Methods of Refining of Crude Benzol
Methods of Refining of Crude Benzol
Benzol or benzole is a mixture of hydrocarbons of the benzene series, in which benzene itself predominates, in association with certain of its homologues and various impurities. Benzol recovered from coke oven gas is called as crude benzol (CB).
Crude benzol contains small quantities of a large number of impurities which consists of the unsaturated and sulphur compounds. These impurities have negative effects on the organic processes. Even insignificant sulphur impurities in benzene and toluene used in organic processes cause fast poisoning of the catalyst, and resinous substances formed as a result of polymerization of unsaturated compounds coat the catalyst surface and deactivate it. Some of these impurities have the proximity of the boiling points of the pure products. To produce pure products from CB, it is necessary to have its preliminary treatment to remove these impurities. Hence, any processing scheme includes preliminary preparation stages which ensure the removal to the required extent the components of sulphurous unsaturated and saturated hydrocarbons.
There are two processes for the refining of the CB. These are (i) sulphuric acid process, and (ii) hydro-refining process.
Sulphuric acid refining
The refining of the CB comprises the purification of the CB and its subsequent separation into the desired commercial products. This entails chemical washing and fractional distillation. The crude benzol is washed with concentrated sulphuric acid which sulphonates the more undesirable compounds and allows easier separation and recovery of the benzene, toluene and xylene (BTX) fractions. The sulphuric acid washing removes unsaturated compounds, olefins, dienes etc. together with any pyridine and some sulphur compounds. The olefins, dienes and other unsaturated compounds are removed by the sulphuric acid, partly by polymerization and resinification, and partly by solution in the acid.
The sulphuric acid method of refining of CB is popular for the by-product plants of the coke ovens. Issues regarding production of superior grades of benzene and utilization of wastes associated with the process have been largely overcome due to implementation of continuous cleaning using static mixers and the application of two-stage treatment with sulphuric acid and with additives of unsaturated compounds instead of treatment oleum, the development of recycling technologies of acid tar and still bottoms as air-oil emulsions supplied to a batch. Currently, the sulphuric acid refining is applied in two options for treatment a BTX fraction and when preparing benzene for synthesis.
Treatment is performed in a continuous process in the system of static mixers (one option is vapour mixers). Sulphuric acid in this process acts as a catalyst. The consumption of concentrated sulphuric acid reaches 70 kg/t of raw materials. The sulphuric acid treatment produces a thick, dark brown residue which is known as ‘acid tar’. Acid tar is heated in various ways by steam. As a result there is recovery of dilute sulphuric acid which can be used in ammonium sulphate production. The remaining material is a hard, carbonaceous solid residue, which is much more easily disposed of than the original acid tar.
The chemical washing with the sulphuric acid is followed by washes of water, dilute caustic soda, and again water, in order to free the benzol of the excess of sulphuric acid and any tar acids which may be present.
If it is desired to eliminate thiophene (C4H4S) and thiotoluene, which is necessary for producing pure benzene, a subsequent wash with concentrated acid is necessary. This effectively removes the sulphur containing compounds by sulphonation. Since it is costly and requires the use of additional acid, various, other processes for removing thiophene have been devised. Probably the most successful is the patented process, in which the result is attained by chlorination. The calculated quantity of chlorine is used, either as a gas or dissolved in water, and the chloro-compound formed with the thiophene is of comparatively high boiling point, and therefore remains in the residue from the subsequent fractional distillation.
Carbon disulphide (CS2) is one of the chief constituent of the benzol. It occurs mixed with low-boiling paraffins and some benzene. For its removal benzol is washed, for two hours and at a temperature of 65 deg C to 68 deg C, with a 10 % solution of caustic soda (NaOH), the washer being provided with a steam coil and agitator. The carbon disulphide combines to form sodium thiocarbonate according to the reaction 6NaOH + 3CS2 = Na2CO3, + 3H2O + 2Na2CS3. The spent wash-liquor is revived with lime or by other means, and is found more active after revival than when new. The hydrocarbon is rewashed with this revived liquor, and 95 %, of the disulphide is removed.
