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Guidance sheet – technical report

Word count:

Up to a maximum of 5,000 words.

Objective:

To demonstrate a depth of theoretical and practical knowledge of chemical engineering and its underlying principles.

Your technical report should demonstrate that your chemical engineering knowledge and understanding is of a high standard, showing how you make effective use of chemical engineering science and principles in creating designs and solving problems. This should be done through an ordered and critical description of aspects of chemical engineering practice where you have made a significant contribution.

You may submit as your technical evidence one of the following:

1. a report written specifically for the application, which addresses the objectives of the technical report;

2. a pre-existing technical report, accompanied by a narrative which details how the report demonstrates that your knowledge and understanding is at the required academic level;

3. a collection of published academic papers along a common theme, accompanied by a narrative which details how the papers demonstrate that your knowledge and understanding is at the required academic level.

You should state on the report’s title page which option you have chosen and identify the narrative if it is required.

Cover sheet:

Please complete the document ‘Cover sheet: technical report or academic papers’ to assist assessment, and use as an aid in structuring your report. The cover sheet must include a signed confirmation from someone familiar with your work, that the report is a true demonstration of your knowledge and understanding of chemical engineering.

Guidance:

The purpose of a technical report is to provide evidence that your knowledge and understanding is of the minimum required level for a Chartered Member (MIChemE).

Your technical report should:

1. be a substantive work of your choice that is sufficiently complex to challenge your chemical engineering capabilities;

2. be an original work on which you have made a significant contribution;

3. be focused on provision of evidence of your knowledge and understanding of the principles that underpin your chemical engineering competence;

4. aim to provide evidence and confidence that your overall chemical engineering knowledge is at the minimum required level for assessment as a Chartered Chemical Engineer;

5. include key calculations, engineering drawings, flow sheets, diagrams and appropriate references;

6. show objectives, problems or development aims addressed by the selected project;

7. include analytical and quantitative processes used in reaching and evaluating a specified solution;

8. be concise, not exceeding the guidance on word count;

9. clearly reference any material which is copied into the report, you should show the application of such material to demonstrate your knowledge of chemical engineering principles, and their application, at the required advanced level.

Your report is not:

10. a review of your training, employment and experience;

11. a report on your managerial or organisational competency and experience.

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Important guidance notes:

IChemE’s review process is such that you will be assessed by chemical engineering professionals; Chartered Members and Fellows of IChemE who volunteer their time to uphold the standards and support professional development of members.

Please submit your report in the best and most complete format possible. This will help IChemE to use the valuable time of our volunteer assessors effectively and to obtain a decision on your application as soon as possible.

■■ ensure that the information contained in the report is sufficient to demonstrate that your chemical engineering knowledge and understanding meets the minimum required level for Chartered Member;

■■ please complete and submit the ‘Cover sheet: technical report or academic papers’ to refer the assessors to the relevant information within your report and as an aid to structure;

■■ if your technical report:

– includes papers which were not wholly authored by you, or

– if you are referencing group activity,

please include a narrative which details precisely what your contribution to the submitted work was, referencing the relevant sections, and clarifying what was undertaken by others;

■■ practical and common sense solutions are a useful contribution but must be supported by science or appropriate theory – the report must demonstrate a level above ‘catalogue engineering’;

■■ acronyms and sector or company specific terms must be defined, either within the document if it is written for IChemE assessment, or with use of a glossary;

■■ the use of imported or adopted information, standards or guides must be clearly referenced as such;

■■ your report must be in English. IChemE is sympathetic to minor errors where English is not your first language, however, please ask someone to review your report, or use spell check software carefully to ensure your meaning is always clear;

■■ do not submit a report in excess of 20% more than the stated word count;

■■ ensure that you will remember the work in your report well enough to discuss at interview.

Focus your report on: Avoid:

✓ technical content ✗ description

✓ personal involvement ✗ observation

✓ explanation ✗ discussion

✓ analysis ✗ opinion

✓ evaluation ✗ commercial detail

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Cover sheet: technical report or academic papersPlease complete and submit with your technical report or academic papers at the time of application in order to assist IChemE with assessment.

Complete the page reference column to refer the assessors to the corresponding elements in your work and supporting document(s). Where possible, please use this structure as a template.

nb: You must also refer to the relevant IChemE guidance document on writing a technical report.

Family name: ............................................................................................................................................

Given name: ............................................................................................................................................

Affiliation (company/university or other organisation): ………………………………………………….................................

Membership no.: ......................................................................................................................................

Title of work: ............................................................................................................................................

You may submit as your technical evidence one of the following – please tick the appropriate box below:





ADVANCINGCHEMICALENGINEERINGWORLDWIDE

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1. Introduction Page reference(s)

Provide a clear introduction to the activity, giving an explanation of the methodologies used and/or background information.

2. Main content

Required elements

nb: Outputs from computer programmes used during the described work should not be included without explanation of the chemical engineering principles and equations used within those programmes.

Objectives or problems addressed by your work.

Evidence of your knowledge and understanding of chemical engineering principles as applied in the described work.

nb: Not all of the principles below need to be included, but are listed as examples which may be used. A range of applications of chemical engineering should be included from this list:

Conservation laws for mass and energy

Fluid flow

Heat and mass transfer

Thermodynamics of chemical and power systems

Reactor design

Analysis of process systems

Safety and environmental aspects

Other (please specify)

Analytical and quantitative processes used in reaching and evaluating the described solution.

Key calculations, to include:

How the basis of the calculation was selected and validated

How the calculation inputs were selected and validated

Relevant equations

3. Summary

Provide a summary and discussion of results

Provide conclusion(s)

4. Attestation

Please sign as the author:

Signature:

Job title:

This report is a true account of my professional activities:

Date:

5. Confirmation

Please also obtain the signature of someone familiar with your work at the time of activity and ideally a chemical engineer:

Family name: Given Name: Title:

Job title:

Affiliation (company/university or other organisation):

Relation to applicant:

Email:

To the best of my knowledge, this is a true account of the applicant’s activities

Signature:

Date:

6. Appendices

Provide inputs to computations and adequate explanation of background or supporting theory to worksheets or programmes.

Flowsheets, for example, process flow diagram (PFD), process and instrumentation diagrams (P&ID)

Engineering drawings

Diagrams

References to published texts, standards or papers which were used during the work described


Other (please specify):

Check that you need to submit a technical report at www.getchartered.org

Chartered Member (MIChemE)Review process for technical report

Technical report assessment (including referral back to you if

clarification or revision is required)

Technical interview (usually held on same day)

Complete your technical report, ensuring it meets the requirements outlined in the guidance, and submit along with your Competence and Commitment report* to apply for

Chartered Member

Competence and Commitment report assessment (including

referral back to you if clarification or revision is required)

Professional review interview (usually held on same day)

If required, you may be invited to a second interview on one or both elements

IChemE will review your application in full and inform you of the outcome the following week

Read the relevant guidance document

IChemE will acknowledge receipt of your application


IChemE will contact your referees in writing for a reference

* See ‘Review process for competence and commitment (C&C) report’ for further details

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Cover sheet: technical report or academic papersPlease complete and submit with your technical report or academic papers at the time of application in order to assist IChemE with assessment.

Complete the page reference column to refer the assessors to the corresponding elements in your work and supporting document(s). Where possible, please use this structure as a template.

nb: You must also refer to the relevant IChemE guidance document on writing a technical report.

Family name: ............................................................................................................................................

Given name: ............................................................................................................................................

Affiliation (company/university or other organisation): ………………………………………………….................................

Membership no.: ......................................................................................................................................

Title of work: ............................................................................................................................................

You may submit as your technical evidence one of the following – please tick the appropriate box below:






www.icheme.org

1. Introduction Page reference(s)

Provide a clear introduction to the activity, giving an explanation of the methodologies used and/or background information.

2. Main content

Required elements

nb: Outputs from computer programmes used during the described work should not be included without explanation of the chemical engineering principles and equations used within those programmes.

Objectives or problems addressed by your work.

Evidence of your knowledge and understanding of chemical engineering principles as applied in the described work.

nb: Not all of the principles below need to be included, but are listed as examples which may be used. A range of applications of chemical engineering should be included from this list:

Conservation laws for mass and energy

Fluid flow

Heat and mass transfer

Thermodynamics of chemical and power systems

Reactor design

Analysis of process systems


Other (please specify)

Analytical and quantitative processes used in reaching and evaluating the described solution.

Key calculations, to include:

How the basis of the calculation was selected and validated

How the calculation inputs were selected and validated

Relevant equations

3. Summary

Provide a summary and discussion of results

Provide conclusion(s)

4. Attestation

Please sign as the author:

Signature:

Job title:

This report is a true account of my professional activities:

Date:

5. Confirmation

Please also obtain the signature of someone familiar with your work at the time of activity and ideally a chemical engineer:

Family name: Given Name: Title:

Job title:

Affiliation (company/university or other organisation):

Relation to applicant:

Email:

To the best of my knowledge, this is a true account of the applicant’s activities

Signature:

Date:

6. Appendices

Provide inputs to computations and adequate explanation of background or supporting theory to worksheets or programmes.

Flowsheets, for example, process flow diagram (PFD), process and instrumentation diagrams (P&ID)

Engineering drawings

Diagrams

References to published texts, standards or papers which were used during the work described


Other (please specify):

Get Chartered – knowledge and understanding

Demonstrating knowledge and understanding

The purpose of a technical report is to assess your knowledge and understanding against the high standard expected of a Chartered Member (MIChemE).

When you submit a technical report for review you are providing an illustrative sample of your underpinning knowledge and understanding of chemical engineering principles through the specific context of your chosen project (addressing an objective, problem or development aim).

The subject of the report is left open for you to choose. You do not have to undertake a new piece of work for the purpose of your technical report. Typically it will be based on a chemical engineering project you have completed recently, as you will be required to expand on the detail and context of the report at interview.

You must ensure that the report includes an explanation of the chemical engineering principles you have applied and is not simply a commentary or description of the work completed.

If you have been advised through the self diagnosis tool or by an IChemE staff member to focus your technical report on further learning or on design, please read the guidance specific to the relevant report type.

Further information

The criteria IChemE refers to when assessing an individual’s knowledge and understanding against the level required for a Chartered Member are those used to characterise a degree for accreditation to Masters level, where we look for attainment of knowledge and understanding against defined standards. IChemE’s expectations for delivering this standard are set out in our accreditation guide.

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Guidance

The required level

IChemE upholds a minimum level of knowledge and understanding required for qualification as a Chartered Member (MIChemE).

As well as core chemical engineering knowledge, this includes:

■■ an advanced breadth and depth of chemical engineering knowledge and its application; and

■■ a sufficient capability to aggregate learning into a complex open ended project (such as design or research work).

The main areas of knowledge and understanding IChemE looks for when assessing an individual, or accrediting a degree against the standard required for a Chartered Chemical Engineer, are:

■■ underpinning maths and science;

■■ core chemical engineering;

■■ engineering practice;

■■ design practice;

■■ wider aspects eg sustainability, process safety, transferable skills, awareness of commercial implications;

■■ major design and/or research work;

■■ advanced depth and breadth of chemical engineering.

IChemE does not publish a definitive guide of the precise knowledge and understanding required because the content of degree programmes and/or an individual engineer’s learning experience will differ and yet may attain outcomes of an equally and acceptably high standard.

Check

Please use the online self diagnosis tool www.getchartered.org/selfdiagnosis to check what to provide as evidence of your chemical engineering knowledge and understanding in support of an application to get chartered and read any suggested further guidance or information.

If further clarification is required, a chargeable preliminary appraisal service is available.

Preliminary career appraisal

As an alternative or addition to the self diagnosis tool, IChemE offers a preliminary appraisal service, for which an administration fee is charged.

The following information is required:

■■ any academic qualifications including non-chemical engineering degree(s):

– degree title(s);

– university or college name(s);

– start and end dates of study (MM/YY – MM/YY);

■■ design/research project(s) details if relevant;

■■ relevant work-based training;

■■ work experience;

■■ current employment.

IChemE then checks:

■■ all accredited and recognised degree programmes;

■■ case history of degree programmes of which we have previous knowledge;


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and reviews:

■■ knowledge and understanding achieved against defined learning outcome standards.

It may be helpful to know that because national education systems vary across the world, our appraisal is not based on:

■■ degree title;

■■ number of years studied;

■■ country of study;

■■ method of learning.

Following this preliminary appraisal we will confirm whether any additional learning is advised, what evidence to submit in support of your application (eg a technical report), or whether we have sufficient existing evidence.

If the preliminary appraisal indicates a need for further learning in one or more specific areas, we will advise on how to address the gap(s). This may be:

■■ a recommendation of relevant formal learning modules; or

■■ approximate guidance on a particular area and duration of work experience you would yet need to gain in a chemical, biochemical or process engineering environment.

Please note that a career appraisal will only determine your application requirements and is not an application in itself. If you are able, we therefore strongly recommend that you follow the guidance provided through the self diagnosis tool.

Find out how to request a career appraisal www.icheme.org/careerappraisal

Providing evidence

You will need to submit evidence that your chemical engineering knowledge and understanding meet the minimum level for a Chartered Chemical Engineer, as indicated in the self diagnosis. The main ways in which we will ask you to provide evidence are as follows:

1. relevant degree certificate(s);

2. course transcript(s);

3. technical report;

4. technical report focused on demonstrating the ability to integrate knowledge into design or similar complex project eg research;

5. technical report focused on demonstrating further learning in industry;

6. a combination of two or more aspects.

