Final Assignment

ASSIGNMENT NO.1

SUBJECT:

INTRODUCTION TO VISBREAKING

AND

GAS CONCENTRATION UNIT

MAJOR UNIT OPERATIONS

TROUBLESHOOTING

SUBMITTED TO:

MR.HAFIZ IMRAN SHAHZAD

MR.SOHAIL AKHTAR HASHMI

SUBMITTED BY:

AMIR SHABIR

INTERNEE P.NO 441

PAK ARAB REFINERY LIMITED

VISBREAKING UNITVisbreaking is a non-catalytic thermal process that converts atmospheric or vacuum residues via thermal cracking to gas, naphtha, distillates, and visbroken residue. Atmospheric and vacuum residues are typically charged to a visbreaker to reduce fuel oil viscosity and increase distillate yield in the refinery.

PROCESS OBJECTIVES:

The objectives of visbreaking are:

Reduce the viscosity of the feed stream: Typically this is the residue from vaccum distillation unit of srude oil.

Reduce the amount of residual fuel oil produced by a refinery: Residual fuel oil is generally regarded as low value product.

Increase the proportion of middle distillates in the refinery output

ADVANTAGES OF THE VISBREAKING PROCESS ARE:

High sulfur fuel oil reduction Reduces distillate cutter stock requirements, making the distillate available for more

valuable transportation distillate fuels Low cost upgrading

With the exception of the coking process, formation of coke in a petroleum refinery is undesirable because coke fouls equipment and reduces catalyst activity. However, in the coking process, coke is intentionally produced as a byproduct of vacuum residue conversion from low value fuel and asphalt into higher value products.

Vacuum residue is fed to the coker fractionator to remove as much light material as possible. Bottoms from the fractionator are heated in a direct fired furnace to more than 900˚F (480˚C) and discharged into a coke drum where thermal cracking is completed. High velocity and stream injection are used to minimize coke formation in furnace tubes. Coke deposits in the drum and cracked products are sent to the fractionator for recovery. Coke drums typically operate in the 25-50 psi (2-4 bar) range while the fractionator operates at a pressure slightly above atmospheric in the overhead accumulator. Fractionator bottoms are recycled through the furnace to extinction.

There is a tradeoff between furnace temperature and residence time for visbreaking operations.Longer residence time leads to lower furnace outlet temperatures. In general, operations are conducted between 800-930˚F (425-500˚C). Material is quenched with cold gas oil.

There are two types of visbreaking technology that are commercially available:

Coil furnace type Soaker process

In the coil process, conversion is achieved by high temperature cracking for a predetermined, relatively short period of time in the heater.

In the soaker process,which is a low temperature/high residence time process, the majority of conversion occurs in a reaction vessel or soaker drum, where the two-phase heater effluent is held at a lower temperature for a longer period of time.

In a ‘coil’ type operation, charge is fed to the visbreaker heater where it is heated to a high temperature, causing partial vaporization and mild cracking. The heater outlet stream is quenched with gas oil or fractionator bottoms to stop the cracking reaction. The vapor-liquid mixture enters the fractionator to be separated into gas, naphtha, gas oil and visbroken resid (tar). The visbroken bottoms are then blended with lighter materials (cutter stock) to meet fuel oil specifications. The fractionated visbreaker gas oil is often used as the cutter stock.

SOAKER VISBREAKING: In soaker visbreaking, the bulk of the cracking reaction occurs not in the furnace but in a drum located after the furnace called the soaker. Here the oil is held at an elevated temperature for a pre-determined period of time to allow cracking to occur before being quenched. The oil then passes to a fractionator. In soaker visbreaking, lower temperatures are used than in coil visbreaking. The comparatively long duration of the cracking reaction is used instead.

SOAKER VISBREAKING VS COIL VISBREAKING:

From the standpoint of yield, there is little or nothing to choose between the two approaches. However, each offers significant advantages in particular situations:

DE-COKING: The cracking reaction forms petroleum coke as a byproduct. In coil visbreaking, this lays down in the tubes of the furnace and will eventually lead to fouling or blocking of the tubes. The same will occur in the drum of a soaker visbreaker, though the lower temperatures used in the soaker drum lead to fouling at a much slower rate. Coil visbreakers therefore require frequent de-coking. This is quite labour intensive, but can be developed into a routine where tubes are de-coked sequentially without the need to shutdown the visbreaking operation. Soaker drums require far less frequent attention but their being taken out of service normally requires a complete halt to the operation. Which is the more disruptive activity will vary from refinery to refinery.

FUEL ECONOMY: The lower temperatures used in the soaker approach mean that these units use less fuel. In cases where a refinery buys fuel to support process operations, any savings in fuel consumption could be extremely valuable. In such cases, soaker visbreaking may be advantageous.

