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INTEGRATED PROCESSING FOR HEAVY CRUDE OIL

A thesis submitted to The University of Manchester for the degree of

Master of Philosophy

In the Faculty of Engineering and Physical Sciences

2014

Yadira López Morán

Centre for Process Integration

School of Chemical Engineering and Analytical Science

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Table of Content

Table of Content ................................................................................................. 2

List of Tables....................................................................................................... 7

Abstract ............................................................................................................... 9

Declaration ........................................................................................................ 10

COPYRIGHT STATEMENT .............................................................................. 11

Chapter 1 Introduction ................................................................................... 12

1.1 Background .......................................................................................... 12

1.2 Objectives of the research ................................................................... 17

1.3 Overview of the Thesis ........................................................................ 18

Chapter 2 Literature review ............................................................................ 20

2.1 Introduction .......................................................................................... 20

2.2 Heavy crude oil .................................................................................... 20

2.3 Transportation of heavy crude oil ......................................................... 22

2.4 Heavy crude oil upgrading ................................................................... 25

2.5 Economic evaluation ............................................................................ 36

2.5.1 Capital cost analysis ...................................................................... 36

2.5.2 Operating cost analysis ................................................................. 38

2.5.3 Discounted cash flow analysis ...................................................... 39

2.6 Chapter summary ................................................................................ 40

Chapter 3 Upgrading processes .................................................................... 42

3.1 Introduction .......................................................................................... 42

3.2 Dilution process ................................................................................... 43

3.2.1 Simulation framework of dilution process ...................................... 43

3.2.2 Simulation of upgrading processes ............................................... 47

3.2.3 Simulation framework of crude distillation unit .............................. 48

3.2.4 Simulation of HTU, HCU and DCU for the upgrader ..................... 55

3.3 Upgrader economic evaluation ............................................................ 59

3.4 Capital cost analyses ........................................................................... 59

3.5 Operating cost analyses ...................................................................... 60

3.6 Discounted cash flow analyses ............................................................ 61

3.6.1 Economic analyses of the upgrader .............................................. 62

3.6.2 Sensitivity analyses of the economic indicators ............................ 63

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3.7 Chapter Summary ................................................................................ 65

Chapter 4 Refining processes ........................................................................ 67

4.1 Introduction .......................................................................................... 67

4.2 Simulation of refinery processes .......................................................... 67

4.3 Refinery economic evaluation .............................................................. 80

4.4 Capital cost analyses ........................................................................... 81

4.5 Operating cost analyses ...................................................................... 83

4.6 Discounted cash flow analyses ............................................................ 84

4.6.1 Economic analyses of the refinery................................................. 84

4.6.2 Sensitivity analyses of the economic indicators ............................ 85

4.7 Chapter summary ................................................................................ 88

Chapter 5 Integration of an Upgrader with a Petroleum Refinery ................... 90

5.1 Introduction .......................................................................................... 90

5.2 Strategies for integration of an upgrader to a refinery .......................... 91

5.2.1 Proposed processing scheme 1 .................................................... 91




5.3 Economic analysis of the integration between upgrader and refinery 101

5.4 Chapter summary .............................................................................. 104

Chapter 6 Integration between upgrader, refinery and integrated gasification

combined cycle (IGCC) ................................................................................... 107

6.1 Introduction ........................................................................................ 107

6.2 Integrated gasification combined cycle .............................................. 107

6.2.1 Simulation framework for IGCC ................................................... 109

6.2.2 Validation of IGCC model ............................................................ 115

6.3 Schemes for integration between upgrader, refinery and IGCC ........ 117

6.3.1 Maximum power and hydrogen production from IGCC ............... 121

6.3.2 Economic analysis for integration schemes ................................ 131

6.4 Chapter summary .............................................................................. 141

Chapter 7 Conclusions and future work ....................................................... 143

7.1 Conclusions ....................................................................................... 143

7.2 Future work ........................................................................................ 144

References...................................................................................................... 146

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Appendix A: Properties of Crude Oils ............................................................. 151

Appendix B: Economic Evaluation for Upgrader and Refinery Integration ...... 154

Cost Analysis Results ..................................................................................... 154

Appendix C: Petroleum refining correlations ................................................... 204

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List of Figures

Figure 1.1: Total energy supply, 2010 and 2035 (Source: OPEC, 2013) .......... 12

Figure 1.2: Shares of the world primary energy sources (Source: EIA, 2013) .. 13

Figure 1.3: Total world oil reserves (Source: Schlumberger, 2006) .................. 15

Figure 2.1: Dilution process and upgrader integration flowsheet ...................... 24

Figure 2.2: Upgrader flowsheet ......................................................................... 25

Figure 2.5 Typical low to medium density (°API) upgrader configurations

(PDVSA, 2013) ................................................................................................. 27

Figure 2.3: Heavy crude oil upgrading technology of patent US 2006/018602129

Figure 2.4: Heavy crude oil upgrading technology of patent US2006/0042999.33

Figure 3.1: Simulation flowsheet of the dilution process ................................... 43

Figure 3.2: Effect of diluents flow rate and type on diluted crude oil density ..... 46

Figure 3.3: Heavy oil upgrader scheme. (Mass flow is in kt/d) .......................... 47

Figure 3.4: General crude distillation unit flowsheet .......................................... 51

Figure 3.5: Effect of diluent flow rate on the specific gravity (sg) of the straight

run naphtha and diesel using naphtha and light crude as diluents ................... 52

Figure 3.6: Effect of syncrude production over the upgrader income and

operating cost ................................................................................................... 63

Figure 3.7: Effect of syncrude production over the upgrader NPV and NPVI .... 64

Figure 3.8: Syncrude and heavy oil prices ........................................................ 64

Figure 3.9: Sensitivity analysis of NPVI with crude oil price .............................. 65

Figure 4.1: General flowsheet selected for refinery .......................................... 68

Figure 4.2: ASTM D-86 curves for different feedstock ...................................... 69

Figure 4.3: Refinery distillation unit simulation flowsheet .................................. 71

Figure 4.4: Effect of light crude oil flowrate over the refinery income and

operating cost ................................................................................................... 86


operating cost ................................................................................................... 86

Figure 4.6: Effect of light crude flowrate over the refinery NPV and NPVI ........ 87

Figure 4.7: Effect of syncrude flowrate over the refinery NPV and NPVI .......... 88

Figure 5.1: Proposed processing scheme 1 ...................................................... 91



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Figure 6.1: General flowsheet for IGCC process. ........................................... 108

Figure 6.2: Simulation flowsheet for coal IGCC unit ........................................ 111

Figure 6.3: Alternative schemes for syngas processing (Hydrogen or CHP) .. 121

Figure 6.5: Hydrogen and power production as a function of fraction to power.

Feedstock: coke and VB product from refinery fed with light crude oil ............ 123


Feedstock: coke and VB product from refinery fed with syncrude .................. 125


Proposed processing scheme 5 ...................................................................... 126







Figure 6.11: NPVI vs. fraction to power. Upgrader-IGCC integration .............. 131

Figure 6.12: Effect of syncrude production on NPVI. Upgrader-IGCC integration

........................................................................................................................ 132

Figure 6.13: Effect of CO2 tax over the NPVI. Upgrader-IGCC integration ..... 132

Figure 6.14 : NPVI vs. fraction to power. Refinery- IGCC integration ............. 133

Figure 6.15: Effect of feedstock flow rate on NPVI. Refinery- IGCC integration.

........................................................................................................................ 134

Figure 6.16: Effect of CO2 tax over the NPVI. Refinery-IGCC integration ....... 135

Figure 6.17 : NPVI vs. fraction to power ......................................................... 136

Figure 6.18: Effect of feedstock flow rate on NPVI. ......................................... 137

Figure 6.19: Effect of CO2 tax over the NPVI. ................................................. 138

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List of Tables

Table 3.1: Properties of heavy crude oil and diluents ....................................... 45

Table 3.2: Controlled and manipulated variables for sensitivity analyses ......... 50

Table 3.3: Modelling details, design considerations ......................................... 50

Table 3.4: Mass balance of dilution process and crude distillation unit of the

upgrader ........................................................................................................... 54

Table 3.5: Feedstock for simulation of upgrader units ..................................... 55

Table 3.6: Upgrader internal unit products ....................................................... 56

Table 3.7: Syncrude blending components and ASTM distillation data ........... 57

Table 3.8: Upgrader main products properties .................................................. 58

Table 3.9: Upgrader utilities and hydrogen demand ......................................... 58

Table 3.10: Summary of upgrader capital cost ................................................. 60

Table 3.11: Summary of upgrader operating cost ............................................. 61

Table 3.12: Assumptions for economic analyses .............................................. 62

Table 3.13: Economic evaluation for the upgrader ............................................ 62

Table 4.10: Refinery main products, utilities and hydrogen demand. Light crude

oil feedstock. ..................................................................................................... 79