A technical report would be a piece of written work, normally based on a project completed in industry. Academic faculty can present a collection of relevant academic papers. Further guidance is provided on the requirements for each report type.

You are strongly advised to use the self diagnosis tool to determine what specific evidence is required according to your individual circumstances.

Assessment

If a technical report is required this should be submitted as part of a full application (including evidence of professional experience www.GetChartered.org/professional_experience), except in the case of a technical report focused on further learning in industry. A technical report focused on further learning should be submitted for review and interview during your early career and before a full application to get chartered.

0618_12


Revamp of Saturated Gas Concentration Unit (SGCU)


Example A Technical Report

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Contents

1. Definitions .................................................................................................................................... 4

2. Introduction .................................................................................................................................. 5

3. Setup ............................................................................................................................................. 6

4. The process in SGCU ................................................................................................................... 7

5. The problem ................................................................................................................................. 9

6. My involvement ......................................................................................................................... 10

7. Analysis of the problem ............................................................................................................. 10 7.1. Operating Pressure........................................................................................................... 10

7.2. Equipment constraints in the low pressure operation mode ........................................ 11

8. Solution to the problem of low LPG recovery ........................................................................... 12 8.1. Operating Pressure........................................................................................................... 12

8.2. Column Performance ....................................................................................................... 14

8.3. Exchanger Performance .................................................................................................. 20

9. Conclusion .................................................................................................................................. 21 9.1. Equipment Performance .................................................................................................. 21

9.2. LPG Recovery ................................................................................................................... 21

9.3. Commissioning.................................................................................................................. 21

9.4. Economics .......................................................................................................................... 22


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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


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1. Definitions

ATU Amine Treating Unit

CDU Crude Distillation Unit – a part of the Crude and Vacuum Distillation Unit

CFC Clean Fuels Complex

DHDS Diesel Hydrodesulphurisation

FEED Front-End Engineering Design

HP High Pressure operation

HT Hydrotreater

H&MB Heat and Material Balance

ISBL Inside Battery Limit

KBPSD Kilo (Thousands of) Barrels Per Stream Day

LCO Light Cycle oil

LP Low Pressure operation

LPG Liquefied Petroleum Gas

MMTPA Mega (Millions of) Metric Tonnes Per Annum

P&ID Piping and Instrumentation Diagram

PFD Process Flow Diagram

SGCU Saturate Gas Concentration Unit

VDU Vacuum Distillation Unit – a part of the Crude and Vacuum Distillation Unit

VGO Vacuum Gas Oil


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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 2006-07 has been estimated at 10.57 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 2007-08 to 11.05 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 86.95 mole %. In view of the great commercial advantage, the task was to improve this recovery to >90 mole%.


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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/cm2g. However, in order to accommodate for the new streams from CFC, which are operating at a lower pressure of 9 kg/cm2g, 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/cm2g. The feed composition to the system is given in Attachment 1.

HP Cooler S01 A/B

C05 Ovhd Condenser

Rich Oil from Primary Absorber (C02)

145

161

CW

CLEAN FUELS COMPLEX

WILD NAPHTHA FROM CDU

WILD NAPHTHA FROM VGO HTU

WILD NAPHTHA FROM DHDS 1


WILD NAPHTHA FROM LCO HC

FEED PREPARATION

101

104

110

107

102

OFF-GAS FROM VGO HTU

103

111

106

112

108 OFF-GAS FROM DHDS 1

OFF-GAS FROM DHDS 2 HTU

OFF-GAS FROM LCO HC

CLEAN FUELS COMPLEX

Overhead vapour from Stripper (C05)

309

305 C05 OVHD Receiver (V12)

117


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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 H2S 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.


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HOT DIESEL FROM CDU

CIRCULATING DIESEL TO CDU

S14 304

307

323

S04 CW

181 180

229

225

145 P03 A/B

312

CW S12

187

313

Naphtha to OSBL

212

233

310

209

S09/S10

P06 A/B

MP Steam

Rich oil to CDU

182

142

P05 A/B

S03

309

210

S13

204

203

202

170

301

161

303 S05 S15

231

302 122

S08 A/B

V01

S11

A02

S06

CW

S01 A/B

C05

Feed Preparation

Condensate

C02

C03

P01 A/B

CW

CW

Fuel Gas to ATU

Lean oil from CDU

CW

LPG

C06

Fig 2, SGCU Process Schematic

CO2:- PRIMARY ABSORBER CO3:- SPONGE ABSORBER C05:- STRIPPER C06:- STABILIZER


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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 4.

5. The problem

The C3 & C4 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 C3 & C4 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 C3 & C4 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 C3 & C4 content of the various feed streams (including that carried in with the lean oil) is 789.35 kg-moles/hr. The C3 & C4 which is recovered from this content (with streams 187, 212, 316) is 686.05 kg-moles. Hence, the recovery is only 86.92 mole%. Due to introduction of the CFC vapour streams at a lower operating


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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/cm2g while that in the new refinery it is only 8.1 kg/cm2g.

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.


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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.)


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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/cm2g. As per ASME Section VIII, Div 1 and API 520 requirement, the highest operating pressure for V01 can be 14.85 kg/cm2g. 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/cm2a.

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/cm2a. This was assuming a suction pressure of 9.8 kg/cm2a. After, considering appropriate pressure drops in the system, the operating pressure of the V01 was found to be 15.2 kg/cm2a. The revised feed system would be:


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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 181.9 kg-mole/hr, which decreased to 107.9 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.

HP Cooler S01 A/B

C05 Ovhd Condenser

Rich Oil from Primary Absorber (C02)

145

161

CW

CLEAN FUELS COMPLEX

WILD NAPHTHA FROM CDU

WILD NAPHTHA FROM VGO HTU



WILD NAPHTHA FROM LCO HC

FEED PREPARATION

101

104

110

107

102

OFF-GAS FROM VGO HTU

103

111

106

112

108A OFF-GAS FROM DHDS 1

OFF-GAS FROM DHDS 2 HTU

OFF-GAS FROM LCO HC

CLEAN FUELS COMPLEX

HP RECEIVER V01

121

302 To Stripper

To Primary Absorber

CW

Overhead vapour from Stripper (C05)

309

305 C05 OVHD Receiver (V12)

307 301

117

CFC Off Gas Compressor

Fig 3, Feed Schematic in HP operation


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Columns Design Pressure Operating Pressure

(LP operation)

Operating Pressure

(HP operation)

Kg/cm2g Kg/cm2a Kg/cm2a

C02 16.5 8.75/9.05 14.8/15.1

C03 16.3 8.5/8.6 14.6/14.7

C05 17.7 16.1/16.4 16.5/16.8

C06 15.0 10.5/10.8 13.0/13.3

8.2. Column Performance

8.2.1. Primary Absorber (C02)

Primary absorber tray loadings for LP operation is given below.

NET FLOW RATES HEATER

TRAY TEMP PRESSURE LIQUID VAPOR FEED PRODUCT DUTIES

DEG C KG/CM2 KG-MOL/HR M*KCAL/HR

1 54.8 8.75 812.7 684.0L 861.4V 2 56.5 8.78 830.3 990.1 3 56.6 8.81 834.9 1007.7 4 56.2 8.84 836.4 1012.3 5 55.1 8.87 605.5 1013.8 231.8P 6 52.6 8.9 855.2 1014.7 231.8P -0.18437 53.5 8.93 861.8 1032.5 8 53.9 8.96 866.6 1039.2 9 53.9 8.99 870.9 1044

10 52.9 9.02 875.8 1048.2 11 49.8 9.05 1053.2 1068.0V 890.7L

0.1W

The operation at higher pressure level is given below



DEG C KG/CM2 KG-MOL/HR M*KCAL/HR1 60.8 14.8 852.1 684.0L 787.0V 2 63.3 14.83 879.9 955.1 3 63.7 14.86 887.7 982.8 4 63.4 14.89 890.5 990.6 5 62 14.92 649.3 993.5 242.7P 6 58.4 14.95 915.1 994.9 242.7P -0.2417 59.2 14.98 922.2 1018.1 8 59.7 15.01 927.1 1025.1 9 59.9 15.04 932.4 1030.1

10 59.2 15.07 940.7 1035.4


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11 55.6 15.1 1043.6 1073.5V 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 C03.

8.2.2. 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:




1 52.4 8.5 238.1 191.2L 771.4V 2 58 8.55 253.7 818.3 3 61.8 8.6 833.9 861.4V 281.2L

The summary of loads for high pressure operation is as follows:

---------- NET FLOW RATES HEATER


DEG C KG/CM2 KG-MOL/HR

M*KCAL/HR

1 50.9 14.6 244 191.2L 693.7V 2 56.3 14.65 257.4 746.5 3 62.5 14.7 759.9 787.0V 284.5L

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 1 1 191.19 0.7346FEED 310 VAPOR 3 0 861.38 2.7552PROD 143 VAPOR 1 771.41 2.1055PROD 313 LIQUID 3 281.15 1.3842

The summary of the column flows for the HP operation is given below


HEAT RATES

TRAY TRAY FRAC KG-MOL/HR M*KCAL/HRFEED 182 LIQUID 1 1 191.19 0.744FEED 310 VAPOR 3 0 786.96 2.317PROD 143 VAPOR 1 693.66 1.665PROD 313 LIQUID 3 284.49 1.396


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8.2.3. 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’

COLUMN SUMMARY NET FLOW RATES HEATER



1 109.5 16.1 4541.6 4437.6L 993.4V

2 110 16.12 4575 913.2 3 110.2 16.14 4594 946.6 4 110.5 16.16 4613.6 965.6 5 110.9 16.17 4640 985.1 6 111.5 16.19 4678.2 1011.6 7 112.4 16.21 4732.7 1049.8 8 113.5 16.23 4806.3 1104.3 9 114.8 16.25 4897.9 1177.9

10 116.3 16.27 5001.4 1269.5 11 117.9 16.29 5107.6 1372.9 12 119.6 16.31 5209.7 1479.2 13 121.5 16.33 5310 1581.3 14 124.3 16.34 5418.7 1681.6 15 129.6 16.36 5545.1 1790.3 16 141.1 16.38 5651.1 1916.6

17R 168.3 16.4 2022.7 3628.4L 16.0269

The summary of tray loads for high pressure operation is as follows: COLUMN SUMMARY NET FLOW RATES HEATER


DEG C KG/CM2 KG-MOL/HR

M*KCAL/HR

1 92.7 16.5 4845.9 4715.3L 1481.6V 2 93.3 16.52 4886.6 1154.2 3 93.7 16.54 4909.5 1194.9 4 94.1 16.56 4933.5 1217.8 5 94.7 16.58 4967 1241.7 6 95.5 16.59 5017.6 1275.3 7 96.6 16.61 5093.2 1325.9 8 98.3 16.63 5200.6 1401.4 9 100.3 16.65 5339.7 1508.9

10 102.5 16.67 5500 1648 11 104.9 16.69 5662.8 1808.3


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12 107.2 16.71 5813.8 1971.1 13 109.9 16.73 5951.6 2122.1 14 113.9 16.74 6086.8 2259.8 15 121 16.76 6222.6 2395 16 135.4 16.78 6289.3 2530.9

17R 166.8 16.8 2597.5 3691.7L 20.0952

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.

8.2.4. 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’

COLUMN SUMMARY NET FLOW RATES HEATER



1C 47 9.6 2299.8 679.1L -13.05392 66.2 10.5 2488 2978.8 3 71.9 10.51 2487 3167 4 75.3 10.52 2467.4 3166 5 77.9 10.53 2430.5 3146.5 6 80.6 10.54 2380.9 3109.5 7 83.8 10.55 2325.8 3059.9 8 87.5 10.56 2274.6 3004.8 9 91.4 10.57 2234.3 2953.7

10 95 10.58 2206.1 2913.4 11 98.2 10.59 2186.7 2885.2 12 100.8 10.6 2170.7 2865.7 13 102.9 10.61 2152.2 2849.7 14 104.9 10.62 2122.5 2831.2 15 107.2 10.63 2059.7 2801.6 16 110.9 10.64 1897.4 2738.7 17 119.5 10.65 1527.5 2576.4 347.2V

18 140.4 10.65 4808 1859.33281.2L

19 142.3 10.66 4889.1 1858.6 20 143.8 10.67 4952.5 1939.8 21 145.2 10.68 5004.8 2003.1 22 146.5 10.69 5051.5 2055.4 23 147.7 10.7 5096.3 2102.2


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24 149 10.71 5140.9 2146.9 25 150.3 10.72 5186 2191.5 26 151.6 10.73 5231 2236.6 27 153.1 10.74 5274.5 2281.6 28 154.9 10.75 5312.9 2325.1 29 157.4 10.76 5336.4 2363.5 30 162.1 10.77 5316.8 2387 31 172.3 10.78 5202.8 2367.4

32S 181.3 10.79 5262 102.4 2949.54L 33R 197.1 10.8 3111 2151 16.1217

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 17 0 347.19 3.585 FEED 170L LIQUID 18 1 3281.23 26.539 PROD 210 LIQUID 1 679.1 1.0413 PROD 316 WATER 1 0 PROD 225 LIQUID 32 2949.5 32.1505

OTHER PRODUCT STREAMS

TYPE STREAM PHASE FROM TO LIQUID FLOW RATES HEAT RATES

TRAY TRAY FRAC KG-MOL/HR M*KCAL/HR NET 209 LIQUID 1 2299.86 3.5267 TOTAL 203 VAPOR 2 2978.89 17.6218

The column summary in HP operation is given below:

Rigorous Column ‘321_C06’, ‘SGCU Stabiliser’ COLUMN SUMMARY NET FLOW RATES HEATER



