iocl section 1 - process specfication

INDIAN OIL CORPORATION LIMITED

IOCL Section 1 – Process Specification.doc February, 2005


NAPHTHA CRACKER PROJECT

800,000 TPA ETHYLENE PLANT

PANIPAT, INDIA

BASIC ENGINEERING PACKAGE

NAPHTHA CRACKER UNIT SECTION 1 PROCESS SPECIFICATIONS


IOCL Section 1 – Process Specification.doc Page 1 February, 2005



800,000 TPA ETHYLENE PLANT PANIPAT, INDIA

PROCESS SPECIFICATION

for

NAPHTHA CRACKER UNIT

LUMMUS PROJECT 12605 REVISION 1 ISSUE DATE FEBRUARY 28, 2005 BY RPT





800,000 TPA ETHYLENE PLANT

PANIPAT, INDIA

NAPHTHA CRACKER UNIT

PROCESS SPECIFICATION

INDEX

ARTICLE I PLANT ELEMENT

1.0 WHAT CONSTITUTES THE PLANT 2.0 GENERAL DESCRIPTION OF THE NAPHTHA CRACKER UNIT

2.1 ISBL Facilities 2.2 ISBL Supporting Facilities 2.3 Capacity and Mode of Operation 2.4 Process Description

ARTICLE II PLANT SPECIFICATIONS

1.0 BASIS FOR PROCESS DESIGN 1.1 Plant Function and Capacity 1.2 Feedstock Properties 1.3 Product Specifications 1.4 Utility Specification 1.5 Site Location and Climatic Conditions 1.6 Battery Limit Conditions 1.7 Environmental Specifications

2.0 BASIS FOR PROCESS PERFORMANCE 2.1 General 2.2 Overall Material Balance 2.3 Once-Through Cracking Yields 2.4 Utility Consumption 2.5 Specific Energy Consumption 2.6 Catalyst and Chemical Consumption 2.7 Pyrolysis Heater Performance 2.8 Assumed Efficiencies 2.9 Heat Leak Factors



ARTICLE I – PLANT ELEMENTS

1.0 WHAT CONSTITUTES THE PLANT The naphtha cracker project for Indian Oil Corporation Ltd. located in Panipat, India consists of ISBL Process units and supporting facility designed to produce polymer grade ethylene, polymer grade propylene, 1,3 Butadiene, benzene, C7-C8 cut, partially hydrogenated C9+ stream and pyrolysis fuel oil. The plant consists of the following process units:

• Naphtha Cracker Unit (NCU) NCU uses naphtha as feedstock to produce polymer grade ethylene, polymer grade propylene, raw C4 mix, raw pyrolysis gasoline and pyrolysis fuel oil. Ethane, propane, hydrogenated C4s, C5s and C6 are recycled to extinction. Recycles from polymer plants, ethylene BOG from atmospheric storage and the FCC streams are the additional feeds to the NCU recovery unit.

• Butadiene Extraction Unit (BEU) BDEU recovers 1,3 Butadiene from the raw mixed C4s stream produced in the NCU. BD raffinate is sent to the C4 hydrogenation unit.

• C4 Hydrogenation Unit C4 hydro is designed to fully hydrogenate the BD raffinate in normal operation. Hydrogenated C4s are recycled to the NCU cracking heaters. In alternate operation, when BEU is not operating, C4 hydrogenation unit will process raw C4 mix. Butadiene and a portion of the butenes are hydrogenated in this mode of operation.

• Pyrolysis Gasoline Hydrogenation Unit (PGHU) PGHU processes raw pyrolysis gasoline from naphtha cracker in a two-stage hydrotreating unit to produce a C6-C8 heart cut that is sent to the Benzene Extraction Unit for benzene recovery, a fully hydrogenated C5 cut that is recycled to the cracking heaters and a partially hydrogenated C9+ product.

• Benzene Extraction Unit (BZEU) BEU recovers benzene from the C6-C8 heart cut. Non-aromatic C6 stream is recycled to the NCU cracking heaters and C7-C8 cut is sent as product to storage.

Block diagram of the process units is presented in Figure 1.1.



This Basic Engineering Package (BEP) for Naphtha Cracker Unit includes total C4 hydrogenation and pyrolysis gasoline hydrogenation units. BEP for Butadiene Extraction Unit and Benzene Extraction Unit are provided as separate documents.



2.0 GENERAL DESCRIPTION OF THE NAPHTHA CRACKER UNIT The naphtha cracker is designed to produce polymer grade ethylene and polymer grade propylene by thermal cracking of naphtha and hydrogenated C4/C5/C6 recycle streams from C4 hydrogenation, pyrolysis gasoline hydrogenation and benzene extraction units respectively. Ethane and propane are recycled to extinction. . Recycles from LLDPE (swing plant), HDPE and PP plants, ethylene boil-off gas (BOG), and FCC feeds are the additional feeds to the naphtha cracker recovery unit. The naphtha cracker unit is designed to produce the following by-products:

• Hydrogen • Methane offgas • Raw mixed C4 • C4 LPG (provision to draw part of hydrogenated C4 to LPG pool) • Hydrogenated C6-C8 • Partially hydrogenated C9+ product • Pyrolysis fuel oil (combined PGO + PFO)

2.1 ISBL Facilities The ISBL facilities of the naphtha cracker unit consists of the following processing areas:

• Cracking Heaters Feed System • Pyrolysis Module • Charge Gas Oil Quenching • Gasoline Fractionation, Pyrolysis Gas Oil and Pyrolysis Fuel Oil

Stripping • Charge Gas Water Quenching • Process Water Stripping and Dilution Steam Generation • Charge Gas Compression, Gasoline Stripping, Condensate Stripping

and Condensate Stripper Bottoms Drying • Acid Gas Removal via Caustic Wash • Spent Caustic Pretreatment via Gasoline Wash • Charge Gas Drying and Dryer Regeneration Facilities • Cracked Gas Chilling • Demethanization and Methane Refrigeration • Methanation and Hydrogen Purification • Deethanization, Acetylene Hydrogenation, C2 Green Oil Removal and

Ethylene Drying • Ethylene Fractionation



• Ethylene Product System • Depropanization and Debutanization • MAPD Hydrogenation and Propylene Fractionation • FCC C3s Treatment • C4 Hydrogenation • Pyrolysis Gasoline Hydrogenation (1st and 2nd stage) • Propylene and Ethylene Refrigeration

2.2 ISBL Supporting Facilities

The supporting facilities include the following systems inside the battery limit: • Fuel gas system including LPG vaporizing facility and the fuel gas mix

drum, distribution and controls • Chemical storage facilities, considered necessary for smooth operation

of the plant. Specifically these shall be limited to the day tanks of caustic, DMDS, wash oil and methanol. Chemical injection facilities are included

• Spent caustic pre-treatment • Boiler feed water (1) • Steam generation from the TLEs (1) • Condensate collection, treatment, polishing and deaeration (1) • Flare load summary, including ISBL knock-out drums, vaporizer and

liquid pumping (1) • Plant air and instrument air (2) • Nitrogen (2) • Cooling water (2) • Electrical (2) Notes: (1) Process design of these facilities is included in FEED package. (2) The BEP does not include process design of these supporting

facilities, but does include user summaries defining the requirements for supporting facilities.

2.3 Capacity and Mode of Operation Base Capacity

The plant shall be designed to produce 800,000 TPA of polymer grade ethylene, 504,834 TPA of polymer grade propylene in 8,000 operating hours a year when cracking design naphtha and ethane, propane and hydrogenated C4/C5/C6 streams. The production rates are net without processing polymer plants recycles, ethylene BOG and the FCC feeds.



Case 1 This case includes recycles from LLDPE and PP planets and ethylene BOG from atmospheric storage. The plant shall produce 830,632 TPA of polymer grade ethylene, 584,352 TPA of polymer grade propylene in 8,000 operating hours a year when cracking design naphtha at flow rate (293,201 kg/h) corresponding to base ethylene and propylene capacity and processing LLDPE (swing plant) and PP recycles and ethylene BOG vapors. FCC feeds are not processed in this case. Ethane, propane and hydrogenated C4/C5/C6 are recycled to cracking heaters as in the base case. Case 2 This case includes two FCC streams as additional feed namely, 1) Treated FCC dry gas C2s and 2) Raw FCC C3s. The plant shall produce 857,408 TPA of polymer grade ethylene, 666,544 TPA of polymer grade propylene in 8,000 operating hours a year when cracking design naphtha at flow rate (293,201 kg/h) corresponding to base ethylene and propylene capacity and processing recycle streams from LLDPE and PP plants, ethylene BOG and FCC feeds. Ethane, propane and hydrogenated C4/C5/C6 are recycled to cracking heaters.

The capacity and mode of operation is summarized below: Design Case Feeds Ethylene

Capacity, TPA

Propylene Capacity,

TPA

Remarks

Base Case • Naphtha 800,000 504,834 -- • Hydrogenated

C4/C5/C6 recycle

• Case 1 • Naphtha 830,632 584,352 Detailed Material

Balance Case • Hydrogenated

C4/C5/C6

• LLDPE recycle

• PP recycles • Ethylene

BOG



Design Case Feeds Ethylene Capacity,

TPA

Propylene Capacity,

TPA

Remarks

Case 2 • Naphtha 857,408 666,544 Detailed Material

Balance Case • Hydrogenated

C4/C5/C6 recycle

• LLDPE recycle

• PP plant recycles

• Ethylene BOG

• FCC feeds 2.3.1 Turndown The plant shall be designed for turndown to 70% of the design capacity. Desired 50% minimum turndown will be achieved by over refluxing/reboiling the towers and recycling. 2.3.2 Future Capacity The plant shall be expandable to 1,000,000 TPA, preinvested for those equipment which are necessary and those that cannot be modified in 30 days mechanical shutdown of the plant. Only critical piping will be designed for 1,000,000 TPA, all other piping will be designed for 800,000 TPA. 2.4 Process Description This section describes the process flow scheme selected for this project and as shown on the Process Flow Diagrams. Cracking Heaters Feed System Naphtha feed is received from storage tanks outside battery limits at pressure and ambient temperature. The naphtha is filtered and mixed with hydrogenated C4, C5 and C6 recycle streams and then preheated against quench water to 60°C prior to being sent to the SRT VI liquid cracking heaters. Sulfur injection (DMDS) is provided into the naphtha/C4/C5/C6 feed upstream of the preheater. DMDS injection in liquid feed is done only when sulfur content in naphtha feed is low.



Recycle ethane, after vaporization against charge gas and reheat in the cold box is mixed with vaporized propane recycle. The combined ethane/propane stream is superheated with quench water to 60°C before being sent to the recycle cracking heater. DMDS injection into the ethane/propane recycle is provided downstream of the superheater. Raw C4 mix or BD raffinate is co-cracked with naphtha/ C4/C5/C6 during regeneration of C4 hydro reactor catalyst Pyrolysis Module Six SRT VI and one SRT III type cracking heaters are provided. Five SRT VI heaters are normally in service cracking liquid feedstock (naphtha/C4/C5/C6). One SRT VI heater is available as a spare to provide operating continuity during decoke operations.

The SRT III heater is in operation cracking ethane/propane recycle. When recycle heater is being decoked, ethane/propane recycle is cracked in one of the two liquid cracking heaters (11-H-0200/11-H-0300).

All the SRT VI and SRT III heaters operate using fuel gas only. The primary fuel gas is methane rich offgas produced in the naphtha cracker unit. Make-up/ back-up fuel is C3/C4 LPG supplied from adjacent refinery or RLNG vapor. As the primary fuel gas is not sufficient to fully meet the heater firing demand, the small shortfall is met by make-up LPG fuel. Combustion air for all the heaters is 100% ambient air. Each heater is equipped with an ID fan for directing the flue gas to the heater stack. All heaters are designed for steam/air decoking of the radiant coil with the decoking effluent being sent directly to the heater firebox for combustion.

For the SRT VI liquid heaters, the liquid feedstock is flow controlled to six sets of radiant coils per liquid cracking heater. The liquid feed is first heated in the feed preheat section of the heater, with the liquid feed being partially vaporized prior to steam injection. The dilution steam injection is added to each coil on flow ratio control. Dilution steam is first superheated in the dilution steam superheat section of the heater which is located below the lower feed preheat section prior to injection into the partially vaporized feed. The total mix is further heated in the heater’s mixed preheat section before entering the radiant coils.

For ethane/propane recycle processing when 11-H-0100 is out of service, the preheated ethane/propane vapor feed is flow controlled to the six sets of radiant coils using an alternate feed control valve set and the dilution steam is added to each coil on flow ratio control. The total feed enters at the inlet of the upper mixed feed preheat section.

1



The crossover pipe from each convection section coil feeds a single radiant coil manifold at the top of the radiant section. Each radiant coil consists of a 20-4 inlet-outlet pass configuration (i.e., five 4-1s). The flow from one radiant coil inlet manifold is distributed to 20 inlet tubes of the first pass using critical flow venturis. The flow is combined to the 4 tubes of the outlet pass. The 4 outlet radiant tubes enter the primary TLE inlet head (bathtub type). There are six TLEs per heater. The effluent from the liquid cracking heaters primary TLEs are combined in two transfer lines.

For the SRT III gas cracking heater, ethane/propane feed is first heated in the feed preheat section and then mixed with dilution steam on flow ratio control prior to entering the secondary TLE (shell-side) where the preheated mix feed is further preheated against heater effluent. The mixture is then split on flow control to the six preheat coils, where it is further heated in the convection section before entering the radiant section. The crossover pipe for each convection coil feeds a single radiant inlet manifold at the inlet to the radiant section.

The recycle heater radiant coil is a six-pass design employing 4-2-1-1-1-1 inlet-outlet pass configuration. Four (4) parallel tubes, which split from the radiant coil inlet manifold comprise the first pass. These tubes join in pairs via a wye piece to two (2) parallel tubes, which are the second pass of the coil. The two second-pass tubes are joined via another wye piece to the third pass. A single tube is used for the third, fourth, fifth and sixth passes.

The effluent from the ethane recycle heater radiant coils is combined in pairs and sent to primary TLEs, three per cracking heater. The effluent from the primary TLEs are further combined and sent to a single secondary TLE which cools the effluent to 255°C. The heat exchanged in the secondary TLE is used to preheat the mixed hydrocarbon feed after partial heating of the hydrocarbon in the convection section and after injection of dilution steam.