PLANT DUTY: The visbreaking process unit is designed to process 15,560 BPSD(103.1 m3/hr) of 550oC+

vaccum residue from either 100% Arabian light crude oil (Case 1) or a 70/30 volume percent blend of Upper Zakum And Murban Crude oils (Case 2).

PROCESS DESCRIPTION:The main objective of this process is to break down long chain hydrocarbons to smaller ones by the application of heat hence reducing the viscosity. The process is known as thermal cracking. The Visbreaking Process Unit is designed to produce visbroken bottoms (fuel oil component),

unstabilized naphtha, and sour overhead vapors.The unit consists of the following equipment systems

Three Visbreaker Heaters One Visbreaker Fractionator with a gas oil Stripper Facilities for economical steam generation Facilities for economical feed preheat Facilities for product delivery to battery limits

In this unit long chain hydrocarbons are heated by means of Visbreaking Unit heaters.Heating of material and rapid quenching of furnace outlet,facilitates the thermal cracking and altered the viscosity necessary for further processing.

FEEDS: Vacuum residue fed to a visbreaker can be considered to be composed of the following:

Asphaltenes: large polycyclic molecules that are suspended in the oil in a coloidal form Resins: also polycyclic but of a lower molecular weight than asphaltenes Aromatic hydrocarbons: derivatives of benzene, toluene and xylenes Parafinic hydrocarbons: alkanes

Visbreaking preferentially cracks aliphatic compounds which have relatively low sulphur contents, low density and high viscosity and the effect of their removal can be clearly seen in the change in quality between feed and product.

MAJOR PRODUCTS: The three major products are;

Overhead tail gas Naphtha accumulator stream Combines bottoms

THEORY OF PROCESS:

THEORY OF CRACKING AND VISBREAK PROCESS: When the hydrocarbon is heated and decomposed under thermal cracking condition, it is broken up into two or more free radicals. A portion of the compound disassociates to form free radicals,

C10H22 → C8H17* + C2H5*

The highly reactive radicals do not appear in the thermally cracked product effluent, but dependind upon size and environment:

a) React with other hydrocarbonsb) Decompose to olefins

c) Combine with other radicalsd) React with metal surfaces

In general, small radicals are more stable than larger radicals, and will more readily react with other hydrocarbons by capturing a hydrogen atom,for example:

C2H5 + C6H14 → C2H6 + C6H13*

Large radicals are unstable and decompose to form olefins and smaller radicals, for example:

C6H13* → C5H10 + CH3*

The polymerization and condensation reactions that occur at thermal cracking conditions to give aromatic tars:

xC4H8 + yC4H8 +zC3H6→ multi-aromatic-ring

DESIGN BASIS FOR FEED AND PRODUCTS

FEEDS:

Case1 Case2TBP Cut point, oC 550 550Gravity, oAPI(Sp.gravity) 7.80(1.0158) 8.79(1.0086)Total Sulfur wt% 3.90 3.30Conradson Carbon, wt% 16.800 19.895Kinetic Viscosity @ 50oC, cst 20,500 85,066Kinetic Viscosity @ 100oC, cst 480 1,030Nitrogen, wt ppm 3,200 3,100Normal pentane Insolubles, wt% 11.3 11.7Ni + V, wtppm 89 101Sodium, wtppm 10 max 10 max

PRODUCTS: The visbreaking Process Unit shall be designed to produce the following products:

Off gas shall be routed to Gas Concentration Process unit. Unstabilized naphtha(IBP – 150oC) shall S be routed to Gas Concentration Process unit. Visbroken bottoms residue (150 + oC) shall be routed to Fuel Oil blending(Normal Flow)

or Refinery Fuel Oil(Normal no flow).

DESIGN CONSIDERATIONS:

The Visbreaking process unit is a coil type visbreaking operation.The design incorporate on-line steam/air decoking facilities and allow for hydraulic pigging for the visbreaker heater.The

Visbreaking Heaters uses a combination fuel gas-fuel oil fired burners.The Visbreaker Heaters is specified as three separate heaters,each sized for 50% of the total required heat absorption.

EFFLUENTS INFORMATION: Effluents of Visbreaking process unit are sour water,Heater flue gas, and blow down water.

SOUR WATER:It is produced from the fractionators overhead drum at a rate of 3,239 kg/hr and 3,220 kg/hr for Case1 and Case2 respectively.Sour water is then sent to the sour water stripper at Amine Treating Unit.