The purified fraction (washed CB) after separation from the spent acid, acid tar, neutralization, and steam stripping is then subjected to fractional distillation process. Fractional distillation is carried out with steam and it separates washed benzol into saleable products i.e. benzene, toluene, a mixture of xylenes and solvent naphtha. The produced benzene meets the specification requirements for nitration grade benzene.
For increasing the efficiency of rectification of CB by sulphuric acid, treatment for the isolation of thiophene concentrates and non-aromatic hydrocarbons, CB is subjected to repeated treatment with sulphuric acid with additives of unsaturated compounds and extracting rectification to decrease the content of saturated hydrocarbon impurities.
The rectification process naturally has some loss in volume of the liquid, and this varies from around 8 % to 10 %. Around two-thirds of this quantity being due to the operations of washing, and the remainder lost during distillation. The yield of rectified products obtained (as well as the amount of loss) varies considerably with the quality of the CB. The yield of clean products under normal operational conditions is around 92 % from the amount of CB processed.
The disadvantages of sulphuric acid refining are (i) excessive losses of clean products, (i) the generation of production wastes in the form of tar acids, and (iii) the generation of spent (re-generable) acid, the use of which for the production of ammonium sulphate is difficult.
Hydro-refining of crude benzol
Crude benzol consists of benzene, toluene, xylene, solvent naphtha, non-aromatics and residue. Hydro-refining treatment method of crude benzol is based on the hydrogenation reactions of sulphur and unsaturated compounds admixtures. During the hydro-refining of CB, treatment of CB is carried out in the presence of hydrogen (H2) gas over a catalyst under pressure. H2 gas is used as a hydrogenation agent. H2 gas is separated from coke oven gas by pressure swing adsorption (PSA) method. The process parameters (temperature, pressure, hydrogen/raw materials molar ratio, contact time, catalyst type) are selected in such a way so to ensure almost hydrogenation of entire amounts sulphur, unsaturated, oxygen- and nitrogen-containing impurities but by avoiding the hydrogenation reactions of aromatic hydrocarbons. H2 gas is mainly consumed for destructive hydrogenation of thiophene and CS2 and hydrogenation of cyclopentadiene and styrene.
Initially, the CB is purified from sulphur, non-aromatics and other compounds to produce BTXS (benzene, toluene, xylene, and solvent naphtha) raffinate for processing in the extractive distillation unit. This unit consists of three sections namely (i) de-fronting section, (ii) reaction section, and (iii) purification section. Typical flowsheet of hydro-refining process is at Fig 1.
Fig 1 Typical flowsheet of hydro-refining process
In this section, CS2 (boiling point 45 deg C) is removed from the CB and the product is called as de-fronted CB. CB from the storage tank is pumped to a surge tank, which is meant for intermediate storage. CB from the surge tank is pumped to the distillation column through feed pre-heater. The feed enters the column at the specified rate and at 70 deg C. The temperatures which are maintained at the top and the bottom of the column are 55 deg C to 65 deg C and 105 deg C respectively. Pressure in the column is maintained at 0.5 kg/sq cm. Sulphur content of the feed is in the range of 1800 ppm (parts per million) to 2000 ppm.. This is decreased to around 1200 ppm in the distillation column.
Distillation column consists of a number of bubble cap trays of which around the middle tray is the feed tray. Steam is fed into the reboiler, which heats the bottom product recycled to the column. The remaining bottom called de-fronted CB is fed to the reaction section through feed pre-heater. The sulphur is removed in the form of CS2. Simple distillation is carried out and due to heating, CS2 vapours rise to the top and these are condensed in a water condenser. Condensed CS2 is collected in CS2 vaporizer. Part of it is fed to the column as reflux and the other part is stored. The produced defronted CB is at 70 deg C. It is sent to the intermediate storage.
This section consists of reactors and evaporators. Here the hydro-refining takes place in the reactors provided which remove the oxygen, nitrogen and sulphur content in the CB.