For the SRT III gas cracking heaters and the SRT VI liquid cracking heaters, the heater firing is controlled to maintain the average COT of all the coils and the flow to each coil is adjusted to maintain a constant delta between the coil COT and the average COT.

The primary TLEs generate steam at 128.5 kg/cm²g via a thermosyphon system connected to a common steam drum for each heater. The boiler feed water to the steam drum is preheated in the convection section of the cracking heater. The steam generated is superheated to about 525°C in a superheater coil integrated with the convection section of the heater. To control the superheater outlet temperature, the coil design allows for phosphate free boiler feed water injection to the partially superheated steam. After the BFW injection, the steam is returned to the convection section for final superheating to the desired temperature.



The effluent from the liquid feed heaters is oil quenched in two quench fittings. The effluent from the recycle heater is not oil quenched, but is cooled via the secondary TLE. After the secondary TLE, the recycle heater is split with a portion sent to the quench tower and the remainder going to the midsection of the pyrolysis feed oil stripper where it is utilized as stripping vapor to maintain and control the viscosity of the circulating quench oil.

Gasoline Fractionation and Charge Gas Quenching In the Gasoline Fractionator, the cracked effluent gases are further cooled, pyrolysis fuel oil (PFO) is separated as a bottoms product, a side stream product-pyrolysis gas oil (PGO) is withdrawn from the Fractionator, gasoline and lighter materials are taken as an overhead vapor. Heat removed by circulating quench oil from the tower bottom is recovered via dilution steam generation and process water preheat going to the dilution steam drum. The tower is refluxed with gasoline condensed in the Quench Tower.

The bottom pyrolysis fuel oil from the Gasoline Fractionator is sent to the Pyrolysis Fuel Oil Stripper. This tower has two stripping sections, the top section where most of the components boiling between 280°C and 370°C are stripped out using part of the ethane recycle heater effluent and the bottom section where PFO from the top section of the Pyrolysis Fuel Oil Stripper is steam stabilized to achieve a fuel oil product of acceptable flash point. Pyrolysis fuel oil product is mixed with heavy C9+ cut from the return tower of the Pyrolysis Gasoline Hydrogenation unit and is cooled to 90°C prior to being sent to OSBL storage. A side stream product withdrawn from the Fractionator is sent to the PGO Stripper where it is steam stabilized using superheated LP steam injection from the steam system. A portion of the stabilized PGO is filtered and used as purge oil for instruments, the rest is blended with fuel oil and cooled before being sent to OSBL storage.

The TLE effluent quenching and pyrolysis fuel oil stripping is accomplished in such a fashion as to reduce the viscosity of the circulating quench oil stream. The composition of the quench oil is changed by increasing the concentration of relatively lighter components. This is accomplished by stripping the quench oil entering the Pyrolysis Fuel Oil Stripper with part of the recycle heater effluent. A relatively high percentage of components boiling between 280°C and 370°C are stripped out. As a result, these components do not readily leave the system with the stripped pyrolysis fuel oil and, therefore, concentrate in the circulating quench oil. The concentration of this mid-boiling range material will maintain quench oil viscosity to within the desired limits. Filtering of the quench oil leaving the gasoline fractionator, the gasoline fractionator bottoms pump discharge and the PFO stripper feed and product are provided to prevent the buildup of coke fines and allow continuity of operation between plant turnarounds. The level of filtering varies from using strainers at the



gasoline fractionator bottoms and cartridge filters on the quench oil pump discharge, cartridge/basket type for PFO stripper feed and basket type on the PFO stripper product suction filters.

Overhead vapor from the Gasoline Fractionator is cooled and partially condensed by direct countercurrent contact with recirculating water in the Quench Tower. The hot recirculating water from the Quench Tower supplies low level heat to various process users, namely the Naphtha Feed Preheater, Recycle Ethane/Propane Preheater, Charge Gas Heater, Deethanizer Reboiler, MAPD Converter Feed Preheater, Propylene Fractionator Reboilers and Recycle Propane Vaporizer. Quench water is also used to cool the PGO/PFO product to 90°C in the PGO/PFO Product Cooler. The quench water is cooled further by cooling water. The overhead vapor from the Quench Tower is sent to the Charge Gas Compressor. The gasoline condensed in the Quench Tower is separated from the condensed dilution steam in the Quench Water Settler. A part of the condensed hydrocarbons is returned as reflux to the Gasoline Fractionator. The net gasoline is sent to the Gasoline Stripper. Process Water Stripping and Dilution Steam Generation

The dilution steam condensed in the quench tower is sent to the process water stripper, where it is stripped with live dilution steam and LP steam to the bottom of the tower, to remove acid gases and volatile hydrocarbons. The vapor leaving the process water stripper is sent to the quench tower.

The process water from the PW Stripper is preheated by quench oil prior to entering the dilution steam drum where it is then vaporized against circulating quench oil and medium pressure steam in the dilution steam generators. The steam generated is superheated against medium pressure steam before reuse as dilution steam in the cracking heaters, stripping steam to the process water stripper and as purge steam to the heaters.

Provisions for dilute caustic injection to dilution steam drum feed, process water stripper feed and quench tower bottoms are provided for pH control.

To prevent a build-up of non-volatiles, a blowdown stream from the dilution steam drum is drawn, cooled to 40°C against cooling water and sent to OSBL waste treatment facility. Charge Gas Compression, Gasoline and Condensate Stripping

The quench tower overhead vapors are compressed from 0.47 to 38.67 kc/cm2g in a five-stage centrifugal compressor with interstage cooling to about 40°C. Recycle streams from downstream swing PE plant, HDPE plant and PP plant



and off spec ethylene vapors from HP ethylene storage as well as internal recycle streams of the naphtha cracker are also reprocessed in the Charge Gas Compressor system. The charge gas compressor design shall incorporate dry mechanical seals.

Between the third and fourth stages, the charge gas is treated by caustic wash to remove acid gases generated in the cracking heaters. Spent caustic from the caustic/water wash tower is washed with gasoline and sent to OSBL spent caustic treatment facility.

The condensate from the third stage discharge drum is recycled to the third stage suction drum; the condensate from the third stage suction drum is recycled back to the second stage suction drum where hydrocarbon and water separation takes place. Water condensed in the second stage suction drum is recycled to the first stage suction drum and then pumped to the quench tower. Hydrocarbon condensed in the second stage suction drum is sent to the gasoline stripper which operates just above the 1st stage suction drum pressure. In this tower, the hydrocarbon liquid is stripped of butane and light material by vapor generated from the steam heated reboilers. A portion of the bottom liquid is used as wash gasoline in the spent caustic pretreatment section. The net bottoms is combined with the debutanizer bottoms, cooled and sent to the Pyrolysis Gasoline Hydrogenation unit or OSBL storage. The overhead vapor is recycled back to the quench tower. Provisions are available to send a portion of the bottom liquid to the bottom of the quench tower to maintain adequate hydrocarbon inventory for the gasoline fractionator reflux. The 5th stage discharge, after water cooling is sent to the 5th stage discharge drum. The vapor is cooled further by reheating lower deethanizer feed and then by two levels of propylene refrigerant to 15.6°C. The partially condensed liquid/vapor mixture flows to the dryer feed drum. Hydrocarbon and water is separated in the dryer feed drum. The hydrocarbon is pumped to the 5th stage discharge drum, and the water is sent to the 5th stage suction drum. Condensed hydrocarbon and water from the 5th stage discharge drum is sent to the 5th stage suction drum, where hydrocarbon and water separation takes place. Water from the 5th stage suction drum is sent to the 4th stage suction drum. Hydrocarbon from the 5th stage suction drum is sent to the condensate stripper. The condensate stripper operates just above the 4th stage suction drum pressure. In the condensate stripper, the hydrocarbon is stripped of ethane and lighter materials by the vapor generated by a steam heated reboiler. The overhead vapor is recycled to the fourth stage suction drum. Water and hydrocarbon condensed in the fourth stage suction drum is recycled to the quench tower. The stripper bottoms is cooled against cooling water and sent to the condensate stripper bottoms filter/coalescer. The purpose of the



coalescer/filter package is to remove free water from the stripper bottoms before being sent to the dryers. Condensate stripper bottoms, after passing through the coalescer is dried in a liquid-phase two-dryer molecular sieve drying system. One dryer is operating while the other is being regenerated in a cyclic operation. The dried stripper bottoms is sent to the depropanizer. Any water collected in the condensate stripper is withdrawn from an intermediate tray and recycled back to the 3rd stage suction drum.

The Charge Gas Compressor fouling is controlled by injecting wash oil into the suction of each compressor stage. Common boiler feed water and wash oil injection nozzles are provided for all wheels as an option for the future. The source of wash oil can be light cycle oil or partially hydrogenated C9+ stream from the Pyrolysis Gasoline Hydrogenation unit. Acid Gas Removal The 3rd stage discharge drum vapor is preheated to 45°C against quench water prior to entering the caustic/water wash tower. In the caustic/water wash tower the acid gases (H2S and CO2) are removed from the charge gas stream by contacting with caustic solution. 50% caustic is received from OSBL which is diluted to 20% before use in the caustic tower. The tower has two caustic circulating sections in addition to the water wash section. The water wash section has bubble cap trays while the two caustic circulating sections have valve trays. After scrubbing in the two caustic sections, the charge gas essentially free of H2S and CO2 is water washed using blowdown from the steam system in the top water wash section to prevent caustic carryover into the downstream equipment. The majority of the spent wash water is used to dilute the make-up caustic to the upper caustic circulation. The net excess spent wash water is sent to OSBL spent caustic treatment. The charge gas from the overhead of the caustic/water wash tower is returned to the 4th stage suction of the charge gas compressor. The spent caustic blowdown is sent to the spent caustic pretreatment section for removal of free oils via gasoline wash prior to being sent to final OSBL spent caustic treatment.

Any polymeric oil (yellow or red oil) which may accumulate in the tower bottom is recombined with spent caustic stream and sent to the spent caustic/gasoline wash system. The polymeric oil leaves with the spent gasoline from the spent gasoline coalescer and is sent to the quench tower.

Spent Caustic Pre-treatment

The spent caustic solution from the caustic/water wash tower cannot be discharged to the environment without further treatment. The spent caustic solution contains sodium carbonate, sodium sulfide and a small percentage of free (not reacted) sodium hydroxide. In addition, the solution may contain dispersed hydrocarbons. The dispersed hydrocarbons in the spent caustic may



cause considerable fouling in the downstream OSBL treating unit and are therefore removed with a gasoline wash.

Spent caustic from the caustic tower is mixed with wash gasoline from the gasoline stripper in the feed line to the spent caustic coalescer. The spent caustic/wash gasoline mixture is first degassed in the coalescer and then is passed through a packed coalescing pad where the oil/caustic separation takes place. The spent gasoline is then mixed with wash water drawn from the dilution steam generator feed pump in the dilution steam system in the feed line prior to entering the spent gasoline coalescer where the gasoline/water mixture is passed through a packed bed for removal of entrained spent caustic. After separation, the spent gasoline is returned to the quench tower. The spent gasoline requires water washing to reduce the quantity of entrained caustic which can cause a pH problem in the quench tower. The gasoline washed spent caustic from the spent caustic coalescer is sent to OSBL for further treatment.

Charge Gas Drying

The charge gas at 15.5°C from the dryer feed drum is dried in a two bed molecular sieve drying system. One bed is on stream while the other bed is being regenerated in a cyclic operation. Methane offgas from the recovery section is heated by high pressure steam and used to regenerate the desiccant. The regeneration gas is cooled and sent to the regeneration gas K.O. drum for removal of water prior to entering the fuel gas system. Any condensed water from the regeneration gas K.O. drum is returned to the quench tower.

Charge Gas Chilling The dry charge gas is progressively chilled against the process (deethanizer feed, ethylene fractionator side reboiler, ethane recycle, demethanizer bottom and side reboilers) and propylene and ethylene refrigeration to –72°C. The condensate is separated in the demethanizer feed separator No. 1 and fed to the demethanizer as two streams (demethanizer feed No. 1 and 2), after heat exchange against itself.

The charge gas from the demethanizer feed Separator No. 1 is chilled against off gases and coldest level of ethylene refrigerant to –97.9°C. The condensate is separated in separator No. 2 and sent to the demethanizer as feed No. 3.

The charge gas from demethanizer feed separator No. 2 is further chilled in the cold box to –131°C by heat exchange with hydrogen and methane offgas and by evaporation of liquid methane. The condensate is separated in demethanizer feed separator No. 3 and sent to the demethanizer as feed No. 4. The coldest feed to the demethanizer is obtained from this drum and the residual gas is hydrogen of approximately 75 mol% purity.



The hydrogen rich stream is further upgraded to 95+ mol% purity in a two stage expansion system. Here, hydrogen purification takes place in an adiabatic heat exchange system in which the refrigeration required is derived from expansion of liquefied methane (Joule-Thompson Expansion). In the first Joule-Thompson Expansion, the 75 mol% purity stream is chilled to -143 °C, condensing a portion of the contained methane and essentially all of the C2s. Refrigeration is provided by expanding the condensed liquid to demethanizer overhead pressure. After reheating, the methane rich stream is blended with the demethanizer overhead, reheated further and sent to the dryer regeneration/fuel gas system as HP methane. In the second Joule-Thompson Expansion, the overhead vapor which is now about 83.5 mol% purity hydrogen is further chilled to -165 °C. The refrigeration for this step is provided by expanding the condensed liquid to fuel gas system pressure. Liquid is expanded, vaporized, reheated and sent to the fuel gas system as LP methane.

The second JT expansion yields approximately 95.7 mol% hydrogen. This hydrogen is reheated in the chilling train and sent to the methanator to meet the carbon monoxide specification. Demethanizer The condensed liquids from the charge gas chilling train along with the vent gas from ethylene fractionation and propylene fractionation, and light gas recycle from PP plant are sent to the appropriate feed locations of the demethanizer. This tower is operated at a pressure just high enough to permit using the overhead methane product for dryer regeneration and still be at fuel gas pressure. Provision for reprocessing off spec ethylene is also provided in the demethanizer. Based on selected technology, there will be no recycle from PP plant to the demethanizer as shown on the process flow diagram.