Table 4.11: Refinery main products, utilities and hydrogen demand. Syncrude

feedstock .......................................................................................................... 80

Table 4.12: Summary of refinery capital cost. Light crude oil ............................ 81

Table 4.13: Summary of refinery capital cost. Syncrude ................................... 82

Table 4.14: Summary of upgrader operating cost. Light crude oil ..................... 83

Table 4.15: Summary of refinery operating cost. Syncrude .............................. 84

Table 4.16: Economic evaluation for the refinery feed with light crude oil ......... 85

Table 4.17: Economic evaluation for the refinery feed with syncrude ............... 85

Table 5.1: Refinery products to be processed in upgrader. .............................. 93

Table 5.2: Upgrader units balance. Proposed processing scheme 2 ................ 94

Table 5.3: Upgrader main products after integration with refinery. Proposed

processing scheme 2 ........................................................................................ 95

Table 5.4: Refinery main products, utilities and hydrogen balance after

integration with upgrader. Proposed processing schemes 2 ............................. 96

Table 5.5: Mass flows of the products (kt/d) for proposed schemes 3 and 4 .... 99

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Table 5.6: Gasoline and diesel properties for proposed processing schemes 3

and 4. .............................................................................................................. 100

Table 5.7: Utilities and hydrogen demand for proposed schemes 3 and 4 ..... 101

Table 5.9: Economic evaluation for proposed processing schemes 2 ............ 103

Table 5.10: Proposed processing schemes 3 and 4 ....................................... 104

Table 6.1: Unit Operation models for IGCC. Base case .................................. 115

Table 6.2: Coal composition............................................................................ 116

Table 6.3: Validation of IGCC. Syngas composition ....................................... 117

Table 6.4 Feedstocks to IGCC ........................................................................ 118

Table 6.5 Feedstock amounts to gasifier. Upgrader-IGCC ............................. 118

Table 6.6 Feedstock amounts to gasifier. Refinery-IGCC ............................... 118

Table 6.7 Feedstock amounts to gasifier. ....................................................... 119

Proposed processing scheme 5 and 6 ............................................................ 119

Table 6.8 Feedstock amounts to gasifier. ....................................................... 120

Proposed processing scheme 7 and 8 ............................................................ 120

Table 6.9 Syngas composition for different Feedstock ................................... 120

Table 6.10: Power and H2 requirements and production. Integration upgrader

and IGCC ........................................................................................................ 122

Table 6.11: Power and H2 requirements and production. Integration refinery

and IGCC ........................................................................................................ 124

Table 6.12: Power and H2 requirements and production. Proposed processing

scheme 5 and 6 .............................................................................................. 127

Table 6.13: Power and H2 requirements and production. Proposed processing

scheme 7 and 8 .............................................................................................. 129

Table 6.14: Economic evaluation for upgrader – refinery and IGCC integration

........................................................................................................................ 139

Table 6.15: Economic evaluation comparison for all schemes studied ........... 140

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Abstract

Energy based on non-renewable resources such as gas, oil, coal and nuclear fission, even with their serious problems of pollution, contributes to 86% of the global energy consumption. Oil will remain the dominant transport fuel: about 87% of transport fuel in 2030 will still be petroleum-based.

Discoveries of conventional sources of light easy-to-access crude oil are becoming less common and current oil production levels are struggling to match demand, it is necessary to develop new non-conventional sources of oil in order to supplement conventional oil supply, whose demand is increasing continuously. A possible clue to solve this situation could be to take advantage of the extensive reserves of heavy crude oils existing in different places around the world, which could be an excellent source of more valuable hydrocarbons.

In this context, some facilities called upgraders are used to process theses heavy crude oils to both increase the hydrogen-carbon ratio and improve their quality, reducing their density (increasing ºAPI) and decreasing their viscosity, sulphur, nitrogen and metals.

The main objective in this work is to study the heavy crude oil upgrading processes in order to identify new operation schemes which explore different opportunities of integration between the upgraders and other processes or new schemes for upgraders that can sustain on its own through the production of a wide range of products.

Each design alternative has been modelled with state-of-the-art commercial software packages. The crude oil dilution process was evaluated using naphtha and a light crude oil as diluents. Sensitivity analyses were done with the purpose of selecting the type and flow rate of diluent. Once the best diluent was selected, the integration of an upgrader to a refinery was studied. Heavy ends from both the upgrader and the refinery were taken as feedstocks to an integrated gasification combined cycle (IGCC). The best operation schemes for IGCC, in order to achieve the requirements of power and hydrogen for the upgrader and the refinery was determined. Different schemes for heavy crude oil processing to produce transportation fuel instead of syncrude were proposed, too. Finally, economic evaluation of all the schemes was performed to find the best solution for heavy crude oils. The best results for the dilution process of heavy crude oils were obtained when naphtha was used as diluent. The configuration proposed for the upgrader allows producing a synthetic crude oil with 35.5 °API. The integration of the upgrader to a refinery allows the treatment of the heavy streams of the refinery and transforms them into products of higher qualities. The integration of the IGCC to the upgrader and the refinery permits a complete elimination of the heavy residues produced in these units and produces hydrogen and power to be used in the site or to export. Economic evaluation shows that all the proposed processing schemes studied are economically attractive. The proposed processing schemes chosen include the integration between upgrader refinery and IGCC unit with CCS.

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Declaration

No portion of the work referred to in this thesis has been submitted in support of

an application for another degree or qualification of this or any other university

or other institution of learning.

Yadira López Morán

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COPYRIGHT STATEMENT

Copyright in text of this thesis rest with the Author. Copies (by any process)

either in full, or of extracts, may be made only in accordance with instructions

given by the Author and lodged in the John Rylands University Library of

Manchester. Details may be obtained from the Librarian. This page must form

part of any such copies made. Further copies (by any process) of copies made

in accordance with such instructions may not be made without the permission

(in writing) of the Author.

The ownership of any intellectual property rights which may be described in this

thesis is vested in The University of Manchester, subject to any prior agreement

to the contrary, and may not be made available for use by third parties without

the written permission of the University, which will prescribe the terms and

conditions of any such agreement.

Further information of the conditions under which disclosures and exploitation

may take place is available from the Head of School of Chemical Engineering

and Analytical Science.

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Chapter 1 Introduction

1.1 Background

World energy demand will continue to grow over the next 25 years. According to

the Organization of Petroleum Exporting Countries (OPEC) (World Oil Outlook

2013) there is a clear expectation that this will increase by more than 52% by

2035 as showed in Figure1.1. Additionally, the International Energy Agency

(IEA, 2013) reported similar tendency in its long term international energy

outlook 2013, they reported that energy will increase 56% approximately from

2010 and 2040.

Figure 1.1: Total energy supply, 2010 and 2035 (Source: OPEC, 2013)

The key drivers of world energy demand are mainly the rising incomes and

population, especially of China and India. China and India’s share of the world

energy will increase from 11% in 2010 to 34% by 2040 (OECD, 2013; EIA,

2013).

World primary energy consumption is projected to grow at an average of 1.6%

per year over the period 2010 to 2040, which means that the primary energy

consumption will add 39% to the global consumption by 2040.

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Figure 1.2 shows the shares of the world primary energy sources. It can be

noticed that fossil fuels (oil, gas and coal) are converging on a market share of

23-28% each and non-fossil fuels (hydro and nuclear) on a market share of 7-

15% each.

In spite of the growing contribution of non-fossil fuels, they grow fast but from

low base, it becomes even clearer in the decade from 2020 to 2040, plus the

earnings of efficiency registered in the last thirty years and the decline of oil

consumption compared with previous periods, energy based on non-renewable

resources such as gas, oil, coal and nuclear fission, even with their serious

problems of pollution, will continue to satisfy the major share of world’s energy

needs, it contributes to 86% of the global energy consumption approximately.

Oil will remain the dominant transport fuel: about 87% of transport fuel in 2040

will still be petroleum-based (OPEC, 2013; EIA, 2013).

Renewable or clean energy sources can complement the conventional sources

of energy for the production of electricity, but no other source is so far good

enough to substitute oil as the main source for transportation fuel

Figure 1.2: Shares of the world primary energy sources (Source: EIA, 2013)

Overall demand for oil is expected to continue to grow by at least 1.0% per year.

It will add up to 16 million barrels additionally per day by 2030, where

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approximately 40% of the oil demand will come from Asia. This significant

increase in oil consumption will carry serious consequences. The International

Energy Agency, which compiles information of the oil producers worldwide,

foresees a maximum level of world oil production between 2013 and 2037.

From the moment it arrives the maximum level, the agency predicts that the oil

production will diminish 3 % annually (IEA, 2013).

Discoveries of conventional sources of light easy-to-access crude oil are

becoming less common and current oil production levels are struggling to match

demand. The statistical review of world energy (BP, 2011) reports an oil

production and consumption of 82095 and 87387 thousand barrels per day

respectively, with the decline in production of conventional oils and the need to

replenish reserves; oil companies are increasingly interested in the heavy crude

oils.