The particles which are retained at the edges of the slots are scraped off. If the pressure difference between the inlet and the outlet streams increases and becomes too high, then the concerned filter is opened and cleaned. The filtered defronted CB is stored in the surge drum. The drum is set to around 2 kg/sq cm pressure. This pressure is controlled by feeding nitrogen and venting of the gases. From surge drum, the defronted CB is fed to pre-vaporizer at a pressure of 30 kg/sq cm using centrifugal pumps.
Pre-vaporizer is provided with turbulence promoters in the tube side to achieve high turbulence so that more heat exchange can occur and no scale formation is attained. This arrangement is provided as the feedstock in partial vapour stage (gas-liquid stage) to prevent rapid fouling of the tubes. This arrangement also provides easy cleaning of the tubes by simply pulling the turbulence promoters.
Similarly the liquid from the second stage flows to the first stage. This liquid is pre-heated in re-boiler and mixed with 85 % of the recycle gas in first mixing nozzle and again fed to the first stage. The temperature at the bottom of the evaporator is maintained at 210 deg C. Due to heating of the feed, the vapours are sent to the top and any residue or polymers in the feed are collected at the bottom. Part of the liquid from the first stage is fed to the residue flash drum, from where it is recycled to benzol distillation plant. The lighter vapours from the flash drum are fed to the surge drum. This residue is nearly 3 % to 4 % of the total feed.
Pre-vaporizer – It consists of a vertically mounted shell and tube heat exchanger. The feed is mixed with a part of recycles gas (containing H2 gas around 15 % of the total gas) before it is fed to the vaporizer. This feed is heated to around 160 deg C to 165 deg C by means of the main reactor effluent passing through shell side. The heated feed which is at a temperature of 160 deg C to 165 deg C is fed to the third mixing nozzle of stage evaporator.
The defronted CB is pumped to the defronted storage tank through a filter. The filter is provided to remove the solid particles and polymers, which can be present in the CB. The CB filter is an edge type filter and consists of a slotted tube inside a shell with a specified filter fineness, which is determined by the slots and scrappers. This is agitated by a hard crank.
Stage evaporator – The stage evaporator is a long cylindrical vessel which has three stages, separated by two plates. Demister pads are provided at the top of the evaporator. Each stage has a mixing nozzle. There are two re-boilers for the second and first stage respectively. A gas pre-heater is also provided in which the recycle gas (85 % of the total gas) is pre-heated to 210 deg C from the main reactor effluent. The two re-boilers are heated by hot oil through tubes to a temperature of 250 deg C. Recycle gas mixed with feed is passed through the shell side. Downcomers are placed so that the liquid in the third stage enters the second and from second to first. Pressure inside the stage evaporator is around 20 kg/sq cm.
The CB mixed with 15 % of the recycle gas is fed at the third mixing nozzle of the evaporator. The vapours coming from the second stage and the feed are mixed thoroughly and fed to the third stage. Lighter vapours are passed through the demister pads and to the pre-reactor. The liquid containing lighter and heavier substance is passed through downcomers to the second stage. Here the feed is mixed with the vapours from the first stage in the mixing nozzle and heated in re-boiler. This is fed to the top of the second stage.
The vaporization of CB in the evaporator is done by reduction of partial pressure of the CB, which is controlled by the addition of the recycle gas. This results in lower operating temperature even at higher pressures. Vaporization of feed in heat exchanger is avoided to reduce fouling of the surfaces.
Pre-reactor – The vapours from the top of the evaporator at the temperature of 180 deg C are heated in a heat exchanger to the temperature range of 190 deg C to 225 deg C by passing main reactor effluent through shell side. The reactor is provided with a bed of catalyst. The most commonly used catalyst is the nickel-molybdenum catalyst. In this pre-reactor such as di-olefins, styrene and CS2 are removed by hydrogenation. Feed enters from the bottom of the reactors through the catalyst bed.