The demethanizer reboiling is carried out with charge gas in both the bottom and side reboilers. The bottoms product is reheated against ethylene and propylene refrigerant in the cold box, after which it is split into two streams. One stream is sent directly as liquid to the deethanizer as the top feed. The other stream is sent to the deethanizer as the second feed, but after further preheat against charge gas in dryer effluent chiller and 5th stage discharge drum overhead vapor in dryer feed cooler. Reflux to the tower is provided by an open-loop methane refrigeration system which utilizes a motor driven centrifugal compressor. The methane refrigeration is condensed by the lowest level ethylene refrigeration. The system provides liquid methane as reflux to the tower as well as vapor/liquid methane refrigeration for the chilling of the coldest demethanizer feed.

The overhead of the demethanizer is split into two streams, one stream is mixed with reflux drum liquid and sent to the cold box as high pressure (HP) methane refrigerant. The other stream is heated and compressed. The compressor discharge is cooled against cooling water, chilled by demethanizer overhead



vapor and ethylene refrigerant and then it is sent to the reflux drum. The vapor from the reflux drum is mixed with part of the demethanizer overhead going to the cold box as HP methane. Part of the liquid from the reflux drum is sent to the demethanizer as reflux. The remainder reflux drum liquid is mixed with demethanizer overhead as discussed above and sent to cold box as methane refrigerant.

A small portion of the methane refrigerant compressor discharge is sent to OSBL as ballast gas for downstream MEG plant. Methanation and Hydrogen Drying

The methanation section takes raw hydrogen (95 + mol% hydrogen) generated in the hydrogen methane separator No. 2 and prepares it for use in the downstream hydrogenation process. This involves two primary processing steps:

- Methanation is the conversion of the carbon monoxide (CO) in the

hydrogen to methane (CH4) and water. CO is a catalyst poison in the downstream hydrogenation reactions.

- Drying of the hydrogen is required for the hydrogenation reactors as water is a poison to these catalysts.

The hydrogen stream from the cold box is first heat interchanged with effluent from the methanator and then further heated using high pressure steam to the reaction temperature of 288°C. This temperature is required to initiate the reaction. The conversion of the CO is an exothermic reaction, therefore, the methanator temperature should be monitored closely as runaway reactions can result.

Reactor effluent is used to heat the feed and is then cooled against cooling water. The hydrogen stream is further cooled to about 18.6°C by propylene refrigerant. Condensed water is separated in the hydrogen dryer K.O. drum. The hydrogen leaving the drum goes to the hydrogen dryers. Two hydrogen dryers are provided to accommodate the periodic regeneration required. The dry hydrogen is used in the acetylene, MAPD, C4 hydrogenation and pyrolysis gasoline hydrogenation reactors. Excess hydrogen is sent as product to the OSBL refinery hydrogen header. In addition, any shortage of hydrogen will be supplied from the OSBL refinery header via the hydrogen booster compressor package. Deethanization, Acetylene Hydrogenation and Ethylene Fractionation

The demethanizer bottoms product, which is split into two streams as described in the demethanizer section, feeds the deethanizer. The deethanizer reflux is supplied by condensing overhead vapors with –27.0°C propylene refrigerant. The



column is reboiled with quench water. A spare reboiler using LP steam is provided to insure continuity of operation.

Deethanizer also processes FCC dry gas C2s and FCC C3 streams received from adjacent refinery. Both these streams require treatment before being sent to the deethanizer as follows:

- The FCC dry gas C2 stream contains significant impurities like H2S, CO,

H2, CO2, N2, oxygen, acetylene which must be removed. The FCC dry gas treatment facility will be provided OSBL to produce a stream suitable for addition to the ethylene plant for hydrocarbon recovery. Provision for supplying C3 wash stream from NCU to the OSBL FCC dry gas treatment facility is kept in the design. The C3 wash stream is drawn from MAPD converter recycle cooler outlet.

- FCC C3 stream is saturated with water and contains impurities like CO2, COS, Mercaptans, H2S, amines, arsine, phosphine, and halogens. which must be removed prior to processing in the deethanizer. The treatment facility for this stream is provided in NCU. The treatment facility is discussed in later sections.

Treated FCC dry gas C2s feed is combined with the 2nd feed to the deethanizer. The treated FCC C3s feed is sent to the deethanizer as the third feed.

Acetylene is removed from the net deethanizer overhead product by selective hydrogenation to ethylene and ethane in a two bed acetylene converter with intercooling. The acetylene converter uses a silver promoted palladium catalyst which greatly reduces the need for CO as a catalyst moderator and the production of C2 green oil. A spare converter with spare intercooler is provided so that the catalyst can be regenerated (with a mixture of superheated steam and air) without interrupting the continuity of operation. Hot hydrogen/nitrogen from the MAPD reduction heater is used for catalyst reduction at initial start-up and after each catalyst regeneration. Provisions have been kept (a bypass flow controller line around the methanator) to allow for CO addition, if needed. Plants typically operate between 5-9 months at selectivity levels above 50% minimum acetylene to ethylene.

The net deethanizer overhead is preheated against the converter effluent, mixed with the required hydrogen and further heated by low pressure steam, and passed over the first catalyst bed. The temperature rise is proportional to the percentage of hydrogen added to the feed. A safety monitor is provided to shut off the hydrogen in the event that the reactor temperature becomes excessive. The effluent from the first catalyst bed is mixed with more hydrogen, cooled by cooling water, and then passed over the second catalyst bed for further acetylene conversion. The effluent from the second bed, containing less than 1



ppm of acetylene relative to ethylene, is cooled by cooling water and by heat exchange with the converter feed.

During hydrogenation, a small portion of the acetylene is converted to a polymer called green oil. This material interferes with the proper drying of the ethylene fractionator feed which is essential to avoid icing problems. This green oil is removed by contacting the converter effluent with a slip stream of ethylene/ethane stream taken as a side draw from the ethylene fractionator. This ethylene/ethane stream is mixed with the converter effluent in a static mixer prior to entering the C2 green oil K.O. drum. The C2 green oil K.O. drum bottoms liquid, containing the green oil, is recycled back to the deethanizer. The contained green oil leaves with the deethanizer bottoms to the depropanizers and ends up in the raw pyrolysis gasoline. The overhead vapor from the C2 green oil K.O. drum passes to the ethylene fractionator via the ethylene dryer. The ethylene dryer consists of a single molecular sieve bed. The desiccant is regenerated using hot methane offgas from the dryer regeneration system.

The ethylene fractionator has one bottom reboiler and two side reboilers permitting the maximum cold recuperation from this tower. One side reboiler is used to chill charge gas and the other is used to condense ethylene refrigeration compressor discharge. The main reboiler supply heat to the tower by condensing –6.7°C propylene refrigerant vapor. The reflux for the tower is condensed with –40.3°C propylene refrigerant. A pasteurization section on top of the ethylene fractionator is used to remove the residual methane and hydrogen from the ethylene product. Vent gas from the reflux drum is recycled back to the demethanizer after passing through the vent condenser which uses –63°C ethylene refrigerant to minimize the vent recycle rate.

The ethylene product is withdrawn as a side draw from the tower and sent to the ethylene product surge drum along with excess ethylene liquid from the ethylene refrigerant accumulator corresponding to ethylene boil-off gas. Liquid ethylene product drawn from the tower is also used as makeup for the ethylene refrigeration system. From the surge drum, the ethylene product is handled as follows:

- One stream is chilled against –63°C, -75.0°C and –101.0°C ethylene

refrigeration in succession to –98°C and delivered to low pressure OSBL cryogenic storage. In normal operation, 1200 kg/h of ethylene consistent with ethylene boil-off gas rate is sent to low pressure cryogenic storage in normal operation. Provision is kept to send up to 30,000 kg/h ethylene to this storage.

- The second steam is pumped, heated, vaporized and superheated against liquid propylene and delivered as MP ethylene vapor product at 28.0 kg/cm2g and 30°C.



- The third stream is pumped, heated, vaporized and superheated against liquid propylene and delivered as HP ethylene vapor product at 52.0 kg/cm2g and 30°C.

In normal operation, liquid ethylene up to 10 t/h (10%) is recycled from atmospheric storage to keep storage pumping system active. Ethane is withdrawn from the ethylene fractionator bottom, vaporized against charge gas and after further reheating against propylene refrigerant in the cold box and is sent to the recycle cracking heaters or to the fuel gas system in the event cracking heaters are unavailable.

Provisions are made for polymerization inhibitor injection into the condensate stripper, reboilers, condensate stripper bottoms, C2 green oil bottoms feed to deethanizer, deethanizer reboilers, depropanizer feed from deethanizer and depropanizer reboilers. Depropanization

The purpose of the depropanizer is to make a sharp separation between C3 components and the C4 and heavier components in the deethanizer bottoms and condensate stripper bottoms.

Deethanizer bottoms, condensate stripper bottoms and a C3 green oil recycle stream from the propane recycle drum are the feed to the depropanizer. The tower overhead is condensed against -6.0°C propylene refrigerant. Part of the distillate is used as reflux and net product and is pumped to the MAPD converter system. The tower is reboiled by low pressure steam. The bottom product containing C4s and heavier material is sent to the debutanizer.

MAPD Hydrogenation

In this section methyl acetylene (MA) and propadiene (PD) contained in the depropanizer overhead are removed by selective hydrogenation to propylene and propane in a single-bed reactor.

The net overhead stream from the depropanizer is pumped to the MAPD converter. The MAPD reactor utilizes the Süd-Chemie G-68HX hydrogenation catalyst. This catalyst achieves a selectivity to propylene of about 50% for an outlet MAPD concentration of 200 mol. ppm (30% selectivity used for design). The MAPD converter system consists of two reactors; one operating, and one spare. As a result, the reactor can be regenerated while the plant is on-stream.

The C3 feed stream is mixed with recycle liquid from the MAPD Converter Effluent Separator and the hydrogen from the methanator section and passed through the converter. The MAPD hydrogenation is an exothermic reaction and



the heat of reaction is removed by vaporizing a portion of the feed with only a moderate temperature rise. The fresh feed is diluted with liquid recycle to limit vaporization, improve the overall selectivity and minimize green oil formation.

Liquid leaving the reactor flows to the MAPD Converter Effluent Separator. Vapor leaving the reactor is cooled in a water-cooled vent condenser. Liquid from the vent condenser is returned to the separator drum. Vapor leaving the vent condenser is returned to charge gas compressor to recover H2, CH4 and C3 hydrocarbons. Liquid recycle leaving the drum is pumped, cooled and returned to the reactor inlet for diluting the fresh feed. Converter feed heater is provided for achieving end of run feed temperature during reduced catalyst activity. The net C3 product from the separator drum is combined with C3 recycle from PP plant and sent to the Propylene Fractionator No 2. Based on selected technology, there will be no recycle from the PP plant to the propylene fractionator.

During hydrogenation, a small amount of the MAPD is converted to a polymer called C3 green oil. Contained green oil in the reactor effluent gets recycled from the propane recycle drum to the depropanizer which eventually ends up in the gasoline product.

Propylene Fractionator

The propylene fractionator uses a two-tower system to separate the feed into a distillate product of polymer grade propylene and a bottom product containing primarily propane. The tower operates at a pressure which permits total condensation of propylene fractionator No. 2 overhead against cooling water. Reboiler heat to the towers is supplied by circulating quench water.

The polymer grade propylene product is withdrawn as a sidedraw from tower no. 2. The product is pumped, cooled to 38°C and delivered to OSBL downstream plants. Provision for diverting off spec propylene to OSBL HP storage is also kept. A portion of the sidedraw product can be chilled at fractionator pressure against -27°C and -40.3°C propylene refrigerant and -63°C ethylene refrigerant in succession and delivered to low pressure OSBL cryogenic storage at -45oC.

The propane-rich bottoms product is withdrawn as bottoms product from propylene fractionator No. 1 and sent to the propane recycle drum which is equipped with a thermosyphon reboiler using quench water as the heating medium. The recycle propane is vaporized and mixed with ethane recycle and sent to the recycle cracking heater. A small liquid stream is withdrawn from the propane recycle drum (to remove the green oil formed in the MAPD converter) and returned to the depropanizer.

A pasteurization section on top of the propylene fractionator is used to remove the residual methane and hydrogen from the propylene product. Vapor from the



reflux drum is cooled in the vent condenser against 15.6°C propylene refrigerant to minimize propylene carryover. Vent gas from the vent condenser is recycled to the demethanizer.

Debutanizer

The bottoms product from the Depropanizer flows to the debutanizer where the raw C4s product is separated. The debutanizer operates at a pressure which permits total condensation of the overhead vapor against cooling water. The bottoms reboil heat is provided by low pressure steam. The debutanizer net overhead product, consisting of mixed C4s, is pumped to the butadiene extraction unit, C4 hydrogenation unit, or to OSBL storage for further processing.

The bottoms product is pumped and combined with gasoline from the gasoline stripper to make up the total raw pyrolysis gasoline product. After cooling against cooling water, the gasoline product is sent to pyrolysis gasoline hydrogenation unit. Pyrolysis gasoline should always flow directly from debutanizer to pyrolysis gasoline hydrogenation unit. Routing to intermediate storage provided between the NCU and PGH units should be done during PGH catalyst regeneration only.

FCC C3s Treatment The FCC C3 stream is passed through the dryer/treater and arsine treater and filter before being sent to the deethanizer. The dryer/treater is regenerable, whereas the arsine treater is non-regenerable. Two dryer/treaters are provided, one dryer/treater on-stream while the other is being regenerated in a cyclic operation. A compound bed of three types of adsorbents is used in the regenerable dryer/treater to remove water, CO2, COS, H2S, amines, halogens and mercaptans. The bottom bed removes water and halogens, the middle bed removes water and the top bed removes the remaining impurities. The arsine treater contains a single bed of copper oxide/aluminum oxide based adsorbent. The treater removes arsine and phosphine from the FCC C3 stream. The treated stream is passed through a filter for removing any carried over adsorbent particles before being sent to the deethanizer as third feed. C4 Hydrogenation

C4 hydrogenation unit is designed to operate in two modes; 1) total hydrogenation of the BD raffinate in normal operation, and 2) partially hydrogenate the raw mixed C4 (butadiene saturation only) when butadiene extraction unit is not in operation.



