Monday, 28 March 2011


Raw Material

Full range naphtha


Specific gravity…………………….0.713(60/60oF)
API gravity………………………….67o
Boiling Range………………………32oC---171oC
Watson Characterization factor (K)…..12.3

Composition (wt %)

(By PONA analysis)

Average Molecular weight…105Kg/Kgmol

Cracked Gas Composition

CH4……………………………. 15%
C2H2…………………………… 0.8%
FUEL OIL………………………..6%

Cracking Conditions

Steam dilution………………………..0.5Kg/Kg of naphtha

Manufacture By Thermal Cracking

Thermal Cracking of Hydrocarbons

Thermal cracking of hydrocarbons is the principle route for the industrial production of olefins. In thermal cracking valuable by products including propylene butadiene and benzene are also produced. Commercially less valuable methane and fuel oil are also produced in significant proportions. An important parameter in Design of commercial thermal cracking furnaces is selectivity to produce the desired products.
The large-scale manufacture of olefins from hydrocarbon raw materials is carried out in plants which are among the most complex and expansion facilities of the hydrocarbon processing industry .An olefins manufacturing facility often forms the core of an entire petrochemical complex designed to produce a variety of petrochemical products.  
All olefins manufacturing process incorporate a great number of unit operations; catalytic and non catalytic reactions, absorption and adsorption, fractionation, compression, heat exchange, and phase separation are arranged in a complex array to achieve the production of high-purity ethylene and by-products in the most efficient manner.
Certain equipment, notably the pyrolysis furnaces and hydrogenation reactors, have to be regenerated periodically to remove carbon or polymer deposit hazards.
The typical olefins plant contains well over 400 pieces of major equipment, i.e., fired heaters and heat exchangers, reactors, drums, fractionating towers, compressors, and pumps. 

Pyrolysis and Primary Waste Heat Recovery

The pyrolysis section is considered the heart of an olefins plant, in this section all the products of the plant are produced, while the other sections serve to separate and purity the products. Through the pyrolysis yields pattern, feedstock consumption, and efficiency of fuel utilization, the pyrolysis section has the great influence on the economics of the plant.
Each furnace contains an even number of multiple pyrolysis coils in the radiant section. The exact number depends on the coli design concept of the furnace designer and the desired capacity. Due to the high fire box temperature, more than 50% of the fired duty has to be absorbed in the convection section of the pyrolysis furnaces. The vaporization and the preheating of the feedstock require only a portion of the convection section duty. The remainder is utilizes to preheat boiler feed water and to superheat high-pressure steam. The final steam superheat temperature is controlled through injection of boiler feed water, as is standard practice for high-pressure boilers.
The feed stocks are usually first preheated externally to the pyrolysis furnaces and then in the uppermost coil of the convection section. Liquid feed stocks are partially vaporized and then mixed with dilution steam in the weight ratio of between 0.5 to 1.0 steam to hydrocarbons. The dilution steam has the dual function of lowering the hydrocarbon partial pressure and reducing the carburization rate of the pyrolysis coils. The steam-hydrocarbon mixture is further heated in the convection section to the crossover temperature between the convection and the radiant sections. This temperature has to be carefully selected for any given feed stocks to obtain maximum heat absorption without reaching temperature high enough to initiate cracking which could result in detrimental coke deposit in the convection section. Depending on the feed stocks, this temperature is usually in the range of 550 to 700°C .Some furnace designers provide a bank of steam-cooked shock tubes for the lowest part of the convection to protect the process convection coils.
The steam-hydrocarbon mixture then enters the pyrolysis coil where it is further heated and cracked into products. For the pyrolysis reactions, hydrocarbon partial pressure and temperature as function of coil length and time are of particular importance. The total residence time the gas spends in the pyrolysis coil of a modern furnace is generally in the range of 0.2 to 0.6 s. The gas outlet temperature is in the range of 770 to 870°C depending upon feedstock and furnace design. The product gas is sent to closely connected transfer line exchangers in whom the temperature is rapidly reduced by several hundred degrees by generating high-pressure saturated steam at approximately 12.5MPa. Given the extreme process gas temperature and the high steam pressure, the transfer line exchangers have to be of very special design.
Two or more transfer line exchanger together with a common steam drum is usually built right into the furnace structure. One exchanger serves one or more often, two pyrolysis coils. Boilers feed water circulation from the steam drum through the transfer line exchanger is by thermosyphon action with typical water to steam ratio of 10 or higher.
Pyrolysis furnaces and their transfer line exchangers form a pyrolysis module. The close integration is further enhanced by preheating of the boiler feed water for and the superheating of the saturated steam from the transfer line exchangers in the convection of the pyrolysis furnace
The principal reason for generating steam of such high pressure is to maintain the tube wall temperature as high as possible to minimize the condensation of tars.