The temperature at the feed at the inlet of the reactor is ranging from 190 deg C to 225 deg C. The life cycle of the catalyst is sensitive to the temperature. Because of the exothermic reaction, the outlet temperatures are in the temperature range of 200 deg C to 235 deg C. Due to continuous operation of the catalyst bed, coke like polymerization products deposit on the catalyst bed resulting in the lower efficiency. This can be overcome by increasing the inlet temperature of the reactor. Catalyst activity can be determined by the temperature difference between inlet and outlet, which is required to be more than 10 deg C. The catalyst can be regenerated by heating the bed with steam and air. The reactions taking place in the pre-reactor are given below.
Di-olefins (CnH2n-2) + H2 —> Mono-olefins (CnH2n)
Cyclopentadiene (C5H6) + H2 —> Cyclopentane (C5H8)
Styrene (C8H8) + H2 —> Ethyl benzene (C9H10)
CS2 + H2 —> CH4 (methane) + H2S
Main reactor – In the main reactor, treated pre-reactor effluent is hydrogenated on a molybdenum sulphide catalyst. The main reactor consists of two beds of catalysts. The makeup gas i.e. pure H2 from the compressor at a pressure of around 18 kg/sq cm is used for hydrogenation and complete saturation of olefin hydrocarbons. The inlet temperature is around 270 deg C and the outlet temperature is around 330 deg C. The temperature increase is due to the exothermic reaction. Main reactions which are taking place are given below.
Mono olefins + H2 —> Paraffins
Ethyl mercaptans (C2H6S) + H2 —> C2H6 (ethane) +H2S
Thiophene (C4H4S) + H2 —> C4H10 (butane) + H2S
Coumarone (C8H6O) + H2 —> C8H10 (ethylbenzene)
Pyridine (C5H5N) + H2 —> C5H10 (pentene)
Benzene + H2 —> C6H12 (cyclohexane)
Toluene + H2 —> C7H14 (methyl cyclohexane)
In the main reactor, desulphurization, densification and olefin saturation of the feedstock takes place. The H2 gas is fed through a distributor below the first bed of catalyst. The oxygen content in the H2 gas is to be very low so that no polymerization occurs in the reactor. Hydrogenation of aromatics is to be prevented by optimizing the temperature. Deactivation of the catalyst is normally determined by the amount of thiophene content at the outlet of the reactor. If this content increases, then there is hydrogenation of aromatics and coke formation. In this case, either the temperature of the reactor is to be increased or the regeneration of the catalyst is to be done. Hence, there is necessity to have a heater to which a part of the effluent is passed, heated and fed to the main reactor for the maintenance of the temperature. Coke oven gas is generally used as fuel in the heater.
The effluent, which is at 330 deg C, is collected at the bottom. The effluent is passed through several heat exchangers and finally cooled in the water cooler. This condenser effluent is fed to the separator. Before water cooler, hot water is dosed into the effluent. This dissolves the deposits of salts such as NH4HS2 and NH4Cl. The cooled effluent at 50 deg C is fed to the separator. A water leg is to separate the dosed water. The water-free effluent is fed to the stripping column. The gases which consist of the unreacted hydrogen gas and other gases are sucked by recycle gas compressor and are recycled. Part of the gas is purged out through a vent provided.
The purification section consists of a stripping column in which the sulphur content is removed as H2S and also the dissolved gases in the defronted crude benzol removed. The process of stripping is described below.
The liquid part from the separator is fed to the stripping column through a pre-heater, which is heated by BTXS fraction from the stripping column. The temperature of the feed to the column is around 135 deg C. The column consists of sieve trays and has a top temperature which is in the range of 125 deg C to 135 deg C and bottom temperature which is at around 150 deg C. Pressure is around 4.3 kg/sq cm. Re-boiler is provided which supplies the required heat to the column. Medium pressure steam is fed to the shell side of the re-boiler. The gas from the column contains H2S. It is condensed in the condenser with water. This condensate which is at a temperature of around 70 deg C is fed to the reflux drum. Part of the condensate is recycled to the column. Moisture present in the gas is removed from the water leg and the off gasses are fed to the off gas mains. The bottom product called BTXS raffinate is passed through the pre heater where it is cooled and finally raffinate is further cooled in the raffinate cooler which is cooled by water. This is stored in intermediate storage.