Raffinate from BD plant or raw mixed C4 from debutanizer or OSBL storage flows to the feed surge drum. In the feed surge drum, any entrained water, if present, is separated. The feed is then pumped to reactor pressure and mixed with liquid recycle stream prior to entering the reactor.

A single reactor is provided. When end of run conditions are reached, catalyst is regenerated in-situ using a conventional steam and air procedure. Expected regeneration cycle varies from approximately 1 year when processing raw C4s to over 2 years when processing BD raffinate. Provision is made to inject CO rich hydrogen to control the activity of the catalyst during partially hydrogenation of raw C4s.

The recycle liquid from the HP flash drum is heated against liquid effluent from the reactor prior to mixing with the feed. This recycle liquid is required to limit vaporization and temperature rise across the reactor. The reactor inlet temperature is controlled by adjusting the split of recycle through the feed/effluent exchanger. At start-of-run, the tube side of the feed/effluent exchanger is completely bypassed. A feed preheater using LP steam is provided for start-up only.

When processing BD raffinate, the entire hydrogen flows to the top bed on flow control (reset by pressure control). When processing raw C4s, the major part of the hydrogen is added to the top bed on flow control and the remainder is injected to the lower bed on pressure control. The mixed phase feed passes downward through the catalyst beds where butadiene and butenes are hydrogenated. The exothermic heat of reaction causes a temperature rise and partial vaporization in the reactor. Liquid effluent from the reactor is cooled in a feed/effluent exchanger. After combining with the vapor effluent from the reactor, the mixture is partially condensed against the cooling water. Effluent from the condenser passes to the HP flash drum where liquid recycle is separated and returned to be mixed with fresh feed to the reactor. HP flash drum vapors are chilled against 3°C propylene refrigerant in a vent condenser. Liquid from the vent condenser flows back to the flash drum and non-condensables are recycled to the charge gas compressor 3rd stage discharge drum for recovery. The net liquid product is withdrawn from the HP flash drum and sent to the C4 stabilizer.

In the C4 stabilizer, dissolved hydrogen and methane are stripped from the C4 product. The stabilizer overhead is partially condensed against cooling water and reboiled using LP steam. The C4 product is withdrawn from the C4 stabilizer bottom, cooled in a water-cooled exchanger and recycled to cracking heater or pumped to OSBL LPG storage. The vent gases containing unreacted hydrogen and methane and equilibrium C4s are chilled against 3°C propylene refrigerant in a vent condenser. Liquid from the vent condenser flows back to the reflux drum and offgases are recycled to the charge gas compressor 3rd stage discharge drum along with offgases from HP flash drum.



Pyrolysis Gasoline Hydrogenation

The raw pyrolysis gasoline from the debutanizer and gasoline stripper, after cooling and filtering, is sent to feed surge drum where entrained water, if any, is separated and drained. Normally, the feed does not contain entrained water.

DPG First Stage

The pyrolysis gasoline is pumped from the surge drum to the first stage reactor. The fresh liquid feed is mixed with recycle liquid from the bottom of the reactor before entering the top bed of the reactor. Catalyst activity decreases with time, requiring raising of the reactor inlet temperature. The reactor inlet temperature is set by controlling the split of recycle flow through the recycle air cooler. The quantity of recycle is controlled to limit the exothermic temperature rise across the reactor. A MP steam exchanger is provided for startup.

The hydrogen enters the top of the first stage reactor under pressure control and travels down through the catalyst bed in intimate contact with the total liquid feed. The principal hydrogenation reactions occurring are the conversion of diolefins to olefins and styrene to ethylbenzene. Less than 10% of the diolefins and olefins are converted to paraffins.

The reactions take place in the liquid phase. The exothermic temperature rise across the reactor is controlled by adjusting the flow of recycle liquid.

Reactor temperature is raised as required to maintain product quality. The inlet temperature of the reactor may vary from about 60°C at start of run to 133°C at end of run. When the upper limit of operating temperature or reactor pressure drop is reached or product quality can no longer be maintained, the catalyst is regenerated in-situ using a conventional steam and air procedure.

The DPG1 reactor consists of one reactor. As a result, when reactor is regenerated, the raw pyrolysis gasoline feed from the debutanizer section is sent to OSBL storage. The storage capacity should be adequate for the regeneration period. Net liquid effluent is cooled against cooling water and is pumped to the first stage HP flash drum where hydrogen-rich gas is separated from the liquid product. Hydrogen-rich gas from the flash drum is sent to the second stage.

Stabilizer

The liquid from the flash drum is sent to the Stabilizer. In the Stabilizer, the light ends are separated from the treated gasoline. The tower is reboiled using MP steam and the overhead is condensed against cooling water. The vent gas from the reflux drum is sent to the charge gas compression section. The tower bottoms flow to the Tailing Tower.



Tailing Tower The Tailing Tower separates the treated gasoline from the stabilizer bottoms into a C5-C8 cut overhead product and a C9+ bottoms product. The tower is reboiled using MP steam and the overhead is condensed against air cooling. Tower pressure is controlled by varying the pitch of the air coolers or by nitrogen injection. The C9+ bottoms product is sent to the rerun tower and the C5-C8 cut is sent to the DPG second-stage reactor.

Rerun Tower The purpose of the rerun tower is to remove any heavy ends (200°C) material from the C9+ stream. Gums present in the first stage DPG product end up with the heaviest cut. As a result, the gums are concentrated in the 200°C plus stream exiting the tailing tower bottom.

C9 plus stream from the tailing tower feeds the rerun tower. The rerun tower is reboiled using MP steam and the tower overhead is condensed with cooling water. The tower is operated under vacuum to minimize fouling in the tower bottom. The reboiler circuit operates in forced circulation mode for zero vaporization to minimize fouling. Non-condensable offgases (i.e., air leakage) are boosted using an ejector with MP steam as the motive fluid to a pressure high enough to send offgas to the heater firebox.

The heavy cut concentrated with gums is separated in the tower bottoms and is mixed with pyrolysis fuel oil, cooled and sent to OSBL storage. DPG Second Stage

The C5-C8 cut from the DPG first stage is the feed to the DPG second stage.

First stage treated C5-C8 heart cut is mixed with hydrogen-rich recycle gas and recycle liquid. The two-phase mixture is completely vaporized and heated by heat exchange against the second stage reactor effluent to the required reactor inlet temperature. A super high-pressure steam exchanger is provide for start-up.

In the second stage reactor, the olefins present in the feed are hydrogenated, and the sulfur compounds are converted to hydrocarbons and H2S over the catalyst. The reaction takes place in the vapor phase. The reactions are exothermic resulting in a temperature rise across the reactor which is moderated by recycle liquid. Catalyst activity decreases with time, requiring raising of the reactor inlet temperature. The inlet temperature of the reactor may vary from 220°C at start of run to about 250°C at end of run. When the upper limit of the operating temperature range is attained, the catalyst is regenerated in-situ, using



the conventional steam/air procedure. The reactor effluent preheats the reactor feed before finally being cooled and partially condensed against cooling water.

The cooled vapor/liquid mixture is then separated in the high-pressure flash drum. The major portion of the vapor leaving this drum is mixed with the remainder of the make-up hydrogen required in the second stage and sent to the recycle gas compressor suction drum. Vapor from this drum is compressed in the recycle gas compressor and mixed with recycle liquid and first stage treated C5-C8 cut. A small net vent flow from the high-pressure flash drum is sent to the charge gas compression section. A portion of the liquid product (recycle liquid) is pumped from the high-pressure flash drum to be mixed with reactor feed. The net liquid product is sent to a low-pressure flash drum. The liquid from the low-pressure flash drum is sent as feed to the Depentanizer Tower.

Depentanizer

In the depentanizer, the C5 stream is separated from the C6 plus stream. The tower is reboiled using MP steam and the overhead is partially condensed in the air cooler and trim condenser against cooling water. The vent gas from the reflux drum plus the vapor from the low pressure flash drum are sent to the charge gas compression section. The tower bottoms is sent to battery limits as the C6-C8 Heart Cut. The C5 product is recycled to the cracking heaters to be co-cracked with the naphtha feedstock. Propylene Refrigeration The propylene refrigeration system is a closed, four stage system utilizing a steam turbine driven centrifugal compressor. The system provides refrigeration at four levels: -40.3°C, -27°C, -6.0°C and 15.6°C. Cooling water is used to condense the compressor discharge vapors.

Refrigeration is recuperated against offgas reheating, ethane recycle heating, deethanizer feed heating, ethylene fractionator reboiling and ethylene product vaporizing and superheating in the chilling train.

The propylene refrigeration system is provided with suction drums at each stage to provide surge volume for the various system users and to minimize liquid entrainment into the compressor at each stage as well as an accumulator at the compressor discharge to provide a liquid seal to ensure refrigerant condensing at the compressor discharge.

The propylene compressor design incorporated dry mechanical seals. The system is also designed to provide refrigeration for two BD plant users; namely C4 acetylene condenser and butadiene product cooler.



Ethylene Refrigeration

The ethylene refrigeration is a closed, three stage system utilizing a steam turbine driven centrifugal compressor. The system provides three levels of refrigeration, i.e., -101.0°C, -75.0°C and –63°C. The compressor discharge vapors are cooled first against cooling water and then partially desuperheated against 15.6°C propylene refrigerant. The discharge is further desuperheated condensed in the ethylene fractionator side reboiler. A back-up condenser against –27°C propylene refrigerant is provided for start-up.

Refrigeration is recuperated by subcooling liquid ethylene refrigerant from the accumulator against offgases, deethanizer feed in the chilling train.

The ethylene refrigerant system is provided with suction drums at each stage to provide surge volume for the various system users and to minimize liquid entrainment into the compressor at each stage as well as an accumulator at the compressor discharge to provide a liquid seal to ensure refrigerant condensing at the compressor discharge.

The ethylene compressor design incorporated dry mechanical seals which prevents any compressor seal oil from entering the system and freezing. This eliminates the need for an oil mist eliminator in the compressor discharge.

BOG vapors from the OSBL low temperature atmospheric ethylene storage is reprocessed in the NCU via the ethylene refrigeration compressor. BOG vapors flow to the first stage of the ethylene refrigeration compressor and get condensed in the compressor discharge condenser. This liquid from the ethylene refrigerant accumulator is sent to the ethylene product surge drum and then chilled in ethylene product rundown chillers before being sent to OSBL cryogenic ethylene storage.

Fuel Gas System Primary fuel gas is a mixture of HP methane offgas from dryer regeneration system and LP methane offgas from chilling train. Makeup/backup fuel is C3/C4 LPG supplied from the adjacent refinery. Provision to use revaporated liquefied natural gas (RLNG) as makeup/backup fuel in the future is also provided. In normal operation, primary fuel gas is not sufficient to fully meet the heater firing demand. The shortfall is met by makeup LPG fuel.

Primary fuel gas is received in fuel gas mix drum. Ethane recycle vapor, BD unit offgas and excess hydrogen can also be received in this drum. Any heavy liquid condensed in the drum is drained to wet flare knockout drum. RLNG vapor will be received in the mix drum under pressure control.



Low NOx burners require fuel gas to be free of rust particles and liquid droplets. Therefore, fuel gas from mix drum passes through filter/coalescer before being sent to the heaters. C3/C4 LPG is received in the fuel gas vaporizer drum under level control. It is vaporized with LP steam. The LPG vapor is superheated and sent to the fuel gas mix drum under pressure control during normal, start-up and upset conditions.



ARTICLE II – PLANT SPECIFICATION 1.0 BASIS FOR PROCESS DESIGN 1.1 Plant Function and Capacity Plant Function The naphtha cracker unit is designed to produce polymer grade ethylene and polymer grade propylene by thermal cracking of naphtha and hydrogenated C4/C5/C6 recycle streams from C4 hydrogenation, pyrolysis gasoline hydrogenation and benzene extraction units respectively. Ethane and propane are recycled and cracked to extinction. Recycles from polymer plants, ethylene BOG and the FCC streams are the additional feeds to the Naphtha Cracker Recovery unit. In addition to ethylene and propylene, the plant will produce the following by-products:

Naphtha Cracker products • Hydrogen • Methane offgas • Raw mixed C4s (1) • C4 LPG (provision to draw part of hydrogenated C4 to LPG pool) • Hydrogenated C6-C8 (2) • Partially hydrogenated C9+ • Pyrolysis Fuel Oil (combined PGO + PFO)

Notes: (1) Raw mixed C4s is sent to Butadiene Extraction unit for butadiene

recovery. BD raffinate is sent to the C4 hydro unit. (2) Hydrogenated C6-C8 is sent to Benzene Extraction unit for benzene

recovery. Non-aromatic C6 stream is recycled to the NCU cracking heaters and C7-C8 cut is sent as product to OSBL storage.

Considering all process units, overall product pattern is as follows:

• Hydrogen • Methane offgas • Polymer Grade Ethylene • Polymer Grade Propylene • Butadiene • Benzene • Hydrogenated C7-C8 cut



• Partially hydrogenated C9+ cut • Pyrolysis fuel oil (combined PGO + PFO)

Plant Capacity The plant is designed to produce 800,000 TPA of polymer grade ethylene and 504,834 TPA of polymer grade propylene in 8,000 operating hours a year. The production rates are net without processing polymer plant recycles, ethylene BOG and the FCC feeds. The plant capacity with recycles from polymer plants, ethylene boil-off gas and FCC feeds is as follows:

Case Ethylene Capacity, TPA

Propylene Capacity, TPA

1 (without FCC feeds) 830,632 584,352 2 (with FCC feeds) 857,408 666,544

1.2 Feedstock Properties 1.2.1 Design Naphtha

Density at 15.5°C, kg/m3 702.03 ASTM D-86 Distillation (vol%) °C 1% 51 5% 60 10% 62 30% 74 50% 97 70% 114 90% 147 95% 151 98% 178 EP 182 PONA (wt%) N-Paraffins * 34.1 I-Paraffins 34.1 Naphthenes 21.7 Aromatic 10.1

* assumed 50/50 iso/normal paraffin split



1.2.2 Check Naphtha

Density at 15.5°C, kg/m3 700.00 ASTM D-86 Distillation (vol%) 1% 45 5% 60 10% 65 30% 79 50% 100 70% 117 90% 138 98% 146 EP 155 PONA (wt%) N-Paraffins * 36.5 I-Paraffins 36.5 Naphthenes 19.0 Aromatic 8.0

* assumed 50/50 iso/normal paraffin split

Contaminants in design and check naphtha are as follows:

Contaminants Unit Value Sulfur wt. ppm 500 max. Total Chloride wt. ppm 3 max. Lead wt. ppb 5 max. Arsenic wt. ppb 1 max. Mercury wt. ppb 1 max.