Secondary Waste Heat Recovery and Pyrolysis Fuel Separation

The pyrolysis gas leaving the transfer line exchangers is at temperature that range from approximately 375 to 500°C in the case of naphtha pyrolysis. The outlet temperature depends on the amount of carbon deposit in the transfer line exchangers. Most processes also recover lower level heat through direct injection of quench oil into the pyrolysis gas either directly downstream of the individual transfer line exchanges or into the transfer line immediately upstream of the primary fractionator. The first location is preferred for gas oil-based plants; the second location for naphtha based plants.
The transfer line can be made of normal carbon steel .it has the disadvantages that the transfer line must be designed for a very large mixed phase flow of pyrolysis gas and quench oil. Pyrolysis gas and quench oil are separated in the based of the primary fractionator. The quench oil is cooled to approximately 185°C in the dilution steam generation or lower pressure steam generators (0.4 to 0.5 MPa) and in heat exchange with other process streams. The maximum temperature of the quench oil in the bottom of the primary fractionator has to be carefully controlled because of the instability of this material. Certain components of which tend to polymerize with an attendant increase of the viscosity. The quench oil of gas oil-based plants exhibits higher thermal stability than does the material in naphtha-based plants. In the latter the material is entirely synthetic since the feedstock does not contain any components boiling in the pyrolysis fuel oil range. As the fuel oil yield from naphtha is significantly lower than from gas oil, the residence time of the quench oil in the system is proportionately longer, resulting in a greater accumulation of polymerization products. Quench oil often contains coke particles which can cause flow problems in the quench oil circuit.
The function of the primary fractionator is to make a sharp separation between pyrolysis gasoline and pyrolysis fuel oil. This tower must assure the distillation end point of the pyrogasoline. Reflux is provided from gasoline condensed in the quench tower which, in effect, represents the reflux condenser of the primary fractionator. The overhead temperature is generally set to avoid the condensation of steam which, due to the presence of acidic components could cause corrosion in the system. Since the bottom temperature is selected for optimum heat recovery in the quench oil circuit, a side stream withdrawal is provided as an exit for fuel oil components that are too volatile to be retained in the bottom and too heavy for the pyrogasoline. The net fuel oil products is withdrawn from the bottom of the primary fractionator and sent to the fuel oil stripper.

Tertiary Waste Heat Recovery and Heavy Gasoline Separation   

The gas leaving the primary fractionator has a temperature of between 100 and 110°C and is further cooled in air-or waste-cooled heat exchangers or in a direct contact cooled quench tower. The latter has the inherent advantage of low-pressure drop and recovers the residual heat of the pyrolysis gas through absorption in hot quench water. This hot quench water is usually available at a temperature of about 80°C for various low-level heating services throughout the plant. Since most of the latent heat of the dilution steam is released in the lower section of the quench tower, the bulk of moderately cooled quench water is added to that section, while a smaller stream is cooled as low as possible and fed to the top section of the lower the temperature of the pyrolysis gas as much as possible before sending it to the main compressor.
In the base of the tower, hot quench water is separated from heavy pyrogasoline which is condensed along the dilution steam. The bulk of the hydrocarbons are pumped to the primary fractionator to serve as reflux, while the net product is sent to the heavy pyrogasoline stripper for the removal of the C4 and lighter hydrocarbons. The hot quench water is sent to various heat exchangers and, in particular, the Reboiler of the propylene fractionator if the plant produces high-purity propylene. Residual heat of the quench water is rejected to the atmosphere in a combination of air and water coolers.
The net dilution steam condensed is withdrawn from the quench water and, after a second more effective water-hydrocarbon phase separation, is sent to the process water stripper for the removal of dissolved hydrocarbons and gases. The treated process water is then pumped to the dilution steam generators which are heated with quench oil and medium-pressure steam (1.2 to 1.5MPa).The net process water discharge from the system is minimal since only the small quantity of live stripper steam fed to the fuel oil strippers represents a net inflow.          The quench water, due to its absorption of acidic compounds, is potentially corrosive. Various neutralization agents and other types of corrosion inhibitors are added to the quench water to protect the equipment. Careful control of pH value is important as high pH values have been found to lead to foaming and the formation of emulsion which are difficult to break.