Hot oil system – The heat needed during the hydro-refining process is supplied by a separate hot oil system. The hot oil is used as a heating medium for several heat exchangers in the hydro-refining unit and extractive distillation unit. A horizontal furnace is used to heat the oil. Coke oven gas is normally used as a fuel for the furnace. Hot oil is pumped into the coils of the furnace where its temperature increases to around 340 deg C to 350 deg C. The heated oil is pumped to the hydro-refining plant.
Pressure swing adsorption (PSA) unit – H2 gas needed for the hydro-refining plant is separated from coke oven gas in this unit. The clean coke oven gas after benzol recovery is sent to a filter at a pressure of around 800 mm WC (water column). Moisture and carbon particles present in the gas are filtered and the filtered coke oven gas is sent to a reciprocating compressor, which compresses the gas to pressure of around 2.5 kg/sq cm. The compressed gas is further compressed in another compressor to a pressure of 6.5 kg/sq cm. The gas is then further filtered to remove the moisture in the gas. The filtered gas is sent to the pressure swing adsorption unit. Pressure swing adsorption is employed to separate hydrogen gas from the coke oven gas under pressure as determined by the molecular characteristics of the hydrogen gas and its affinity for adsorbent materials.
The PSA process works at basically constant temperature and uses the effect of alternating pressure and partial pressure to perform adsorption and desorption. Since heating or cooling is not required, short cycles within the range of minutes are achieved. Adsorption is carried out at high pressure (and hence high respective partial pressure) typically in the range of 10 kg/sq cm to 40 kg/sq cm until the equilibrium loading is reached. At this point of time, no further adsorption capacity is available and the adsorbent material is to be regenerated. This regeneration is done by lowering the pressure to slightly above atmospheric pressure resulting in a respective decrease in equilibrium loading. As a result, the impurities on the adsorbent material are desorbed and the adsorbent material is regenerated. The amount of impurities removed from a gas stream within one cycle corresponds to the difference of adsorption to desorption loading. After termination of regeneration, pressure is increased back to adsorption pressure level and the process starts again from the beginning.
A PSA plant consists basically of the adsorber vessels containing the adsorbent material, tail gas drum(s), valve skid(s) with interconnecting piping, control valves and instrumentation and a control system for control of the unit. The PSA process has four basic process steps namely (i) adsorption, (ii) depressurization, (iii) regeneration, and (iv) repressurization.
To provide continuous supply of H2 gas, minimum 4 adsorber vessels are required. Adsorption of impurities is carried out at high pressure being determined by the pressure of the feed gas. The feed gas flows through the adsorber vessels in an upward direction. Impurities such as water, heavy hydrocarbons, light hydrocarbons, CO2, CO and nitrogen etc. are selectively adsorbed on the surface of the adsorbent material. Highly pure hydrogen exits the adsorber vessel at the top. After a defined time, the adsorption phase of this vessel stops and regeneration starts. Another adsorber takes over the task of adsorption to ensure continuous supply of H2 gas.
The regeneration phase consists of basically five consecutive steps namely (i) pressure equalization, (ii) provide purge, (iii) dump, (iv) purging, and (v) repressurization. These steps are combined so as to minimize losses 0f H2 gas and consequently to maximize the recovery rate of the H2 gas in the PSA system.
In the pressure equalization step, depressurization starts in the co-current direction from bottom to top. The H2 gas still stored in the void space of the adsorbent material is used to pressurize another adsorber having just terminated its regeneration. Depending on the total number of adsorbers and the process conditions, one to four of these so-called pressure equalization steps are performed. Each additional pressure equalization step minimizes losses of H2 gas and increases the H2 gas recovery rate.
The ‘provide purge step’ is the final depressurization step in cocurrent direction providing pure H2 gas to purge or regenerate another adsorber.