The American Petroleum Institute (API) classifies crude oils by theirs API

gravity. It is a measure of the material’s gravity or density at 15.5 °C (60 °F) and

used to classify the crude oils as light, medium or heavy. Light crude oil is

defined as having a API gravity higher than 31.1°API. For medium crude oil the

API gravity is between 22.3 and 31.1 °API. The heavy crude oil is defined as

having 22.3 °API or less. Oils with 10 °API or less are considered extraheavy,

superheavy or ultraheavy oil. Conventional oils such as Brent or Texas

Intermediate have gravities from 38 to 40 °API (Gary et al., 2007; Ancheyta ,

2013; Speight , 2011; Speight , 2013).

Figure 1.3 is a representation of total world oil reserves. This figure shows that

conventional oil is only 30% of world oil reserves; the remaining oil is not

conventional oil, including heavy and extraheavy oils, oil sands and bitumen.

The latter shares some attributes with heavy oil, although denser and more

viscous. Natural bitumen is oil with a viscosity greater than 10000 cP.

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Figure 1.3: Total world oil reserves (Source: Schlumberger, 2006)

According to Schlumberger (2006), the world reserves of heavy crude oils were

6 to 9 trillion barrels, approximately 70% of word oil reserves. Black et al. (2009)

indicated that the global reserves of heavy crude oil, extraheavy crude oil and

bitumen are about 1.2 trillion barrels.

Since conventional oil production is diminishing, it is necessary to develop new

non-conventional sources of oil in order to supplement conventional oil supply,

whose demand is increasing continuously. A possible clue to solve this situation

could be to take advantage of the extensive reserves of heavy crude oils

existing in different places around the world, which could be an excellent source

of more valuable hydrocarbons.

Years ago, heavy crude oils were discarded as an energy source due to the

complex production process and associated cost; however nowadays, with the

progressive depletion of conventional oil supply, heavy oil reserves have

attracted interest from oil companies and governments all around the world.

Most of the international energy outlooks show how unconventional oil and gas

are playing a major in meeting global demand (IEA, 2013; OPEC, 2013; BP,

2013).

So far the discussion has focused on the future of energy demand , how

conventional oils is diminishing and the role of unconventional oil as part of the

solution to the global energy demand, but the question is how to process these

heavy crudes oils?

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Heavy crude oils are difficult to process in conventional refineries due to most

refineries were designed to deal with light to medium crude oils. As a

consequence, a special design or changes in the process for upgrading heavy

oil are needed. Additionally, environmental legislation is more demanding,

especially in transportation fuels; and the market for fuel oils is decreasing.

These all represent a challenge that refinery industries are facing nowadays.

Therefore, there is a need in the refinery industry to increase its processing

capacities of heavy crude and residues. (Gary et al., 2007; Ancheyta , 2013;

Speight , 2011; Speight , 2013).

One solution to this problem is the installation of plants for heavy oil upgrading

before sending this raw material to a refinery. These upgrading plants will

convert heavy oil to medium/light oil with reduced amounts of impurities and

high content of valuable distillates. There are several upgrading processes

reported in the literature that are based on two main principles: (1) carbon

rejection and (2) hydrogen addition.

The main technology in the first category is the delayed coking process, which

is the most widely used in the refining industry. Catalytic hydrotreating belongs

to the second category and is the second largest process of industrial

application

Heavy crude oil produced is diluted and received as feed in the upgraders. The

type and amount of diluent and its effect over the upgrader main products

properties are very important variables which must be studied (Gary et al., 2007;

Ancheyta , 2013; Speight , 2011; Speight , 2013).

Most upgraders are designed to produce a synthetic crude oil called syncrude

and are exploited as the upstream facilities to current refining processes. The

integration of these upgraders with other industries has not been studied in

detail so far. Additionally, to design an upgrader that can sustain on its own,

through the production of a wide range of products is a challenge.

In the same way, these technologies can also be used to convert vacuum

residues into valuable lighter products. However, they cannot completely

eliminate the “bottom of barrel” (e.g. asphaltenes, coke) and involve important

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emissions of greenhouse gases. A good alternative to overcome these

environmental problems is to use gasification processes. The main advantage

of this technology is that the full amount of asphaltenes and coke or any heavy

residues can be converted into valuable products, such as hydrogen, syngas

and energy with almost zero emissions. The integration of upgraders with

gasification processes can be a good solution, but definitely economic

justification will be the key for sustaining these industries (Sadhukhan et al.,

2002; Higman and Van Der Burgt, 2003, Gadalla et al., 2009; Ng et al, 2010;

Domenichini et al.,2010).

1.2 Objectives of the research

The main objectives in this work are the study of the heavy crude oil upgrading

processes in order to identify new operation schemes which explore different

opportunities of integration between the upgraders and other processes or new

schemes for upgraders that can sustain on its own through the production of a

wide range of products and the economically comparison for all the proposed

schemes.

The economic analysis plays an important role in this study because it provides

quantitative evaluation of the economic worthiness of all the proposed schemes.

The economic analyses for the proposed schemes were calculated via

discounted cash flow analysis (DCF). Additionally, the profitability of

investments was also evaluated with the net present value index (NPVI).

The understanding of the heavy crude oil upgrading processes will help refiners

to face current and future challenges such as how to deal with heavy feedstock,

how to meet environmental regulations and how to evaluate and optimise the

processes. This understanding need to be accomplished either by experimental

or modelling studies. It is really difficult to do experimental studies. The

simulation of the upgrader and refineries schemes developed in this work not

only provide a base for future heavy crude oils processes projects but also

provide an useful tool to evaluate the upgraders installed simply using the

conditions and feedstocks for a particular upgrader or refinery.

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To reach this main objective, the state of the art in heavy crude oil upgrading

processes is studied. Different opportunities for integration between an

upgrader, a conventional refinery and an integrated gasification combined cycle

(IGCC) facilities are proposed and evaluated. Additionally, a new scheme for

upgraders is proposed and compared with the traditional ones. The economic

analysis for the proposed schemes will help to find the best solution for

processing heavy crude oil.

1.3 Overview of the Thesis

The thesis is structured into seven chapters. Chapter 1 presents an introduction

to the heavy crude oils problem, the evidences of the depletion of conventional

crude oil sources and the reserves of heavy crude oil. The objective of this

research has been raised in this chapter.

Chapter 2 consists of a review of the existing literature related to this work. This

chapter is focused on what has been done in upgrading heavy crude oil to

produce synthetic crude oil or different products, including integration

technologies in upgraders or between upgraders and other industries or

processes.

Chapter 3 shows an analysis of the proposed schemes for upgrader as a base

of this study. The first part of this chapter is the study of dilution process, how

the nature and the ratio of crude/diluent affect the properties of products.

Simulation and sensitivity analysis for heavy oil distillation unit using different

diluents, as well as the details for the upgrader simulation framework have also

been provided in this chapter. Finally the upgraders economic analysis was

performed with different sensitivity analysis to see how profitable is upgrading

heavy crude oil.

Chapter 4 presents an analysis of the proposed schemes for refinery dealing

with conventional light to medium crude oil as base for comparison and

syncrude produced in the upgraders plants. The details for the refinery

simulation framework are provided in this chapter. Finally the refinery economic

analysis was performed with different sensitivity analysis to see how profitable

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is refining light to medium crude oil compared with the refinery fed with

syncrude.

Chapter 5 presents an analysis of the proposed schemes for integration of an

upgrader to a refinery. This chapter includes different parts: the first part deals

with the integration between upgrader and refinery, here the feedstock to

refinery is the syncrude from the upgrader, taking the upgrader as the feed pre-

treatment for the refinery. The second part is the integration of syncrude and

some heavy products streams between upgrader and refinery, here different

scenarios are studied. The economic evaluation for all cases for integration

between upgrader and refinery are presented and sensitivity analyses are

performed to do the comparison with the refinery and upgrader without

integration. Finally the most profitable scheme is selected.

Chapter 6 shows a complete study of Integrated Gasification Combined Cycle

(IGCC). The integration of IGCC to an upgrader and a refinery is considered as

the solution for the complete disposing of bottom of barrel from these sites and

production of valuable product to supply the site requirements. A methodology

for IGCC simulation is presented and validated. Different strategies to produce

hydrogen and utilities to meet the upgrader and refinery requirements are

proposed. The economic analyses for all the schemes proposed are compared

In the last chapter, conclusions and future works are presented. Details of the

future work are presented and explained.

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Chapter 2 Literature review

2.1 Introduction

As discussed in the previous chapter, worldwide crude oils are becoming

heavier while the reserves of conventional crude oils are decreasing. Extensive

reserves of heavy crude oils existing in different places around the world, could

be a source of valuable refinery products, for this reason, nowadays there is a

need in the refinery industry for processing heavy crude oils and residues.