FLUE GAS (from heater stack): 25,500 kg/hr

BLOW DOWN WATER (from V12,V13): 200 kg/hr

BATTERY LIMIT CONDITIONS:

Origion/Destination Temp oC Pressure kg/cm2gFeedsVacuum residue Vacuum distillation unit 232 6.33Start up fuel gas Offsite 60 5.7Start up FLO Offsite 50 11.5ProductsOff gas GCU 49 3.52Unstabilized Naphtha

GCU 49 16.52

Visbroken bottoms To blender 175 5.3To refinery fuel oil tank 175 3.5

Sour water Amine treating unit 49 3.5

OPERATING CONDITIONS AND CONTROL:

PROCESS VARIABLES: The yield and quality of the products from VBU are related to four process variables:

Furnace Outlet Temperature Pressure Feed Stripping Steam

VISBREAKER HEATERS:

The heater supply the heat of reaction for visbreaking.The mode of unit operation and product formation depends on the furnace outlet temperature.The higher the heater outlet

temperature, the higher the conversion to a lighter liquid yield.The lower the furnace outlet temperature, the lower the lighter liquid yield.

If the feed rate to the furnace is reduced, the high pressure steam must be increased to give the feed less residence time ine furnace.This prevents coke laydown in the heater tubes.

TEMPERATURE: The inlet temperature to the heater is 316oC.The temperature leaving the heater is 470oC with a maximum outlet temperature of 496oC prior to quenching. Following quenching, the temperature going into the Visbreaker Fractiionator, 130-V1 is 385oC.

PRESSURE:

The pressure drop across the heaters when they are clean is 14.0 kg/cm2. The pressure drop across the heaters when they are dirty is 21.0 kg/cm2.

VISBREAKER FRACTIONATOR AND GAS OIL STRIPPER: The operating variables of the fractionators are:

Reflux Furnace outlet temperature Feed rate Stripping steam

The operating variables for the gas oil stripper, 130-V2, are stripping steam rate. The normal stripping steam rate to the side stream stripper is based on 2.9 kg of steam to 100 kg of gas oil product.The fractionators stripping steam rate is based on 1.9 kg of steam to 100 kg of the combined residue feed quench and fractionator bottoms rate.

EFFECT ON PRODUCT QUALITY:

Once the unit is on stream with products meeting specifications, operating conditions (temperatures, pressures, and flows) are to be kept constant.If it is necessary to change an operating condition such as the heater charge rate, it is to be done in small increments.At the same time product rates are to be changed proportionally to maintain the temperatures and pressures of all streams.

PROCESS FLOW AND CONTROL: Vacuum Distillation Unit Residue enters the Feed Surge Drum 130-V-4, pumped by Heater Charge Pumps P1 A/B followed by feed preheating through E1-A/B/C/D on shell side to a temperature of 316 0C before entering into Visbreaker heaters.

VISBREAKER HEATERS:The unit consists of three identical heaters H1-A/B/C.Two of the heaters are in operation while the third one remains as spare. Each heater takes about 50% load of total visbreaking process. The feed enters through flow control valves FV-035A/B/C into the respective heaters.

Heater coils consist of two main sections; radiant section and convection section. Radiant section is located at the top of furnace which is directly heated by the rising burner flames two passes at the sides of the furnace walls which is known as convection section.

At each inlet to heater velocity stream is injected at a specific rate to control the resident time of heater charge. Each Visbreaker has a steam superheat coil which superheats medium pressure steam to be used in the Gas Oil Stripper 130-V-2 and Visbreaker Fractionator 130-V-1.

There are eight burners provided with both Fuel Oil and Fuel Gas for burning. A pilot burner is provided to each main burner as well.The outlet temperature of heaters determines the extent of conversion. A higher outlet temperature means more lighter products and vice versa.

The furnace effluent is controlled at 470 0C and the stream passes through Quench Valve PV-063 is quenched by two in coming streams named as Gas Oil Quench through FV-064 and Residue Quench through FV-067. The temperature drops down from 470 0C to 385 0C.The objective of quenching is to suppress the cracking process any further. The stream followed by quenching enters the Visbreaker Fractionator at a temperature not more than 3850C.

VISBREAKER FRACTIONATOR:The quenched stream flashes across the fractionator column and vapor-liquid separation takes places. Superheated steam from furnace enters near the bottom. Lighter ends rise up the column.There are three main streams that leave the column, Overhead vapors, Gas Oil Pump Around and Fractionator Bottoms.

1-FRACTIONATOR BOTTOM Fractionator bottom stream is filtered through ME-8 A/B strainers and pumped by P-3 A/B directly to the tube side of heat exchangers E-1 A/B/C/D. Then to MP steam generators E-3A/B and temperature controlled by TV-127A/B.From here the fractionator bottoms stream splits into three parts: Heater Residue Quench, V-1 Bottom Quench flow controlled by FV-087 and the net bottom product level controlled by LV-082 passing through E-4 and E-5 and finally to RFO through FV-037 and to Fuel Oil Blending.