At a certain point of time, the remaining pressure needs to be released in counter-current direction to prevent break-through of impurities at the top of the adsorber. The dump step is the first step of the regeneration phase when desorbed impurities leave the adsorber at the bottom and flow to the tail gas system of the PSA plant.
In the regeneration step, final desorption and regeneration is performed at the lowest pressure of the PSA sequence. Highly pure H2 gas obtained from an adsorber in the provide purge step, is used to purge the desorbed impurities into the tail gas system. The residual loading on the adsorbent material is reduced to a minimum to achieve high efficiency of the PSA cycle.
Repressurization step is explained here. Before restarting adsorption, the regenerated adsorber is to be pressurized again. This is accomplished in the pressure equalization step by using pure H2 gas from adsorbers presently under depressurization. Since final adsorption pressure cannot be reached with pressure equalization steps, repressurization to adsorption pressure is carried out with a split stream from the H2 product line. Having reached the required pressure level again, this regenerated adsorber takes over the task of adsorption from another vessel having just terminated its adsorption phase.
The H2 gas is collected from the top of the bed and is fed to the makeup gas compressor. This is usually a vertical reciprocating compressor of the double stage. The H2gas is compressed to a pressure of around 30 kg/sq cm. The recycle gas from the gas separator is fed to the recycle gas compressor, which is usually a horizontal single stage compressor. The H2 gas thus collected is sent to the makeup gas compressor.
Extractive distillation unit
In this unit, the BTXS raffinate is processed to separate benzene, toluene, xylene, and solvent. Further benzene and toluene are also separated. NFM is used as solvent for the separation of BTX into BT (benzene and toluene) and X (xylene). Non aromatic compounds present in BTXS are removed by pressure distillation as solvent which is recovered in solvent recovery column. Benzene and toluene are separated in BT separation column.
The objective of the extractive unit is to separate non aromatics from benzene and toluene. This unit produces pure benzene, pure toluene, xylene, and light solvent oil along with still bottom oil from BTXS raffinate. Typical flowsheet of extractive distillation unit is at Fig 2.
Fig 2 Typical flowsheet of extractive distillation unit
The extractive distillation unit consists of three main sections, which are pre-distillation, extractive distillation and solvent regeneration. In the pre-distillation section, the feed (consisting of benzene and non-aromatics) is separated into toluene and benzene fractions. The benzene fraction enters the extractive distillation column, and is recovered using the N-formylmoropholine (NFM) solvent having formula O(C2H4)2NCHO. In fact, adding NFM as solvent alters the vapour pressure, facilitating paraffins and naphthalene (i.e., non-aromatics) removal by distillation from aromatics. The vapours at the overhead of the extractive distillation column, containing non-aromatics and a small amount of benzene and solvent, are fed into the solvent recovery column, where solvent and non-aromatics are separated. The bottom product of the extractive distillation column, consisting of solvent, benzene and a small amount of non-aromatics, is conveyed to the stripper column, in which pure benzene is obtained as the overhead product by vacuum distillation (solvent regeneration unit). The stripped hot solvent from the bottom of the stripper column is pumped through several heat exchangers to the top of the extractive distillation column.
The total heat required for the extractive distillation unit is supplied from various means. Pressure distillation receives heat from hot oil, while the aromatic separation column and solvent recovery column receives heat from the vapours of the pressure distillation column. The BT column receives heat from the low pressure steam.
The unit consists of various sections namely (i) pressure distillation section, (ii) extractive distillation column, (iii) solvent recovery section, (iv) aromatic stripper, (v) BT separation section, and (vi) batch distillation section
Pressure distillation section – This section consists of a distillation column in which the BT and XS are separated by simple distillation from the BTXS raffinate. The BTXS raffinate from the purification section is pumped to the feed surge drum. The drum is a horizontal tank provided with a vane and a pipe from reflux drum which carries vapours to this drum. The BTXS raffinate from the surge drum is pumped to pressure distillation column through four heat exchangers in series. The first two are heated by bottom product i.e. XS fraction, while the remaining two are heated by BT fraction. The column consists of the required number of bubble cap trays of which the middle is the feed tray. Column pressure is around 15 kg/sq cm. A re-boiler is provided to the column through which hot oil passes through shell side. This supply the heat required for the column. The BT vapours from the top of the column are condensed and collected in reflux drum. This condensed BT fraction is collected in reflux drum. Some of it is reflux to the column and the remaining is passed to the last two heat exchangers where BT vapours are cooled and further condensed by using water.
