the sulphur sealing pot, the process gas flows to a downstream gas heater.

Using produced high-pressure steam of the high-pressure boiler, the process gas is indirectly reheated up to conversion temperature for the Claus reaction. After reheating the process gas is led to the second stage of the Claus reactor to convert more hydrogen sulphide / sulphuric acid and to produce further sulphur according to above mentioned reaction. Then, in the second stage of the sulphur condenser the gas is indirectly cooled while low pressure steam is generated. After cooling the process gas and sulphur precipitation in the downstream, the process gas (tail gas) with residual contents of sulphur di-oxide and hydrogen sulphide, is directed to the crude gas collection main in front of the primary gas cooler.

Leaving the separator, the sulphur flows into the sulphur sealing pot. The produced liquid sulphur, firstly collected in the sulphur sealing pot, is then discharged into a sulphur drain tank and from there pumped to the sulphur storage tank. Both, the sulphur-drain and the storage tanks are steam heated by coils. Periodically, the liquid sulphur is fed into transport vessels for distribution to further disposal by a loading station. The surplus steam, produced in high-pressure / low pressure boiler and sulphur condenser, is led to the steam distribution network for further utilization. Beside on a higher efficiency by using a two-stage Claus reactor, also the composition of the tail gas needs to be considered. For avoiding additional organic based sulphur in effluent, the Claus unit is designed as described below.

By using of a double stage Claus unit, the cracking of ammonia and the production of liquid sulphur from hydrogen sulphide / ammonia vapour feed are realized. Also, the hydrolysis reactions take place as per the reactions (i) COS + H2O = CO2 + H2S, and (ii) CS2 + H2O = CO2 + 2 H2S. Contrary to operating a double stage Claus unit, during operating a single stage Claus unit above mentioned hydrolysis reactions are not executed. COS (carbonyl sulphide) and CS2 (carbon di-sulphide) is not converted in a single stage Claus unit and is recycled through tail gas line to primary gas coolers and downstream plant components. These gas components react in downstream scrubber and distillation columns and cause through effluent the thiocyanate problem in the biological effluent treatment plant.

Results – The impurities of crude coke oven gas (depending on the coals used) in front of the gas treatment plant are tar – 15 g/N cum, hydrogen sulphide – 8 g/N cum, ammonia – 9 g/N cum, naphthalene – 8 g/N cum, and BTX – 33 g/ N cum. After leaving the gas treatment plant the impurities of the cleaned coke oven gas are tar – less than 0.02 g/N cum, hydrogen sulphide – less than 0.5 g/N cum, ammonia – less than 0.04 g/N cum, naphthalene – less than 0.1 g/N cum, and BTX – less than 5 g/N cum.

Comments on Post (1)

  • Mahesh

    very nice & detail information

    • Posted: 19 January, 2014 at 06:41 am
    • Reply

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