Pyrolysis Gas Compression, Acid Gas Removal and Drying

The pyrolysis gas leaving the quench tower usually has a temperature of 35 to 40°C and a pressure slightly above atmospheric pressure. Most processes call for compression to a pressure of approximately 3.5 MPa which appears optimal for the subsequent cryogenic treatment. The pyrolysis gas, in particular, if derived from liquid feed stocks, contains appreciable quantities of highly reactive diolefins and acidic components. The diolefins tend to polymerize at compression system. In fact, polymer formation is of concern throughout the entire process wherever elevated temperature and moderate to high concentration of diolefins exist. The acidic components raise the specter of stress corrosion.
Modern ethylene plants employ radial centrifugal compression exclusively; the future may possible see the introduction of more efficient axial compressors for the early stages of pyrolysis gas compression in very large capacity plants. Most operators and designers of ethylene plants limit the maximum discharge temperature from the individual stages of the pyrolysis gas compressor to minimize the danger to polymer deposits. The temperature limits largely determine the number of composition stages. Plant based on gaseous feed stocks generally employ four stages while many naphtha-and gas oil-based plants employ five stages of pyrolysis gas compression. However, even a five-stage compression system will not yield discharge temperatures as low as 85°C with typical cooling water temperature available at most plant locations. If lower compressor discharge temperatures are desired, cooling with water followed by high-level propylene refrigeration is required which then permits the use of a four-stage pyrolysis gas compressor. Although the five-stage pyrolysis gas compression system is more common, four-stage systems are in operation with and without propylene intercooling.
Water and hydrocarbons condensed between stages are separated from the pyrolysis gas in interstage separators. The water is returned to the quench water system while the hydrocarbons formed in the first three stages is sent to the heavy pyrogasoline stripper while the condensate from the last one or two stages is sent to the condensate deethanizer.
Hydrogen sulfide and carbon dioxide are removed from the pyrolysis gas between the third and fourth stages of the compression system. This location is optimum as the actual gas volume has been significantly in the first three stages of compression while the acidic components are still present in the gas stream and have not contaminated and products that are separated ahead
Plants processing feed stocks with sulfur contents of a few hundred ppm, scrubbing with a dilute caustic soda solution have proven most economical. A two- or sometimes three-stage scrubbing tower is generally provided to obtain maximum utilization of the caustic soda. The solution circulation over the top section of the scrubbing tower typically contains between 8 and 12% free sodium hydroxide while the solution in the tower section will contain between 2 and 5% free caustic.
Relatively weak caustic solution is preferred to avoid the precipitation of sodium salts and to minimize the formation of organic sodium complexes. The pyrolysis gas leaving the caustic scrubber contains less than 1 ppm of acid gases and thus assures that the final products of the plant will meet specifications in this respect. The spent caustic represents environmentally the most troublesome liquid effluent of ethylene plants and requires substantial subsequent treatment before it can be safely discharged.
Depending on the environmental requirements at specific plant locations, the acid gases leaving the regenerative acid gas removal and spent caustic neutralization systems either burned and the combustion products sent to flue gas, or, as is becoming more common, are treated in a Claus unit for conversion of the hydrogen sulfide to elemental sulfur. Complete removal of water from the pyrolysis gas is required in preparation for the cryogenic treatment.
Most processes, however, employ a single adsorptive drying system located immediately after the final stage of pyrolysis gas compression. Molecular sieves are the preferred desiccants because of their selectivity. Of two dryers normally provided, one is in normal operation while the other is regenerated, recooled, or held in a standby position. High-pressure methane heated with steam or in a direct-fired heater to a temperature of approximately 225°C is the preferred regeneration medium. The dryers are designed for on-stream times of 24 to 48 h between successive regenerations.

