Revamp of Saturated Gas Concentration Unit (SGCU)

Transcription

1 Revamp of Saturated Gas Concentration Unit (SGCU) ADVANCING CHEMICAL ENGINEERING WORLDWIDE

2 Contents 1. Definitions Introduction Setup The process in SGCU The problem My involvement Analysis of the problem Operating Pressure Equipment constraints in the low pressure operation mode Solution to the problem of low LPG recovery Operating Pressure Column Performance Exchanger Performance Conclusion Equipment Performance LPG Recovery Commissioning Economics of 33

3 Attachments: 1. Feed composition to SGCU 2. Summary of streams for LP operating case 3. Summary of streams for HP operating case 4. Product Specifications 5. Economic Evaluation 3 of 33

4 1. Definitions ATU CDU CFC DHDS FEED HP HT H&MB ISBL KBPSD LCO LP LPG MMTPA P&ID PFD SGCU VDU VGO Amine Treating Unit Crude Distillation Unit a part of the Crude and Vacuum Distillation Unit Clean Fuels Complex Diesel Hydrodesulphurisation Front-End Engineering Design High Pressure operation Hydrotreater Heat and Material Balance Inside Battery Limit Kilo (Thousands of) Barrels Per Stream Day Light Cycle oil Low Pressure operation Liquefied Petroleum Gas Mega (Millions of) Metric Tonnes Per Annum Piping and Instrumentation Diagram Process Flow Diagram Saturate Gas Concentration Unit Vacuum Distillation Unit a part of the Crude and Vacuum Distillation Unit Vacuum Gas Oil 4 of 33

5 2. Introduction Over the last few years, the price of the crude oil has escalated manifolds. One of the main reasons for this price rise is the increased demand of distillates from the developing countries like China, India and Brazil, which is mainly spurned by the great strides made by these countries economically. In order to meet this ever increasing demand of distillates in the world, a number of refineries are being built around the world. In one of these new refineries, there are two crude columns. The crude distillation units produce many distillate cuts, which are mainly blended into various products. The overhead from the atmospheric distillation column is condensed in a partial condenser. The uncondensed vapor is either routed to the fuel gas header or to the flare header. A part of this condensate is routed back to the crude column as reflux, while the rest of it is further processed in the Saturated Gas Concentration Unit (SGCU). This condensate is the unstabilised naphtha, which is containing dissolved LPG fraction. In the proposed refinery, there are two Saturate Gas Concentration Units, Unit 3 and 4 (SGCU 3 & 4), catering to the two crude columns. One of the main requirements of the SGCU is to extract this valuable liquefied petroleum gas (LPG). The LPG has a very high commercial value in many countries. For example, in India, LPG is imported regularly to meet the domestic demand and any increased production of LPG is highly profitable. The LPG demand in India during the has been estimated at million tons, whereas the indigenous availability during this period is 8.64 million tones. The shortfall of 1.93 million tones will have to be imported. It is also expected that the demand for LPG will increase by approximately 4.5% during the to million tones. The cost of LPG in the international market is US$467 / ton. LPG is sold at a subsidized price to the domestic consumers in India. The LPG is sold to Indian consumers in a 14.5 kg cylinder. The cost of each cylinder is US$ 6.55 (US$ 452/ton) and this selling cost includes the processing and retail marketing costs also. It is estimated that Indian nationalized oil firms are loosing US$3.4 for the sale of each 14.5 kg cylinder, while the government is providing a subsidy of US$0.50 per cylinder. It can easily be seen from the above figures that if LPG is imported into India, the shortfall in price for the imported LPG would be much higher. For the present design of the new refinery, the LPG yield from the SGCU 3 & 4 is as low as mole %. In view of the great commercial advantage, the task was to improve this recovery to >90 mole%. 5 of 33