The study presented in this thesis is mainly focused on upgrading processes for

heavy crude oils, particularly on the integration between these plants with

refineries an integrated gasification combined cycle (IGCC) to produce from

heavy oils valuable commercial products or utilities to be used in the site.

2.2 Heavy crude oil

Heavy crude oil is defined as oil with 22.3 °API (920 kg/m3) or higher density.

The American Petroleum Institute gravity, or API gravity, is a measure of how

heavy or light a crude oil is compared to water. The oils of 10° API (1,000 kg/m3)

or higher density are known as extraheavy, ultraheavy or superheavy because

they are denser than water. Comparatively, conventional crude oils, such as

raw Brent or West Texas Intermediate, have densities that range between 38

and 40 °API (835 and 825 kg/m3). (Gary et al., 2007; Ancheyta , 2013;

Speight , 2011; Speight , 2013).

The fluid property that affects production and recovery more than density is the

viscosity of the oil. The more viscous the oil, the more difficult it is to recover it.

No standard relation exists between density and viscosity, but the terms "heavy"

and "viscous" for heavy oils tend to be used interchangeably, because heavy

oils tend to be more viscous than conventional oils. The viscosity of

conventional oils can range between 0.001 Pa.s (1 cP) and approximately 0.01

Pa.s (10 cP). The viscosity of heavy and extraheavy oils ranges from less than

0.02 Pa.s (20 cP) to more than 1,000 Pa.s (1,000,000 cP). The most viscous

hydrocarbon, bitumen, is solid at ambient temperatures and is softened easily

21

when it warms up (Gary et al., 2007; Ancheyta , 2013; Speight., 2011; Speight ,

2013).

The main characteristics of heavy crude oils are: high density, low

hydrogen/carbon ratio, high residues of coal, and high asphaltenes, heavy

metals, sulphur and nitrogen contents. Heavy crude oils need specialized

treatment to produce more useful oil fractions, such as naphtha, kerosene and

diesel (Gary et al., 2007; Ancheyta , 2013; Speight , 2011; Speight, 2013).

As heavy crude oils are less valuable, more difficult to produce and more

difficult to refine than conventional oils, the following question arises: what

motivates petroleum companies to spend resources to extract and process it?

Firstly, in the current situation, many deposits of heavy crude oils now can be

exploited in profitable form; and secondly, these resources are abundant.

According to OPEC (OPEC, 2013) current estimates, the total of resources of

oil of the world is of approximately 1477 billion barrels and according to EIA

(EIA, 2013) report the total world oil reserves are 1525.957 billion barrels. Other

sources like Oil and gas journal in his worldwide look at reserves and

production reports 1637.9 billion barrels (Oil and gas journal, 2012).

Conventional oils represent only approximately 30% of this total; the rest is

heavy crude oil, extra heavy crude oil and bitumen.

Heavy oil promises to play a very important role in the future of the petroleum

industry and many countries are starting to increase its production, to estimate

reserves, to develop new technologies and to invest in infrastructure, in order to

exploit resources of heavy oil.

The Orinoco belt, located in Venezuela, has the biggest source of heavy crude

oil in the world (265.1 billion barrels), a total area of 55,314 km2 and an area of

current exploitation of 11,593 km2 (PDVSA, 2013; EIA, 2013).

In Canada, heavy oil reserves are approximately 179 billion bbl, mainly

non-conventional oil which can be extracted from the bituminous sands.

Canada’s production is approximately 1 million barrels per day, projected to rise

to 3.5 millions in 2025 (EIA, 2013); this resource makes Canada the most

22

important oil supplier for the United States. More than 99% of the crude oils that

Canada exports, are sent to the USA (EIA, 2013).

In the United States, at the end of the 19th century, explorers discovered oil in

California drilling shallow deposits of heavy crude oil and tar near to the earth

surface. Three oil fields of California are fields of heavy crude oil: Midway-

Sunset, Kern River and South Belridge, having each one already produced

more than 160 millions of m3 (1,006.4 billion barrels) of heavy crude oils.

2.3 Transportation of heavy crude oil

Once a heavy crude oil is produced, it is necessary to take the crude oil to the

sites of separation, treatment and storage inside the boundaries of the oil field.

Then it needs to be transported to refineries nearby or far away. Finally, large

volumes of oil products need to be transported to the points of consumption.

Transportation of the heavy and extra heavy oils is difficult due to the low

mobility and flowability of the crude and wax and asphaltene deposition on

pipeline.

Due to the characteristics of heavy crude oils, the transport of these types of

crude by pipelines, sometimes an option is to keep them at a certain

temperature to lower the viscosity and to facilitate pumping. This last implies

also the possibility of having additional stations for warming in the route to keep

the viscosity low. Another alternative to reduce the viscosity and to facilitate the

pumping of heavy and extra heavy crude oil is to mix them with another lighter

crude oil (diluent) (Martínez et al., 2011).

For the case of heavy crude oils, a treatment is needed to facilitate its

transportation. The treatments for heavy crude oils are group in three categories:

viscosity reduction, drag minimization and in-situ oil upgrading (Martínez R. et

al., 2011).

Reduction in the viscosity includes: dilution with other substance, formation of

an oil-in-water emulsion, increase or conserve temperature and depress crude

oil pour point (Martínez et al., 2011). Drag minimization is the reduction of

23

friction between the pipeline and the heavy oil, it can includes addition of

substance that can reduce friction and developing of different type of flow. The

upgrading of heavy crude oil is the reduction of metals and increases the API

gravity (Martínez et al., 2011).

Thermal treatment, since the viscosity diminishes significantly with an increase

in temperature, is an attractive method to improve the flow properties of heavy

crude oils. A well-documented example is the pipeline Alyeska in Alaska, which

transports the crude oil at approximately 50 °C. Nevertheless, the design of a

heated pipeline is not easy since it requires many considerations, including the

expansion of the gas in the pipeline, the number of pumping/heating stations,

and the pipeline heat losses. Other important considerations are the high costs

and the high rate of corrosion inside the pipeline due to high temperature.

With regard to the dilution process, an advanced method of improving the

transport of a crude oil is mixing it with a less viscous hydrocarbon such as

natural-gas condensate, naphtha, kerosene, or lighter crude oil. Guevara et al.

(1997) affirm that an exponential relation exists between the resultant viscosity

of the mixture and the volume of diluent, making the dilution a very efficient

method for viscosity reduction.

Nevertheless, in order to reach the acceptable viscosity limits for the transport,

a fraction of up to 30% in volume of diluent is necessary and impacts on the

capacity of the oil pipeline (Crandall and Wise, 1984). Problems might also arise

from the availability and cost of the diluent. However recycling it can be a

solution, for example recycling straight run naphtha from the atmospheric

distillation unit.

The method of emulsion consists of the dispersion of the crude oils in the water

in the shape of drops stabilized by surfactants, reducing the viscosity

significantly. This method was applied in Venezuela for the commercialization of

Orimulsion, emulsion that was sold as fuel for power plants.

Partial upgrading, involves modifying the composition of the crude oil in order to

reduce the viscosity. Upgrading technologies, such as hydrotreating processes

traditionally used in refineries, could be also considered for this application.

24

The annular flow of water can be an attractive way for the transport of viscous

crude oil. In this method water film around the oil, acts as a lubricant in order to

make the pressure of pumping necessary for the lubricated flow, comparable to

that of water (Martínez et al., 2011; Guevara et al., 1997). The fraction of water

is normally in the range of 10 to 30% by volume.

As mentioned above, different methods can be applied to transport heavy crude

oils from the reservoir to the upgrader facilities. In this work, the dilution method

was selected for transportation of heavy crude oil because it is an effective

method (Martínez et al., 2011; Guevara et al., 1997), and has been used in the

upgraders installed in the countries with the larger reserves of heavy crude oils

in the world, such as Venezuela and Canada (PDVSA, 2013; Suncor, 2013;

Syncrude,2013).

Different diluents were studied. A detailed explanation of this study appears in

Chapter 3. Figure 2.1 shows a flowsheet diagram of the diluents process

integrated with an upgrader. Naphtha as diluent is recycled from the distillation

unit in the upgrader.

Figure 2.1: Dilution process and upgrader integration flowsheet

25

2.4 Heavy crude oil upgrading

Upgrader is a specialized upstream processing facility for heavy crude oil to

produce low sulphur and lighter synthetic crude oils (syncrude) for refineries.

Figure 2.2 presents a general scheme for upgrading heavy crude oil. Here the

heavy crude oil with low API gravity is mixed with diluent to produce a diluted

crude oil to be fed to the upgraders. The upgrading processes include catalytic

processes, such as hydrotreating, hydrocracking and fluidised catalytic cracking

processes, thermal processes, such as visbreaking and delayed coking and

solvent deasphalting. Finally, from the blending of all the products the synthetic

crude is obtained with improved characteristics in comparison to those of the

heavy crude oil fed, that is, lower density (higher API gravity), viscosity, amount

of sulphur, nitrogen, metals and Conradson carbon (Speight., 2013; Flint , 2004).