2- GAS OIL PUMP AROUND The objective is to draw off heat from the fractionator column. The stream is taken off

and splits into two parts one enters the Gas Oil Stripper V-2 level controlled by LV-118, counter-flows across superheated steam which takes out any lighter ends back to V-1 and the second one is pumped by P-2 A/B to MP steam generator E-7 and back into V-1 as wash oil flow controlled by FV-086.From here the main stream divides into two parts one as Gas Oil Quench and the other as Gas Oil Pumparound. Gas Oil Quench passes through Gas Oil Quench Cooler EA-1 and then enters as Gas Oil Quench stream at the heater effluent through FV-064 while a

portion of it mixes with Gas Oil Pumparound stream through TV-107B at the down stream of TV-107A. Both streams combine and flow through FV-085 to the top of Pumparound section to cool overhead vapors.The gas oil product from V-2 is pumped by P-4A/B through FV-121 to the net bottom products which flow to the battery limit.

3- OVERHEAD VAPORS The fractionator overhead vapors are condensed in Fractionator Overhead Condenser EA-2 and then into Fractionator Overhead Trim Condenser E-6.The vapor-liquid stream then flows down to Fractionator Overhead Drum V-3.

The vapors flow through FV-096A to GCU a provision is made to direct the vapors to flare through FV-096B under abnormal conditions.

Sour Water is pumped by P5-A/B controlled by LV-093 from the V-3 boot to the Amine Treating Unit.

Unstabilized Naphtha is pumped by two pumps; by P7-A/B to the GCU flow controlled by FV-101 and by P6-A/B back to the Visbreaker Fractionator V-1 as naphtha reflux flow controlled by FV-084.

UNIT OPERATIONS: The major unit operations are

Visbreaker Fractionator column Gas oil Stripper Pumps Heat exchangers

VISBREAKER FRACTIONATOR: The purpose of this column is the separation of the heater outlet into its various components.This column have 22 valve trays, 5 shed trays, 2 sccumulator trays.This column is protected by PSV-131.

GAS OIL STRIPPER: The purpose of Gas oil stripper is to remove the light end components and

returning it to the fractionators with stripping steam.Column draw off rate is controlled by LIC-121 and product rate is controlled by LIC-118.

PUMPS: Equipment used for the transportation of incompressible fluid is called pump.Pumps are installed in oil refineries and petrochemical complexes for a number of reasons.Due to nature of the materials being processed and the unit operations involved, following actions need to be carried out on liquids in the system.

Item No. Service Type Capacity (m3/hr)130-P1A/B Visbreaker charge pump Centrifugal 142130-P2A/B Gas oil pump around pump Centrifugal 137130-P3A/B Fractionators bottom pump Centrifugal 230130-P4A/B Gas oil product pump Centrifugal 15130-P5A/B Fractionator sour water pump Centrifugal 4130-P6A/B Fractionator reflux pump Centrifugal 7130-P7A/B Unstabilized naphtha pump sundyn 60

HEAT EXCHANGER:

A heat exchanger is a device built for efficient heat transfer from one medium to another.The media may be separated by a solid wall,so that they never mix, or they may be in direct contact.

GAS CONCENTRATION UNIT The Gas Concentration Process Unit is designed

to recover C3/C4 components known as LPG to produce a stabilized naphtha feed for Naphtha Hydrotreating Unit.

Gases are compressed in compressor section that come from various process units and passed through an absorption section to maximize the recovery of LPG. Off gas leaving the adsorption system is sent to Fuel Gas System.

After passing through the distillation section Stabilized naphtha and LPG are separated from Unstabilized naphtha that comes from various process unit. The recovered LPG is sent to LPG Merox Unit and stabilized naphtha is sent to Naphtha Hydrotreating Unit.

PROCESS DESCRIPTION: The Gas Concentration Unit consists of following major sections;

The wet gas compressor, The primary absorber, The sponge absorber, The stripper, The debutanizer. This equipment recovers by absorption and separates by

distillation the lighter components.