6 3. Setup The Saturated Gas Concentration Units 3 and 4 (SGCU 3 & 4), has been designed as a modified repeat design based on SGCU 1 & 2 of an existing refinery belonging to the same owners. The new SGCUs were envisaged to process about 10% higher throughput. In the new refinery there is a Clean Fuels Complex (CFC) having diesel hydrodesulphurization units (DHDS 1 & 2) and a LCO hydrocracker. This complex produces low sulfur diesel and other distillates which are sold in the international market at a premium. The light ends from this complex are needed to be further processed in a Light Ends Recovery Unit (LERU). The modified repeat design of the SGCU 3 & 4 has been made to accommodate these additional loads from the CFC - wild naphtha from DHDS 1, 2 and LCO Hydrocraker. This inclusion has eliminated the need for a new Light Ends Recovery Unit (LERU). A simple schematic of the SGCU feed scheme is given below (kindly refer to fig 1). The SGCU 1 & 2 of the existing refinery does not have the streams from the CFC. The operating pressure of the HP receiver for the existing unit (SGCU 1 & 2) is 14.4 kg/cm 2 g. However, in order to accommodate for the new streams from CFC, which are operating at a lower pressure of 9 kg/cm 2 g, in the SGCU 3 & 4, the overall operating pressure of the whole unit has to be decreased. The operating pressure of the HP receiver is considerably decreased to 8.1 kg/cm 2 g. The feed composition to the system is given in Attachment 1. WILD NAPHTHA FROM CDU 101 WILD NAPHTHA FROM VGO HTU OFF-GAS FROM VGO HTU WILD NAPHTHA FROM DHDS FEED PREPARATION WILD NAPHTHA FROM DHDS 2 WILD NAPHTHA FROM LCO HC CLEAN FUELS COMPLEX 309 Rich Oil from 145 Primary Absorber (C02) OFF-GAS FROM DHDS 1 OFF-GAS FROM DHDS 2 HTU Overhead vapour from Stripper (C05) C05 OVHD Receiver (V12) CW C05 Ovhd Condenser HP Cooler of OFF-GAS FROM LCO HC CLEAN FUELS COMPLEX

7 4. The process in SGCU The feed streams from various upstream units are mixed together and the resultant 2 phase stream is then cooled in the HP cooler (S01 A/B). The cooled stream is separated in the HP separator (V01). The vapour phase from V01 is then fed to the Primary Absorber, C02. C02 is having valve trays as internals. A naphtha stream from the stabiliser, C06, is used as the lean oil for absorption. In C02, the heavy ends are absorbed by the lean oil. The rich oil from the bottom of the C02 is routed to V01 via the HP cooler, S01 A/B. The overhead from the C02 is fed to a Sponge Absorber, C03. In the sponge absorber, which is a packed column, external lean oil (stream # 182) from the crude column is fed as the absorbing oil. This lean oil is actually circulating diesel oil. The heat from the hot lean oil is given to the rich oil (coming from the bottom of C03) in Lean oil Rich oil exchanger, S03. The rich oil is then routed back to the crude unit. The lean oil is partially cooled in the exchanger S03, and is further cooled in air cooler A01 & trim cooler S04. The unabsorbed gas from the top of the Sponge Absorber is routed to the ATU, where H 2 S is removed from this gas by amine absorption. It is finally routed to the fuel gas system of the refinery. Any LPG which slips through the C02 & C03 will be lost into the fuel gas system. A simple schematic of the system is given below in fig 2. 7 of 33

8 Feed Preparation S Fuel Gas to ATU CO2:- PRIMARY ABSORBER CO3:- SPONGE ABSORBER C05:- STRIPPER C06:- STABILIZER CW V P03 A/B Lean oil from CDU 301 S01 A/B C S08 A/B CW P05 A/B S03 C CW S04 C 0 5 Condensate MP Steam S06 C S09/S10 A02 CW P06 A/B S13 Naphtha to OSBL 312 CW 122 P01 A/B 302 S S S CIRCULATING DIESEL TO CDU HOT DIESEL FROM CDU S11 CW 212 LPG Rich oil to CDU Fig 2, SGCU Process Schematic 8 of 33