Figure 2.2: Upgrader flowsheet

To date the oil industry has implemented upgrading as a separate step from

refinery to produce a syncrude. In existing upgrading plants, coking has

26

prevailed historically as the choice for primary upgrading (PDVSA 2013; Flint,

2004; Speight , 2013; Ancheyta , 2013). Practically the totality of the production

of Suncor Canada Ltd and Syncrude Canada Ltd and all the upgraders in

Venezuela employs coking as primary upgrading technology (Flint, 2004,

Suncor, 2012; Syncrude, 2012, PDVSA, 2013).

Hydroconversion processes need special management of feedstocks with high

content of solids and do not convert the asphalt completely to zero residues.

The Scotford upgrader in Alberta, Canada and Syncor and Ameriven upgraders

in Venezuela use hydroconversion for primary upgrading (Flint, 2004; PDVSA

2013). The Husky upgrader converts asphalt in a catalytic bed and the residue

is sent to a coking process. These two primary processes (hydroconversion and

coking) produce a boiling range in the product comparable with that of light

crude oils, but the concentration of impurities, such as sulphur and nitrogen, are

higher. The secondary processes of hydrotreating eliminate these impurities, to

produce sweet oil fractions to be a blending component of the synthetic crude

oil. In the hydrotreating processes, substantial conversion does not occur;

boiling range is essentially controlled by the primary upgrading (Flint, 2004). All

the upgraders in Venezuela use hydrotreating as secondary processes (PDVSA,

2013).

Generally the upgrading systems are divided in two categories: low to medium

and medium to high API gravity.

Low to medium API upgraders produce a syncrude with API gravity

approximately between 16° to 20°. The process includes distillation unit,

hydrotreating and delayed coking for vacuum residue and some auxiliaries

processes for sulphur recovery (PDVSA, 2013). Petromonagas (former Cerro

Negro) (16 °API) and Petroanzoategui (former Petrozuata (20 °API) in

Venezuela are some good examples in this category. Figure 2.5 shows a

general scheme for these upgraders.

27

Delayed Coking Unit

Coke

Crude

(Density 8 °API)

CDU

Hydrotreating

Syncrude

(Density 16-22 °API)

Figure 2.5 Typical low to medium density (°API) upgrader configurations

(PDVSA, 2013)

Medium to high API upgraders include additional hydrocracking processes, as

shows Figure 2.6. Petropiar (25.9 °API) (former Ameriven) and Petrocedeño

(former Sincor) (32.5 °API) in Venezuela are examples of this configuration

(Hamaca Project, 2007).

Delayed Coking UnitCoke

Crude

(Density 8 °API)CDU

Hydrotreating

Syncrude

(Density 25-32 °API)

Hydrocracking

Figure 2.6. Medium to high density (°API) typical upgrader configuration

(PDVSA, 2013)

28

Many researchers have been developed processes for upgrading heavy crude

oils to produce syncrude and some by-products (Carrillo and Corredor , 2013;

Castañeda et al., 2012; Marchionna et al., 2006; Iqbal et al., 2006).

Carrillo et al. (2013) compared various alternatives of producing synthetic crude

from Castilla heavy crude oil. The raw materials were crude oil free of lights

(199 °C+), reduced crude (370 °C+) and vacuum bottoms from a Castilla heavy

crude oil. The technologies used in the studied scheme were visbreaking,

delayed coking, solvent deasphalting, hydrotreating and distillation. The

selection of these technologies was based in the well-known technologies

applied for the heavy crude oil upgrading in both Orinoco belt (Venezuela) and

Alberta province (Canada), the greater source of synthetic crude. Of all

alternatives studied, they found that the visbreaking of the vacuum bottoms is

the most economical and innovative alternative.

Castañeda et al. (2012) proposed various integrations of upgrading processes,

which included deasphalting, gasification, delayed coking, RFCC, ebullated-bed

hydrocracking, slurry-phase hydrocracking and fixed-bed hydrotreating. The

main advantages of the integrated process schemes were highlighted in terms

of product yields, quality of products, and elimination of low-value by-products

and reduction of impurities.

They concluded that the decision of which approach is the best depends mainly

on the properties of petroleum, the target regarding quality of the upgraded oil,

prices of oil, and products demand. The combination of more than one process

for upgrading of heavy oils seems to be a good choice but technical and

economical studies are crucial to make a decision about the suitability of

integrating various process technologies.

Marchionna ( 2006) with SNANPROGETTI in its US patent 2006/0186021

published a process for the conversion of heavy charges such as heavy crude

oils, tar and distillation residues into syncrude. Figure 2.3 shows a simplified

diagram of this process.

29

Hydrotreater Unit

H2

Asphalt

CDU

Heavy Residue (TAR)

Lighter fraction

Distilled products

DAO

Heavy Crude Oil/

SolventDeasphalter Unit

Tars/Distillation

Residues

Figure 2.3: Heavy crude oil upgrading technology of patent US

2006/0186021

In this scheme a fraction of the heavy crude oil is sent to the deasphalting unit

to recover deasphalted oil (DAO) from asphalt, which is mixed with hydrogen

and sent to the hydrotreater unit. The other part of the feed is optionally mixed

with the asphalt product from the deasphalting unit and sent to the hydrotreater

unit. The hydrotreated stream is sent to distillation or flash units where the

separation between lighter fractions and distillation residue or liquid stream from

flash unit takes place. Approximately 60-95% of the residue or liquid from flash

unit is recycled to the deasphalting unit. In this process 15% naphtha, 17 % light

gas oil and 68% deasphalted oil and heavy gasoil are obtained.

The common characteristic for the works discussed so far is that all the

schemes produce only syncrude. The researches have proposed different

schemes for the upgraders to produce a synthetic crude oil to be sent as a feed

to the refineries.

Others relevant works in heavy oil upgraders include the integration with

gasification process in order to get more valuable products such as steam,

power, syngas and hydrogen to satisfy the internal demand or to export.

30

The Integrated gasification combined cycle (IGCC) is an extremely useful and

flexible arrangement for converting heavy ends into power, heat, Syngas and

hydrogen (Higman and Van Der Burgt, 2003, Ng et al, 2010; Gadalla et al.,

2009; Domenichini et al.,2010).

Gasification was mainly used to produce power from Syngas. But nowadays the

process is being used to produce not only power but also hydrogen and steam,

such as in Texaco gasification process (Higman and Van Der Burgt, 2003; Zhen

et al, 2005; Gadalla et al, 2009; Ng et al, 2010; Coca, 2003). A wide range of

feedstock from natural gas to heavy oil residues and coal, as well as waste

streams and biomass, can be processed with IGCC (Higman and Van Der Burgt,

2003; Coca, 2003).

IGCC can be a choice for economical and environmentally friendly new

generation fuel requirements for load growth, repowering old coal plants and

existing combined cycle plants which cannot be operated due to high natural

gas prices (Coca 2003). There are different technologies for IGCC classified

according to the gasifier configurations and the flow geometry (Michener, 2005).

The four major commercial gasification technologies are (in order of decreasing

capacity installed): Sasol-Lurgi, GE (originally developed by Texaco), Shell and

Conoco Phillips E-gas (originally developed Dow) ( Zheng et al., 2005).

The heavy products from refineries, liquid residues and petroleum coke have an

increasingly limited market; for these reasons the integration of IGCC with

refineries has been studied as a possible solution for disposal of heavy ends

into value products.

Many studies have been done related to the integration of a refinery or upgrader

with a gasification process. Gasification is the key to the conversion of low-

value feedstock to high-value fuels. An integrated gasification combined cycle

allows the heavy residues from refineries or upgraders to be converted into a

mixture of hydrogen and carbon monoxide (syngas) to be used in the

production of power, steam and/or hydrogen.

31

Some examples of refineries which have invested in gasification technology are

Motiva refinery in Delaware-USA, which uses Texaco technology to process

2000 tons per day of petroleum coke to produce 240 MW of power and vapour

(Coca, 2003). Shell Pernis refinery in Rotterdam-Holland, uses Shell technology

(Shell, 2000) to produce 127 MW of power, H2 and vapour from visbreaking

residue. Orissa refinery in India produces 180 MW of power and vapour from

petroleum coke with Shell technology. Exxon Mobil in Singapore produces 180

MW of power, H2 and vapour from heavy oil, and Pascagoula refinery in Los

Angeles-USA, uses Chevron Texaco technology to produce 570 MW of power

and vapour (Coca, 2003).

Integration of refineries and gasification has been considered as a clean,

efficient and economical solution for disposing of heavy ends (Sadhukhan et al.,

2002).

IGCC plants to meet the refinery needs of hydrogen and electrical power have

been presented for Arienty et al. (2006). In this study, alternative refinery

schemes for asphalt or petroleum coke to be fed in IGCC plant are proposed. In

both schemes the hydrogen produced is sent to the hydrocracking unit to

generate more valuable products.