FEEDS:The gas feeds to the Gas Concentration Process Unit are as follows:

Gas from the stripper receiver at the DieselMax Unit Gas from the flash fractionator receiver at the DieselMax Unit Gas from the fractionator receiver at the Visbreaking Unit Gas from the stripper at the Naphtha Hydrotreating Unit Gas from the net gas compressor at the Crude Distillation Unit

The liquid feeds to the Gas Concentration Process Unit are as follows:

Unstabilized naphtha from the Crude Distillation Unit Unstabilized naphtha from the fractionator receiver at the Visbreaking Unit Unstabilized naphtha from the product fractionator receiver at the DieselMax Unit Unstabilized naphtha from the flash fractionator receiver at the DieselMax Unit Unstabilized liquid from the stripper receiver at the Naphtha Hydrotreating Unit

(normally no flow)

PRODUCTS: A mixed C3/C4 LPG stream is recovered and sent to the LPG Merox Unit. Debutanized naphtha is sent to the Naphtha Hydrotreating Unit. Off gas from the Absorber is routed to the fuel gas treater at the Amine Treating Unit.

PROCESS CONTROLS: The Gas Concentration Unit can be divided into three sections,

1. The compressor section,2. The absorber section and 3. The distillation section.

In the compressor section, off gas from several units is pressurized to facilitate the operation of subsequent two sections.

In the absorber section, ethane and lighter hydrocarbons and hydrogen sulfide are removed. It has the high pressure receiver, primary absorber, sponge absorber.

The distillation section consists of a stripper and a debutanizer.

COMPRESSOR SECTION: The feed source to the compressor section of the Gas Concentration Unit consists of:

Gas from Stripper at Dieselmax Unit Gas from Flash Fractionator at Dieselmax Gas from Fractionator at Visbreaking Unit Gas from CDU Gas from Stripper at Naphtha Hydrotreating Unit Unstabilized Naphtha from CDU Unstabilized Naphtha from Product Fractionator at Dieselmax Unit Unstabilized Naphtha from Fractionator at Visbreaking Unit Unstabilized Naphtha from Flash Fractionator at Dieselmax Unit Unstabilized Stripper Overhead Liquid from Naphtha Hydrotreating Unit

Gas streams from different processes combine and flow through the compressor suction cooler 411-E1 and compressor suction drum 411-V1 into centrifugal compressor 411-C1 which has two stages. Condensate accumulated in the compressor suction drum is sent to the CDU.

First stage pressurizes the gases from 3.7 kg/cm2 to 8.2 kg/cm2

Gas from the Naphtha Hydrotreater Stripper combines with first stage discharge and enters the interstage cooler 411-E2, then the compressor interstage suction drum 411-V2 and finally into second stage of the compressor.

Gas leaving the second stage is pressurized to 17.1 kg/cm2

Gas from Second stage mixes with

1. Stripper (411-V6) overhead vapor2. Primary Absorber bottoms pumped by 411-P4 A/B3. Interstage Suction Drum condensate pumped by 411-P1 A/B4. Combined Liquid Stream from various Process Units is cooled in the high

pressure cooler 411-E3 and sent to the high pressure receiver 411-V3.

ANTISURGE CONTROL: During normal operation, the compressor suction 411-V1 pressure controller 411-PIC-004A maintains drum pressure by adjusting the turbine speed. The compressor spillback valve is closed.For a decreasing pressure in the 411-V1, 411-PIC-004A output will decrease. This will decrease the turbine speed. If at any time the first stage antisurge controller, 411-XIC-006 or the second stage antisurge controller, 411-XIC-019 detects that the compressor operating point is crossing the surge control line, the antisurge controller will open the first or second stage spillback valve to prevent compressor surge. A decoupling signal is communicated to the process controller to insure stable pressure control during spillback operation.

For an increasing pressure in the 411-V1, 411-PIC-004A output will increase. This will increase the turbine speed. If the turbine speed is at maximum, or for any reason the 411-V1 pressure reaches the pressure override controller 411-PIC-004B set point, the pressure control valve 411-PV-004 will open to maintain pressure in the 411-V1.

ABSORBER SECTION: Absorption takes place in two of the columns namely:

Primary Absorber 411-V4 Sponge Absorber 411-V5

Both performs the same function of absorption. The difference being in their internal structure. Primary absorber contains trays while the sponge absorber has internal packings to ensure intimate contact between liquid and gas flows. Stabilized naphtha is used in Primary Absorber to absorb C3/C4 contents of the off-gas leaving the High Pressure Receiver while Sponge Absorber uses Circulating Diesel instead.

Absorption of C3/C4 takes place when Rich Gas flows from the high pressure receiver to the primary absorber (411-V4) where it meets stabilized naphtha as an absorbing media. The gas flows upward through the primary absorber contacting a down flow of stabilized naphtha from the debutanizer bottoms recycle pumps (411-P6A/B). The counter-current flow of gas and liquid streams causes an intimate contact which maximizes the absorption rate. The debutanizer

bottoms is introduced into the primary absorber at top tray. The heat generated during absorption is removed by pumping rich naphtha from the column through the intercooler (411-E4) and back to the column. At the bottom of the primary absorber, the rich naphtha is pumped on level control through the high pressure cooler into the high pressure receiver by primary absorber rich oil pump (411-P4).