1C 47 12.1 2632.6 726.5L -14.91082 73.7 13 3038.7 3359.1 3 79.9 13.01 3037.3 3765.2 4 83.8 13.02 3016.4 3763.8 5 86.7 13.03 2974 3742.9 6 89.6 13.04 2915 3700.5 7 92.9 13.05 2848.2 3641.6

8 96.7 13.06 2784.6 3574.7 9 100.7 13.07 2733.7 3511.1


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10 104.6 13.08 2697.6 3460.2 11 108 13.09 2672.6 3424.1 12 110.9 13.1 2652.1 3399.1 13 113.4 13.11 2628.1 3378.6 14 115.6 13.12 2588.5 3354.7 15 118.2 13.13 2503.5 3315 16 122.3 13.14 2292.2 3230 17 131.4 13.15 1850.4 3018.7 435.4V

18 151.7 13.15 5105.3 2141.53256.3L

19 153.9 13.16 5202.4 2140.1 20 155.6 13.17 5279.6 2237.1 21 157.1 13.18 5342.9 2314.4 22 158.5 13.19 5397.5 2377.7 23 159.8 13.2 5447.2 2432.3 24 161 13.21 5494.2 2482 25 162.2 13.22 5539.4 2529 26 163.5 13.23 5582.6 2574.2 27 164.9 13.24 5622.7 2617.4 28 166.7 13.25 5655.7 2657.4 29 169.4 13.26 5670.8 2690.5 30 174.3 13.27 5638.7 2705.6 31 184.9 13.28 5515.8 2673.4

32S 194.1 13.29 6029.6 100.5 2965.2L 33R 208.7 13.3 3579.5 2450.1 17.311

FEED AND PRODUCT STREAMS


HEAT RATES

TRAY TRAY FRAC KG-MOL/HR M*KCAL/HR FEED 170V VAPOR 17 0 435.44 4.7184 FEED 170L LIQUID 18 1 3256.29 28.9406 PROD 210 LIQUID 1 726.51 1.1042 PROD 316 WATER 1 0

PROD 225 LIQUID 322965.2

2 34.9549

OTHER PRODUCT STREAMS

TYPE STREAM PHASE FROM TO LIQUID

FLOW RATES

HEAT RATES

TRAY TRAY FRAC KG-MOL/HR M*KCAL/HR

NET 209 LIQUID 1 2632.7 4.0013TOTAL 203 VAPOR 2 3359.3 20.0164

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


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changes in the loads.

8.3. Exchanger Performance

While performing the study, I found that a few of the exchangers were limiting and hence did not allow any manipulation.

8.3.1. 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 1st exchanger. Also, this stream was at a low temperature of 39oC and hence the bypass was possible. But, after the introduction of the compressor, this compressed gas will come in at around 73oC (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 49oC. 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 S15.

8.3.2. 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 49oC,) 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 16.03 MMKcal/hr. After all the changes, I could decrease the stripper overhead temperature from C05 to 92oC from 109oC. The heat required for S06 is provided by MP Steam.

8.3.3. 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 m2 of effective heat transfer area (compared to 1546 m2 for S01 A/B). Within the


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limits of the exchanger, I have maximised the process flow through the exchanger to 18220 kg/hr from 16220 kg/hr. With increasing the flow further, I found that there were vibration problems in the shell side (process side).

8.3.4. 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.

9.2. 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.

9.3. 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


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without any lengthy downtime.

9.4. Economics

For the high pressure operation case, 38550.8 kg/hr of LPG is recovered as against 36536.7 kg/hr for the LP operating case. This means that by increasing the operating pressure, there was an increased recovery of 2014.1 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$ 890.59 / 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.


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Feed Composition Attachment 1 Stream Name 101 102 103 104 106 107 108 110 111 112 113

Phase Liquid Liquid Mixed Liquid Mixed Liquid Mixed Liquid Mixed Mixed Liquid Total Stream Rate KG-MOL/HR 2,052.8 277.4 245.0 165.5 257.2 368.7 539.8 160.9 153.6 129.1 807.0 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 288.2 29.5 15.9 22.5 14.1 49.3 31.4 19.8 9.1 8.2 98.6 Temperature C 51 39 54 73 40 39 39 39 40 40 39 Pressure KG/CM2 12.02 14.02 11.03 17.23 10.03 14.02 9.88 14.02 10.03 10.03 14.02 Molecular Weight 97.89 65.52 28.29 92.60 22.71 92.03 25.27 82.20 27.84 27.29 80.96 Mole Fraction Liquid 1.000 1.000 0.011 1.000 0.012 1.000 0.015 1.000 0.008 0.037 1.000 Sp. Gravity 0.7 0.6 0.4 0.7 0.4 0.7 0.4 0.7 0.5 0.4 0.7 Vapor Rate KG-MOL/HR 242.26 254.13 531.95 152.41 124.34 KG/HR 6,664 5,536 12,922 4,171 3,152 M3/HR 583.0 655.5 1378.9 386.1 316.6 Molecular Weight 27.51 21.78 24.29 27.37 25.35 Liquid Rate KG-MOL/HR 2,052.8 277.4 2.8 165.5 3.0 368.7 7.9 160.9 1.2 4.7 807.0 KG/HR 200,947 18,174 268 15,329 303 33,935 717 13,225 106 370 65,334 M3/HR 301.4 30.8 0.4 24.4 0.4 50.8 1.1 20.5 0.2 0.6 102.0 Molecular Weight 97.89 65.52 96.15 92.60 100.37 92.03 90.82 82.20 87.43 78.39 80.96 Total Molar Comp. Rates KG-MOL/HR H2O 4.45 0.36 2.06 0.78 1.57 0.42 3.48 0.19 1.04 0.87 0.96 AIR 5.64 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 H2 0.79 1.37 76.73 0.00 89.30 0.99 187.99 0.44 45.33 53.35 2.80 NH3 0.00 0.02 0.01 0.00 0.00 0.00 0.03 0.00 0.00 0.03 0.02 H2S 21.75 1.41 20.80 0.00 42.39 20.90 77.03 11.36 32.58 2.06 33.67 METHANE 1.41 1.37 24.59 0.00 45.34 3.29 72.01 0.88 15.50 11.17 5.54 ETHANE 26.45 5.91 48.60 0.00 38.30 14.13 62.66 3.84 14.14 10.22 23.88 PROPANE 120.75 42.52 24.17 0.00 23.36 27.88 72.78 21.36 25.29 24.14 91.76 PROPENE 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 IBUTANE 43.65 47.05 23.88 0.44 3.65 9.90 24.33 16.96 8.92 11.76 73.91 BUTANE 180.42 47.65 12.04 1.14 7.82 29.76 23.86 18.47 7.09 8.94 95.88 1BUTENE 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 IPENTANE 108.90 34.68 5.82 29.44 0.00 0.00 4.35 9.52 1.58 2.77 44.20 PENTANE 181.34 21.38 3.75 28.92 0.00 0.00 2.03 4.63 0.62 1.41 26.01


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C6+ 1357.24 73.66 2.60 104.82 5.43 261.47 9.30 73.23 1.54 2.34 408.37

Summary of streams for LP operating case Attachment 2

Stream No. 101 102 103 104 106 107 108 110 111 112 117 121 143 180 187 212 301 312

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 2,053 277 245 166 257 369 540 161 154 129 4,812 876 771 191 281 679 5,502 2,265

KG/HR 200,947 18,174 6,931 15,329 5,839 33,935 13,640 13,225 4,277 3,523 400,686 23,188 16,563 43,000 49,001 36,399 423,255 243,14

7 Temperature C 51 39 54 73 40 39 39 39 40 40 62 39 52 268 148 42 39 41 Pressure KG/CM2 12.02 14.02 11.03 17.23 10.03 14.02 9.88 14.02 10.03 10.03 10.21 9.10 7.73 11.00 6.45 16.82 9.10 10.24 Total Molar Comp. Rates KG-MOL/HR

H2O 4.45 0.36 2.06 0.78 1.57 0.42 3.48 0.19 1.04 0.87 13.86 6.97 7.21 0.51 0.61 0.00 18.21 0.00 AIR 5.64 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 6.94 5.43 5.57 0.01 0.09 0.00 6.98 0.00

H2 0.79 1.37 76.73 0.00 89.30 0.99 187.99 0.44 45.33 53.35 89.86 266.33 267.67 0.00 0.64 0.00 277.90 0.00

NH3 0.00 0.02 0.01 0.00 0.00 0.00 0.03 0.00 0.00 0.03 0.07 0.05 0.05 0.00 0.00 0.00 0.10 0.00

H2S 21.75 1.41 20.80 0.00 42.39 20.90 77.03 11.36 32.58 2.06 265.59 136.38 140.38 0.00 11.68 1.20 361.40 0.00 METHANE 1.41 1.37 24.59 0.00 45.34 3.29 72.01 0.88 15.50 11.17 55.10 98.98 102.17 0.00 1.38 0.00 127.80 0.00 ETHANE 26.45 5.91 48.60 0.00 38.30 14.13 62.66 3.84 14.14 10.22 243.08 145.71 152.94 0.00 7.79 0.86 318.24 0.00 PROPANE 120.75 42.52 24.17 0.00 23.36 27.88 72.78 21.36 25.29 24.14 457.88 120.42 77.11 0.06 11.72 220.71 562.54 0.00 PROPENE 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 IBUTANE 43.65 47.05 23.88 0.44 3.65 9.90 24.33 16.96 8.92 11.76 217.97 26.24 2.60 0.03 1.06 160.14 256.20 2.44 BUTANE 180.42 47.65 12.04 1.14 7.82 29.76 23.86 18.47 7.09 8.94 403.25 35.25 3.20 0.20 1.82 290.60 450.39 17.94 1BUTENE 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 IPENTANE 108.90 34.68 5.82 29.44 0.00 0.00 4.35 9.52 1.58 2.77 258.08 9.81 5.41 0.19 12.37 5.08 272.47 170.05 PENTANE 181.34 21.38 3.75 28.92 0.00 0.00 2.03 4.63 0.62 1.41 326.78 9.51 4.07 0.40 14.37 0.34 340.10 223.68 C6+ 1357.24 73.66 2.60 104.82 5.43 261.47 9.30 73.23 1.54 2.34 2473.49 14.88 3.00 189.77 217.63 0.14 2509.79 1851.4

LPG content 344.81 137.2

3 60.09 1.58 34.83 67.53 120.9

7 56.79 41.30 44.84 1079.10 181.9

1 82.91 0.29 14.60 671.4

5 1269.13 20.39 kg-mole/hr LPG in with feed 789.00 (With streams 101+102+103+104+107+108+110) LPG with Lean oil Stream 180 0.29 Total LPG in 789.29

LPG out 212 671.45 LPG recoveredout 212 671.45 kg-mole/hr

143 82.91 187 14.60 kg-mole/hr 312 20.39 686.05 kg-mole/hr 187 14.60


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789.35 % LPG recovered 86.92%

Stream No 126 145 161 170 182 202 203 204 209 225 229 231 233 302 303

Phase Liquid Liquid Vapor Liquid Liquid Mixed Vapor Liquid Liquid Liquid Liquid Liquid Liquid Liquid Liquid Total Stream

Rate KG-MOL/HR 4,622 891 993 3,629 191 3,629 2,979 2,979 2,300 2,949 2,949 684 684 4,622 4,622

KG/HR 399,988 81,321 47,044 352,958 43,000 352,958 159,674 159,674 123,275 316,559 316,559 73,412 73,412 399,988 399,988 Temperature C 110 50 109 168 44 148 66 47 47 181 123 123 43 40 57 Pressure KG/CM2 19.95 9.05 16.10 16.40 9.20 11.00 10.50 9.70 17.12 10.79 10.49 10.49 10.09 20.65 20.25 Total Molar Comp. Rates KG-MOL/HR

H2O 6.81 1.15 6.81 0.00 0.51 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 6.81 6.81 AIR 1.55 0.28 1.55 0.00 0.01 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 1.55 1.55

H2 11.57 1.96 11.57 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 11.57 11.57

NH3 0.05 0.01 0.05 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.05 0.05

H2S 225.02 42.71 223.84 1.20 0.00 1.20 5.25 5.25 4.05 0.00 0.00 0.00 0.00 225.02 225.02 METHANE 28.82 4.67 28.82 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 28.82 28.82 ETHANE 172.53 31.66 171.67 0.86 0.00 0.86 3.79 3.79 2.93 0.00 0.00 0.00 0.00 172.53 172.53 PROPANE 442.12 76.11 221.45 220.72 0.06 220.72 968.21 968.21 747.50 0.01 0.01 0.00 0.00 442.12 442.12 PROPENE 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 IBUTANE 229.96 32.42 66.66 163.32 0.03 163.32 702.48 702.48 542.32 3.18 3.18 0.74 0.74 229.96 229.96 BUTANE 415.14 47.46 101.21 313.96 0.20 313.96 1274.77 1274.77 984.14 23.36 23.36 5.42 5.42 415.14 415.14 1BUTENE 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 IPENTANE 262.66 46.00 36.21 226.47 0.19 226.47 22.29 22.29 17.21 221.39 221.39 51.34 51.34 262.66 262.66 PENTANE 330.59 61.18 39.05 291.56 0.40 291.56 1.51 1.51 1.16 291.21 291.21 67.53 67.53 330.59 330.59 C6+ 2494.92 545.11 84.49 2410.48 189.77 2410.48 0.63 0.63 0.48 2410.34 2410.34 558.97 558.97 2494.92 2494.92