9 The liquid stream (unstabilised naphtha) from the bottom of V01 (stream 122) is pumped to a Stripper (C05). This stream is heated against the hot outlet stream in Stripper Feed/Bottoms Exchanger, S05 and further against the stabiliser bottom stream in S15. The purpose of the stripper (C05) is to strip off the absorbed light vapours in stream 122. The heat requirement for this column is provided by a steam heated reboiler, S06. The vapour is routed back to the feed preparation section and passing through the HP cooler is routed to the V01 (kindly refer to fig 1). The naphtha from the bottom of the Stripper (C05) is then passed through Stripper Feed/Bottoms Exchanger, S05 (where it provides heat to the inlet stream to C05) and routed to a Stabilizer, C06. The C06 is a distillation column where LPG faction is separated from the stabilised naphtha faction. The overhead from this column is passed through a total condenser. A part of this condensate is fed to the column as reflux while the other part is taken out as a LPG product. The detailed stream summary of this section is given in Attachment 1. The process schematic is shown in fig 2. The product properties are listed in Attachment The problem The C 3 & C 4 brought in with the feed streams to SGCU are either routed to the LPG faction (stream 212), or to naphtha from C06 bottom (stream 312), or with the rich oil going to the crude column (stream 187) or with the fuel gas going to the ATU (stream 142). The desired routing of the C 3 & C 4 content is to LPG faction (stream 212). However, if it is routed with the rich oil (stream 187) and sent back to the crude column, then also it is recoverable from the crude unit. However, when the C 3 & C 4 content is routed to the ATU then it is lost to the fuel gas system. Also, when it is routed with the naphtha then also it is lost to the naphtha stream. The C 3 & C 4 content of the various feed streams (including that carried in with the lean oil) is kg-moles/hr. The C 3 & C 4 which is recovered from this content (with streams 187, 212, 316) is kg-moles. Hence, the recovery is only mole%. Due to introduction of the CFC vapour streams at a lower operating 9 of 33

10 pressure, and resultant lower operating pressure in the overall unit, the LPG recovery is way below the desired recovery of 90 vol.%. Hence, the refinery will be loosing a substantial amount of money in a costly product. 6. My involvement I was involved in the process of accommodating the new CFC loads within the existing framework. I was given the additional task to evaluate the system and make a study to provide a solution to improve the LPG yield from the system. While doing the study, I have extensively used the process simulator ProII to evaluate the process conditions and feasibility of making changes. I have also used the programmes HTRI Xchanger Suite, and KG towers to evaluate the exchanger performances as well as the column performances in the changed operating conditions. 7. Analysis of the problem 7.1. Operating Pressure In the existing refinery, the recovery is above 90 vol. %. An obvious difference between the existing refinery and the new refinery was the addition of the liquid wild naphtha streams (stream nos. 102,107,110) and the vapour streams (stream nos. 106,111,112) from the CFC. This has resulted in a lower operating pressure of the feed section V01 and hence C02 & C03. In the existing refinery, the operating pressure of the V01 is 14.4 kg/cm 2 g while that in the new refinery it is only 8.1 kg/cm 2 g. Obviously, V01 is one of the critical sections of the whole recovery process because higher the operating pressure of V01, more the LPG faction will dissolve in the liquid phase and more will be routed to C05 & C06. In the existing refinery the mole% of LPG carried away with the vapour phase (stream 121) is 11.3 mole%. However, in the new design this loading is 20.8 mole%. When this relatively excess amount of LPG goes into the vapour phase, the extraction possibility of this faction in the absorbers decreases appreciably. I started my analysis from this point. I could easily see that the best solution for this problem is to increase the operating pressure of the system. Since, most of the procured equipment for the new refinery has been copied from the existing refinery (to save fabrication time) the design pressures of these equipment are quite high. Hence, my first task was to arrive at new operating pressures within the framework of the existing design pressures. 10 of 33

11 7.2. Equipment constraints in the low pressure operation mode In addition to the design pressures, another crucial area for this revamp study was the available heat transfer area of the exchangers. During the design of the low pressure operation, it was seen that the heat transfer area for the S01 A/B was limited. This again was because the S01 A/B was a repeat design from the existing refinery. The overall process flow through S01 A/B in the existing refinery is 303,100 kg/hr. This has now increased to 400,686 kg/hr (stream 117), which is a net 32% increased flow by weight. Now, in addition to this increase, the new streams from the CFC were also to be accommodated. It was seen that if this stream is added to stream 117, then due to the high vapour flow the overall heat transfer coefficient was low and vibration problems were expected in the S01A. Hence, the stream was added to a cooler stream after S01A (before S01B) (kindly refer to fig1). The vapour load from V01 was very high. In the existing refinery, the flow is 8,746 kg/hr while that in the new design is 23,188 kg/hr. This 2.65 times increase in the vapour load in C02 & C03 was checked against the old columns. The new valve trays in C02 were designed to handle the new increased loads. C03 is a packed tower and with increasing vapour load the performance of the tower was acceptable. The liquid stream from the V01 in the old design was 294,200 kg/hr. In the new design, this flow is 399,988 kg/hr. This increased load obviously affected the tray design of the C05 & C06. For C05, the tray vendor proposed a new type of valve trays which would be able to handle the increased liquid load within the existing diameter of the tower. However, for the C06, the trays were found to be limited in the capacity. It was decided that high capacity Superfrac trays from Koch Glitz would be used, instead of conventional valve trays to handle increased loads. The pumps were checked against the existing refinery. It was found that new pumps will be required for P01 A/B, P03A/B and P05 A/B. Only P02 A/B may be repeated from the existing design. The pumps P02 A/B, P04 A/B will require new impellers in the existing casing. (Kindly refer to fig 2 for details.) 11 of 33