Off-gas leaving the top flows to the bottom of the sponge absorber(411-V5) for final absorption of any C3/C4 components where it is contacted with a down flow of lean sponge oil ( cold circulating diesel) which entered the column on flow control above the packing. The off gas leaves the unit on pressure control by 411-PV-049 and directed to the Amine Unit for treating.

Circulating diesel from the Crude Unit provides heat to the debutanizer (411-V7) and lean oil to the sponge absorber. After the debutanizer reboiler (411-E11) a stream of diesel is cooled by heat exchange (411-E5) with rich oil from the bottom of the sponge absorber and a fin-fan cooler (411-EA1) and trim cooler (411-E6). The diesel is pumped on flow control into the top of the tower above the packing by (411-P5 A/B). The lean sponge oil flows downward through the sponge absorber contacting the upward flow of gas. From the sponge absorber bottom, the rich oil flowing on level control is heated by the lean oil in the sponge absorber lean oil - rich oil exchanger (411-E5) before joining the main circulating diesel and back to the Crude Distillation Unit.

DISTILLATION SECTION:

Liquid stream from the high pressure receiver is pumped by 411-P2A/B to the stripper (411-V6) where hydrogen sulfide (H2S), ethane and lighter components are stripped out. Feed to the stripper is heated using the total (recycle plus net) debutanizer bottoms stream in the stripper

feed exchanger (411-E7). Stripper charge enters the column above the top tray. Stripping action is provided via reboilers (411-E8/E9) at the bottom of the column through which heat is supplied from the debutanizer bottoms stream and MP steam. The steam flows through the tube side of the stripper reboiler with its condensate flow regulated by a flow controller which is reset by the stripper overhead vapor rate. The overhead vapor is returned to the high pressure cooler. Bottoms material is pressured to the debutanizer (411-V7) on level control.

The heating oil to the reboiler is the circulating diesel from the Crude Unit. Heat is supplied via a reboiler (411-E11).The feed to the debutanizer is the stripper bottoms stream. The diesel is regulated by a flow recording controller. The total flow of debutanizer bottoms is directed to one of the stripper reboilers and to the stripper feed exchanger. After exiting the stripper feed exchanger, the debutanizer bottoms flow is divided into two flows. The first flow is directed on flow control to the primary absorber as lean oil or sent to the stabilized naphtha storage tank. The second flow or net production of stabilized naphtha is pumped on level control to the Naphtha Hydrotreating Unit. Overhead vapors (butane, propane and propylene) are condensed by a fin-fan cooler condenser (411-EA2), and then collected in the debutanizer receiver (411-V8). A portion of the total overhead is pumped back to the column as reflux. The quantity of reflux is regulated by a flow recording controller which is reset by the level indicator controller on the receiver. The net overhead liquid product is pumped out on flow control which is reset by the Tray 10 temperature indicator controller. The debutanizer overhead product is pumped to the LPG Merox Unit for sulfur removal. The debutanizer column pressure is controlled by a pressure recording controller on the overhead vapor line which controls the amount of vapor directed through the condenser. The pressure differential controller regulates a valve in a bypass line around the cooler. This allows a small amount of vapor to bypass the condensing system to regulate the surface area available to condense the overhead vapor, therefore regulating the rate at which the vapor is removed from the column.

UNIT 0PERATIONS:

Absorption Distillation Compressor Heat exchanger Pumps

ABSORPTION SECTION:

The main operating variable is the flow recorder controller in the stripper overhead vapor line which controls the flow of steam condensate from the stripper reboiler tubes.This vapor rate is the primary control in the recovery section of the GCU.Increasing this vapor flow wili decrease the H2S content of the stripper bottoms but will increase the propylene and propane content of the lean gas from the sponge absorber.The vapor flow should be adjusted to give a minimal content of H2S in the debutanizer overhead with a minimal content of propylene and propane in the sponge absorber off-gas.

A sudden increase in the H2S level of LPG stream should not be construed as being a result of lower stripping alone.

Since our objective is to maximize recovery of C3 and minimize the H2S content of LPG stream.

1. LEAN OIL FLOW TO THE ABSORBERS: PRIMARY ABSORBER

Recycled stabilized naphtha from the debutanizer bottoms can be adjusted to increase C3/C4 recoveries.