Cryogenic Treatment and Demethanizer

The cryogenic treatment involves the successive chilling of the pyrolysis gas in heat exchange with boiling refrigerant and with cold process streams that have to be vaporized and/ or reheated. As the pyrolysis gas temperature is reduced, more and more condensate is formed
The entire pyrolysis gas is, after some cooling and condensation, sent to a high-pressure demethanizer which accomplished the essentially complete removal of methane from the condensate but loses some ethylene in the hydrogen –methane overhead products. Ethylene refrigerant boiling in the reflux condenser at slightly above atmospheric pressure controls the dew point temperature of the overhead products. The ethylene contained in the overhead product is usually recovered either through condensation and recycling to the main compressor or through condensation with refrigeration generated by a turbo expander. Most process, even those designed for the separation of C4 and heavier hydrocarbons in the compression system, contain an arrangement in which only the condensates are sent to the demethanizer. Operating at a pressure of approximately 302 MPa, the tower is designed for essentially complete separation of methane from ethylene and heavier components. Condensing propylene refrigerant supplies the necessary heat to the Reboiler, and vaporizing ethylene condenses reflux and some of the over head products.
Most modern ethylene plants also produce a high-purity hydrogen stream (95mole % being typical) and a concentrated methane-rich off gas stream (95 mol% methane with the remainder mostly hydrogen
In those processes in which the entire pyrolysis gas stream is sent to the demethanizer, the hydrogen stream is withdrawn downstream of the demethanizer.
The cryogenic treatment systems of all ethylene processes contains inherent thermodynamic irreversibilities which as the isenthalpic throttling of the demethanizer overhead and fractionators with large temperature differences between top and bottom, more than 100°C in the case of the high-pressure demethanizer. Side vaporizers for fractionators can significantly reduce the irreversibilities of fractionation systems and are, in fact, applied in ethylene plants to save energy. Interest in a three-stage propylene-ethylene-methane refrigeration cascade, which was used in early small-scale ethylene plants [84], has recently revived. Provision of methane refrigeration allows the design of an efficient low-pressure demethanization system but requires the reintroduction of reciprocating compressors.
The practice of using multicomponent refrigerants for natural gas liquefaction plants has led to an investigation of the suitability of this concept for the cryogenic treatment of pyrolysis gas. The condensates formed can be used as multicomponent refrigerants with the main compressor also serving as the refrigerant compressor.
Brazed aluminum plate-fin exchangers are generally used for the multipass cryogenic heat transfer services and are sometimes even used for the refrigerated chillers with thermosyphon circuits for refrigerant circulation. Much of the cryogenic equipment is also installed in a rectangular carbon steel container, commonly called cold box, with its void spaces filled with perlite or rockwool for insulation. Considering the complexity of the system, this mechanical design is characterized by surprising compactness and minimum space requirements.