12 8. Solution to the problem of low LPG recovery 8.1. Operating Pressure The LPG recovery of appx. 87mole % is not acceptable to the client. However, the plant can be started up with the design scenario (in low pressure operation mode) but recovery is to be increased for continuous operation. As I found out through a number of simulations, the basic tool to improve the LPG recovery was in increasing the operating pressure of the whole system. As the unit was designed as per the design pressures of the existing refinery, there was limited scope to increase the operating pressure if other constraints were tackled well. Now, the most optimum operating pressure would be governed by V01 pressure. The design pressure of V01 is 16.5 kg/cm 2 g. As per ASME Section VIII, Div 1 and API 520 requirement, the highest operating pressure for V01 can be kg/cm 2 g. This is assuming a 10% margin between operating and design pressure. Hence, in order to achieve this higher operating pressure, the first part of the problem was to bring in the feedstocks at higher pressures. The liquid streams may easily be provided at a higher pressure if the pumps were suitably rated. I have checked that a couple of the procured pumps may be limiting in providing the higher discharge pressure. But, this can easily be achieved by changing the impellers of the pumps within the existing casing. Also, I found out that a booster pump would be required for the CFC liquid streams. The off-gas from the VGO hydrotreater can be provided at a higher pressure as the high pressure separator inside VGO hydrotreater unit (from where this off-gas is routed) is operating at a pressure of 18.6 kg/cm 2 a. The only problem is to bring in the CFC gases at higher pressure. I proposed that a compressor need to be procured and provided ISBL the SGCU. It would be a common compressor for both the SGCU (SGCU 3 & 4). The discharge pressure for this compressor was decided to be 16.2 kg/cm 2 a. This was assuming a suction pressure of 9.8 kg/cm 2 a. After, considering appropriate pressure drops in the system, the operating pressure of the V01 was found to be 15.2 kg/cm 2 a. The revised feed system would be: 12 of 33

13 WILD NAPHTHA FROM CDU 101 WILD NAPHTHA FROM VGO HTU OFF-GAS FROM VGO HTU WILD NAPHTHA FROM DHDS FEED PREPARATION CFC Off Gas Compressor WILD NAPHTHA FROM DHDS 2 WILD NAPHTHA FROM LCO HC CLEAN FUELS COMPLEX 309 Rich Oil from 145 Primary Absorber (C02) 108A OFF-GAS FROM DHDS 1 OFF-GAS FROM DHDS 2 HTU Overhead vapour from Stripper (C05) C05 OVHD Receiver (V12) CW C05 Ovhd Condenser HP Cooler S01 A/B OFF-GAS FROM LCO HC CLEAN FUELS COMPLEX CW 121 To Primary Absorber HP RECEIVER V To Stripper Fig 3, Feed Schematic in HP operation I estimated using ProII simulation that due to higher operating pressure of the V01, the vapour load decreased from 23,188 kg/hr to 17,300 kg/hr. Most importantly, the LPG content of the V01 vapour was kg-mole/hr, which decreased to kg-mole/hr in the high pressure operation mode. The operating pressures of the various columns in the LP operation vis-à-vis HP operation is summarised below. This higher operating pressure meant that the actual vapour load in the columns was lesser and the columns operated better hydraulically. Of course, this increase in the operating pressure in the HP operating case could be achieved because the unit was a repeat unit and had enough margins the design pressure. 13 of 33