Note: Plant design will be based on design naphtha. However, plant is capable

of processing check naphtha with production rates same as with design naphtha.

1.2.3 Recycle Streams from Polymer Plants Naphtha cracker is designed to process recycle streams from polymer plants. Flow rates and composition of these streams are presented in this section. As polymer plants licensor selection was not completed at the design basis meeting stage, preliminary information of the potential licensors was given as the basis for material balances. Therefore, for each recycle, data is given for two cases:



i) Flow rate and composition used in material balances, ii) Flow rate and composition based on selected licensors.

Ethylene Purge Stream from LLDPE/HDPE (Swing Plant) Based on selected licenser

Ethylene 93-95 wt% Ethane 2-3 wt% Hydrogen 0.5 wt% max Nitrogen 0.5 wt% max Methane 1.0 wt% max Acetylene 120 wt ppm max CO 15-30 wt ppm max CO2 0.25 wt% max Propylene 0.06 wt% max Butene-1 3.0 wt% max Butene-2 0.1 wt% max Flow rate, kg/h Normal 2000-2800

Max 3800 Lummus Material Balance Basis

Composition, wt% Hydrogen 1.5 Ethylene 25.0 Ethane 7.0 Propane 65.5 Flow rate, kg/h 1,750 Phase Vapor

Recycle from PP Plant Based on selected licenser

wt% Propylene 70 - 80 Propane 17.5 - 20 Butene-1 0 - 13 Hydrogen 0 - 0.05 Ethylene 0 - 3.5 Ethane 0 – 0.1 Flow rate, kg/h 1800 - 2400

(Traces of CO, CO2,, methanol and iso-propanol are present)



Lummus Material Balance Basis Two recycle streams from PP plant are considered in material balances.

(i) C3 Recycle

Component kg/h Ethylene 1 Ethane 0 Propylene 8,397 Propane 452 Total 8,850 Phase Liquid

(ii) Light Gas Recycle

Component kg/h Nitrogen 7 Hydrogen 8 Methane 3 Ethylene 1,244 Ethane 18 Propylene 1,244 Propane 63 Total 2,587 Phase Vapor

HDPE Plant Recycle Based on selected licenser Energy rich stream (low N2 concentration)

wt% H2 0 – 5.0 N2 1 – 7.5 C2s 60 – 90 C4s 0 – 0.5 C6 5.0 - 20 Flow rate, kg/h Normal: 270

Peak: 650

Note: The stream is water saturated at about 45°C and 4.08 kg/cm2g



Lummus Material Balance Basis Recycle HDPE was not envisioned at design basis meeting stage, therefore, it is not considered in the material balances. 1.2.4 Ethylene Boil Offgas from Atmospheric Storage

Flow Rate 1200 kg/h (estimated) Composition Same as of ethylene product

1.2.5 FCC Feeds The plant processes two FCC streams received from adjacent refinery.

• Treated FCC dry gas C2s • Raw FCC C3s

Flow rates and composition of FCC feeds are as follows:

(i) Treated FCC Dry Gas C2s

Treated FCC Dry Gas C2s

Gross Feed including C3 Wash Stream

Component kg/h Ethylene 1,449.50 1,449.50 Hydrogen 0.18 Methane 1.52 Ethane 634.40 634.61 MAPD 0.25 Propylene 144.00 1,942.46 Propane 44.10 229.95 1-butene 41.00 41.58 Butadiene 0.03 N-butane 55.80 55.80 C5s 116.12 116.12 C6s 0.92 Total 2,484.92 4,472.92



(ii) Raw FCC C3s

Component kg/h Methane 0.90 Ethylene 11.71 Ethane 123.35 Propylene 9,688.03 Propane 2,588.31 Total 12,412.30

Contaminants CO 0.1 vol ppm max CO2 5 wt ppm max. O2 2 vol ppm max. COS 5 wt ppm ma. H2S 1.0 wt ppm max Mercaptan Normally nil In upset condition:

4 times per month for 8-10 hrs, mercaptan level can reach 50 wt ppm

Arsine 100 wt ppb max. Phosphine 100 wt ppb max. Mercury Nil Amine 3-4 vol ppm max. Fluoride 1 wt ppm max. Chloride 1 wt ppm max. Water Saturated at 43°C (no free water)

1.2.6 Hydrogen Import from Adjacent Refinery

Purity PSA quality, 99.5 mol% hydrogen



1.3 Product Specifications 1.3.1 Polymer Grade Ethylene

Component Specification Unit Test Method Ethylene 99.95 mol% min. ASTM D2505 Methane + Ethane 500 mol ppm max. ASTM D2505 Ethane 300 mol ppm max. ASTM D2505 Acetylene 1 wt. ppm max. ASTM D2505 Propylene 10 wt. ppm max. ASTM D2505 Hydrogen 2 mol ppm max. ASTM D2504 Carbon Monoxide 0.03 mol ppm max. ASTM D2504 Oxygen 1 mol ppm max. O2 meter Carbon Dioxide 5 ** mol ppm max. ASTM D2505 Water 1 Mol ppm max. Dew Point Meter Ammonia 1 mol ppm max. Colorimetric Method Phosphine 0.03 mol ppm max. Sulfur as H2S 1 mol ppm max. GC Oxygenated solvents* 1 wt. ppm max. GC

* Methanol, ethanol, propanol, acetone, acetaldehyde ** Expected CO2 in ethylene product is 3 mol ppm 1.3.2 Polymer Grade Propylene

Component Specification Unit Test Method Propylene 99.5 wt% min. 100-impurities Propane 0.4 wt% max. ASTM D2163 Methylacetylene 2 wt. ppm max. ASTM D2712 Propadiene 3 wt. ppm max. ASTM D2712 Methane 200 vol ppm max. ASTM D2712 Ethane 300 vol ppm max. ASTM D2712 Ethylene 10 wt. ppm max. ASTM D2712 Acetylene 1 wt. ppm max. ASTM D2712 1,3-Butadiene* 3 wt. ppm max. ASTM D2712 Total C4s 10 wt. ppm max. ASMT D2712 C5 & heavier 10 wt. ppm max. ASTM D2712 Hydrogen 1 wt. ppm max. ASTM D2504 Carbon monoxide 0.03 wt. ppm max. ASTM D2504 Oxygen 1 wt. ppm max. O2 meter COS 20 wt. ppb max. GC Carbon dioxide 1 wt. ppm max. ASTM D2505 Water 2 wt. ppm max. Dew Point Meter Ammonia 0.2 wt. ppm max. Colorimetric Method Total sulfur as H2S 0.5 wt. ppm max. ASTM D3246 Arsine 20 wt. ppb max. UOP 834/UOP 387 Phosphine 0.03 wt. ppm max. Methanol 5 wt. ppm max. GC



Component Specification Unit Test Method Oxygenated solvents** 10 wt. ppm max. GC

* 1,3 Butadiene expected to be nil ** Methanol, ethanol, propanol, acetone, acetaldehyde

1.3.3 H2 Product

Component Hydrogen 95.68 mol% Methane 4.32 mol% Carbon monoxide 1 mol ppm CO2 + Nitrogen Traces Chloride content 1 vol ppm

1.3.4 Methane Rich Offgas (expected)

Component, mol% Case 2 Hydrogen 14.32 Methane 84.85 Total C2s 0.41 Carbon monoxide 0.42 100.00

1.3.5 Raw Mixed C4

C3s and lighter 0.3 wt% max. C5s and heavier 0.3 wt% max.

Expected Composition, wt% Case 2 MAPD 0.13 Propylene 0.06 Propane 0.02 1,3 Butadiene 28.80 1,2 Butadiene 0.17 Vinyl Acetylenes 0.29 Ethyl Acetylenes 0.10 1-butene 11.58 Cis-2-butene 3.26 Trans-2-butene 3.99 Isobutene 24.44 n-butane 12.66 iso-butane 14.20 C5s 0.30 Total 100.00



Impurities, wt ppm Water Nil Total sulfur 10 wt ppm max.

1.3.6 Butadiene Product

Sr. No. Property Specification Unit Test Method 1. 1-3 Butadiene 99.6 min. wt% GC Technique 2. 1-2 Butadiene 20 max. wt ppm GC Technique 3. Propadiene 5 max. wt ppm GC Technique 4. Total Acetylene 20 max. wt ppm GC Technique 5. Carbonyls as

Acetaldyhyde 10 max. wt ppm GC Technique

6. Sulfur Note 1 wt ppm ASTM D5453 7. NMP/Solvent 5 max. wt ppm GC Technique 8 Butadiene Dimer 50 max. wt ppm GC Technique 9. Butane + Butene 0.4 max. wt% GC Technique

10. Methyl Acetylene 5 max. wt ppm GC Technique 11. TBC content Note 2 wt ppm ASTM D1157

Notes:

1. This specification depends upon the sulfur content (mercaptan) of the mixed C4 stream. The source of the Mercaptans in C4 product is the unreacted mercaptans contained in the naphtha feedstock. Since feed concentrations are not available and the kinetics are not precisely known, it is difficult to predict the sulfur content of the mixed C4 stream. Typical sulfur specification for butadiene product is 10 wt ppm.

2. TBC is injected into the butadiene product to inhibit polymer formation

during storage. Typical TBC dosage in butadiene product is 50-75 wt ppm.

1.3.7 Raw Pyrolysis Gasoline

C4 and lighter 0.5 wt% End Point (ASTM) 204°C

Expected Composition, wt% Case 2 C4s 0.31 C5 Saturates 8.76 C5 Olefins 3.08 C5 Diolefins 10.53 C6 Saturates 6.34 C6 Olefins 4.33 C6 Diolefins 4.77 Benzene 16.37



C7 Saturates 1.65 C7 Olefins 1.81 C7 Diolefins 1.87 Toluene 15.50 C8Saturates 0.75 C8 Olefins 0.52 C8 Diolefins 0.52 Ethylbenzene 1.89 Xylenes 6.71 Styrene 2.52 C9 – 204°C 11.77 Total 100.00 Sulfur, wt ppm 130

1.3.8 Hydrotreated C6-C8 Heart Cut to Benzene Extraction Unit

C5s 1.0 wt% max. C9-204°C 1.0 wt% max.

Expected Composition, wt% Case 2 Cyclopentane 1.00 C6 Paraffins 10.45 Methyl cyclopentane 9.98 Cyclohexane 2.80 Benzene 24.56 C7 Paraffins 4.92 1,1 dimethyl cyclo C5 0.17 cis 1,3-dimethyl cyclo C5 0.37 trans 1,3-dimethyl cyclo C5 0.30 cis 1,2-dimethyl cyclo C5 0.13 trans 1,2-dimethyl cyclo C5 0.54 Methyl cyclohexane 1.13 Ethyl cyclopentane 0.64 Toluene 23.34 C8 Paraffins 1.50 Total dimethyl cyclohexane 0.30 Ethyl cyclohexane 0.03 n-propyl cyclopentane 0.09 Isopropyl cyclopentane 0.06 Total trimethyl cyclopentane 0.52 Ethylbenzene 6.67 Xylenes 9.81 C9 plus 0.69 Total 100.00



Bromine Index, mg/100g 300 (max) Water No free water Total Sulfur, wt ppm 0.2 (max)

1.3.9 Benzene Product

Property Specification Unit Test Method Solidification Point 5.4 °C min. ASTM D852 Benzene Purity 99.9 wt% min. ASTM D4492 Total Sulfur 0.5 wt ppm max. ASTM D4045 Thiophene 0.5 wt ppm max. ASTM D1685 Toluene 100 wt ppm max. ASTM D4492 Toluene + MCH 150 wt ppm max. ASTM D5713 Xylene 10 wt ppm max. ASTM D4492 Non-Aromatics 400 wt ppm max. ASTM D2360 Acid Wash Color 1 wt ppm max. ASTM D848 Acidity Nil ASTM D847 Copper Strip Corrosion

Pass 1a or 1b ASTM D849

Bromine Index 10 mg Br/100 gm max.

ASTM D1492

Appearance Clear sediment free

Visual

Color 10 Alpha max. ASTM D1209 Relative Density 0.882-0.886 At 15.56/15.56 C ASTM D4052 Distillation Range <1 incl. 80.1 At 760 mm Hg. °C ASTM D-850 Total Chloride, as Cl 1 wt ppm max. ASTM D5194 H2S/SO2 Nil wt ppm max. ASTM D853 Nitrogen-organic bound

1 wt ppm max. ASTM D4629

Water 200 wt ppm max. ASTM D1364 Solvent 1 wt ppm max. ASTM D-

3961 Dioxane 1 wt ppm max. UOP 921

Method



1.3.10 C7-C8 Product

Expected Composition, wt% Case 2 Cyclohexane 0.02 Benzene 0.06 C7 Paraffins 7.38 1,1 dimethyl cyclo C5 0.05 cis 1,3-dimethyl cyclo C5 0.34 trans 1,3-dimethyl cyclo C5 0.34 cis 1,2-dimethyl cyclo C5 0.24 trans 1,2-dimethyl cyclo C5 0.62 Methyl cyclohexane 2.22 Ethyl cyclopentane 1.29 Toluene 46.79 C8 Paraffins 3.32 Total dimethyl cyclohexane 0.66 Ethyl cyclohexane 0.07 n-propyl cyclopentane 0.20 Isopropyl cyclopentane 0.13 Total trimethyl cyclopentane 1.15 Ethylbenzene 13.30 Xylenes 19.79 C9 plus 2.03 Total 100.00

Related Specs Bromine Index, mg/100g 10 (max) Total Sulfur, wt ppm 0.2 (max) Estimated RON 100 Estimated MON 94 Density 844 kg/m3 at 15.5°C Benzene Content 0.054 to max 0.9 vol% Aromatics 82 wt% Olefins 60 wt ppm max



1.3.11 C9+ Product

Estimated composition for Case 2.