Deethanizer and Acetylene Hydrogenation

The deethanizer is a conventional tray-type fractionator operating at a pressure of 204 to 208 MPa to separate the demethanizer bottom products from an overhead stream containing all three C2-hydrocarbons and a bottom product of C3 and heavier hydrocarbons. Essentially complete removal of C2-hydrocarbons from the bottom products is more important than the purity of the overhead product. Any C2-hydrocarbons contained in the bottom product will eventually reach the propylene product which often has a 10-ppm specification on the ethylene content. Minor amounts of C2-hydrocarbons in the overhead will be concentrated in the ethane products from the ethylene fractionator which is either recycled to pyrolysis or sent to the fuel with neither utilization being steam heats the rebolier and propylene refrigerant condenses the reflux.
Treatment of the C2-stream can take various routes, depending on the acetylene content and on whether or not high-purity acetylene is to be recovered. The system for moderate acetylene contents when the recovery of acetylene is not of interest. The deethanizer gross overhead is heated to a temperature of between 20 and 100º C, hydrogen is added in the molar ratio of approximately 2 referred to acetylene, and the mixture is passed over a fixed bed of palladium catalyst. The heat of reaction is removed in a water cooler and the reactor effluent is further cooled by heat exchange with the feed and partially condensed by heat exchanger with propylene refrigerant. The liquid is pumped as reflux to the deethanizer while the vapor flows through the guard dryer to the ethylene-ethane fractionator. The reactor effluent typically contains less than 1 ppm of acetylene but is contaminated with traces of hydrogen and methane which represent the major disadvantage of this location of the acetylene hydrogenation system. Yet this location has the advantage over the front end hydrogenation in that it allows very accurate control of the hydrogen concentration and reaction temperature, resulting in good selectivity of the reaction. Under normal conditions as much acetylene is hydrogenated to ethylene to ethane with no net loss of ethylene. It has been found that traces of carbon monoxide in creases the selectivity of the catalyst.
Some undesirable polymerization side reactions occur in the hydrogenation reactors, resulting in the formation of olefins with carbon numbers ranging from four to over twenty. The side reactions in the hydrogenation reactors cause a slow deactivation of the catalyst which calls for gradually rising inlet temperature to maintain adequate activity.
Ultimately, after several months of operation, regeneration of the catalyst is required. Controlled combustion of the deposits with a steam-air mixture at about 750º C completely restores the activity of the catalyst. Acetylene recovery facilities can be located between the deethanizer and the ethylene-ethane fractionator or down stream of the latter.
Location of acetylene recovery facilities down stream of the main ethylene –ethane fractionators eliminates this problem but involves higher concentrations of acetylene in this tower, resulting in partial pressure that, however, remain well below established safety limits.

Ethylene-Ethane Fractionation

The ethylene-ethane mixture leaving the hydrogenation system may contain traces of hydrogen and methane. A “pasteurizing “section is occasionally provided in the top of the fractionator for the removal of hydrogen, carbon monoxide, and methane. This process element permits these contaminants to be reduced by factors approximately equal to the equilibrium constants of the contaminants under conditions as they exist at the point of ethylene withdrawal. A small stream containing most of the contaminants but still rich in ethylene is reheated and recycled to the main compressor.

Purification of Propylene and Heavier Products

The bottoms products of the condensate stripper and deethanizer are processed in the depropanizer for a sharp separation of the C3-hydrocarbons from C4 and heavier hydrocarbons. The bottoms product of this tower has a high content of butadiene and heavier diolefins which tend to polymerize at moderate temperatures. Deposits of a rubber like polymer on the inner surfaces of the Reboiler tubes and the internals of the tower can severely curtail the capacity of this fractionation system. Mechanical cleaning may thus be required every few months.
The depropanizer bottoms are further processed in the debutanizer for separation of C4- product from the light pyrogasoline. The debutanizer, operating at a moderate pressure of 0.4 to 0.5 MPa, is a conventional fractionator with a steam-heated Reboiler and a water-cooled condenser. A spare Reboiler is often provided although deposition of polymers is much less of a problem in this tower than in the depropanizer.
The overhead product of the depropanizer can be sent directly to the propylene-propane fractionator. However, it will require hydrogenation of the contained methyl acetylene and propandiene if the bottoms product of the fractionator is to meet LPG specifications or if only so-called chemical-grade propylene is to be produced. The latter requires only stripping for removal of hydrogen and methane, and no fractionation of the hydro treated depropanizer overhead.
Treatment of the depropanizer overhead for hydrogenation of the methyl-acetylene and propadiene over a palladium catalyst is typically carried out in a vapor-phase system quite similar to the acetylene hydrogenation. The hydrogenation can also be accomplished in proprietary mixed-phase isothermal systems.

Due to the low relative volatility of propylene to propane, the fractionation of propylene and propane is even more difficult than the fractionation of ethylene and ethane. As the depropanizer overhead product normally contains well in excess of 90% propylene, most propylene-propane fractionation systems are designed for a propylene content of 10 to 40% in the propane product to keep the number of fractionation trays in a practical range. Design for lower propylene concentration of the propane product will add only little to the recovery of high-purity propylene.

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