14 Columns Design Pressure Operating Pressure (LP operation) Operating Pressure (HP operation) Kg/cm 2 g Kg/cm 2 a Kg/cm 2 a C / /15.1 C / /14.7 C / /16.8 C / / Column Performance Primary Absorber (C02) Primary absorber tray loadings for LP operation is given below. NET FLOW RATES HEATER TRAY TEMP PRESSURE LIQUID VAPOR FEED PRODUC T DUTIES M*KCAL/H DEG C KG/CM2 KG-MOL/HR R L 861.4V P P V 890.7L 0.1W The operation at higher pressure level is given below NET FLOW RATES HEATER TRAY TEMP PRESSURE LIQUID VAPOR FEED PRODUC T DUTIES DEG C KG/CM2 KG-MOL/HR M*KCAL/HR L 787.0V P P of 33

15 V 971.4L 0.4L 0.5W As can be expected the pumparound cooling duty for C02 has increased for the HP case 9as compared to the LP operating case. The net vapour product from C02 has come down (stream 141), which would in turn decrease the load in C Sponge Absorber (C03) The sponge absorber is a packed column. It has been formulated in the calculation as a 3 tray column. The summary of loads for low pressure operation is as follows: NET FLOW RATES HEATER TRAY TEMP PRESSURE LIQUID VAPOR FEED PRODUC T DUTIES M*KCAL/H DEG C KG/CM2 KG-MOL/HR R L 771.4V V 281.2L The summary of loads for high pressure operation is as follows: NET FLOW RATES HEATER PRODUC T DUTIES TRAY TEMP PRESSURE LIQUID VAPOR FEED KG- DEG C KG/CM2 MOL/HR L 693.7V V 284.5L M*KCAL/H R The summary of the column flows for the LP operation is given below FEED AND PRODUCT STREAMS LP Operation TYPE STREAM PHASE FROM TO LIQUID FLOW RATES HEAT RATES TRAY TRAY FRAC KG-MOL/HR M*KCAL/HR FEED 182 LIQUID FEED 310 VAPOR PROD 143 VAPOR PROD 313 LIQUID The summary of the column flows for the HP operation is given below TYPE STREAM PHASE FROM TO LIQUID FLOW RATES HEAT RATES TRAY TRAY FRAC KG-MOL/HR M*KCAL/HR FEED 182 LIQUID FEED 310 VAPOR PROD 143 VAPOR PROD 313 LIQUID of 33

16 Stripper (C05) The stripper is a column with valve trays. The column internals has been designed for the LP operation case. However, for the HP operation case, the liquid flows are much higher. The summary of tray loads for low pressure operation is as follows: Rigorous Column 321_C05, SGCU Stripper COLUM N SUMMARY NET FLOW RATES HEATER TRAY TEMP PRESSURE LIQUID VAPOR FEED DEG C KG/CM2 KG-MOL/HR PRODUC T L 993.4V DUTIES M*KCAL/H R R L The summary of tray loads for high pressure operation is as follows: COLUM N SUMMARY NET FLOW RATES HEATER PRODUC T DUTIES TRAY TEMP PRESSURE LIQUID VAPOR FEED KG- DEG C KG/CM2 MOL/HR L V M*KCAL/H R 16 of 33

17 R L Although the pressure level of the HP operation for the Stripper is not much higher than the LP operation case, but the higher pressure in V01 has resulted in higher liquid loads in this column Stabilizer (C06) The stabilizer is a column with valve trays. However, in view of the heavy loads, the internal trays have been changed to proprietary trays from Koch Glitsch called the Superfrac trays. These trays are used where the vapour and the liquid loads are appreciably high. The column summary in LP operation is given below: Rigorous Column 321_C06, SGCU Stabiliser COLUM N SUMMARY NET FLOW RATES HEATER TRAY TEMP PRESSURE LIQUID VAPOR FEED PRODUC T DUTIES M*KCAL/H DEG C KG/CM2 KG-MOL/HR R 1C L V L of 33

18 S L 33R FEED AND PRODUCT STREAMS TYPE STREAM PHASE FROM TO LIQUID FLOW RATES HEAT RATES TRAY TRAY FRAC KG-MOL/HR M*KCAL/HR FEED 170V VAPOR FEED 170L LIQUID PROD 210 LIQUID PROD 316 WATER 1 0 PROD 225 LIQUID OTHER PRODUC T STREAMS TYPE STREAM PHASE FROM TO LIQUID FLOW RATES HEAT RATES TRAY TRAY FRAC KG-MOL/HR M*KCAL/HR NET 209 LIQUID TOTAL 203 VAPOR The column summary in HP operation is given below: Rigorous Column 321_C06, SGCU Stabiliser COLUM N SUMMARY NET FLOW RATES HEATER TRAY TEMP PRESSURE LIQUID VAPOR FEED PRODUC T DUTIES M*KCAL/H DEG C KG/CM2 KG-MOL/HR R 1C L of 33