Extractive distillation column – In this column the separation of non-aromatics contained in the feed is carried out, which is not possible under normal distillation conditions. The non-aromatics originally with boiling points higher than aromatics, are converted into low boiling non-aromatics which can be withdrawn at the top of the column while the aromatic substances dissolve in the NFM is yielded at the bottom of the column.
The cooled BT fraction is conveyed by steam pressure as feed to the extractive distillation column. The feed is introduced on the tray at the middle of the column. The N-formylmoropholine solvent is introduced on the top tray of the column at the physically required conditions (flow ratio of 56 kg NFM per ton of) at 92 deg C. The NFM temperature is regulated for obtaining low level of aromatics in the non-aromatics.
The NFM which is fed at the top of the column promotes the scrubbing of aromatics out of ascending vapours. The non –aromatic vapours are dissolved only to a slight extent in the NFM. Extractive distillation column has a number of bubble cap trays and the middle has the feed plate trays. The top temperature is 110 deg C, while the bottom temperature 150 is deg C. Top pressure in the column is 0.8 kg/sq cm, and the bottom pressure in the column is around 0.4 kg/sq cm.
Solvent recovery column – The solvent recovery column is used for separation of the non-aromatics present at the top of the extractive distillation column from the residual solvent. A portion of the bottom is fed to the solvent separator in order to recover NFM. The solvent separator splits the liquid into two liquid phases. The non-aromatic-rich phase is returned to the ‘solvent recovery column’, while the NFM rich phase is directed to the ‘extractive distillation column’ (lean solvent circuit). The solvent recovery column in combination with the solvent separator reduces the use of NFM. This section is controlled by pressure, temperature and level controllers.
NFM recovered at the bottom of the column is returned to the extractive distillation column. The solvent recovery column minimizes the NFM losses by means of extensive recalculation of NFM flow inevitably leaving the top of the extraction column.
Stripping column – The product obtained at the bottom of the extractive distillation column consists of NFM in which the extracted aromatic substances are dissolved. The non-aromatic content is low (usually in parts per million) because of there is pressure drop. This flow is conveyed into the aromatics column, which is operated under vacuum. In this column the pure aromatics are separated from the NFM. NFM depleted from the aromatic substances is collected as the bottom product. It is cooled in the heat exchanger system prior to be returned to the extractive distillation column. NFM is sent to the extractive distillation column through (i) centre re-boiler on the extractive distillation, (ii) re-boiler on solvent recovery column, and (iii) NFM re-boiler on the stripper column. Solvent cooler serves as a trim cooler for NFM. The bottom of the stripper column is heated by means of the two continuous re-boilers, which are heated by BT vapours and heated by means of NFM. The reflux to the stripping column serves to remove the solvent in the lower section of the column. The vapour liquid mixture discharges from the re boiler is fed below the chimney tray in the stripping column. Traces of the solvent flash are washed back by the aromatics reflux and directed into the lower part of the column. There are a number of trays and the feed tray is the fifth tray. The temperature at the top is 56 deg C while the temperature at the bottom is 119 deg C. Pressure is around 0.36 kg /sq cm.
Benzene toluene separation column – BT separation is a normal two-phase distillation process for the production of the pure aromatics. The BT fraction is routed using a reflux pump from reflux drum via exchanger to separation column. The heat required for distillation is supplied to the system via re-boilers by means of low pressure steam. Pure benzene is taken out at the top while the pure toluene is taken out from the bottom. There are a number of bubble cap trays with the middle tray as the feed tray. The pressure is around 1.2 kg/sq cm. Benzene with purity of 99.97 % and toluene with the purity of 99.95 % is produced.
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