Sadhukhan and Zhu (2002) proposed a methodology for integrating gasification

with an overall refinery. They developed a four-stage optimisation programme

for the integration. Their methodology explores all the integration opportunities

incorporating hierarchical decomposition into mathematical programming. The

approach includes screening and scoping analysis for preliminary study

followed by site level optimisation where the refinery margin is maximised. Then

process level optimisation incorporates the capital cost and the objective

function is the minimisation of this variable and finally the integration of the two-

level optimisation stages to maximise the refinery margin. This programme has

the advantage that design decisions are based on minimum data generation.

The results of this work have shown that the integration of gasification following

this methodology not only solves the bottom of the barrel problem but also

enhances the throughput and improves the overall refinery operation by relaxing

32

the existing bottlenecks. The capital investment for gasification was reduced by

integrating it effectively with the refinery.

Sadhukhan and Zhu (2002) concluded that all the benefits of gasification are

only possible if the integration is considered in the context of the overall refinery

operation.

Guo (2005) proposed the integration of a refinery hydroprocessing unit with a

syngas stream, a hydrocarbon synthesis unit and a utility generation unit. This

process satisfies the need to increase refining hydroprocessing unit hydrogen

purity, to maximise the desirable and environmentally acceptable product from

petroleum coke and refining residues. Guo (2005), consider that the integration

has an advantage in that it uses membrane and PSA to achieve more efficient

integration of IGCC, gas to liquid (GTL) plant and refining processes and save

on capital and operation costs. High-quality hydrogen is sent to refinery and

sulphur removal cost is less with integration than without integration by sharing

acid gas removing system (AGR) between gasification and refining units.

A process for the conversion of heavy hydrocarbon feedstock to distillates was

developed by Delbianco and Panariti (2009), in which the integration between

upgrader and gasification is considered for the production of hydrogen for the

hydrotreating process.

Here, the heavy feedstock is sent to a first distillation unit, the light fraction is

sent for hydrotreating and the residue to solvent deasphalting process (SDA).

The effluent from hydrotreating is sent to a second distillation unit where the

distillates are separated and the residue is recycled to the first distillation unit or

SDA. The deasphalted oil DAO is sent to hydrocracking unit (HCU) and the

effluent from HCU is sent to hydrotreating unit (HTU) or to the second distillation

unit. Asphalt is sent for gasification and to a gas separation unit to obtain the

hydrogen to be used in HTU and HCU.

The process proposed by Delbianco and Panariti (2009), allows the production

of a completely deasphalted and demetallised ‘light syncrude’ (atmospheric and

vacuum distillates), also upgraded in terms of density, viscosity, and CCR

sulphur content.

33

Few works have been published where upgrader and gasification integration

have been investigated. Some researchers have investigated the integration of

upgraders and gasification to produce the hydrogen necessary for hydrotreating

(Montanari et al., 2006; Eilers et al., 2008). Others like Selment et al. (2007)

and Iqbal et al. (2008) studied the gasification to produce not only hydrogen but

also steam or power for processing or exporting, for example Figure 2.4 depicts

a flow diagram for the process in this category where an upgrader is integrated

with IGCC plant for disposing heavy residues. This process is described in the

US patent 2006/0042999 A1 developed by Abdel and Subramanian (2002),

where the conversion of a heavy crude oil feed to valuable lighter compounds

with very low metal content and without asphaltene has been developed.

Figure 2.4: Heavy crude oil upgrading technology of patent

US2006/0042999.

Firstly, the heavy oil or bitumen is sent to deasphalting unit (SDA) to eliminate

the asphaltenes and produce deasphalted oil (DAO). Then, the DAO stream

free of asphaltenes and with a reduced metal content is fed to a fluid catalytic

Gasification

Unit

FCC

Unit

DAO

Solvent

Deasphalter

Unit

Steam

Fuel

Gas

Heavy Oil/Bitumen from reservoir

Hydrotreater

Unit

Power or H2

Asphalt

Fuel

Coker

Unit

Synthetic Crude

To Oil recovery process To export

34

cracking unit to obtain an effluent with reduced metal content, which is sent to

hydrotreating unit to produce a low sulphur syncrude. Additionally, the

asphaltenes are converted to steam, power, and H2 for the hydrotreating unit, to

produce heavy oil or bitumen from reservoir and/or to export. This process can

optionally include a delayed coking unit to treat part of the asphaltenes. The

products can be sent to hydrotreating unit along with the FCC unit effluent.

From the discussion of refinery and IGCC integration and upgrader and IGCC

integration it can be observed that although the development in these areas,

there are some important issues that have not been considered in deep, for

example the refineries could have serious limitations in the future when only

heavy oils will be available, because most of refinery were designed to process

light to medium crude oils. Additionally as mention before in chapter 1, the

spare refining capacities of 1990s have been absorbed due to global economic

growth.

The design of the refinery in the future will need to deal with heavy crude oils; it

has to take the upgrader as part of the pre-treatment in the total site or to

design a highly integrated configuration that can be fed directly with heavy or

extra heavy crude oil.

Few studies have been published in the area of integration between an

upgrader and a refinery.

The VEBA – COMBI – CRACKING technology (Niemann and Wenzel, 1993)

was the first work published in this area; it is a process for integration of heavy

fraction processing technologies and refinery units. The feed to the upgrader is

a refinery vacuum stream from virgin vacuum residue as well as visbreaker

vacuum residue, and the synthetic crude oil produced is sent to the refinery.

This technology allowed the production of a sweet and light syncrude to the

refinery and improved the refinery economics.

The research developed by Sadhukhan and Smith (2007) is one of the works in

this area; they developed a methodology for the design of industrial systems

based on their differential value analysis. Sadhukhan and Smith (2007) applied

this methodology for the integration of an oil upgrading system, consisting of

35

crude distillation units, hydrocracking, hydrotreating and solvent deasphalting

units, with an existing refinery, but no details about the impact of the integration

in term of product quality is presented.

Allison and Munson (2008) developed a program for an integrated system

design where a heavy oil upgrader supplies a tailor-made synthetic crude oil to

a target refinery. Before this work, no one had proposed an upgrader for

producing synthetic crude tailored to the needs of a particular refinery.

The integrated process, described in the work of Allison and Munson (2008),

includes a data communication system between the upgrader and a target

refinery for the selection of upgrading process conditions in order to produce

synthetic crude for the target refinery. This tailored synthetic crude oil is created

by altering operating conditions and the different conversion units that can be

used in the upgrader, so the synthetic crude can be produced with the

composition (TBP or ASTM assays) and properties (e.g. API gravity, sulphur

and nitrogen content) that the refinery requires.

A recent works of Aguilar et al. (2012) presented a superstructure for the

simulation of a petroleum refinery. This work is an example of integration of

different refining units to process heavy oil and produce directly transportation

fuels. The process modelled were hydrotreating of naphtha, jet fuel, kerosene,

light gas oil and FCC, catalytic reforming and catalytic cracking. Delayed coking,

visbreaking and gasification for the processing of bottom-of-barrel were

considered. The superstructure is capable to select those processes that will

meet the products specifications. The processing units have operating variables

which affect the product flow rates and properties. The objective function is to

maximize the profitability of the entire refinery, optimizing the operating

variables of the units and vacuum residue flow rate sent to each upgrading

process. The best process scheme resulted with the combination of delayed

coking and gasification.

All the studies presented in this last part are examples of the integration of an

upgrader to a refinery but detailed studies of how all the units are simulated, the

advantages and disadvantages of this integration, and its economic analysis

36

were not presented. In addition, information about the utility system is not

explored. Additionally in these studies the dilution of the heavy crude oil is not

considered. In the present work, all these subjects are explored and the best

scheme for the heavy oil processing is presented after an economic evaluation.

2.5 Economic evaluation

An economic analysis is essential to evaluate alternatives for heavy crude oil

processing and plays an important role in providing quantitative evaluation of

the economic worthiness of a particular project (Sadhukhan et al., 2008; Solli et

at., 2009).

2.5.1 Capital cost analysis

The estimation of capital investment for a specific process may vary from a pre-

design, where little information is necessary, to a detailed estimate prepared

from complete drawings and specifications. Between these two extremes,

various estimates can appear. These estimates vary in accuracy depending

upon the stage of development of a project. The following five categories

represent the accuracy range that is normally used:

1. Order-of-magnitude estimate, which is based on similar previous cost data;

the accuracy of estimate is over +/- 30 percent.

2. Study estimate (factored estimate) based on knowledge of major items of

equipments; the accuracy of estimate is up to +/- 30 percent.

3. Preliminary estimate (budget authorization estimate; scope estimate) based

on sufficient data to permit the estimate to be budgeted; its accuracy of estimate

is within +/- 20 percent.