19 V L S L 33R FEED AND PRODUCT STREAMS TYPE STREAM PHASE FROM TO LIQUID FLOW RATES HEAT RATES TRAY TRAY FRAC KG-MOL/HR M*KCAL/HR FEED 170V VAPOR FEED 170L LIQUID PROD 210 LIQUID PROD 316 WATER 1 0 PROD 225 LIQUID OTHER PRODUCT STREAMS TYPE STREAM PHASE FRO M TO LIQUID FLOW RATES HEAT RATES TRAY TRAY FRAC KG-MOL/HR M*KCAL/H R NET 209 LIQUID TOTAL 203 VAPOR The reboiler duty in the HP case is higher as expected. The reboiler has been checked and is found to be adequate for the higher load. The vapour load in the rectifying section is also higher. The liquid load is higher in the overall column. The tray supplier was contacted and they have confirmed that the Superfrac trays can accommodate such 19 of 33

20 changes in the loads Exchanger Performance While performing the study, I found that a few of the exchangers were limiting and hence did not allow any manipulation High Pressure Cooler (S01 A/B) The high pressure cooler is a 2 shell shell & tube heat exchangers connected in series. The cooling water side also in a series configuration and the cooling water flow is countercurrent to the process flow. The exchanger is a repeat design from the existing refinery. When the CFC gas was to be introduced, it was found that due to the high vapour load there was extensive vibration problem. Hence, the vapour was added after the 1 st exchanger. Also, this stream was at a low temperature of 39 o C and hence the bypass was possible. But, after the introduction of the compressor, this compressed gas will come in at around 73 o C (stream 108A). I proposed that this gas should be introduced upstream both the exchangers. I found that there was no vibration problem because the compressed gas was having lesser actual volume. The main limitation which I found for this arrangement was that the cooling water return from the exchanger S01 A was quite high. The maximum return temperature is 49 o C. When I analysed the high temperature, I could see that the main contributor to this was the overhead gas from the stripper (stream 309). In order to decrease this load, I aimed at decreasing the temperature of the feed to C05 (stream 126). I analysed that I could achieve this decrease by increasing the load on the reboiler S06 and in return decreasing the duty of S Stripper reboiler Exchanger (S06) In order to decrease the load from S15, (which decreased the overhead temperature from C05, which in turn brought down the load in S01A/B, decreasing the cooling water return temperature to 49 o C,) I utilised the available area of the reboiler to the maximum. I moved the required heat input for operation of the column, from that carried with the inlet stream to the reboiler. I increased the duty for the reboiler to 20.1 MMKcal/hr from MMKcal/hr. After all the changes, I could decrease the stripper overhead temperature from C05 to 92 o C from 109 o C. The heat required for S06 is provided by MP Steam Stripper overhead condenser (S14) A small cooler, S14 and V12 is available to decrease the load on V01 and S01 A/B. I have tried to maximise the load through S14. However, the S14 exchanger is a small exchanger with 109 m 2 of effective heat transfer area (compared to 1546 m 2 for S01 A/B). Within the 20 of 33

21 limits of the exchanger, I have maximised the process flow through the exchanger to kg/hr from kg/hr. With increasing the flow further, I found that there were vibration problems in the shell side (process side) Stabilizer reboilers (S09/S10) The heat required for S09/S10 is provided by hot diesel stream from the crude unit. The S09/S10 heat transfer area was another limiting factor. I used HTRI calculation to get the maximum possible duty from these reboilers. The S09 / S10 are parallel heat exchangers. The duty required by the column C06 was to be provided by either S09/S10 or had to be carried along with the feed, which was preheated by S05. I distributed the heat between S05 and S09/S10 after finding the limiting duty for S09/S10. It was a challenging job to distribute the overall heat load in the system so that I could utilise the maximum available area of the various exchangers and perform the desired function. 9. Conclusion 9.1. Equipment Performance I could successfully rerate the unit for the higher operating pressure. All the exchangers were checked for predicted performance (using HTRI program) and they were found to be adequate for the HP operation. The separators were also checked. Since, the operating pressures were higher; the separator diameters did not cause concern, as expected. The alarm set points of the liquid levels in the separators were re-checked and some of them were changed. After checking the predicted column performance, I could confirm that the HP operation would work within the purchased equipment LPG Recovery After the re-rating of the unit, the LPG recovery improved to 92.9 mole%, which was higher than the desired recovery. The brief stream summaries of the feeds and the product are given in Attachment 2 for the LP operating case & Attachment 3 for the HP operating case Commissioning The plant will be commissioned in the LP operating case. However, the plant shall be designed in a fashion that the operation will be switched over to the HP operating case 21 of 33