LPG content 1087.22 155.99 389.3

2 698.00 0.29 698.00 2945.47 2945.47 2273.96 26.55 26.55 6.16 6.16 1087.22 1087.22


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Stream No 304 308 309 310 313 Phase Vapor Vapor Vapor Vapor Mixed

Total Stream Rate 343 1,068 651 861 281 16,220 30,474 30,824 22,565 49,001 Temperature 109 39 109 55 62 Pressure 16.10 9.05 16.10 8.60 6.80 Total Molar Comp. Rates

H2O 2.35 8.45 4.46 7.30 0.61 AIR 0.53 5.92 1.02 5.64 0.09

H2 3.99 270.26 7.58 268.31 0.64

NH3 0.02 0.07 0.03 0.06 0.00

H2S 77.18 194.76 146.66 152.06 11.68 METHANE 9.94 108.22 18.89 103.55 1.38 ETHANE 59.19 192.40 112.48 160.73 7.79 PROPANE 76.35 164.87 145.10 88.76 11.72 PROPENE 0.00 0.00 0.00 0.00 0.00 IBUTANE 22.98 35.32 43.68 3.64 1.06 BUTANE 34.90 46.86 66.32 4.82 1.82 1BUTENE 0.00 0.00 0.00 0.00 0.00 IPENTANE 12.48 12.25 23.72 17.59 12.37 PENTANE 13.46 11.69 25.59 18.04 14.37 C6+ 29.13 17.00 55.36 30.86 217.63 LPG content 134.23 247.05 255.09 97.22 14.60


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Attachment 3 Summary of streams for HP operating case

Stream No. 101 102 103 104 106 107 108A 110 111 112 117 121 143 180 187A 212 301 Phase Liquid Mixed Mixed Liquid Mixed Liquid Vapor Liquid Mixed Mixed Mixed Vapor Vapor Liquid Mixed Liquid Mixed Total Stream Rate KG-MOL/HR 2053 277 245 166 514 369 532 161 307 258 5824 755 694 191 285 727 5936 KG/HR 200947 18174 6931 15329 11679 33935 12912 13225 8555 7045 429617 17301 13104 43000 48358 38556 436131 Temperature C 51 39 54 73 40 39 73 39 40 40 65 39 49 268 150 42 39 Pressure KG/CM2 16.5 16.5 16.5 17.23 10.03 16.5 15.944 16.5 10.03 10.03 15.944 15.2 8.43 11 12.55 11.8 15.3 Total Molar Comp. Rates KG-MOL/HR H2O 4.45 0.36 2.06 0.78 3.13 0.42 3.47 0.19 2.08 1.75 18.17 3.58 4.07 0.51 0.67 0.00 18.45 AIR 5.64 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 8.49 5.28 5.50 0.01 0.16 0.00 8.52 H2 0.79 1.37 76.73 0.00 178.61 0.99 187.96 0.44 90.67 106.70 289.92 264.78 267.08 0.00 1.21 0.00 289.97 NH3 0.00 0.02 0.01 0.00 0.00 0.00 0.03 0.00 0.00 0.06 0.14 0.05 0.05 0.00 0.01 0.00 0.14 H2S 21.75 1.41 20.80 0.00 84.79 20.90 76.63 11.36 65.16 4.12 534.77 125.02 130.44 0.00 21.19 1.27 551.88 METHANE 1.41 1.37 24.59 0.00 90.68 3.29 71.94 0.88 31.01 22.33 153.48 95.34 100.97 0.00 2.52 0.00 154.06 ETHANE 26.45 5.91 48.60 0.00 76.60 14.13 62.43 3.84 28.27 20.45 453.42 134.74 146.90 0.00 13.97 0.52 464.35 PROPANE 120.75 42.52 24.17 0.00 46.71 27.88 71.97 21.36 50.58 48.28 620.84 76.37 33.83 0.06 10.19 264.70 644.00 PROPENE 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 IBUTANE 43.65 47.05 23.88 0.44 7.30 9.90 23.68 16.96 17.84 23.52 241.72 13.61 0.38 0.03 0.34 162.11 250.79 BUTANE 180.42 47.65 12.04 1.14 15.64 29.76 23.02 18.47 14.19 17.89 422.16 17.95 1.26 0.20 1.67 292.01 437.52 1BUTENE 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 IPENTANE 108.90 34.68 5.82 29.44 0.00 0.00 4.01 9.52 3.15 5.55 267.81 5.17 1.43 0.19 10.61 5.35 275.11 PENTANE 181.34 21.38 3.75 28.92 0.00 0.00 1.83 4.63 1.24 2.82 333.89 4.99 1.00 0.40 11.63 0.45 342.13 C6+ 1357.24 73.66 2.60 104.82 10.86 261.47 4.83 73.23 3.07 4.67 2479.02 7.81 0.85 189.77 210.33 0.16 2498.71 Total 2052.79 277.39 245.04 165.54 514.31 368.72 531.81 160.88 307.25 258.13 5823.83 754.68 693.76 191.19 284.50 726.56 5935.65 LPG content 344.81 137.23 60.09 1.58 69.66 67.53 118.67 56.79 82.60 89.68 1284.72 107.93 35.48 0.29 12.20 718.82 1332.32 Kg-mole/hr

LPG in with feed 786.70 (With streams 101+102+103+104+107+108A+110) LPG with Lean Oil Stream 180 0.29 LPG in 786.99 LPG out 212 718.82 LPG recovered 212 718.82 143 35.48 187 12.20 312 20.53 Total 731.02 187 12.20 787.03 % LPG recovered 92.89%


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Stream No. 126 145 161 170 182 202 203 204 209 225 229 231 233 302A 303 Phase Mixed Liquid Mixed Liquid Liquid Mixed Vapor Liquid Liquid Liquid Liquid Liquid Liquid Liquid Liquid Total Stream

Rate KG-MOL/HR 5173 971 1482 3692 191 3692 3359 3359 2633 2965 2965 684 684 5173 5173

KG/HR 418694 83810 62573 356114 43000 356114 178264 178264 139708 317559 317559 73254 73254 418694 418694 Temperature C 92 56 92 167 44 161 74 47 47 194 135 135 47 40 40 Pressure KG/CM2 18.40 15.10 16.50 16.80 9.20 13.65 13.00 12.20 12.10 13.29 12.99 12.99 19.17 29.85 18.70 Total Molar Comp. Rates

KG-MOL/HR

H2O 7.37 1.23 7.37 0.00 0.51 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 7.37 7.37 AIR 3.24 0.55 3.24 0.00 0.01 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 3.24 3.24 H2 25.18 3.79 25.18 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 25.18 25.18 NH3 0.10 0.02 0.10 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.10 0.10 H2S 426.87 80.26 425.58 1.27 0.00 1.27 5.87 5.87 4.60 0.00 0.00 0.00 0.00 426.87 426.87 METHANE 58.72 8.36 58.72 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 58.72 58.72 ETHANE 329.61 58.79 329.08 0.52 0.00 0.52 2.39 2.39 1.87 0.00 0.00 0.00 0.00 329.61 329.61 PROPANE 567.63 97.52 302.87 264.71 0.06 264.71 1223.86 1223.86 959.16 0.01 0.01 0.00 0.00 567.63 567.63 PROPENE 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 IBUTANE 237.18 25.50 71.47 165.70 0.03 165.70 749.53 749.53 587.42 3.59 3.59 0.83 0.83 237.18 237.18 BUTANE 419.58 35.62 104.47 315.09 0.20 315.09 1350.10 1350.10 1058.09 23.09 23.09 5.33 5.33 419.58 419.58 1BUTENE 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 IPENTANE 269.95 49.29 36.90 233.04 0.19 233.04 24.75 24.75 19.40 227.69 227.69 52.52 52.52 269.95 269.95 PENTANE 337.15 64.54 38.79 298.34 0.40 298.34 2.07 2.07 1.62 297.90 297.90 68.72 68.72 337.15 337.15 C6+ 2490.91 545.98 77.85 2413.04 189.77 2413.04 0.72 0.72 0.56 2412.88 2412.88 556.60 556.60 2490.91 2490.91


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Stream No. 304 306 307 308 309 310 313 Phase Mixed Mixed Mixed Mixed Mixed Vapor Mixed Total Stream Rate 431 320 112 1075 1050 787 285 18220 11716 6504 29017 44353 18461 48358 Temperature 92 45 45 39 92 61 63 Pressure 16.50 16.25 16.25 15.20 16.50 14.70 12.90 Total Molar Comp. Rates H2O 2.15 1.87 0.27 5.46 5.23 4.23 0.67 AIR 0.94 0.91 0.04 6.19 2.30 5.64 0.16 H2 7.33 7.29 0.04 272.07 17.85 268.28 1.21 NH3 0.03 0.03 0.00 0.07 0.07 0.06 0.01 H2S 123.92 106.87 17.05 231.89 301.66 151.63 21.19 METHANE 17.10 16.52 0.58 111.86 41.62 103.49 2.52 ETHANE 95.82 84.93 10.89 219.67 233.26 160.87 13.97 PROPANE 88.19 65.11 23.09 141.48 214.68 43.96 10.19 PROPENE 0.00 0.00 0.00 0.00 0.00 0.00 0.00 IBUTANE 20.81 11.75 9.06 25.36 50.66 0.69 0.34 BUTANE 30.42 15.08 15.34 33.02 74.05 2.73 1.67 1BUTENE 0.00 0.00 0.00 0.00 0.00 0.00 0.00 IPENTANE 10.74 3.45 7.29 8.62 26.15 11.85 10.61 PENTANE 11.30 3.06 8.24 8.04 27.50 12.22 11.63 C6+ 22.67 2.98 19.69 10.79 55.18 21.41 210.33


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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

10.75 mol% 5.11 mole%

2) LPG ex Sat Gas Unit (Stream 212)



LPG Recovery from feed Target > 90% 89.2 mole% 92.9 mole%

LPG Vapour Pressure @ 37.8oC (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. 1056 ppm wt. 1054 ppm wt.

3) Full Range Naphtha ex. Sat-Gas Unit (Stream 312)



C4- content < 1% mol 0.9 mole% 0.9 mole%

D86 ASTM 95% Resultant >155°C, <180°C 156.8oC 157oC


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Economic evaluation Attachment 5

LP operating case Stream No. 101 102 103 104 106 107 108 110 143 180 187 212 312

Phase Liquid Liquid Mixed Liquid Mixed Liquid Mixed Liquid Vapor Liquid Mixed Liquid Liquid Total Stream Rate KG-MOL/HR 2,053 277 245 166 257 369 540 161 771 191 281 679 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 36,399 243,147 Temperature C 51 39 54 73 40 39 39 39 52 268 148 42 41 Pressure KG/CM2 12.02 14.02 11.03 17.23 10.03 14.02 9.88 14.02 7.73 11.00 6.45 16.82 10.24 Total Molar Comp. Rates KG-MOL/HR H2O 4.45 0.36 2.06 0.78 1.57 0.42 3.48 0.19 7.21 0.51 0.61 0.00 0.00 AIR 5.64 0.00 0.00 0.00 0.00 0.00 0.00 0.00 5.57 0.01 0.09 0.00 0.00 H2 0.79 1.37 76.73 0.00 89.30 0.99 187.99 0.44 267.67 0.00 0.64 0.00 0.00 NH3 0.00 0.02 0.01 0.00 0.00 0.00 0.03 0.00 0.05 0.00 0.00 0.00 0.00 H2S 21.75 1.41 20.80 0.00 42.39 20.90 77.03 11.36 140.38 0.00 11.68 1.20 0.00 METHANE 1.41 1.37 24.59 0.00 45.34 3.29 72.01 0.88 102.17 0.00 1.38 0.00 0.00 ETHANE 26.45 5.91 48.60 0.00 38.30 14.13 62.66 3.84 152.94 0.00 7.79 0.86 0.00 PROPANE 120.75 42.52 24.17 0.00 23.36 27.88 72.78 21.36 77.11 0.06 11.72 220.71 0.00 PROPENE 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 IBUTANE 43.65 47.05 23.88 0.44 3.65 9.90 24.33 16.96 2.60 0.03 1.06 160.14 2.44 BUTANE 180.42 47.65 12.04 1.14 7.82 29.76 23.86 18.47 3.20 0.20 1.82 290.60 17.94 1BUTENE 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 IPENTANE 108.90 34.68 5.82 29.44 0.00 0.00 4.35 9.52 5.41 0.19 12.37 5.08 170.05 PENTANE 181.34 21.38 3.75 28.92 0.00 0.00 2.03 4.63 4.07 0.40 14.37 0.34 223.68 C6+ 1357.24 73.66 2.60 104.82 5.43 261.47 9.30 73.23 3.00 189.77 217.63 0.14 1851.37

LPG content KG-MOL/HR 344.81 137.23 60.09 1.58 34.83 67.53 120.9

7 56.79 82.91 0.29 14.60 671.4

5 20.39 kg/hr 18309 7364 3147 92 1693 3526 5997 2995 3729 16 683 35854 1182

Kg-mole/hr kg/hr LPG in with feed 789.00 41429.54 (With streams 101+102+103+104+107+108+110) LPG with Lean oil Stream 180 0.29 16 Total LPG in 789.29 41445.54 LPG out 212 671.45 35854 LPG out 212 35854.11 kg/hr 143 82.91 3729 187 682.57 kg/hr 312 20.39 1182 36536.67 kg/hr 187 14.60 683