Temp., °C IBP 158 10 vol% 170 50 vol% 175 90 vol% 185 EP 195 Total Sulfur <150 wt ppm Aromatics 80 wt% Existent Gum <5 mg/100 ml

Related Specs Estimated RON 103 Estimated MON 94 Density 898 kg/m3 Benzene Content Nil (2 ppb) Olefins 20 wt%

Note: Inhibited with antioxident

Estimated Component Breakdown wt% Cumene 1.65 n-Propyl Benzene 1.65 1-Methyl-3-Ethyl-Benzene 12.62 1-Methyl-4-Ethyl-Benzene 5.19 1,3,5-Trimethyl Benzene 3.30 1-Methyl-2-Ethyl-Benzene 5.78 Meta Methyl Styrene 2.48 1,2,4-Trimethyl Benzene 10.38 C10 Paraffins/Naphthenes 0.47 1,2,3-Trimethyl Benzene 3.19 Dicyclopentadiene (DCPD) 1.30 Indan 8.61 Indene 6.37 1,2-Diethyl Benzene 2.12 1-Methyl 2-Propyl Benzene 1.06 1,3 Diethyl Benzene 1.65 1,3 Dimethyl 5-Ethyl Benzene 2.24 Cymenes 0.47 1-Methyl 4-n-Propyl Benzene 0.35 Dihydro-DCPD 12.64 Tetrahydro-DCPD 5.37



1,4 Dimethyl 2-Ethyl Benzene 1.89 3,5 Dimethyl Styrene 0.35 1,2 Dimethyl 4-Ethyl Benzene 2.71 Methyl-Indenes 3.20 Naphthalene 0.05 C10 Aromatics 2.91 Total 100.00

1.3.12 Pyrolysis Fuel Oil (PGO + PFO Mixed Stream)

Sr. No.

Property Specification Unit Test Method

1. Acidity Nil mg KOH/gm

ASTM D1093

2. Ash 0.1 max. % mass 3. Density @ 15°C To report Gm/cm3 ASTM D287

or ASTM D4052

4. Flash Point 70 min. °C ASTM E502 5. Pour Point Summer: +15

max. Winter: +21 max.

°C ASTM D97

6. Viscosity at 90°C 35 cP ASTM D445 7. Water Content 1 max. % vol 8. Sediment 0.25% max. % mass

Expected BMCI for pyrolysis fuel oil is 110-130. Note: PGO stream is essentially a 204-288°C cut with flash point of 81°C minimum.

1.3.13 Hydrogenated C4 Product for LPG Pool

(i) Raffinate Case

Expected Composition, wt% Case 2 Propane 0.17 Isobutane 52.37 Normal Butane 44.09 Isobutylene 1.90 1-Butene 0.05 t-2-Butene 0.65 c-2-Butene 0.40 Butadiene Nil Pentanes 0.37



Specification Olefins 5.0 wt% max

(ii) Raw C4 Mix Hydrogenation (Butadiene Unit not operating)

Expected Composition, wt% Case 2 Propane 0.05 Propylene 0.13 Isobutane 18.91 Isobutylene 19.22 Normal Butane 22.58 1-Butene 2.61 t-2-Butene 24.57 c-2-Butene 11.62 1,3-Butadiene 0.01 Pentanes 0.06 Pentenes 0.24 Specification Butadiene 0.1 wt% max

1.4 Utility Specification

Utility specifications are provided on the following pages. With the exception of steam and fuel gas, which are generated ISBL, all other utilities are imported from the central utility center.

1.4.1 Steam The following table summarizes the steam header conditions and steam conditions for the equipment design.



Table 1.4.1.1 Steam Header Conditions and Steam Conditions for Equipment Design

Steam Conditions for Equipment Specification

Generation Steam Header

Conditions

Header Conditions

for Equipment Design (1)

At Turbine

Inlet

At Turbine

Extraction (2)

At Heat Exchanger

Inlet Mechanical

Design Min Normal Max Super High Pressure (SHP)

Pressure, kg//cm2g 122.5 122.0 120 119 Temperature, °C 525 520 515 515 High Pressure (HP) Pressure, kg//cm2g 40.0 42.0 44.0 40.0 39.0 41.0 38.0 50/FV Temperature, °C 380 390 400 380 375 375 427 Medium Pressure (MP) Pressure, kg//cm2g 15.1 16.3 17.7 15.1 14.1 15.6 13.6 20.4/FV Temperature, °C 270 285 310 270 265 265 350 Low Pressure (LP) Pressure, kg//cm2g 3.6 4.0 4.5 3.6 4.1 2.7 6.5/FV Temperature, °C 195 200 225 195 190 288 Low Pressure Desuperheated (LPS)

Pressure, kg//cm2g 2.7 6.5/FV Temperature, °C 148.2 288



Notes: 1. Minimum steam header conditions are used for design of

equipment (exchangers and turbines) with suitable pressure drops in lines to the users.

2. Turbine extraction pressure is consistent with minimum header conditions given for equipment design.

3. Steam condensing pressure at turbine exhaust shall be considered as 0.12 kg/cm2a.

4. Import steam from OSBL auxiliary boilers will be available at SHP level. SHP steam import is required for the following cases: − For start-up − During turndown operation − To purge the cracking heater coils during emergency

1.4.2 Cooling Water

Normal Design Supply Pressure, kg/cm2g 5.0 10.0 Supply Temperature, °C 33.0 max. 65 Return Pressure, kg/cm2g 2.2 Return Temperature, °C 45 max.

Note:

1. Design fouling factor for cooling water: 0.0004 h-m2-°C/kcal 2. Cleanliness factor for surface condenser: 0.80

1.4.3 Instrument Air

• Quality oil and dust free • Dew point -40°C at atmospheric pressure • Conditions

Normal Mechanical Design Pressure, kg/cm2g 7.0

6.5 (min) 10.0

Temperature, °C 40 65

1.4.4 Plant Air

• Quality oil and dust free • Conditions

Normal Mechanical Design Pressure, kg/cm2g 6.0 (min) 10.0 Temperature, °C 40 65



1.4.5 Nitrogen

• Quality Parameter Unit Specification Purity vol% (min) 99.99 Oxygen vol ppm 3.0 Carbon dioxide vol ppm 3.0 Carbon monoxide vol ppm Traces Oil Nil Dew Point °C at

atmospheric pressure

(-) 100°C

• Conditions

Normal Mechanical Design Pressure, kg/cm2g 6.0 10.5 Temperature, °C 40 65.0

1.4.6 DM Water (makeup)

Quality suitable for SHP steam generation.

Parameter Units Specification pH @ 25°C 6-8 Conductivity Micro mho/cm 0.2 max. Silica (as SiO2) ppb wt 20 max. Iron (as Fe) ppb wt 10 max. Total copper (as Cu) ppb wt 3 max. Total dissolved solids ppb wt 100 max. Oily matter ppb wt 200 max. Chlorides ppb wt 10 max. Total CO2 ppb wt None detectable Total hardness ppb wt None detectable Bacteria (coliform) None Turbidity NTU None Color Colorless

Conditions at Deaerator

Normal Pressure, kg/cm2g 0.75 Temperature, °C 115.6



1.4.7 Electricity Power for electric drives and lighting shall be:

1. 6600V ± 6%, 3 phase, 50 Hz ± 3% resistance grounded for drives of 160 kW.

2. 415V ± 10%, 3 phase, 50 Hz ± 3% for drives up to 160 kW, neutral is solidly earthed.

3. For instruments and lighting the voltage shall be 240V ± 6%, 50 Hz ± 3%, single phase AC grounded.

4. UPS system shall be 110V AC. 1.4.8 Fuel Gas

Heaters operate using fuel gas only. The primary supply of fuel gas is methane rich offgas produced in the naphtha cracker unit. During normal operation, methane rich offgas produced in NCU is not sufficient to meet the heaters firing demands. During start-up or when there is a deficit, fuel gas will be supplied by vaporizing makeup/backup fuel. Provision to use RLNG as makeup/backup fuel will be kept in the design. (a) Primary fuel gas to heaters

Component, mol% Case 1 Case 2 Hydrogen 14.20 14.32 Methane 84.95 84.85 Total C2s 0.42 0.41 Carbon monoxide 0.43 0.42 100.00 100.00

(b) Makeup/backup fuel

LPG from refinery and/or RLNG vapors will be makeup/backup fuel with the following composition: Composition: (i) LPG

Composition, wt% Ethane 1.11 Propane 40.00 Butanes 58.55 C5+ 0.34 Water <5 wt. ppm



Note: Hydrogenated C4 product from C4 Hydrogenation unit

may also be used as makeup/back-up fuel. (ii) RLNG Vapor

Component Specification Range, mol % Min. Methane Max. Methane Methane 84.50 98.77 Ethane 9.00 0.69 Propane 3.00 0.03 Butane 2.00 0 Pentane and Heavier 0.25 0 Nitrogen 1.25 0.51 100.00 100.00

Other specifications: Cross heating value of gas = approximately 8,900 - 9.900 kcal/SCM Total non hydrocarbons = 2.0 mol% (max) Total sulfur including H2S = 10 wt ppm max Expected H2S = 4 vol ppm max Carbon dioxide = 0 Water content = 112 kg/MM SCM max. Gas shall be reasonably free from dust (max size 5 microns), gum forming constituents and other deleterious solid and/or liquid matter which will cause damage to or interfere with the operation of transport facilities.



1.5 Site Location and Climatic Conditions The naphtha cracker and associated units will be located at Panipat, in the state of Haryana, India.

1.5.1 Meteorological Data

Sl# Parameter Minimum Normal/ Average

Maximum/ Design

(A) METEOROLOGICAL DATA 1 Elevation above mean sea level, m 238 m 2 Barametric pressure, mbar 967.3 978.675 988.4 3 Ambient temperature, °C tmin = (-) 0.7 tnor = tmax = 46.6 4 Relative humidity, % @ tmin 88 @ tnor 95 @ tmax

5 Rainfall data: (a) for 1-hour period 72 mm (b) for 24-hour period 218 mm

6 Wind data (a) Wind velocity at a height of 30m 168 km/h (b) Wind velocity at a height of 10m 152 km/h (c) Wind direction: i) Morning SE to NW ii) Evening NW to SE

(B) DATA FOR EQUIPMENT DESIGN 1 Design dry bulb temperature, °C 39 2 Design wet bulb temperature, °C 27.5 3 Low ambient temperature for MDMT, °C (-) 0.7 4 Design air temperature for air cooled exchangers, °C 45 5 Preferred process temperatures break point for air cooled

exchangers where followed by water cooling, °C 65

6 Preferred process temperatures break point for air cooled exchangers where not followed by water cooling, °C

55

7 Coincident temperature and relative humidity for Air Blower/Air Compressor design.

(a) 40°C (b) 80%*

* Maximum relative humidity is 95% and normal is 88%.

1.5.2 Earthquake Seismic load shall be calculated by Indian Codes and Standards. Panipat lies in the seismic zone IV as per IS1893 (to be confirmed).

1.5.3 Special Consideration Provision for winterization is not required.



1.6 Battery Limit Conditions Pressure

kg/cm2g Temperature,

°C Phase

Feedstocks Naphtha 12.0 min Amb. (5 min) Liquid Hydrogenated C4 Recycle from C4

Hydrogenation Unit 12.0 43 Liquid

Hydrogenated C5 Recycle from PGHU

12.0 43 Liquid

C6 Recycle from Benzene Extraction Unit

12.0 40 Liquid

Downstream Plant Recycle C3 Recycle from PP (Note 5) 20.6 Liquid Light Gas Recycle from PP (Note 5) 8.0 40 Vapor Purge Gas from LLDPE/HDPE (swing

plant) 1.8 40 Vapor

Ethylene Boil-off gas 1.5 -20.0 Vapor Recycle from PP plant (based on

selected technology) 2.0 – 3.0 10 – 15 Vapor

Recycle from HDPE (Note 6) Vapor Other Streams FCC Dry Gas Hydrocarbons 24.5 -20 Liquid Raw FCC C3s 27.0 43.5 Offspec Ethylene 14.0 -34.5 Mixed

phase Off-spec. Propylene 27.5 45 Liquid Products Hydrogen Export 26.0 18.6 Vapor Hydrogen Import from Refinery 31.0 40 Vapor Methane Offgas 3.5 (1) 38-40 Vapor HP Ethylene to downstream plant (2) 52.0 30 Vapor MP Ethylene to downstream plant (2) 35.0

(Revised to 28.0)

30 Vapor

Ethylene Product to LP storage (3) 1.95 -98.0 Liquid Polymer Grade Propylene 28.0 38 Liquid Propylene to Atmospheric Storage (4) 4.0 -45 Liquid Raw Mix C4s to storage 6.0 41.4 Liquid Hydrogenated C4 to LPG pool 20.0 43.0 Liquid Raw Mix C4 from storage to C4 Hydro 7.0 41.0



Pressure kg/cm2g

Temperature, °C

Phase

Raw Pyrolysis Gasoline to storage/PGHU

4.0 40 Liquid

C6-C8 Heart Cut to storage 4.5 40 Liquid C6-C8 Heart Cut to BEU 5.0 89 Liquid C9+ Product to storage 4.0 43 Liquid Benzene Product 4.0 40 Liquid C7-C8 Cut to storage 4.0 40 Liquid Fuel Oil 5.0 90 Liquid Miscellaneous Relief to flare 0.1 normal

1.5 max. 1.7 build-up at PSV inlet

Vapor

Pretreated Spent Caustic (gasoline wash) to OSBL

2.5 74 Liquid

50% Caustic from OSBL 4.5 40 Dilution Steam Drum Blowdown to

OSBL 3.2 40 Liquid

Fuel Gas Export 3.2 37.0 Vapor Methane Ballast Gas 35.0 41.0 Vapor Offspec Ethylene Vapor from HP

Storage 10.5 5 Vapor

Liquid Ethylene Recycle from atmospheric storage

18.0 -72.0 Liquid

Liquid Propylene Recycle from atmospheric storage

24.5 -35.0 Liquid

RLNG Fuel Ambient Vapor Notes:

1. At fuel gas knock out drum. 2. Ethylene is produced and delivered as vapor product to downstream

plants at two pressure levels.