22 without any lengthy downtime Economics For the high pressure operation case, kg/hr of LPG is recovered as against kg/hr for the LP operating case. This means that by increasing the operating pressure, there was an increased recovery of kg/hr (kindly refer to Attachment 6 for details). Now, LPG is priced at US$ 467/ton in the international market. The compressor is expected to be a small compressor with 750 kwh motor. The electricity cost is expected to be US$ 58 per hour (at an estimated cost of Rs.3.5 / kwh of electricity consumed). Hence, there is an increased recovery of US$ / hour. It is considered that there will be 8000 operating hours per year (a service factor of 91%). Hence, the excess money made due to the higher operating case would amount to US$ 7.06 million per year. Now, it is estimated that the installation of the compressor would cost US$ 4 million. I have assumed that the associated additional equipment would cost another US$ 750,000. An additional US$ 250,000 may be assumed for the engineering and piping. Hence, I could show that after the re-rating of the unit for higher operating case, the payback is expected to be less than 1 year. 22 of 33

23 Feed Composition Attachment 1 Stream Name Phase Liquid Liquid Mixed Liquid Mixed Liquid Mixed Liquid Mixed Mixed Liquid Total Stream Rate KG-MOL/HR 2, KG/HR 200,947 18,174 6,931 15,329 5,839 33,935 13,640 13,225 4,277 3,523 65,334 Std. Liq. Rate M3/HR Temperature C Pressure KG/CM Molecular Weight Mole Fraction Liquid Sp. Gravity Vapor Rate KG-MOL/HR KG/HR 6,664 5,536 12,922 4,171 3,152 M3/HR Molecular Weight Liquid Rate KG-MOL/HR 2, KG/HR 200,947 18, , , , ,334 M3/HR Molecular Weight Total Molar Comp. Rates KG-MOL/HR H 2 O AIR H NH H 2 S METHANE ETHANE PROPANE PROPENE IBUTANE BUTANE BUTENE IPENTANE PENTANE of 33

24 C Summary of streams for LP operating case Attachment 2 Stream No Phase Liquid Liquid Mixed Liquid Mixed Liquid Mixed Liquid Mixed Mixed Mixed Vapor Vapor Liquid Mixed Liquid Mixed Liquid Total Stream Rate KG-MOL/HR 2, , ,502 2,265 KG/HR 200,947 18,174 6,931 15,329 5,839 33,935 13,640 13,225 4,277 3, ,686 23,188 16,563 43,000 49,001 36, , ,14 7 Temperature C Pressure KG/CM Total Molar Comp. Rates KG-MOL/HR H 2 O AIR H NH H 2 S METHANE ETHANE PROPANE PROPENE IBUTANE BUTANE BUTENE IPENTANE PENTANE C LPG content kg-mole/hr LPG in with feed (With streams ) LPG with Lean oil Stream Total LPG in LPG out LPG recoveredout kg-mole/hr kg-mole/hr kg-mole/hr of

25 % LPG recovered 86.92% Stream No Phase Liquid Liquid Vapor Liquid Liquid Mixed Vapor Liquid Liquid Liquid Liquid Liquid Liquid Liquid Liquid Total Stream Rate KG- MOL/HR 4, , ,629 2,979 2,979 2,300 2,949 2, ,622 4,622 KG/HR 399,988 81,321 47, ,958 43, , , , , , ,559 73,412 73, , ,988 Temperature C Pressure KG/CM Total Molar Comp. Rates KG-MOL/HR H 2 O AIR H NH H 2 S METHANE ETHANE PROPANE PROPENE IBUTANE BUTANE BUTENE IPENTANE PENTANE C LPG content of 33

26 Stream No Phase Vapor Vapor Vapor Vapor Mixed Total Stream Rate 343 1, ,220 30,474 30,824 22,565 49,001 Temperature Pressure Total Molar Comp. Rates H 2 O AIR H NH H 2 S METHANE ETHANE PROPANE PROPENE IBUTANE BUTANE BUTENE IPENTANE PENTANE C LPG content of 33