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789.35 41448.54

HP operating case Stream Name 101 102 103 104 106 107 108A 110 143 180 187A 212 312 Phase Liquid Mixed Mixed Liquid Mixed Liquid Vapor Liquid Vapor Liquid Mixed Liquid Liquid Total Stream Rate KG-MOL/HR 2053 277 245 166 514 369 532 161 694 191 285 727 2281 KG/HR 200947 18174 6931 15329 11679 33935 12912 13225 13104 43000 48358 38556 244304 Temperature C 51 39 54 73 40 39 73 39 49 268 150 42 42 Pressure KG/CM2 16.5 16.5 16.5 17.23 10.03 16.5 15.9444 16.5 8.43 11 12.55 11.8 12.7103 Total Mol. Comp. Rates KG-MOL/HR H2O 4.45 0.36 2.06 0.78 3.13 0.42 3.47 0.19 4.07 0.51 0.67 0.00 0.00 AIR 5.64 0.00 0.00 0.00 0.00 0.00 0.00 0.00 5.50 0.01 0.16 0.00 0.00 H2 0.79 1.37 76.73 0.00 178.61 0.99 187.96 0.44 267.08 0.00 1.21 0.00 0.00 NH3 0.00 0.02 0.01 0.00 0.00 0.00 0.03 0.00 0.05 0.00 0.01 0.00 0.00 H2S 21.75 1.41 20.80 0.00 84.79 20.90 76.63 11.36 130.44 0.00 21.19 1.27 0.00 METHANE 1.41 1.37 24.59 0.00 90.68 3.29 71.94 0.88 100.97 0.00 2.52 0.00 0.00 ETHANE 26.45 5.91 48.60 0.00 76.60 14.13 62.43 3.84 146.90 0.00 13.97 0.52 0.00 PROPANE 120.75 42.52 24.17 0.00 46.71 27.88 71.97 21.36 33.83 0.06 10.19 264.70 0.01 PROPENE 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 IBUTANE 43.65 47.05 23.88 0.44 7.30 9.90 23.68 16.96 0.38 0.03 0.34 162.11 2.76 BUTANE 180.42 47.65 12.04 1.14 15.64 29.76 23.02 18.47 1.26 0.20 1.67 292.01 17.76 1BUTENE 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 0.00 IPENTANE 108.90 34.68 5.82 29.44 0.00 0.00 4.01 9.52 1.43 0.19 10.61 5.35 175.17 PENTANE 181.34 21.38 3.75 28.92 0.00 0.00 1.83 4.63 1.00 0.40 11.63 0.45 229.18 C6+ 1357.24 73.66 2.60 104.82 10.86 261.47 4.83 73.23 0.85 189.77 210.33 0.16 1856.28 Total 2052.79 277.39 245.04 165.54 514.31 368.72 531.81 160.88 693.76 191.19 284.50 726.56 2281.16 LPG content Kg-mole/hr 344.81 137.23 60.09 1.58 69.66 67.53 118.67 56.79 35.48 0.29 12.20 718.82 20.53 Kg/hr 18308.57 7363.76 3146.84 91.65 3386.10 3526.44 5875.12 2994.91 1583.88 16.00 565.07 37985.70 1190.61

kg-mole/hr kg/hr

LPG in with feed 786.70 41307.28 LPG with Lean Oil Stream 180 0.29 16.00 LPG in 786.99 41323.28 kg/hr LPG out 212 718.82 37985.70 LPG recovered 212 37985.70 143 35.48 1583.88 187 565.07 312 20.53 1190.61 Total 38550.77

187 12.20 565.07 Extra LPG recovered than LP 2014.10 which @ US$467/ton is 940.58 US$/hr


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787.03 41325.26 Cost

A Dynamic Simulation Study: Dehydration Tanks TK-60A and TK-60B

Inter-tank Level Balancing Lines

Summary: The effectiveness of the intended tank balance lines was investigated by performing simulations to determine the effect of upstream transients on the dehydration tank levels. A number of upset operating cases were simulated at both 17% and 35% WC and maximum tank level variations calculated. The results show that the current design is adequate when compared to operating with two extra hypothetical oil and water balance lines.

1. INTRODUCTION

As part of an Onshore Crude Processing Facilities Project, two new dehydration tanks (TK-60A and TK-60B) will be installed. Each tank has a capacity of 28,000 bbl and will be used to separate oil and water phases before further downstream processing. Two 24” booster pump suction lines (crude and water) are currently planned to be used for tank level balancing. A separate 6” emulsion balance line and 20” vapour balance line have also been provided.

Fig 1.Dehydration Tank Balance Lines

2. OBJECTIVE

The objective of this study is to determine the effectiveness and suitability of the current design under different dynamic operating scenario’s. In particular, the variation in tank levels as a result of a lower offshore production rate (80,000 BPD) in one of the two separate process trains will be examined.

3. GENERAL TANK LEVEL EQUATIONS

Well established physical models for a typical gravity flow tank will form the basis for dynamic tank level modelling

purposes

p

cF

AgK

hLg

dtdv

ρ−= (1)

FFdtdhA ot −= (2)

4. MODELLING ROUTE

Dynamic tank level modelling was carried out using the high fidelity hydraulic simulator – TLNET, which is a component of Energy Solutions’ Pipeline Studio V 3.0. The following assumptions are made for the transient simulation:

• Tank internal pressures are the same (i.e. The gas vapour balance line has not been modelled)

• Emulsion and water phases have not been modelled, instead only one oil phase is assumed to be present within the tank.

• Independent tank objects are not available within the TLNET simulation environment, however it is possible to model a tank by allocating one or more inlet and/or outlet nozzles to a predefined tank. Therefore, in order to simulate balancing lines, a number of outlet nozzles at different elevations are implemented. A representation of the oil balance line is shown below, a tank nozzle is depicted by a blue object, flow sinks (pumps, nozzles, etc) are shown as green objects.

Tank60A_Oil_Outlet

Pipe0026

Tank60B_Oil_Outlet

Pipe0027

P-060A/B/D/EPipe0001

Emulsion

• The same flowrate to the desalters is maintained at all times.

• A fixed simulation time-step of 0.05 seconds was used.

• No hydraulic verification has been performed on the inlet lines upstream of the dehydration tank, i.e. it is assumed that the dynamic flowrates are achievable.

5. SIMULATION ENVIRONMENT

The process flow scheme, as shown in the below figure, was configured to represent the dehydration tank crude inlet manifold. The two segregated production trains (A and B) with crossover line valves closed is the normal expected operating scenario. However, other scenario’s, where the crossover valves are open, have also been simulated.

Pipe0004 Pipe0005

Pipe0006 Pipe0007 Pipe0008

Pipe0012

Pipe0013Blkv0001

Pipe0014 Pipe0015

Pipe0016

Pipe0017Blkv0002

Pipe0018

Pipe0019 Regu0001

Regu0002

Pipe0020

Pipe0021

Tank60A_Inlet

Tank60B_inlet

Blkv0003

Blkv0004Pipe0023

Supply0003

Supply0004

Pipe0024 Pipe0025

Tank60A_Oil_Outlet

Pipe0026

Tank60B_Oil_Outlet

Pipe0027

P-060A/B/D/

Tank60A_Water_Outlet

Tank60B_Water_Outlet

Pipe0029

Pipe0030

P-066A/B

Pipe0001

Pipe0003

TK-60B_EmulsionTK-60A_Emulsion

6" Emulsion Balance Line

5.1 Case A – Offshore Flow Disruption-Failure of Emulsion Balance Line - 17% Water Cuts

At steady state conditions, the total offshore rate 365,000 BPD is equally split between two segregated process trains. At the start of the simulation the offshore rate to train A is cut to only 80,000 BPD, train B maintains 182,500 BPD. The initial tank level is 15m. Oil outlet nozzles have been set at 13m, and water outlet nozzles at 3m. The total dynamic simulation run time is one hour and twenty minutes under these condition. The results presented below do not include the use of the 6” emulsion balance line.

13.0

13.2

13.4

13.6

13.8

14.0

14.2

14.4

14.6

14.8

15.0

0 500 1000 1500 2000 2500 3000 3500 4000 4500 5000

Tank Level TK-60A

M

Elapsed Time (seconds)

Fig. 2. Case A: TK-60A dynamic level variation

13.4

13.6

13.8

14.0

14.2

14.4

14.6

14.8

15.0

0 500 1000 1500 2000 2500 3000 3500 4000 4500 5000

TK-60BTank Level Variation

M

Elapsed Time (seconds) F

Fig. 3. Case A: TK-60B dynamic level variation

Under these conditions, the maximum difference in level is approximately 0.4m. Further tank level equalisation could be achieved, by opening one of the crossover line valves and reducing the set-point of one of the PCV’s.

5.2 Case B – Offshore Flow Disruption- Emulsion Balance Line Operational-17% Water Cut

This scenario follows the same operational pattern and dynamic disruption as case A, however it is assumed that the 6” emulsion balance line is now in operation.

TK-60ATank Level Variation

M

Elapsed Time (seconds)0 1000 2000 3000 4000 5000

13.013.213.413.613.814.014.214.414.614.815.0

Fig. 4. Case A: TK-60A dynamic level variation

13.2

13.4

13.6

13.8

14.0

14.2

14.4

14.6

14.8

15.0

0 1000 2000 3000 4000 5000


M


Fig. 5. Case B: TK-60B dynamic level variation

In this case it is observed that the maximum difference in level is 0.3m. The expected flow rate through the 6” line is approximately 80 m3/h.

5.3 Case C – Offshore Flow Disruption- Two Additional Balance Lines - 17% WC

This case follows the same operational scenario as A and B. Two additional 24” balance lines have been provided at 12m and 1m.

13.2

13.4

13.6

13.8

14.0

14.2

14.4

14.6

14.8

15.0

0 1000 2000 3000 4000 5000


M


Fig. 6. Case C: TK-60A dynamic level variation

13.2

13.4

13.6

13.8

14.0

14.2

14.4

14.6

14.8

15.0

0 1000 2000 3000 4000 5000


M


Fig. 7. Case C: TK-60B dynamic level variation

5.4 Case D – Train A to Train B Crossover Valve Open -17% WC

During steady state conditions, the two production trains are running with equal flowrates to both dehydration tanks. Both of the crossover valves are closed and tank levels are equal at 15m. The transient scenario maintains the full offshore rate (365,000 BPD), however 80,000 BPD is sent to train A whilst train B receives the remaining 285,000 BPD. This is simulated by fully opening one of the cross over valves (as would be the case with an MOV) and throttling the PCV upstream of dehydration tank A so that only 80,000 BPD is allowed the pass through the valve.

14.90

14.91

14.92

14.93

14.94

14.95

14.96

14.97

14.98

14.99

15.00

0 500 1000 1500 2000 2500 3000 3500 4000 4500 5000


M


Fig. 7. Case D: TK-60A dynamic level variation

15.00

15.01

15.02

15.03

15.04

15.05

15.06

15.07

15.08

15.09

15.10

0 500 1000 1500 2000 2500 3000 3500 4000 4500 5000


M


Fig. 8. Case D: TK-60B dynamic level variation

5.4 Case E – Train A to Train B Crossover Valve Open – Additional Balance Lines -17% Water Cut

The assumed scenario is the same as Case D, however two additional balance lines at 12m and 1m have been included. The results are shown below:

14.982

14.984

14.986

14.988

14.990

14.992

14.994

14.996

14.998

15.000

0 500 1000 1500 2000 2500 3000 3500 4000 4500 5000


M


Fig. 9. Case E: TK-60A dynamic level variation

15.000

15.002

15.004

15.006

15.008

15.010

15.012

15.014

15.016

15.018

0 500 1000 1500 2000 2500 3000 3500 4000 4500 5000


M


Fig. 10. Case E: TK-60B dynamic level variation

Tank levels equalize rapidly within 500 seconds with a slight and insignificant offset.

5.5 Case F – Train A to Train B Crossover Valve Open – 35% Water Cut

This simulation scenario is based on 35% water cut or a total offshore rate of approximately 467,000 BPD. It is expected that this case represents the worst case conditions in terms of tank level unbalancing. During transient conditions, dehydration tank A receives 80,000 BPD, whilst tank B receives 387,000 BPD. The flow re-routing is carried out in by opening only one MOV and throttling the PCV on the inlet of dehydration tank A.

14.80

14.82

14.84

14.86

14.88

14.90

14.92

14.94

14.96

14.98

15.00

0 500 1000 1500 2000 2500 3000 3500 4000 4500 5000


M


Fig. 11. Case F: TK-60A dynamic level variation

15.00

15.02

15.04

15.06

15.08

15.10

15.12

15.14

15.16

15.18

15.20

0 500 1000 1500 2000 2500 3000 3500 4000 4500 5000


M


Fig. 12. Case F: TK-60B dynamic level variation

As can be seen from the simulation results, the maximum difference in tank level is 0.4m whilst operating under these conditions. The dynamic simulation results show that the level-flow cascade controller averaging these two levels should still send the same signal to the existing flow controllers thus allowing the desalter rate to be maintained.

5.6 Case G – Train A to Train B Crossover Valve Open – Additional Balance Lines -35% Water Cut

The final case, simulates two 24” hypothetical balance lines at 12m and 1m. The implemented dynamic disruption is the same as case F.

14.975

14.980

14.985

14.990

14.995

15.000

0 500 1000 1500 2000 2500 3000 3500 4000 4500 5000


M


Fig. 13. Case G: TK-60A dynamic level variation

15.000

15.005

15.010

15.015

15.020

15.025

0 500 1000 1500 2000 2500 3000 3500 4000 4500 5000


M


Fig. 14. Case G: TK-60B dynamic level variation

6. CONCLUSIONS

The simulation results demonstrate that under certain and non-frequent operating conditions a maximum tank level imbalance of 0.4m may occur. The introduction of two extra 24” balance lines (intended for the oil and water phases) can improve the situation. However, since the tank level difference is not considered to be significant under the worst case conditions, further consideration of additional lines or a flow inlet balancing system is not required.