Ethylene Vapor

Product Flow Rate kg/h User BL Pressure, kg/cm2g Case 1 Case 2 - HDPE and PP Plant 52.0 40,650 40,650 - Swing Polyethylene

and MEG Plants 35.0 Balance Balance

In addition 1,200 kg/h of ethylene product, equivalent to boil-off gas, is sent as liquid ethylene to low pressure storage.



Subsequently, flow rates and BL pressures for ethylene product have been revised as follows: BL Pressure, kg/cm2g Flow Rate, kg/h HP Ethylene 52.0 Normal: 46,000

Max.: 52,000 MP Ethylene 28.0 Balance

(Case 2: 59,976) Revised flow rates and pressure will be used for specification of product pumps and exchangers only. Material balances will not be revised to reflect this change.

3. Up to 240,000 TPA (30,000 kg/h) ethylene may be delivered to atmospheric storage at these conditions. The flow rate is 30% of the base ethylene capacity of 800,000 TPA.

4. Up to 151,450 TPA (18,930 kg/h) propylene may be delivered to

atmospheric storage at these conditions. The flow rate is 30% of the base capacity of 504,834 TPA.

5. Based on selected technology, these steams do not exist.

6. Connected to charge gas compressor first stage suction.

7. All pressures are at grade unless otherwise noted.

1.7 Environmental Specifications This section describes the atmospheric emissions, liquid effluents and solids that will be generated by the naphtha cracker unit. The quantity and characteristics for these environmental streams are summarized in Section 6.3. 1.7.1 Atmospheric Emissions The following gaseous emissions will be produced by the Naphtha Cracker Plant:

• Cracking Heater Flue Gas during Normal Operation • Gas emissions during High Steam Standby/Decoking • Acetylene Converter Regeneration Offgas • MAPD Converter Regeneration Offgas • C4 Hydrogenation Reactor Regeneration Offgas • First Stage DPG Reactor Regeneration Offgas • Second Stage DPG Reactor Regeneration Offgas



1.7.2 Liquid Effluents The following aqueous effluent streams will be produced by the Naphtha Cracker plant:

• Dilution Steam Drum Blowdown • Spent Wash Water from the Caustic Water Wash Tower and

Pretreated (Gasoline Wash) Spent Caustic • Continuous Blowdown from the SHP Steam Drums • Intermittent Blowdown from the SHP Steam Drums and Transfer Line

Exchangers • TLE Hydrojetting Water • Polymeric Oil from the Caustic Water Wash Tower • Slop Oils • Rerun Tower Ejector Condensate

1.7.3 Solids Periodically, the desiccants, absorbents, catalyst, as well as various filter media must be changed. In addition, the following sources of material will be produced by the plant:

• Coke from TLE Hydrojetting • Coke from Quench Oil Pump Suction and Discharge Filters • Coke from Pyrolysis Fuel Oil Stripper Feed and Product Suction Filters

1.7.4 Flue Gas Specification Gaseous emissions will be in the following stipulations.

Particular Matter 150 mg/Nm3 max CO 500 ppm vol max NOx 50 ppmv at 3% O2 (dry) in emissions SOx To be reported

1.7.5 Noise Level Maximum noise level generated by any equipment shall be less than 85 dB(A) at a distance of 1 meter. 1.7.6 Tolerance Limit Tolerance limit for effluent discharged into inland surface water – refer to Table 1.7.6.1.

1



Table 1.7.6.1 Tolerance limits for effluent to be discharged into inland surface water (most stringent between general effluent standards for discharge of environmental pollutants (Part – A) and MINAS). SR No.

Characteristics Unit Requirements (Tolerance Limits)

1 Color and Odor See Note 1 2 Suspended solids Mg/lit max. 100 (20 ppm for MINAS) 3 Particle size for suspended

solids Shall pass 850 micron IS

sieve 4 pH value 5.5 to 9.0 (6.0 to 8.5 for

MINAS) 5 Temperature Deg C max. Shall not exceed 5 deg C

above the receiving water temp.

6 Oil and Grease Mg/lit max. 10 7 Total residual chloride Mg/lit max. 1 8 Ammoniacal Nitrogen (as N) Mg/lit max. 50 9 Total Kjeldahl Nitrogen (as N) Mg/lit max. 100

10 Free Ammonia (as NH3) Mg/lit max. 5 11 Biochemical oxygen demand

(5 days at 20 deg C) Mg/lit max. 30 (15 ppm as per MINAS)

12 Chemical oxygen demand Mg/lit max. 250 13 Arsenic Mg/lit max. 0.2 14 Mercury (as Hg) Mg/lit max. 0.01 15 Lead (as Pb) Mg/lit max. 0.1 16 Cadmium (as Cd) Mg/lit max. 2 17 Hexacalent chromium (as

Cr+6) Mg/lit max. 0.1

18 Total chromium (as Cr) Mg/lit max. 2 19 Copper (as Cu) Mg/lit max. 3 20 Zinc (as Zn) Mg/lit max. 5 21 Selenium (as Se) Mg/lit max. 0.05 22 Nickel (as Ni) Mg/lit max. 3 23 Cyanide (as CN) Mg/lit max. 0.2 24 Chloride (as Cl) Mg/lit max. 1000 25 Fluoride Mg/lit max. 2 26 Dissolved phosphate (as P) Mg/lit max. 5 27 Sulfide (as S) Mg/lit max. 2 (as per MINAS 0.5 ppm

max) 28 Pesticides Absent 29 Phenolic compounds (as

C6H5OH) Mg/lit max. 1

30 Radioactive material (a) Alpha emitters Micro curie/ml max. 1 x 10-7 (b) Beta emitters Micro curie/ml max. 1 x 10-6

Note 1: All efforts should be made to remove color and unpleasant odor as far as

practicable.



2.0 BASIS FOR PROCESS PERFORMANCE 2.1 General The Naphtha Cracker Plant design shall be flexible to accommodate normally anticipated fouling of equipment and catalyst deactivation. The stated process performance is based on all equipment in the clean condition, with compressors operating at or near the efficiencies stated in section 2.8. The exception to this is the transferline exchangers which operate on relatively short cycles and are assumed to be operating in the average fouled condition. Hydrogenation reactors are assumed operating with average selectivities as below: Converter: Acetylene Selectivity to Ethylene: 50% Converter: MAPD 30% Selectivity to Propylene: 2.2 0verall Material Balance When cracking the design feedstocks having the specifications given in Section 1.0, the Plant shall be designed to produce the following terminal material balances for the design cases. Balances are based on 8000 operation hours per year.



Table 2.2.1

Material Balance (includes Naphtha Cracker Unit, C4 Hydrogenation Unit and Pyrolysis

Gasoline Hydrogenation Unit) Case 1 Case 2 Feed, kg/h Naphtha 293,201 293,201 BD Raffinate to C4 Hydro 42,608 42,867 C6 Raffinate from GTC 14,306 14,319 Treated FCC C3s 12,412 Treated FCC Dry Gas 2,485 (1) BOG from ethylene storage 1,200 1,200 C3 recycle from PP plant 8,850 8,850 Light gas from PP plant 2,587 2,587 Purge gas from PE plant 1,750 1,750 BD unit offgas 101 101 Net steam reacted 655 662

Total 365,258 380,434 Products, kg/h Hydrogen 506 584 Methane Offgas 50,220 51,083 Polymer Grade Ethylene 103,829 107,176 Polymer Grade Propylene 73,044 83,318 Raw C4s 59,847 60,256 C6-C8 cut 59,650 59,795 C9+ cut 7,627 7,591 Pyrolysis Fuel Oil 10,432 10,629 Water from Methanator 123 124 Acid Gases 424 428

Total 365,702 380,984

Material Balance Deviation, kg/h 444 550 Recycles, kg/h Ethane 21,407 22,797 Propane 5,471 8,314 C5s 19,560 19,782

Note:

(1) The flow rate excludes 1,988 kg of C3 wash stream.



Table 2.2.2

Material Balance (including Naphtha Cracker, C4 Hydrogenation Unit, Pyrolysis Gasoline

Hydrogenation, Butadiene and Benzene Units) Case 1 Case 2 Feed, kg/h Naphtha 293,201 293,201 Raw FCC C3s -- 12,412 Treated FCC Dry Gas -- 2,485 (2) BOG from ethylene storage 1,200 1,200 C3 recycle from PP plant 8,850 8,850 Light gas from PP plant 2,580 2,587 Purge gas from PE plant 1,750 1,750 Net steam reacted 655 662

Total 308,236 323,147 Products, kg/h Hydrogen 506 584 Methane Offgas 50,220 51,083 Polymer Grade Ethylene 103,829 107,176 Polymer Grade Propylene 73,044 83,318 Butadiene 17,020 17,163 Benzene (1) 14,491 14,589 C7-C8 cut (1) 29,422 29,456 C9+ cut 7,627 7,591 Pyrolysis Fuel Oil 10,432 10,629 Water from Methanator 123 124 Acid Gases 424 428 Waste HC Stream from

Butadiene Unit 122 127 Total 307,260 322,268

Material Balance Deviation, kg/h 976 879 Recycles, kg/h Ethane 21,407 22,797 Propane 5,471 8,314 C4s 42,919 43,177 C5s 19,560 19,782 C6s 14,306 14,319

Note:

(1) Products from Benzene Extraction Unit are adjusted based on final feed rate and composition of treated C6-C8 stream from PGHU shown in Table 2.2.1.

(2) The flow rate excludes 1,988 kg of C3 wash stream.



(3) Case 1 material balance for Butadiene and Benzene Extraction units is based on terminal balance using Case 2 material balance information.

2.3 Once-Through Cracking Yields The once-through cracking yields for all feedstocks plus ethane and propane recycle are presented in tabular form in this section. Please note that these yields are based upon a cracking selectivity level determined by an SRT radiant coil, a coil outlet pressure of 1.76 kg/cm2a and a steam-to-hydrocarbon ratio of 0.50 for Naphtha/C4/C5/C6 and 0.3 for ethane and propane recycle.

IOCL Naphtha Cracker - Once-Thru Yields

Feedstock Naphtha

(1) C4 Recy

(1) C5 Recy

(1) C6 Recy

(1) Naphtha/ C4/C5/C6

Recycle Ethane Stream

(2)

Recycle Propane Stream

(2)

Yields, wt%

Hydrogen 0.75 0.90 0.76 0.80 0.77 3.85 1.48 Methane 12.20 12.75 14.62 15.69 12.53 4.42 25.24 Acetylene 0.39 0.22 0.26 0.59 0.37 0.42 0.55 Ethylene 25.10 14.12 18.63 30.21 23.66 52.04 36.87 Ethane 3.69 1.84 2.71 3.42 3.41 34.65 3.99 Propadiene 0.52 0.90 1.07 1.21 0.62 0.06 0.48 Propylene 15.92 20.59 18.97 17.56 16.69 1.15 12.33 Propane 0.37 0.28 0.50 0.41 0.37 0.17 5.63 Butadienes 5.07 1.23 4.21 5.79 4.60 1.31 4.11 Butenes 5.54 13.43 15.14 6.55 7.02 0.19 0.82 Butanes 0.67 30.91 0.09 0.11 4.14 0.23 0.07 C5's 4.97 1.04 19.99 1.37 5.20 0.30 1.66 C6-C8 NA 6.40 0.50 0.86 6.85 5.43 0.38 0.31 Benzene 4.63 0.61 0.64 2.77 3.87 0.57 3.12 Toluene 4.62 0.28 0.48 0.89 3.74 0.11 0.67 Xylenes + EB 2.61 0.05 0.06 0.20 2.08 0.01 0.10 Styrene 0.73 0.08 0.09 0.16 0.60 0.00 0.52 C9-200oC 2.69 0.09 0.74 1.19 2.23 0.04 1.12 PGO/PFO 3.13 0.18 0.18 4.23 2.67 0.10 0.93 Total 100.00 100.00 100.00 100.00 100.00 100.00 100.00 Notes:

(1) Naphtha and hydrogenated C4, C5 and C6 are co-cracked in SRT-VI heaters. This yield data is for the Naphtha, C4, C5 and C6 part of the blend.



(2) Recycle ethane and propane are co-cracked. Ethane conversion is 65% and corresponding propane conversion is 94.23%. Once through yield data is given for recycle streams.



2.4 Utility Consumption Utility consumptions for Case 1 and 2 are provided in Section 3 of the Basic Engineering Package. The basis for the balances are indicated in each utility section. The utility consumptions are summarized in Table 2.4.1 below. The consumption includes all the process units (Naphtha Cracker, C4 Hydro, Pyrolysis Gasoline Hydrogenation, Butadiene and Benzene units) and the ISBL steam/condensate system including deaerater/BFW pumps.

Table 2.4.1

Utility Consumption Utility Unit Normal Hourly Quantity Case 1 Case 2 Fuel gas MMKcal 629.7 638.6 Steam (SHP) T 0 0 Steam (HP) T (-) 9.20 (-) 8.9 Steam (MP) T 0 0 Steam (LP) T 7.1 9.2 Electric Power kW 15,317 15,754 Cooling Water m3 44,487 46,590 Note: 1. (-) means export 2.5 Specific Energy Consumption The specific energy consumption is a useful measure to evaluate the overall efficiency of a process. When making this comparison, it is extremely important to understand the basis of the specific energy calculation. In general, two techniques are utilized in the industry. The first approach evaluates steam as equivalent SHP steam since useful work may be derived between the various steam levels. When using this approach the fuel equivalent values are assigned as follows:

(1) For SHP steam, the fuel equivalent value is calculated by taking the enthalpy difference between the blend of turbine and LP condensates feeding the deaerater at 100°C and the SHP steam level of 120.0 kg/cm2g and 520°C. SHP steam is raised in the cracking heater at 93.5% efficiency. The resulting value is 761 kcal/kg.