27 Attachment 3 Summary of streams for HP operating case Stream No A A Phase Liquid Mixed Mixed Liquid Mixed Liquid Vapor Liquid Mixed Mixed Mixed Vapor Vapor Liquid Mixed Liquid Mixed Total Stream Rate KG-MOL/HR KG/HR Temperature C Pressure KG/CM Total Molar Comp. Rates KG-MOL/HR H2O AIR H NH H2S METHANE ETHANE PROPANE PROPENE IBUTANE BUTANE BUTENE IPENTANE PENTANE C Total LPG content Kg-mole/hr LPG in with feed (With streams A+110) LPG with Lean Oil Stream LPG in LPG out LPG recovered Total % LPG recovered 92.89% 27 of 33

28 Stream No A 303 Phase Mixed Liquid Mixed Liquid Liquid Mixed Vapor Liquid Liquid Liquid Liquid Liquid Liquid Liquid Liquid Total Stream Rate KG- MOL/HR KG/HR Temperature C Pressure KG/CM Total Molar Comp. Rates KG- MOL/HR H2O AIR H NH H2S METHANE ETHANE PROPANE PROPENE IBUTANE BUTANE BUTENE IPENTANE PENTANE C of 33

29 Stream No Phase Mixed Mixed Mixed Mixed Mixed Vapor Mixed Total Stream Rate Temperature Pressure Total Molar Comp. Rates H2O AIR H NH H2S METHANE ETHANE PROPANE PROPENE IBUTANE BUTANE BUTENE IPENTANE PENTANE C of 33

30 Attachment 4 Product Specifications 1) Off Gas ex Sat Gas Unit (Stream 143) Quality Required Quality Quality Achieved LP Operation HP Operation C5+ Content Target < 1.7% mole 0.39 mol% 0.35 mole% C3 Content Will be determined by the LPG recovery mol% 5.11 mole% 2) LPG ex Sat Gas Unit (Stream 212) Quality Required Quality Quality Achieved LP Operation HP Operation LPG Recovery from feed Target > 90% 89.2 mole% 92.9 mole% LPG Vapour 37.8 o C (ASTM-1267) 520 kpa min 1050 kpa max (Target 800 kpa) 739 kpa 773 kpa C5+ content < 1% mol 0.82 mole% 0.82 mole% H2S 1,200 ppm wt. max ppm wt ppm wt. 3) Full Range Naphtha ex. Sat-Gas Unit (Stream 312) Quality Required Quality Quality Achieved LP Operation HP Operation C4- content < 1% mol 0.9 mole% 0.9 mole% D86 ASTM 95% Resultant >155 C, <180 C o C 157 o C 30 of 33

31 Economic evaluation Attachment 5 LP operating case Stream No Phase Liquid Liquid Mixed Liquid Mixed Liquid Mixed Liquid Vapor Liquid Mixed Liquid Liquid Total Stream Rate KG-MOL/HR 2, ,265 KG/HR 200,947 18,174 6,931 15,329 5,839 33,935 13,640 13,225 16,563 43,000 49,001 36, ,147 Temperature C Pressure KG/CM Total Molar Comp. Rates KG-MOL/HR H 2 O AIR H NH H 2 S METHANE ETHANE PROPANE PROPENE IBUTANE BUTANE BUTENE IPENTANE PENTANE C LPG content KG-MOL/HR kg/hr Kg-mole/hr kg/hr LPG in with feed (With streams ) LPG with Lean oil Stream Total LPG in LPG out LPG out kg/hr kg/hr kg/hr of 33

32 HP operating case Stream Name A A Phase Liquid Mixed Mixed Liquid Mixed Liquid Vapor Liquid Vapor Liquid Mixed Liquid Liquid Total Stream Rate KG-MOL/HR KG/HR Temperature C Pressure KG/CM Total Mol. Comp. Rates KG-MOL/HR H2O AIR H NH H2S METHANE ETHANE PROPANE PROPENE IBUTANE BUTANE BUTENE IPENTANE PENTANE C Total LPG content Kg-mole/hr Kg/hr kg-mole/hr kg/hr LPG in with feed LPG with Lean Oil Stream LPG in kg/hr LPG out LPG recovered Total Extra LPG recovered than LP US$467/ton is US$/hr 32 of 33