Design of a Fuel Gas Treatment and Distribution System


CONTENTS PAGE

ABBREVIATIONS 1.1

1. INTRODUCTION 1.2

2. SCOPE OF WORK AND BASE DATA 2.22.1 Scope of Work 2.22.2 Process Description 2.22.3 Base Data 2.2

3. TECHNICAL ASSESSMENT 3.13.1 Project Constraints 3.13.2 Project Approach 3.1

4. DESIGN AND CALCULATIONS SUMMARY 4.24.1 Assessment of Fuel Gas Demand 4.24.2 Fuel Gas Simulation 4.44.3 Fuel Gas Heater Review 4.54.4 Fuel Gas Scrubber Review 4.64.5 Process Control 4.64.6 Process Safeguarding 4.74.7 Line Sizing 4.10

5. CONCLUSIONS AND RECOMMENDATIONS 5.125.1 Conclusions 5.125.2 Recommendations 5.12

6. REFERENCES 6.13

APPENDICES 6.14

A1: FUEL GAS SYSTEM - PFD 6.14

A2 - FUEL GAS BALANCE 6.15

B1 - CALCULATION: Fuel Gas Scrubber 6.16

B2 - CALCULATION: Pressure Reduction Station 6.17

B3 - CALCULATION: Pressure Safety Valves 6.18

B4 - CALCULATION: Fuel Gas Lines 6.20

C1 - BLOWDOWN VOLUME 6.21

C2 - BLOWDOWN: DISPERSION 6.22

C3 - RADIATION: ISOPLETHS 6.23

.

1.1

ABBREVIATIONS

BDV Blowdown Valve DP Design Pressure EOS Equation of State HTC Heat Transfer Coefficient ID Inside Diameter LFL Lower Flammability Limits LHV Lower Heating Value NB Nominal Bore OD Outside Diameter OP Operating Point PCV Pressure Control Valve PFD Process Flow Diagram PID Proportional, Integral, Derivative Controllers PSV Pressure Safety Valve SP Set Point SRK Soave-Redlich-Kwong TCV Temperature Control Valve UFL Upper Flammability Limits WFL Woodhill Frontier

.

2.2

1. INTRODUCTION

This report presents the design of a fuel gas processing and distribution system for an onshore oil pumping facility operated by XXX Petroleum Co.

The pumping station will be designed for a throughput of 60,000bbls of oil per day and includes processing units such as; storage tanks for oil and water, pumps, water bath heaters, metering skid and other utilities and services such as power generation, plant and instrument air, industrial water systems and drain systems.

Fuel gas is the primary fuel source for power generation and the water bath heaters and is imported by pipeline from a nearby facility.

2. SCOPE OF WORK AND BASE DATA

2.1 Scope of Work

The fuel gas will be required to be treated to ensure that the fuel gas specification requirements advised by the heater and generator vendors is met, such that no power loss occurs due to fuel gas quality. The system in designed for 2.5MMscfd gas.

The scope of work is to design the fuel gas treatment and distribution for the pumping station. This includes review of fuel gas heater and scrubber, heat/material balance, line sizing and associated instruments, etc.

2.2 Process Description

A process flow diagram of the fuel gas system is given in Appendix A1. The incoming fuel gas will be heated to 45o

After pressure reduction, water/hydrocarbon condensates are removed in a scrubber before the gas is heated to about 15

C in a heating coil in a water bath heater to prevent hydrate formation before pressure reduction to the required distribution pressure.

o

The water bath heater will be dual-fired by fuel gas. Diesel will be used for black-start, when fuel gas is not available.

C by a second heating coil in the same heater, prior to distribution. Separated liquids from the scrubber will flow via level control to the produced water tank.

2.3 Base Data

The base data for the fuel gas system is summarised below.

2.3.1 Temperatures and Pressures

• Fuel gas system design pressure (inlet system) 142 barg • Fuel gas system design pressure (distribution) 6.9 barg • Fuel gas system design temperature (inlet) 80 o

• Fuel gas system design temperature (distribution) 100 C o

• Normal operating pressure (inlet system) 100 barg C

• Normal operating pressure (distribution) 5.7 barg • Minimum fuel gas temperature to users 10 o

• Maximum fuel gas temperature to users 50 C

o

C

.

2.3

2.3.2 Fuel Gas Composition and Fluid Properties

The fuel gas is expected to be supplied to the users at 100barg and at 25oC with the following composition listed in Table 1.

Component

Table 1: Fuel Gas Composition

Mol % Mass % N2 0.37 0.55

CO2 1.06 2.46 C1 89.64 75.84 C2 3.52 5.58 C3 2.77 6.44

i-C4 0.68 2.08 n-C4 1.24 3.80 i-C5 0.42 1.60 n-C5 0.07 0.27 C6+ 0.23 1.39

Total 100.00 100.00

Property

Table 2: Fuel Gas Properties

Unit

@ 0.7o

& 5.7barg C @ 25o

& 100barg C

Vapor fraction - 0.997 1.00 Mass density kg/m3 5.75 97.22

Molecular weight - 18.96

HHV MMBtu/gal 15230

LHV MMBtu/gal 13770 Hydrate formation temp o -4.7 C 20.78

HC dewpoint o

20.59 C 22.98 Water dewpoint

o-8.4 C 24.6

Cp/Cv - 1.30 1.79

3.1 IChemE_Technical Report_August 2009

3. TECHNICAL ASSESSMENT

3.1 Project Constraints

The client has provided existing facilities for use for the fuel gas system distribution. These are: • 150lb rated fuel gas Scrubber 1.219m (ID) x 2.438m (H). • 900lb rated fuel gas Indirect (Water Bath) Heater with 2 sets of heating coils (2” and 4” respectively).

Water at 93o

C is the heating medium.

The supplied equipment/instruments are constraints in the process design and the system will be modified to optimally utilize them. The system design capacity is 2.5MMscfd.

3.2 Project Approach

The following approach has been used to assess the fuel gas system requirements and develop an optimal design. 1. Determine gas consumption of each user. 2. Determine overall fuel gas rate required - up to 2.5MMscfd gas available. 3. Perform fuel gas system simulation to obtain heat and material balance for system. 4. Capacity review of the heater - water bath heater with the process fluid heated up within coils. 5. Capacity review of the scrubber - 150lb rated to be capable of handling 2.5mmscfd of gas. 6. Review the process control - i.e. the scrubber pressure control and the heater temperature control. 7. Review the equipment process safeguards - pressure safety valves and blowdown valve. 8. Size associated lines in fuel gas system.


4. DESIGN AND CALCULATIONS SUMMARY

4.1 Assessment of Fuel Gas Demand

The fuel gas consumers are as follows: 2 x Inlet Heaters; 6 x Export Heaters; 1 x Fuel Gas Heater; 3 x Gas Generators. There are also 2 Bi-Fuel Generators which use diesel and can also use gas. The fuel gas flowrate required per consumer is obtained as follows:

[ ])/()/(()( 3smbtuLHVhrmmbtuDutyHeatermmscfdFG ÷=

Heaters:

Eqn 4-1

Where, Duty = 6.30, 5.67 and 1.0MMBtu/hr (for inlet, export and fuel gas heaters) Efficiency = 69%

[ ])/()/(.()( 3smbtuLHVhrmmbtuDutyGenmmscfdFG ÷=

Generators:

Eqn 4-2

and generator duty:

[ ])./(10)(()/(. 6 hrkWbtuptionFuelConsumkWOutputhrmmbtuDutyGen −××= Eqn 4-3

Where, Output = 0.975 and 1.20kW (for the gas generator and bi-generator respectively) Fuel consumption = 10,720 Btu/kW-hr Efficiency = 95%

The bi-fuel generator runs primarily on diesel and is used for start-up and as backup generator. It can also run on fuel gas.

Heaters

Table 3: Flowrate for the individual heaters

Heater Duty Heater Duty FG LHV FG Flowrate FG Flowrate (MMBtu//hr) (MMBtu//hr) (Btu/sm3 (sm) 3 mmscfd /hr) % of Total

@69% LHV Eff. 1 6.30 9.13 36214.67 252.12 0.21 17.92 2 6.30 9.13 36214.67 252.12 0.21 0.43 3 5.67 8.22 36214.67 226.91 0.19

48.37

4 5.67 8.22 36214.67 226.91 0.19 5 5.67 8.22 36214.67 226.91 0.19 6 5.67 8.22 36214.67 226.91 0.19 7 5.67 8.22 36214.67 226.91 0.19 8 5.67 8.22 36214.67 226.91 0.19 1.15 9 1.00 1.45 36214.67 40.02 0.03 1.42

Total 1905.71 1.62


Gas

Table 4: Flowrate for the individual generators

Gen Output Eff. Heating Value Heat Rate FG FG FG Flow FG Design % of

(MW) (%) (BTU/kg) (BTU/kW-hr) (kg/hr) (sm3/hr) (mmscfd) (mmscfd) Total

1 0.975 95% 45,153 10,720 244 302.95 0.2568 0.2568 32.29 3 =generators 0.7703

Consumer

Table 5: Balance Summary

Unit Rate (mmscfd)

No of Units

Total (mmscfd)

Prorated to 2.5mmscfd

Ratio (%)

Inlet Heaters 0.21 2 0.43 0.45 17.92 Export Heaters 0.19 6 1.15 1.21 48.37

FG Heater 0.03 1 0.03 0.04 1.42 Gas Gen. 0.26 3 0.77 0.81 32.29

2.39 2.50 100.00

Figure 1: Summary


4.2 Fuel Gas Simulation

The process was simulated using Aspen HYSYS to obtain the heat and material balance. The SRK property package was selected.

The Heat and Material is given in Appendix A2.

Figure 2: Aspen HYSYS Simulation PFD


4.2.1 Phase Envelope

The fuel gas composition has been used to generate a phase envelope in order to determine the dewpoint conditions at the supply pressure of 5.5barg.

Fuel Gas Envelope

0

20

40

60

80

100

120

-150 -140 -130 -120 -110 -100 -90 -80 -70 -60 -50 -40 -30 -20 -10 0 10 20 30 40

Temp (oC)

Pres

sure

(bar

)

Bubble Pt.Dew Pt.Hydrate

Figure 3: Hydrocarbon Envelope for Composition

The envelope indicates that at a fuel gas supply pressure of 5.5barg the dew point for the fluid is about 15o

4.3 Fuel Gas Heater Review

C. This is the minimum temperature the fuel gas is to be heated to by the water bath heater. The fuel gas scrubber is installed before the heater to knock out any liquids resulting from the pressure drop thus increasing the gas dew point.

The heat and transport properties of the heater are obtained via simulation using the heater characteristics from vendor data (Ref 1, Table 6).

Table 6: Data for Heaters

Inlet Heaters (2x100%)

Export Heaters (6x100%)

Fuel Gas Heater

Duty (MMBtu/hr) 6.30 5.67 1.0 Efficiency (%) 69 69 69 Coil N.D (in) 4 4 2 4 Spec. A-106B A-106B A-106B A-106B Schedule 40 40 160 160 Heating Surface (ft2 1431 ) 2456 - - Coil Length (ft) - - 18 18 Overall HTC (Btu/hr-ft2 40 -F 40 29 29


The fuel gas heater has two coils; the first for preheating the import gas and the second one for heating the gas prior to distribution. A TCV is situated on the outlet of the second coil. The heater is a water bath type with heating coils immersed in water at 93o

The gas is preheated to compensate for any heat loss from the pressure letdown. A reduction in the inlet pressure from 100barg to 5.7barg results in the temperature falling from 25

C.

oC to -26.11o

The target for the heater is to preheat the fuel gas to 45

C.

oC in the first coil thus compensating for the heat loss across the control valves. The hydrate formation temperature for the gas is -5o

The results of the simulation show that when the heater preheats the fuel gas to approximately 45

C.

oC the temperature across the control valves drops to 0.7 o

The heating effect of the second coil is limited to about 15

C. Thus the water bath heater is adequate for the process requirement.

o

4.4 Fuel Gas Scrubber Review

C by the use of a three-way TCV supplied with the heater.

Table 7 gives the design details of the client supplied scrubber.

Diameter, ID (m)

Table 7: Fuel Gas Scrubber

1.219 Length, s/s (m) 2.438

Surface Area (m2 19.0 ) Volume (m3 3.32 )

Operating Pressure (psig) 83.0 Design Pressure (psig) 100.0

Operating Temperature (o 0.7 C) Design Temperature (F) 212 max / -20MMDT

Wall Thickness (in) 0.375 Corrosion Allowance (in) 0.125

Service Sweet Service, (Oil, Gas, Water, Wax) Insulation 2” Fiberglass with 0.020” aluminum clad

Calculations (Appendix B1) show that the size of scrubber required for 2.5MMscfd fuel gas is 0.6m (ID) x 2.15m (H). The scrubber is also provided with a pressure safety valve and a blowdown valve which are adequately sized. Thus the supplied scrubber is adequate for the project requirements.

4.5 Process Control

The letdown station consists of two pressure control valves in series. Refer to Fuel Gas System PFD (Appendix A1). The detailed sizing for the control valves are in Appendix B. The outlet temperature of the second PCV is 0.7oC. This is within acceptable limits of the fluids hydrate formation temperature of -4.7oC.