(2) For HP steam import or export the fuel equivalent value is set by calculating the additional SHP steam that must be produced and let down in the various compressor turbines to produce a fixed quantity of HP steam. This ratio is 0.736 kg of SHP steam per kg of HP steam. Hence, the fuel equivalent value is 560 kcal/kg. In a similar fashion the fuel equivalent value of MP steam is 419 kcal/kg and LP steam is 302 kcal/kg.

(3) For electricity the fuel equivalent value is 2.46 MMkcal per MW of

electric power based on a 35% power plant efficiency. The second approach is to value the export steam at its enthalpy value. This approach will result in the reporting of lower specific energy consumptions. Using this technique the following fuel equivalent values may be assigned:

kcal/kg SHP Steam 122.0 kg/cm2g & 520°C 812.4 HP Steam 40.0 kg/cm2g & 380°C 756.0 MP Steam 15.1 kg/cm2g & 270°C 708.5 LP Steam 3.6 kg/cm2g & 195°C 679.9

Values have been calculated both ways to show the relative impact. The values are presented in the Table 2.5.1. The basis for the specific energy calculation is the normal operating fuel gas, steam, and electric power at 800,000 TPA plant capacity for all ISBL process units (Naphtha Cracker, C4 Hydro, PGHU, BEU and BZEU) and deaerater and BFW pumps. .



Table 2.5.1

Specific Energy Consumption

Normal Operation

Approach 1 Approach 2 Case 1 Case 2 Case 1 Case 2 Fuel Fired Cracking Heaters, MM Kcal/h 629.7 638.6 629.7 638.6 HP Steam Export T/h 9.20 8.90 9.20 8.90 MM Kcal/h 5.15 4.98 6.96 6.73 LP Steam Import T/h 7.1 9.2 7.1 9.2 MM Kcal/h 2.14 2.78 4.83 6.26 Electrical kW 15,317 15,754 15,317 15,754 MM Kcal/h 37.68 38.75 37.68 38.75 Specific Energy MM Kcal/h 664.37 675.15 665.25 676.88 kcal/kg of ethylene 6,398 6,299 6,407 6,316 kcal/kg of ethylene + propylene 3,756 3,544 3,761 3,553 2.6 Catalyst and Chemical Consumption The Catalyst and Chemical Consumptions for Case 2 at normal operation of the plant at the 800,000 TPA Plant capacity are summarized below. 2.6.1 Desiccant/Adsorbents

• Charge Gas Dryers Volume Basis: 24 hour cycle + 6 hour guard bed. Type: UOP molecular sieve Type: 3A-EPG

compound bed (1/4” TRISIV + 1/16” pellets)

Volume: 114.2 m3 per dryer, 228.4 m3 for two dryers Weight: Approx. 142,885 kg for 2 dryers Estimated Life: Projected 5 years



• Condensate Stripper Bottoms Dryers Volume Basis: 48 hour cycle Type: UOP Molecular Sieve Type: 3A-EPG

1/16” pellets Volume: 47.1 m3 per dryer, 94.2 m3 for two dryers Weight: Approx. 60,420 kg for 2 dryers Estimated Life: Projected 5 years

• Hydrogen Dryers Volume Basis: 48 hour cycle + 6 hour below moisture

probe Type: UOP Molecular Sieve Type: 3A-EPG

1/16” pellets Volume: 16.5 m3 per dryer, 33.0 m3 for two dryers Weight: Approx. 20,680 kg for 2 dryers Estimated Life: Projected 5 years

• Ethylene Dryer Volume Basis: 48 hour cycle Type: UOP Molecular Sieve Type: 3A-EPG 1/8” pellets Volume: 15 m3 in single dryer Weight: 9,660 kg for single dryer Estimated Life: Project life 5 years

• FCC C3’s Dryers/Treaters

Volume Basis: 72 hour cycle for first two years of operation 48 hour cycle for third year of operation 24 hour cycle at end of five years Type: UOP Adsorbent Type: D-201, 3A-EPG AZ-300 compound bed Volume: 41.8 m3 for one dryer/treater, 83.6 m3 for two

Weight: 53,070 kgs for two dryer/treaters (4626 kg of D-201; 24222 kg of type 3A-EPG; 24222 kg of Type AZ-300)

Estimated Life: Guaranteed 3 years, projected life 5 years with reduced cycle time



FCC C3’s Arsine Treater Volume Basis: The absorbent quantity determined by

minimum contact time Type: UOP Adsorbent Type: GB-238 Volume: 4.6 m3 in single treater Weight: 3265 kg in single treater Estimated Life: Projected 5 years, guaranteed 3 years

2.6.2 Catalyst

• Acetylene Hydrogenation Volume Basis: To meet 1 mol ppm acetylene spec. in

Ethylene Product. Type. Sud-Chemie Type: G-58C silver promoted palladium catalyst Cycle Length: 5-9 months between regeneration Volume: 52.8 m3 per reactor, 105.6 m3 for two

reactors Weight: 76,800 kg for two reactors Estimated Life Typically 8-12 years, 5 years guaranteed

• Methanator Type: Sud-Chemie Type C13-4-04 nickel catalyst ¼” – 1/8” spheres. Volume: 7.8 m3 in single reactor Weight: 7,020 kg Estimated Life: Guaranteed 5 years

• MAPD Hydrogenation

Volume Basis: To achieve less than 200 mol ppm MAPD in reactor effluent Type: Sud-Chemie Type: G-68HX

1.3 mm CDS Extrusions No. Reactors Req'd. 2 Volume: 15.8 m3 per reactor, 31.6 m3 Total Weight: 19,000 kg for two reactors Expected Cycle Length: Over one year Expected Catalyst Life: Expected over 5 years, guaranteed 3 years



• C4 Hydrogenation Volume Basis: Partially hydrogenate mixed C4’s

butadiene saturation only) when BD unit is not in operation. Butadiene spec. for C4 product is 1,000 wt. ppm max.

No. Reactors Req'd. 1 Volume: Combination of Sud-Chemie G-68HX selective hydrogenation palladium catalyst and T-2464B olefin saturation, palladium catalyst. G-68HX: 13.7 m3, 1.6 mm CDS extrusion T-2464B: 13.7 m3, 1.6 mm CDS extrusion Weight G-68HX – 8,200 kg T-2464B – 6,165 kg Total = 14,385 kg Estimated Cycle Time: 1 year between regeneration, (over 2 years for C4 raffinate case) Estimated Life: 3 years guaranteed

• DPG - Stage I

Volume Basis: To meet C5-C8 Product Spec. No. reactors req'd. 1 Type DPG-1 selective hydrogenation palladium catalyst. The reactor consist of two beds in a single vessel. Volume: 24.6 m3 per bed, 49.2 m3 total Weight 35,424 kg total Estimated cycle time 1 year Estimated Life: Expected 5 years, 3 years guaranteed



• DPG - Stage II Volume Basis: To meet depentanizer product spec. (C5 cut and C6-C8 Product Spec.) No. Reactors Req'd. 1 Type - Sud-Chemie Type OPGH II, Ni-Mo

2,5 mm CDS olefin saturation catalyst - Sud-Chemie Type OPGH II, CO-Mo

1.5 mm CDS hydrodesulphurization catalyst

Volume: OPGH III 37.3 m3 OPGH II 40.4 m3 Weight OPGH III – 22,380 kg + OPGH II - 24,240 kg Total = 46,620 kg Estimated cycle time 12 months after initial start-up Estimated Life: Estimated life is over 5 years, guaranteed 3 years 2.6.3 Chemicals Caustic Soda Caustic/Water Wash Tower Usage: 1000 kg/h as 100% NaOH Concentration: 20 wt.% Quality: Rayon grade Methanol Type: Commercial Grade 99.85% purity Inventory: 20 m3 Consumption: Intermittent



Injection Sulfur Type: DMDS – Technical Grade (68% by weight sulfur) Usage: 100 ppm by wt. sulfur injection in ethane/ propane recycle and naphtha/C4/C5/C6 feed to heaters Consumption: 75 kg/day Note: Above consumption is for ethane/propane recycle feed only. Naphtha/C4/C5/C6

feeds do not require sulfur injection as naphtha feed contains 500 ppm sulfur. Chemical Injection Agents (a) Corrosion Inhibitor for Dilution Steam System

Chemical Dosage Consumption Caustic Soda 4 ppm (as 100% NaOH basis)

to process water/quench water 25 kgs (100% basis) per day

Note: Normal injection point of caustic is to the dilution steam drum feed. Dilute caustic soda of approximately 2 wt.% solution is injected. (b) Polymerization Inhibitor Type: Nalco Actrene EC3214A Usage: 30-50 ppm by wt. dosage of neat process Antifoulant Nalco Actrene EC3214A.

Injection Point

Dosage ppmw

Liter/h max. as 100%

Inhibitor − Condensate Stripper Reboiler 30 3.29 − Condensate Stripper Bottoms 50 5.91 − Recycle from C2 Green Oil Drum or

Deethanizer Reboilers 30 3.10

− Depropanizer Feed from Deethanizer − Bottoms

50 5.91

− Depropanizer Reboilers 30 3.55



(c) Mixed C4 Product Inhibitor

Type: Tertiary Butyl Catechol (TBC) at a Concentratiaon of 10 wt.% in toluene solution. Usage: 50 wt. ppm TBC into mixed C4 product when sent to storage. Consumption: Normal Nil Max, 75 kg/day as 100% TBC

(d) DPG Antioxidant Injection Type: Nalco Energy Services EC3199A, diluted to 50 vol. % antioxidant agent. Treated C6-C8 cut from the depentanizer or C7-C8 cut from the benzene extraction unit can be used as diluent. Usage: - 75 wt. ppm (100% antioxidant) injection into raw pyrolysis gasoline when set to storage 60 wt. ppm (100% antioxidant) injection into - C9 plus product to storage. Consumption: Pyrolysis Gasoline: Normal: 0 Max: 7 kg/h as 100% antioxidant C9 Plus Product: 11 kg/day as 100% antioxidant

(e) Depentanizer Corrosion Inhibitor Type: Nalco Energy Services EC1046A. The

inhibitor is diluted by a small slip-stream of reflux from Depentanizer reflux pumps (11-P-5506A/B) and atomized through a spray nozzle into the overhead vapor just upstream of the Depentanizer condenser.

Usage: 8-15 wt. ppm of 100% corrosion inhibitor Consumption: 13 to 25 kg/day as 100% inhibitor

(f) DPG Presulfiding Agent

Type: DMDS – Technical Grade (68% by wt. sulfur) Usage: As required for presulfiding DPG second stage catalyst. Consumption: 7990 kgs as DMDS



(g) Boiler Feed Water Chemical Treatment Amine, phosphate and oxygen scavengers are injected in BFW for SHP Steam Drum Water quality control. The type of chemical and consumption levels will be provided in the FEED package. (h) Chemicals for Condensate Treatment Caustic soda and acid are required for regeneration of cation and mixed bed polishers. Consumption rates will be provided in the FEED package.

• Compressor Wash Oil and Flux Oil

Source: Partially hydrogenated C9 plus from 1st Stage DPG is used as wash oil in normal operation. Light cycle oil is available as back-up wash oil from adjacent refinery. Flux oil is required during start-up. Imported wash oil or flux oil properties are as follows: Type: Light Cycle oil or equivalent Boiling Range, oC 230 to 340 Sulfur Content, wt.%: 2.5 max. Viscosity @100oC cP: 20 – 30 max. SG @15.6 oC: 0.95 Consumption: 2.5 t/h (500 kg/h/stage)

• Purge Oil

Source: Filtered pyrolysis gas oil from gasoline fractionator Usage: As required for instrument purging in quench oil system. Consumption: 12.0 t/h as PGO from gasoline fractionator Note: Purge oil after instrument flush ends up in gasoline fractionator. Therefore, there is no consumption.

• Propylene Refrigerant

Usage: As required for initial inventory/make-up in propylene refrigeration system. Type: 99.5 wt.% Propylene Water Dew Point – 45oC Inventory: As calculated during detailed engineering



• Ethylene Refrigerant

Usage: As required for initial inventory/make-up in ethylene Refrigeration system. Type: 99.5 mol.% Ethylene Water Dew Point – 105oC Inventory: As calculated during detailed engineering

2.7 Pyrolysis Heater Performance Normally, there will be five operating liquid SRT IV cracking heaters and one SRT III recycle cracking heater. The heater guaranteed run lengths is presented below.

Feedstock Run Length Naphtha + Hydrogenated C4/C5/C6 60 days Ethane/Propane Recycle 60 days Note: The run length of the heater shall be terminated when the maximum tube

wall metal temperature exceeds 1125°C (corrected for tube reflectivity) or the average tube wall temperature reaches 1115°C (corrected for tube reflectivity).

2.8 Assumed Efficiencies The balances provided in this section are based on the following major equipment efficiencies: · Charge Gas Compressor 1st Stage 0.866 2nd Stage 0.871 3rd Stage 0.867 4th Stage 0.866 5th Stage 0.869 Turbine Driver Extract./Cond. 0.77/0.75 · Propylene Refrigeration Compressor 1st Stage 0.849 2nd Stage 0.879 3rd Stage 0.847 4th Stage 0.871



Turbine Driver Extract./Cond. 0.77/0.75 · Ethylene Refrigeration Compressor 1st Stage 0.807 2nd Stage 0.783 3rd Stage 0.790 Turbine Driver Condensing 0.77 · Methane Refrigeration Compressor 1st Stage 0.807 2nd Stage 0.689 Motor 0.88 DPG Recycle Compressor 0.75 Motor 0.88 Pump Motor Turbine EFF EFF EFF Pump BFW Pump 0.8 0.9 0.65 QO Pump 0.85 0.94 0.65 QW Pump 0.84 0.94 0.60

Pumps and motor efficiencies for other process pumps are provided in electrical power consumption tabulations in Section 3, Tab D, of the Basic Engineering Package.

2.9 Heat Leak Factors The propylene and ethylene refrigeration loads are estimated based on the

following heat leak factors:

• For shell and tube or kettle exchangers:

Temperature Level, °C Heat Leak, % 13 0.0 6 0.2

-27 1.2 -40 1.6 -63 2.3 -75 2.7

-101 3.5 -120 4.4 -150 5.0

iocl section 1 - process specfication

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