Arabian light Upper ZakumFeed gas temperature (oC) 39.5 39.5Feed liquid temperature (oC) 39.5 39.5Top temperature (oC) 49.2 49.8Bottom temperature (oC) 44.4 44.7Intercooler flow rates (kg/hr) 35,352 31,180Intercooler return temperature (oC) 39.5 39.5Column overhead pressure (kg/cm2.G) 14.1 14.1

SPONGE ABSORBER

Increasing the sponge oil flow rate tends to increase the absorption efficiencyof the sponge absorber.Changes in sponge oil flow rates have only a slight effect on C3 recovery.

Arabian light Upper ZakumTop temperature (oC) 44.9 45.1Lean oil temperature (oC) 39.5 39.5Lean oil flow rate (kg/hr) 22,181 22,093Rich gas flow rate (kg/hr) 9,131 9,334Column top pressure (kg/cm2.G) 13.9 13.9

2. ABSORBER TEMPERATURE: Absorption is favored by low temperatures.The intercooler is operated at less than 48oC.This operation will not scale the coolers because absorber temperatures are not high enough to precipitatescale from good quality cooling water.

3. HIGH PRESSURE RECEIVER TEMPERATURE:It is held close to 40 oC.Variations of this temperature will result in large variationin the stripping and absorption rates.his inlet temperature is high enough to scale the tubes should the water flow be restricted.

4. SYSTEM PRESSURE:

The sponge absorber pressure control maintains a constant pressure on all vessels in the GCU, except the debutanizer.The system pressureis a sponge absorber control which, can be used to keep necessary stripper vapor rate in a convenient range.The lower system pressure is desired when the unit swings to a high sulfur charge stock because the stripper column will be more efficient.Absorption efficiency decreases as the pressure decreases.Thus, more H2S will be rejected, but the LPG recovery will decrease.

DISTILLATION PROCESS: The distillation section consists of a stripper and a debutanizer.

1) STRIPPER COLUMN:The purpose of the column is to strip off the light ends plus the bulk of hydrogen sulfide that is present in the liquid stream from the high pressure receiver.

OVERHEAD VAPOR RATEIn this section all variables are held constant and this vapor rate is varied to control the H2S content of the LPG stream.

Arabian light Upper ZakumFeed temperature (oC) 60 60Column top temperature (oC) 64.1 64.1Column bottom temperature (oC) 171.7 163.5Column overhead pressure (kg/cm2.G) 15.1 15.1

2) DEBUTANIZER COLUMN:The stripped naphtha from the recovery section has to be stabilized for vapor pressure adjustment.In the debutanizer, a majority of the C3 and C4 components are taken overhead.This is the LPG product stream which can be sent for further treating and fractionated inti C3 and C4 product streams.

TOP TEMPERATUREThis variable is controlled by the 10th tray temperature and if it(10th tray temperature) is varied , then the composition of the overhead material will vary.

PRESSUREPressure is not considered a process variables, the column pressure is maintained by a control valve which controls the amount of overhead flow that is allowed to pass through the condenser and enter the receiver.There is also differential pressure controller that allows hot overhead flow to bypass the condenser and enter the receiver.This flows maintains set pressure differential between the column and the receiver.The pressure is not normally changed, but if there is a change in the pressure, then the composition will change.

BOTTOM TEMPERATURE

The bottoms temperature variables controls the amount of C3/C4 material that is allowed to remain in the bottoms product.This variabl is controlled by the reboiler heat input back to the column.This heat input is controlled by the amount of heating oil allowed to flow through the tube side of the reboiler on flow control.

Arabian light Upper ZakumColumn top temperature (oC) 69.8 69.8Reflux flow rate (kg/hr) 62,220 73,102Column bottom temperature (oC) 192.5 186Column overhead pressure (kg/cm2.G) 11.1 11.1

COMPRESSOR: A gas compressor is a mechanical device that increases the pressure of a gas by reducing its volume. Compressors are similar to pumps: both increase the pressure on a fluid and both can transport the fluid through a pipe. As gases are compressible, the compressor also reduces the volume of a gas.

Centrifugal compressors are machines in which velocity and pressure are given to the air or gas in radial direction by one or more impeller diffuser combination.

The compressor 411-C1 is a 2-stage centrifugal compressor which is driven by back pressure type steam turbine.Each stage has an anti-surge control valve.

All incoming gas streams except for the Naphtha Hydrotreater Stripper gas combine and flow into a two stage centrifugal compressor 411-C1 via the compressor suction cooler 411-E1 and the compressor suction drum 411-V1. Mists separated and accumulated in the compressor suction drum is manually drained out and sent to the CDU.