Table 8: Summary of Results

PCV-1 PCV-2 Cv 1.51 , valve flow coefficient 8.33 Valve Type Globe Globe Valve Body Size (inch) * 1 1

* Size subject to vendor review


Figure 4: Pressure Letdown Station

Two control valves in series are required to reduce the import pressure of 100barg to 5.7barg based on the sizing criteria for control valves and because of the large pressure differential required.

Apart from the pressure letdown station the outlet of the fuel gas coil is controlled to a minimum of 15oC by a three-way TCV. The first coil inlet is uncontrolled. Both of the controls are traditional feed forward PID controllers.

Figure 5: Temperature Control Valve

4.6 Process Safeguarding

4.6.1 Depressurization

All process equipment operating above 7barg or containing at least 4m³ of butane or a more volatile liquid under normal operating conditions shall need to be provided with remotely operated vapour depressurisation valves (Ref. 4).

The BDV is also actuated automatically by a signal from the emergency shutdown system, initiated by fire or gas detection. A restriction orifice is usually used in conjunction with a BDV to restrict flow.

The Scrubber is operated at 83 psig (5.7barg) and the pressure of the inlet stream is controlled by two pressure control valves in series.

Under emergency conditions the fuel gas system (i.e., fuel gas inlet line, heater, PCVs and scrubber) can be isolated. The fuel gas scrubber is blown down to atmosphere via a 2” ball valve (Ref. 13).

Inlet Header Outlet Header

Heater

TC


The scrubber operates at pressures below 7barg and would not ordinarily require depressurization; however the blowdown of the vessel via the provided valve is examined.

The BDV and vent sizing has been based on the total inventory as a worst case. This is estimated to be 40Sm3

Aspen HYSYS is used to determine the maximum vent rate obtainable and minimum temperatures. The requirement is to blowdown this inventory to 2.86barg (50% of the operating pressure of 5.7barg) within 15mins. The depressurisation profiles are attached in Appendix C.

of gas (See Appendix C). Depressurisation calculations have been carried out for the design as it is and also with the use of a restriction orifice.

The minimum temperature obtained is -9.4°C. The piping specification for ASME Class 150 (carbon steel) pipework has a minimum operating temperature of -29°C (Ref. 8). Thus carbon steel piping will be adequate for depressurisation requirements.

4.6.2 Dispersion

Dispersion modelling has been carried out for the vented gases using PHAST. The vented gases comprise a mixture of hydrocarbons which can form a potentially flammable mixture when mixed with air, i.e., between 5% and 15% methane in air; the lower and upper flammability limits (LFL and UFL). The dispersion calculation determines the location of the vent.

PHAST is third-party consequence modelling software used to analyse hazards resulting from leaks and emissions of fluids.

PHAST is used to determine the hazardous area around the vent, i.e., the horizontal and vertical distance from the vent to the edge of the LFL gas cloud (Ref. 12). The scenarios for gas dispersion are given in Table 9.

Case

Table 9: Relief/Blowdown Scenarios – Fuel Gas System

Scenario Flowrate (kg/s) Pressure(barg) Temp. °C) A1 Fuel Gas Scrubber fire case 0.341 8.34 107 A2 Fuel Gas Scrubber closed outlet 0.658 7.58 76.2 A3 Fuel Gas Scrubber control valve failure 0.636 5.73 0.7 A4 Fuel Gas Scrubber blowdown - without relief orifice 2.466 5.73 0.7 A5 Fuel Gas Scrubber blowdown - with relief orifice 0.026 5.73 0.7

Case A4 is the worst case scenario. Modelling was carried out under the worst case weather conditions obtainable on site with the following results:

Horizontal distance to dispersion (m)

Table 10: Fuel Gas Venting Dispersion Distances from Vent Outlet (Case A4)

Vertical distance to dispersion (m)

100% LFL 50% LFL 100% LFL 50% LFL

1.77 4.66 +9.74 +13.37 1.82 4.49 +10.81 +15.17 2.41 6.11 +5.42 +7.29 2.94 7.07 +3.47 +4.56


Graphical output from the dispersion calculations are shown in Appendix C.

4.6.3 Radiation: Fuel Gas Vent

During venting there is a possibility that the vented gases could ignite with resultant damage to personnel/equipment from radiation.

Ref. 4 specifies the allowable radiation levels as a function of time. Solar radiation of 0.9kW/m² is assumed for the location.

Permissible design level

Table 11: Recommended Design Total Radiation

Conditions

1.58 kW/m² (500 Btu/h ft²) Location where personnel with appropriate clothing may be continuously exposed 4.73 kW/m² (1500 Btu/h ft²) Areas where emergency actions lasting several minutes may be required by

personnel without shielding but with appropriate clothing 6.31 kW/m² (2000 Btu/h ft²) Areas where emergency actions lasting up to one minute may be required by

personnel without shielding but with appropriate clothing 9.46 kW/m² (3000 Btu/h ft²) Value at design flare release at any location to which people have access; exposure

should be limited to a few seconds, sufficient for escape 15.77 kW/m² (5000 Btu/h ft²) Structures/areas where operators are not likely to be performing duties and where

shelter from radiant heat is available

The radiation levels from an ignited vent were modelled using FLARESIM - a third-party software used for the design and rating of flare stacks.

The height of the vent will be determined based on the radiation level measured at grade being below 4.73kW/m² in order to protect personnel. Case A4 is used as basis and the vent is a simple 4” NB unimpeded pipe work vent stack venting to atmosphere.

Appendix C contains the radiation contour plots for Case A4. It was determined that a 10.25m high vent would be required. The 4.73 kW/m², 6.31 kW/m² and 15.7 kW/m² radiation contours would sit at their lowest point at approximately 0m, 3m and 7.5 m above grade, respectively. The flame length was estimated to be approximately 23.65 m. The horizontal distances (i.e. the radii) to radiation levels of 4.73kW/m² and 6.3kW/m² were estimated to be approximately 17.1m and 10.2 m (at “head” height).

The distances to these radiation contours may affect personnel on location and thus the vent height may need to be increased. The rate of “decay” of the radiation was also investigated with the following results.

Time (sec)

Table 12: Radiation Decay (Case A4)

Mass Flow (kg/s) Volume Flow (m3/s) Distance(m)

4.7 kW/m² 6.3 kW/m² 9.46 kW/m²

0 2.466 1.921 16.8 9.7 - 5 1.711 1.497 13.4 - -

10 1.202 1.168 9.75 - - 15 0.805 0.9353 - - - 20 0.605 0.7536 - - -

Notes: The distances are the horizontal radii from the centre of the vent stack at head height (2 m above grade).


Table 12 shows that despite the fact that the initial distance to radiation level of 4.7kW/m² and 6.3kW/m² are high, this configuration of the vent stack may still be acceptable because the intensity of the radiation decreases rapidly. The radiation intensity around the vent stack decreases to within acceptable levels within 20 seconds because the vented flow is unrestricted.

Radiation modelling was also carried out for Case A2, Fuel Gas Scrubber closed outlet. It was determined that the 4.73 kW/m², 6.31 kW/m² and 15.7 kW/m² radiation contours would sit at their lowest point at approximately 2.2 m, 3.5 m and 9 m above grade, respectively. The flame length was estimated to be approximately 12.6 m. This result was for the same vent stack of 10.5 m. On further investigation if the stack height were to be reduced to 9m, the 4.73 kW/m², 6.31 kW/m² and 15.7 kW/m² radiation contours would sit at their lowest point at approximately 1.4 m, 3 m and 6.3 m above grade, respectively.

However, the Case A2 vent radiation scenario is only applicable if the relief/vent piping is modified to include a restriction orifice downstream the blowdown valve thus restricting its flow.

4.6.4 Over-Pressure Protection

In addition to the blowdown system there is a PSV on both the Fuel Gas Scrubber inlet line and the scrubber itself, both are set at 100psig.

The over-pressure relief devices (PSVs) protecting the fuel gas system have been sized in accordance with API RP 520 (Ref. 5) and API RP 521 (Ref.4) for the most severe individual relief condition. Table 13 identifies the applicable relief conditions considered for the sizing of the PSVs.

Case

Table 13: PSV Relief Conditions

1. Closed / blocked outlet 2. Control valve malfunction 3. Excess heat input/vapour generation 4. External fire

The relief valve has been selected in accordance with API Standard 526 (Ref. 6).

The operating pressure of the scrubber is 83psig (5.7bar). The PSVs are 3” x 4” and calculations are carried out to determine the suitability of these safety valves for the process. The minimum size of PSV required is 1.5” x 3” thus the pre-installed PSVs are adequate. See Appendix B3 for detailed calculations.

4.7 Line Sizing

The sizing of lines for the project was done as per company practice and principles which were based on engineering standards (Ref 8 and Ref 9). The equations and correlations used are as follows:

Line velocities are estimated using: 4

2dvQ π= Eqn 4-4

Where: Q = Flowrate (m3/s) V = velocity (m/s) D = inside pipe diameter (mm)


The pressure drop for liquid or gas lines are calculated using the Darcy formula:

=∆ 5

2

10062530

dfwP m ρ

Eqn 4-5

Where: ΔP100m = Pressure drop (kPa/100m) W = Mass flow (kg/hr) ρ = Density (kg/m3

The Moody friction factor is a function of the Reynolds number and the surface roughness of the pipe. The Moody diagram (Ref.

) f = Moody friction factor d = internal diameter

8) may be used to determine the friction factor once the Reynolds number is known:

µρdv

=Re Eqn 4-6

Where: Re = Reynolds number ρ = Density (kg/m3

Finally, erosional velocities are calculated as per Ref.

) v = Velocity (m/s) d = internal diameter (mm) μ = Viscosity (cP)

7: 5.0m

eCV

ρ= Eqn 4-7

Where: Ve = Erosional velocity (ft/s) C = Constant (100 for continuous flow) ρm = Gas/Liquid density (lb/ft3

The inlet line to the fuel gas system is 4”, ASME CL 900. From calculations this size is adequate. The inter-connecting lines within the process are also 4” (and ASME CL 150) and the distribution lines are 2” ASME CL 150 lines. The line sizing and calculations are given in Appendix B4.

)


5. CONCLUSIONS AND RECOMMENDATIONS

5.1 Conclusions

1. Fuel gas demand is expected to be 2.39MMscfd, which is within the design rate of 2.5MMscfd.

2. The supplied scrubber is adequate for project requirements.

3. The supplied heater is adequate for project requirements.

4. The minimum temperature on blowdown is -9.4o

5. Dispersion - the 50% LFL from the Fuel Gas Vent was estimated to be 7m horizontally. This is not expected to pose a hazard to personnel or equipment so long as the fuel gas vent is higher than any equipment within 7m of the vent.

C which does not exceed the ASME CL 150 piping specification.

6. Radiation - A Fuel Gas Vent of 10.25m high will expose personnel and equipment to radiation levels of not more than 4.73kW/m². This radiation level has a radius of approximately 17.1 m and will last for less than 20 seconds.

5.2 Recommendations

1. A minimum Fuel Gas Vent height of 10.25m above grade is recommended.

2. Modifying the vent piping downstream of the BDV to include a restriction orifice will mean that the vent height can be reduced to 9 m.


6. REFERENCES

1. Client Scope of Work and Vendor Data

2. Perry’s Chemical Engineering Handbook, Vol. 1 & 2, 7th Ed.

3. Coulson and Richardson’s Chemical Engineering, 5th Ed.

4. American Petroleum Institute: “API RP 521: Guide for Pressure Relieving and Depressurising Systems”, 4th Ed, March 1997.

5. American Petroleum Institute: “API RP 520: Sizing, Selection and Installation of Pressure Relieving Devices”, 7th Ed, January 2000.

6. American Petroleum Institute: “API Standard 526: Flanged Steel Pressure Relief Valves”, 5th Ed, June 2002.

7. American Petroleum Institute: “API RP 14: Design and Installation of Offshore Production Platform Piping Systems, 5th Ed, Oct 1991.

8. GPSA: “Engineering Data Book”, 12th Ed. 2004.

9. Woodhill Frontier Engineering Standards

10. Masoneilan Control Valve Sizing Handbook, bulletin OZ1000, 2000

11. The Centre for Marine and Petroleum Technology (CMPT): “A Guide to Quantitative Risk Assessment for Offshore Installations”, Publication 99/100.

12. Energy Institute: “Model Code of Safe Practice Part 15: Area classification code for installations handling flammable fluids”, 3rd Edition, July 2005.