4. Definitive estimate (project control estimate) based on almost complete data

but before completion of drawings and specifications; probable accuracy of

estimate within +/- 10 percent.

5. Detailed estimate (contractor’s estimate) based on complete engineering

drawings, specifications, and site surveys; probable accuracy of estimate within

+/-5 percent.

37

The cost estimation in this study is based on a pre-design cost estimate that

includes some information to be considered in the followings categories: order-

of-magnitude, study, and preliminary estimates. The pre-design estimates are

extremely important for determining if a proposed project should be given

further consideration and to compare alternative designs (Petters et al, 2003).

For the purpose of preliminary cost estimation for examining the viability of all

the proposed schemes, the use of the correlations to obtain the costs is

adequate. The cost correlation used in this work is showed in equation 6.1,

which was proposed by Kaiser and Gary (2007). The parameters required for

this correlation were obtained from different sources (Kaiser and Gary, 2007;

Meyer, 2004; Klara and Wimer, 2007).

(2.1)

α and β are different parameters for each unit. Typical values of these

parameters are shown in Table 2.1

Table 2.1: Parameters α y β for equation 2.1

Cost (MM$) = αααα*(Capacity)ββββ

Process Units αααα ββββ Units of Capacity

Atmospheric Distillation (ADU) 8.20 0.510 1000 bbl/d

Vacuum Distillation (VDU) 8.34 0.493 1000 bbl/d

Reforming (REF)

Continuous 12.19 0.547 1000 bbl/d

Isomerization (ISO)

Butane 9.57 0.514 1000 bbl/d

Hydrotreating (HTU)

Naphtha Desulfurization (NHT) 4.97 0.524 1000 bbl/d

Distillate Desulfurization (KHT,NHT) 8.62 0.576 1000 bbl/d

Residue Desulfurization (VGOHDS) 8.61 0.834 1000 bbl/d

Hydrocracker (HCU)

3000 scf H2/1000 bbl 26.18 0.714 1000 bbl/d

Fluid Catalytic Cracking (FCC)

Distillate Feed 24.67 0.461 1000 bbl/d

Visbreaking (VBU) 5.80 0.741 1000 bbl/d

Delayed Coking (DCU)

30 bbl feed/ton coke 24.42 0.644 1000 bbl/d

Integrated Gas Combined Cycle (IGCC) 0.29 1.000 t/d

Solvent Deasphalting (SDA) 1.2 1.000 1000 bbl/d

38

Most cost data, which are available for immediate use in a preliminary or pre-

design estimate, are based on conditions at some time in the past, for this

reason some methods have to be used for updating this cost. The cost index is

an index value for a given point in time showing the cost at that time relative to

a certain base time (Petters et al., 2003).

The costs of equipments were obtained for different years and levelised to 2014

by using equation 2.1 and cost indexes. The cost indexes were taken from

Chemical Engineering Plant Cost Index (CEPCI) which is published monthly in

Chemical Engineering (Economic indicators, 2014). The depreciation

considered in CEPCI is the straight-line method, in which the value of the

property decreases linearly with time (Petters et al., 2003).

(2.2)

2.5.2 Operating cost analysis

The operating costs were calculated using the process unit correlations from

Gary and Handwerk (2007) and Meyer (2007). The cost for utilities, such as

steam, electricity, process and cooling water, compressed air, natural gas and

fuel oil, varies widely depending on the amount of consumption, plant location

and source (Petters et al., 2003). The prices for utilities in this work were taken

from different sources (Ulrich and Vasudevan, 2006; Gary and Handwerk, 2007;

Kraiser et al.,2007; EIA, 2012).

Table 2.2 Utilities Prices

Summary of Utilities Prices

Utility Price

Fuel 70,44 $/Gcal

LP Steam 59,84 $/ton

MP Steam 66,12 $/ton

HP Steam 69,41 $/ton

Boiler Feed Water 5,00 $/m3

Cooling Water 0,11 $/m3

Power 0,05 $/kWh

Hydrogen to Sales 1350,00 $/ton

Hydrogen Offsite 2150,00 $/ton

39

2.5.3 Discounted cash flow analysis

DCF analysis uses a cumulative cash flow method based on present value

evaluation. Future value of money needs to be converted into present value.

The calculations of present value from future value were done using equation

2.3.

(2.3)

where n is the number of years and r is a discount rate. The net present value

(NPV) is calculated as shown in equation 2.4.

(2..4)

where Cf is the cash flow in a particular year and TPL is the plant life.

Additionally, the profitability of the investment was also evaluated with the net

present value index (NPVI). The NPVI is the ratio of the net present value of an

investment to its capital expense (CAPEX).

(2.5)

A ratio of more than 1 indicates a profitable investment, while a ratio of less

than 1 indicates one that will likely result in a loss.

The different assumptions for the economic analysis are summarized in Table

2.3. A plant life time of 25 years and 3 years construction time are typical in

most projects reported (IFC, 2014). The CO2 emission tax was taking from

Argus energy (Argus, 2014), A total of 8000 operating hours per year was

assumed (Ng et al., 2010). An international interest rate of 3% was assumed.

40

Table 2.3: Assumptions for economic analysis

Parameter Value

Plant Life 25 years

Construction Period 3 years

CAPEX at -2,-1 and 0 20%, 45%, 35%

Operating Percentage 90.41% (8000 h)

CO2 Emission Tax 13 $/ton

Interest Rate 3%

State + Federal Tax 60% of Income

Taxes may be classified into three types: property, excise and income taxes.

These taxes may be levied by the Federal government, state governments or

local governments (Petters et al., 2003). In this study, the state and federal

taxes are assumed as 60% of the total income.

2.6 Chapter summary

This chapter has presented a literature review about all the basics definitions

about the characteristics heavy crude oils and why it is very important

nowadays. Through the review of the state of the art in heavy crude oil

processing, it is clear that, some information related to upgrader and refinery

and gasification integration is available, but most of this information has

underpinned to the production of syncrude in an upgrader completely separated

from refinery. Few studies have tried to develop some upgraders and refinery

integration, but they do not give enough information about the simulation,

advantages and disadvantages, economic comparison and others interesting

topics. One of the main objectives in the present work is the detailed study of

the integration of an upgrader to a refinery. Chapter 5 give a complete study of

upgrader and refinery integration and chapter 6 presents the study of upgrader,

refinery and IGCC integration.

No current and future project has been planned for the design or operation of an

upgrader to produce directly transportation fuels. Chapter 5 in this work

presents some proposed schemes to produce transportation fuels at the same

41

time that some of its utilities are generated with the integration to an IGCC unit

(Chapter 6).

Most of works up to date have not presented the economical evaluation of the

heavy oil upgrading processes. Each chapter in this work presents an economic

analysis of all proposed schemes to choose the best route for heavy oil

processing.

42

Chapter 3 Upgrading processes

3.1 Introduction

Heavy oils are very difficult to recover because they are extremely viscous with

high flow resistance in oil reservoirs, which makes equally difficult their

transportation by conventional means.

As discussed in Chapter 2, once heavy oil is produced, it needs to be treated at

an elevated temperature or be mixed with diluents or solvent to maintain its low

viscosity for transportation through pipelines to an upgrader installation (Abdel

and Subramanian, 2002; Speight , 2011; Speight , 2013; Ancheyta , 2013).

In most cases, heavy crude oil recovered by steam treatment is then diluted

with naphtha and taken as a feed for upgrader installations (Schlumberger,

2006). In this work dilution process is selected to transport the heavy crude oil

to the upgraders facilities. Usually naphtha used for dilution is recycled from the

distillation unit in the upgraders.

The amount and kind of diluents to be used in diluting heavy crude oil and their

effect on the properties of the main products from the upgrader are important

factors that have not been considered in most of the research so far. Light

crude oils and light to medium products from the upgrader or refinery processes

can be used as heavy crude oil diluents.

In Venezuela and Canada, where the reserves of heavy crude oil are the largest

in the world, a number of different upgraders technologies have been installed

to produce valuable products (PDVSA, 2013 and Farouq, 2003).

This chapter includes two main parts. The first part is the evaluation of the

dilution process which compares different diluents and their effect over the API

density of feedstock. The second part is the evaluation of the upgrader. The

simulations were performed in Aspen Plus (Aspen plus, 2007) for distillation

units and Excel environment for the rest of upgraders units. The economic

analysis based in Net present value for different feed flowrates and different

43

prices of feed and products was performed. The net present value index gives

an idea about the profitability of the project.

3.2 Dilution process

The methodology for the evaluation of the dilution process included the

simulation of dilution of heavy crude oil, in which sensitivity analyses were

carried out in order to select the most suitable type and amount of diluent. The

generic flowsheet for the dilution process is presented in Figure 3.1. It includes

the blending process of heavy crude oil and diluents, this include diluents from

upgraders or from market.