Gas from the Naphtha Hydro treater Stripper combines with the first stage discharge which is pressurized from 3.7 kg/cm2G to 8.2 kg/cm2G, and enters the second stage of the compressor via the interstage cooler 411-E2 and the compressor interstage suction drum 411-V2. Gas pressurized to 17.1 kg/cm2G mixes with the Stripper overhead vapor, the incoming liquid stream, and Primary Absorber bottoms. The combined stream is cooled in the high pressure cooler 411-E3 and sent to the high pressure receiver 411-V3. Liquid separated and accumulated in the compressor interstage suction drum is sent to the combined stream by the compressor interstage suction drum pumps 411-P1A/B. During normal operation, the compressor suction 411-V1 pressure controller 411-PIC-004A maintains drum pressure by adjusting the turbine speed. The compressor spillback valve is closed. For a decreasing pressure in the 411-V1, 411-PIC-004A output will decrease. This will decrease the turbine speed. If at any time the first stage antisurge controller, 411-XIC-006 or the second stage antisurge controller, 411-XIC-019 detects that the compressor operating point is crossing the surge control line, the antisurge controller will open the first or second stage spillback valve to prevent compressor surge. A decoupling signal is communicated to the process controller to insure stable pressure control during spillback operation. For a increasing pressure in the 411-V1, 411-PIC-004A output will increase. This will increase the turbine speed. If the turbine speed is at maximum, or for any reason the 411-V1 pressure reaches the pressure override controller 411-PIC-004B set point, the pressure control valve 411-PV-004 will open to maintain pressure in the 411-V1. The following inputs are used to characterize surge control line for each stage:

HEAT EXCHANGER: A heat exchanger is a device built for efficient heat transfer from one medium to another. The media may be separated by a solid wall, so that they never mix, or they may be in direct

contact. APPLICATIONS: They are widely used in space heating, refrigeration, air conditioning, power plants, chemical plants, petrochemical plants, petroleum refineries, natural gas processing, and sewage treatment. One common example of a heat exchanger is the radiator in a car, in which the heat source, being a hot engine-cooling fluid, water, transfers heat to air flowing through the radiator (i.e. the heat transfer medium).

PUMPS : Equipment used for the transportation of incompressible fluid is called pump.Pumps are installed in oil refineries and petrochemical complexes for a number of reasons.Due to nature of the materials being processed and the unit operations involved, following actions need to be carried out on liquids in the system.

USES OF PUMPS:

Movement of liquid from low levels to higher levels. Movement of liquid from areas of a low to high pressure. Increasing in the rate of fluid flow through a process and its associated equipme

In gas concentration unit mostly the pumps are centrifugal pump so now I discuss this pump.

CENTRIFUGAL PUMPS: Centrifugal Pumps operate by the principle of ‘Centrifugal Force’ e.g. A swinging pail generates a centrifugal force.

WORKING MECHANISM OF CENTRIFUGAL PUMPS:

• A centrifugal pump is one of the simplest pieces of equipment in any process plant. • Its purpose is to convert energy of a prime mover (a electric motor or turbine) first into

velocity or kinetic energy and then into pressure energy of a fluid that is being pumped.• The energy changes occur by virtue of two main parts of the pump, the impeller and the

volute or diffuser. The impeller is the rotating part that converts driver energy into the kinetic energy.

The volute or diffuser is the stationary part that converts the kinetic energy into pressure energy.

TROUBLESHOOTING: It should be an ongoing part of normal operations.Daily logs of operations are recorded reviewed to catch any changes before they become problems.

Problem Cause Solution

Primary absorber(411-V4) gas feed has high C4/C5 content.

Low pressure set point on the sponge absorber(411-V5) OVHD pressure PIC-049.

Increase pressure controller PIC-049 setpoint to design.

Low cooling water flow to high pressure cooler(411-E3).

Increase cooling water flow rate.

High C3 content in the sponge absorber OVHD gas.

Low flow rate of lean oil to the primary and/or sponge absorber.

Increase the flow rate of lean oil flow.

Low pressure set point on the 411-V5 OVHD pressure PIC-049.

Increase pressure controllerPIC-049 setpoint to design

High C2 or H2S levels in the primary absorber rich oil.

High flow rate of lean oil to the primary and/or sponge absorber

Decrease the flow rate of lean oil flow.

Low primary absorber feed gas flow rate.

Reduce primary absorber lean oil flow rate.

Primary absorber has high differential pressure across the column(flooding).

The vapor rate from the stripper OVHD is higher than necessary.

Decrease either the stripper feed preheater or reboiler duty to design.

Stripper bottoms have high H2S levels.

Low stripper feed inlet temperature.

Increase stripper feed temperature.

Insufficient stripper reboiler duty.

Increase stripper reboiler heating medium flow rate.