13. Fisher Vee Ball Rotary Valves, Doc. No. D350004X012/MS11-CD171/4-06


APPENDICES

A1: FUEL GAS SYSTEM - PFD


A2 - FUEL GAS BALANCE

Table 14: Heat and Mass Balance

STREAM 1 2 3

NORMAL DESIGN NORMAL DESIGN NORMAL DESIGN FLUID VAPOUR VAPOUR VAPOUR VAPOUR MIXED MIXED VAPOUR FLOW (kg/hr) 2267 2370 2267 2370 2240 2341 VAPOUR DENSITY (kg/m3 97.44 ) 97.44 85.17 85.55 5.66 5.68 VAPOUR MW 18.96 18.96 18.96 18.96 18.96 18.96 LIQUID FLOW (kg/hr) - - - - 26.11 28.3 LIQUID DENSITY (kg/m3 - ) - - - 701.26 701.51 PRESSURE (barg) 100.0 100.0 100.0 100.0 5.7 5.7 TEMPERATURE (o 25 C) 25 45.6 44.8 1.7 0.7

4 5 6 7

NORMAL DESIGN NORMAL DESIGN NORMAL DESIGN NORMAL DESIGN VAPOUR VAPOUR VAPOUR VAPOUR VAPOUR VAPOUR VAPOUR VAPOUR

2240 2341 2240 2341 2240 2341 0 1552 5.66 5.68 4.63 4.66 5.38 5.38 5.38 5.38

18.79 18.78 18.79 18.78 18.79 18.78 18.79 18.78 0 0.0 - - - - - -

701.26 701.51 - - - - - - 5.7 5.7 5.7 5.7 5.7 5.7 5.7 5.7 1.7 0.7 58.8 57.0 15.0 15.0 15.0 15.0

8 9 10 11

NORMAL DESIGN NORMAL DESIGN NORMAL DESIGN NORMAL DESIGN VAPOUR VAPOUR VAPOUR VAPOUR VAPOUR VAPOUR VAPOUR VAPOUR

1517 1585 1084 1132 401 420 723 756 5.38 5.38 5.38 5.38 5.38 5.38 5.38 5.38

18.79 18.78 18.79 18.78 18.79 18.78 18.79 18.78 - - - - - - - - - - - - - - - -

5.7 5.7 5.7 5.7 5.7 5.7 5.7 5.7 15.0 15.0 15.0 15.0 15.0 15.0 15.0 15.0

12 13 14 15

NORMAL DESIGN NORMAL DESIGN NORMAL DESIGN NORMAL DESIGN VAPOUR VAPOUR VAPOUR VAPOUR VAPOUR VAPOUR MIXED MIXED

723 756 0 592 32 33 0 0 5.38 5.38 5.38 5.38 5.38 5.38 1.25 1.26

18.79 18.78 18.79 18.78 18.79 18.78 99.16 98.73 - - - - - - 25.73 27.9 - - - - - - 707.51 707.84

5.7 5.7 5.7 5.7 5.7 5.7 0.0 0.0 15.0 15.0 15.0 15.0 15.0 15.0 -0.4 -1.5


B1 - CALCULATION: FUEL GAS SCRUBBER

Feed data under normal case and worst case (Fuel gas with heater failure) scenarios.

Fluid Property Unit

Value

Normal case Worst case

Temperature (T) o 0.70 C -26.11 Pressure (P) Kpa 673 673 Liquid Density (ρ l kg/m³ ) 701.6 688.0 Vapor Density (ρv kg/m³ ) 5.68 8.78

Mixed Density (ρm kg/m³ ) 5.75 6.45

Liquid Mass Flow (FL kg/h ) 28.3 59.6 Vapor Mass Flow (FV kg/h ) 2341 2302

Inlet Mass Flow (FM kg/h ) 2370 2362

Liquid Volume Flow (QL m³/h ) 4.03E-02 8.65E-02

Vapor Volume Flow (QV m³/h ) 412.4 366.2

Using Stokes law, Ref 3: 5.0

−=

v

vlV KV

ρρρ

Eqn. B1-1

V

V

VQA = and

ATQL l

l×

= Eqn. B1-2, B1-3

Where: VV max allowable vap vel (m/s) = ρv vap density (kg/m³) =

K = constant A = cross sectional area (m²) ρl liquid density (kg/m³) = QV vap volumetric flowrate (m³/s) =

T = holdup time (mins) Ql liquid volumetric flowrate (m³/s) =

Ll liquid depth (m) = For a vertical separator of height < 3 m,K = 0.037m/s, From (Eqns. 10 - 12), the following is calculated:

Unit Normal case Worst case V m/s v 0.41 0.33 A m 0.28 2 0.31 D mm 597 631 ≈ 600mm T mins 10 10

L m l 2.38E-02 5.10E-02 0.15 (Min) Sketch

1 = 0.4 m 2 = 1 m 3 = 0.60 m 4 = 0.15 m Tan-Tan = 2.15 m

D (1.0m Min) (2)

0.4m Min (1)

Liquid Level 0.15m (Min) (4)

D/2 (0.6m Min) (3)


B2 - CALCULATION: PRESSURE REDUCTION STATION

The control valve sizing is carried out as per GPSA (Ref. 8) and (Ref. 10) Gas lines:

The valve sizing equations used is: XM

ZTYPFN

wCvP

1

18

= Eqn. B2-1

where, q, fluid volumetric flowrate in m3

/hr

w, fluid mass flowrate in kg/hr N8

, numerical constant = 94.8

Fp

, pipe geometry factor

P1

, fluid inlet pressure in bara

P2

, fluid outlet pressure in bara

M, fluid molecular weight T, gas inlet temperature in K Z, compressibility factor

Y, expansion factor is calculated by: Tk XF

XY3

1−= Eqn. B2-2

(Y should not be less than 0.67. Also X should not exceed FkXc for gas)

Where, Fk, ratio of spec. heats is calculated by: Tk XF

XY3

1−= and )/( 1PPXT ∆= Eqn. B2-3, B2-4

Two-phase flows:

The valve sizing equation used is: 26 Yp

fpf

FNwCv

gg

g

ff

f

p γγ ∆+

∆= Eqn. B2-5

where, w, fluid mass flowrate in kg/hr N6 , numerical constant = 27.3 Fp , pipe geometry factor ff , weight fraction of liquid phase fg , weight fraction of vapor phase ∆pf , pressure drop for liquid phase in bara ∆pg , pressure drop for vapor phase in bara γf specific weight (mass density) in kg/m3

(for liquid phase)

γg specific weight (mass density) in kg/m3

(for vapor phase)

Actual pressure drops are used for Dpf and Dpg, but with individual limiting pressures: )( 1

2vFLf pFpFp −=∆ and 1pxFp Tkg =∆ Eqn. B2-6, B2-7

Summary of results: PCV 1 PCV 2 Inlet temperature oC 44.84 9.33 Outlet temperature oC 9.33 0.70 Inlet pressure barg 100 19 Outlet pressure bar 19 5.73 CV - 1.51 8.33 Valve type - Globe Globe


B3 - CALCULATION: PRESSURE SAFETY VALVES

Design Data Operating pressure 83.0 psig 5.7 barg Design pressure 100 psig 6.9 barg Relieving pressure set point 100 psig 6.9 barg Operating temperature 492.6 °R 0.50 °C Fire Non-fire Allowable over pressure 21 10 % Relieving pressure 121 110 psig 8.3 7.6 barg Relieving temperature 684.8 629.2 °R 107.3 76.4 °C 225.1 169.5 °F

Relieving temperature: nn

TPPT

= 1

1 Eqn. B3-1

Where: P1P

= Relieving pressure n

T = Operating temperature

n

= Operating temperature

Closed Outlet Case: Relieving rate (vapour relief)

MTZ

KKPCKWA

cbd 1

= , where 11

12520

−+

+=

kk

kkC Eqn. B3-2, B3-3

W (relief load) 2,341 kg/hr 5,161lb/hr k (ratio of specific heats) 1.255 - C (co-efficient) 342.7 - Kd (co-efficient of discharge) 0.975 - API RP 520 [3.6.2.1.1] Kb (capacity correction factor) 1.0 - API RP 520 [3.6.2.1.1] Kc (rupture disc correction) 1.0 - API RP 520 [3.6.2.1.1] Z (compressibility) 0.9906 - M (molecular weight) 19.0 - T (relieving temperature) 629.2 R 76.4°C P1 (relieving pressure) 124.5 psia A (effective discharge area) 0.711 in²

Closed Outlet Case: Relieving rate (liquid relief)

2138 ppG

KKKKQA

vcwd −= Eqn. B3-4

Q (flow rate) 28.3 kg/hr 6.3bbl/d 0.2USGPM Kd (co-efficient of discharge) 0.65 - API RP 520 [3.8.1.2] Kw (back pressure correction) 1.0 - API RP 520 [3.8.1.2] - assuming P2 / P1 > 0.15 Kc (combination correction) 1.0 - API RP 520 [3.8.1.2] Kv (viscosity correction) 1.0 - API RP 520 [3.8.1.2] - preliminary estimate G (specific gravity) 0.702 - at flowing temperature p1 110 (relieving pressure) psig p2 (back pressure) 3.0 psig (assumed) A (preliminary discharge area) 0.001 in²


( )A

GQµ

×=

2800Re Eqn. B3-5

μ (absolute viscosity) 0.555 cP at flowing temperature A (effective discharge area) 0.11 in² from API Std 526 (standard orifice areas) Re (Reynold's Number) 1.95E+3 - Kv (viscosity correction) 0.941 NRe

A (effective discharge area) adjusted value

0.001 in² Closed Outlet Case: PSV sizing

A (total discharge area) 0.712 in² (vapor + liquid relief) PSV selection

A (maximum discharge area) 0.712 in² Relief orifice designation H Ref. 6 Valve body rating 150 lb Valve body size 1.5 x 3 Minimum size

Note: the PSV was sized for other cases such as fire and control valve failure. The closed case was the largest relief load.


B4 - CALCULATION: FUEL GAS LINES

The lines are numbered as per the PFD – Figure 6

INPUT Stream Number Unit 1 2 3 4 5/6 15 Fluid Vapor Vapor Mixed Vapor Vapor Liquid Pressure (kPa) 10100 10100 674 674 674 674 Temperature (C) 25.00 45.00 0.70 0.70 15.00 0.70 Molecular weight 18.96 18.96 18.96 18.78 18.78 98.73 Vapor Vol. Flow (m³/h) 24.32 27.61 412.29 412.15 435.13 - Vapor Density (kg/m³) 97.44 85.85 5.68 5.68 5.38 - (lb/ft3) 6.08 5.36 0.35 0.35 0.34 Vapor Mass Flow (kg/h) 2370 2370 2341 2341 2341 - Liquid Vol. Flow (m³/h) - - 0.04 - - 0.04 Liquid Density (kg/m³) - - 701.50 - - 701.50 (lb/ft3) - - 43.79 - - 43.79 Mixture density (kg/m³) - - 5.75 - - - Liquid Mass Flow (kg/h) - - 28.28 - - 28.28 Gas Viscosity (cP) 0.015 0.015 0.011 0.011 0.010 - Liquid/Mixed Viscosity (cP) - - 0.4302 - - 0.549 Line SCH'D 120 120 40 40 40 80 Nominal diameter (in) 4 4 4 4 4 2 Inside diameter (in) 3.62 3.62 4.03 4.03 4.03 2.07 (mm) 92.05 92.05 102.26 102.26 102.26 52.50 C 100 100 100 100 100 OUTPUT Fluid Vel. (m/s) 1.02 1.15 13.94 13.94 14.72 0.01 Erosional vel. (Ve (ft/s) ) 40.55 43.20 166.97 167.93 172.55 - (m/s) 12.36 13.17 50.89 51.19 52.59 -

Reynolds no. (Re) 6.01E+05 5.98E+05 1.90E+04 7.70E+05 7.94E+05 3.47E+0

2 Moody fr. factor (f) 0.0175 0.0195 0.0256 0.0175 0.017 0.031 Resis. Co-efficient (K) 19.01 21.18 - 17.11 16.62 59.05 ∆P (bar/100m) 0.01 0.01 0.14 0.09 0.10 5.54E-06 WFL Max velocity (m/s) 45.72 45.72 45.7-61 45.7-61 45.7-61 0.6-1.8 WFL Max ∆P (bar/100m) 0.45-1.13 0.45-1.13 0.11-0.22 0.11-0.22 0.11-0.22 0.09

Note 1. The lines are selected such that they meet the criteria for ∆P (bar/100m) and velocity. The company’s

internal standard for is used. 2. The distribution lines are to be a piping minimum size of 2” (NB)


C1 - BLOWDOWN VOLUME

Section OD WT ID Area l passes Total L in in in m m ft 2 ft m 1 Preheater Inlet 2.375 0.344 1.687 0.0428 0.0014 150 - 150 45.72 2 Preheater 2.375 0.344 1.687 0.0428 0.0014 18 4 72 21.95 3 Preheater Outlet 4.5 0.531 3.438 0.0873 0.0060 20 - 20 6.09

4 Scrubber Inlet 4.5 0.237 4.026 0.1023 0.0082 5 - 5 1.52 5 Scrubber - - 48 1.2192 1.1675 8 - 8 2.44 6 After heater 4.5 0.237 4.026 0.1023 0.0082 18 4 72 21.95 7 Heater Outlet 4.5 0.237 4.026 0.1023 0.0082 150 - 150 45.72

Vol Pressure Vol @ atm @ 6.9 barg m barg 3 m m3

3

0.07 100 6.59 0.96 0.03 100 3.16 0.46 0.04 100 3.65 0.53

0.01 6.9 0.09 0.01 3.32 6.9 22.92 3.32 0.18 6.9 1.24 0.18 0.38 6.9 2.59 0.38

Sections 4 / 5 / 6 / 7

Total: 26.84 3.89 All Sections Total: 40.25 5.83


C2 - BLOWDOWN: DISPERSION

Figure 6: Dispersion Envelope Showing UFL and LFL (Upper and Lower Flammability Limits)

Figure 7: UFL and LFL – Side View


C3 - RADIATION: ISOPLETHS

Figure 8: Radiation Isopleths

Figure 9: Blowdown – Pressure vs Time

0

2

4

6

8

10

0 10 20 30 40 50 60 70 80

Time (seconds)

Pres

sure

(bar

a)

Fuel Gas Inlet Separator Emergency Depressurisation - Pressure Profile

Case A4: 10.5m stack, 2.466kg/s vent rate

zsejmerd pas

Documents

chemical engineering

preexisting technical

engineering drawings

technical evidence

chartered chemical engineer5

technical report2

minimum required level

required advanced level