Figure 3.1: Simulation flowsheet of the dilution process

3.2.1 Simulation framework of dilution process

Aspen Plus (Aspen Plus, 2007), a commercial simulation program, was used to

simulate the dilution process of the heavy crude oil.

The model “MIXER” was selected to represent the mixture between the heavy

crude oil and the different diluents.

The data on heavy crude oil and diluents assays were taken from a library of

the simulation program (Aspen Plus, 2007) and compared with the literature

(Meyer, 2003).These data included the true boiling point distillation curve, it is a

standard batch distillation test for crude oil used to determine the quality of

Diluted crude 16-19 ˚API

Heavy crude 10.1˚API

Diluent from Upgrader or imported

DILUTION PROCESS

44

products (petroleum cuts) (Gary et al., 2007), the API gravity and some other

properties curves, such as aniline point and sulphur content .

The existing light components in the heavy oil and diluents streams were

represented directly and the remaining mixtures were represented using pseudo

components.

The assay data analysis and pseudo component system in Aspen Plus allowed

to enter assay data and to generate a set of pseudo components to represent

the petroleum mixture.

Then, the diluents selected were evaluated through sensitivity analyses, using

the tool “sensitivity analysis” in Aspen Plus. This tool is used for determining

how a process reacts to varying key operating and design variables. In this work

the effect of varying the type of diluent and the ratio of crude/diluent (were

defined as manipulated variables) on the density (ºAPI) of the feedstock to the

crude distillation unit (defined as controlled variable) were carried out and the

results were evaluated.

The criterion for selecting the type of diluents in this study was the boiling range

of diluents. Thus, different cuts can be considered:

1. Light cuts: this category includes light gases, light and heavy naphtha.

They can be taken from distillation units or other processes in the

upgrader or refinery installation or they can be imported.

2. Middle distillates: This category includes diesel, kerosene and light gasoil.

3. Light crude: a crude oil with a lower density (or higher API gravity) than

the heavy crude oil from the reservoir. There are many medium density

crude oils available which can reduce the viscosity and allow the

transportation and processing of heavy crude oils.

In this work middle distillates from upgrader were not considered as diluents

because they are the main blending components for the syncrude.

45

A heavy crude oil with a density of 10.1 ˚API has been used as a reference in

this study. According to the crude classification of American Petroleum Institute

(API), it is considered heavy oil, which is heavier than conventional oil. Two

diluents were selected as examples to evaluate the proposed methodology:

naphtha with a density of 47.5 ˚API and crude oil with a density of 31.4 ˚API.

The main properties of the heavy crude oil and the two diluents selected appear

in Table 3.1. The rest of data for simulating the petroleum assay such as

distillation curves, aniline and density (ºAPI) curves appear in detail in Table A1

in Appendix A and true boiling point curves (TBP) for the light crude oil and

naphtha as diluents are presented in Table A2 in Appendix A.

Table 3.1: Properties of heavy crude oil and diluents

Heavy crude oil Lighter crude oil Naphtha

Density (ºAPI) 10.1 31.4 47.5

wt% sulphur 5.48 0.90 0.25

Viscosity (m2/s) (37.7ºC)

0.00742

7.65e-6 9e-7

The density (ºAPI) of the mixture is the most important decision making involved

in the dilution process. An appropriate density for the diluted heavy crude oil

should be between 16 and 19 ºAPI, in order to be processed through a crude

distillation unit (Schlumberger, 2002).

To calculate the ratio of crude/diluent, a basis of a fixed flow rate of 100,000

bbl/d of heavy crude oil was considered.. The ratio of crude/diluent was varied

between 5:1 (20,000 bbl/d of diluent) and 1:1 (100,000 bbl/d of diluent).

The results from the sensitivity analysis (Density values of the feed to the

atmospheric distillation column in terms of the ratio crude/diluent), are

presented in Figure 3.2. As can be seen, to achieve a feed with a density of 16 -

19 °API, the crude/naphtha and crude/light crude ratios range between 3.64-2.5

and 3.64 - 1.43, respectively. As expected the amount of naphtha as diluent is

46

lower than the amount of lighter crude oil to obtain a diluted crude oil with the

same density, because of this naphtha is the selected diluent in this work.

These results are consistent with those reported in previous work, where a ratio

crude / diluent of approximately 3 is taken as an appropriate ratio

(Schlumberger, 2002) and also the ratio used in the upgraders in Venezuela is

in this range, approximately 70 % of heavy crude oil and 30% of diluent,

specifically naphtha (PDVSA, 2013).

The dilution of heavy crude oil could be an effective method for transportation of

heavy crude oil as reported Martinez in his work (Martinez et al., 2010)

Figure 3.2: Effect of diluents flow rate and type on diluted crude oil

density

Once calculated the range for the crude/diluent ratio for both diluent selected to

get the diluted crude oil with a density between 16 – 19 ºAPI, this diluted crude

oil is taken as a feed to upgrading processes.

The following section shows the simulation of the proposed scheme for the

upgrading process in this- study.

47

3.2.2 Simulation of upgrading processes

Figure 3.3 shows the simplified process scheme for upgrader. The main

process technologies used in the proposed upgrader scheme are: crude oil

distillation units (ADU and VDU), hydrotreater (HTU), hydrocracker (HCU), and

delayed coker (DCU) units. The selection of these technologies is based on the

typical medium to high quality upgrader configuration (PDVSA, 2007). The

mass balances, together with process yield models were used to predict stream

flow rates shown in Figure 3.3.

The product from the upgrader is the syncrude with low sulphur (0.10 wt%) and

maximum density of 31 ºAPI to be processed through the refinery. Naphtha

produced can be recycled without further treatment to dilute the heavy crude oil,

the remaining products are blending components that form the syncrude. Some

heavy ends produced from upgraders can be sent as a feedstock to gasification

units (as will be described in Chapter 6). The balance for utilities and hydrogen

demand in the upgrader was calculated.

Figure 3.3: Heavy oil upgrader scheme. (Mass flow is in kt/d)

48

The simulation of the upgrading process is presented in two sections, the first

one is the simulation for the distillation units where the best diluent and

crude/diluent ratio is selected for processing the heavy crude oil taken as

reference in this work, and the second one is the simulation of the rest of

upgrading units using HPI Consultants petroleum process correlations.

The simulation procedure for the upgrader main units is described below.

3.2.3 Simulation framework of crude distillation unit

The diluted crude oil with the required density (16-19 ºAPI) from the dilution

process is extracted as a feedstock to the crude oil distillation unit. Figure 3.4

depicts the flowsheet for the simulation of the heavy diluted crude oil distillation

unit with naphtha as the diluent; an analogous flowsheet pertains to a lighter

crude oil diluents. The configuration for crude distillation unit is a typical

configuration in the actual upgrader in Venezuela. (PDVSA, 2007).

The mixture of heavy crude oil with diluents (Naphtha in this figure) is sent to

the atmospheric distillation unit (ADU), where three different products (Naphtha,

Diesel and atmospheric gas oil, AGO) are produced alongside light gases. The

heavy residue produced is converted to light vacuum gas oil (LVGO) and heavy

vacuum gas oil (HVGO) in the vacuum distillation tower (VDU). The vacuum

residue can also be sent to other downstream processes. The atmospheric and

vacuum towers were simulated using a PetroFrac model in Aspen Plus (2007).

PetroFrac provides a framework for rigorous simulation of all types of complex

vapour-liquid fractionation operations for the petroleum refining industry.

Pumparound and side stripper conditions can also be manipulated.

The thermodynamics method selected for the crude distillation unit simulation

was BK10 property method (Braun K-10 method), this method is appropriate for

most refining applications involving heavy petroleum fractions and low

pressures (Aspen Plus, 2007).

49

The procedure to get the best configuration for crude oil distillation unit was set

through sensitivity analyses. The boiling range for main products (Naphtha,

diesel, and atmospheric gasoil), the density of diluted crude oil, the Reid vapour

pressure for naphtha and the cetane number for diesel were the quality

variables to be controlled as can be seen in Table 3.2 (Gary et al., 2007). The

manipulated variables also indicated in Table 3.2 were:

- The crude/diluent ratio between the range calculated in section 3.2.1,

between 3.64-2.5 using naphtha as diluent and between 3.64 - 1.43

when using lighter crude.

- The range of number of trays in the main and side columns was selected

according to the literature (Gary et al., 2007). For each side product, 4 or

5 trays were considered. These specific cases used three side products

(naphtha, diesel and atmospheric gas oil) plus 4 trays above and 4 trays

below the feed tray. Thus, the number of trays for atmospheric distillation

tower will be 20 to 26 trays.

- For side columns, the number of trays considered was 4 to 10. In

addition, 2 pumparounds were considered for CDU and VDU to

necessary gradient.

- The steam was calculated with the feed flow to the distillation unit

following suggestions from the literature (Gary et al., 2007) and ensuring

that the amount was the minimum required.

Table 3.2 presents the specific values for controlling and manipulated

variables to carry out sensitivity analyses.

50

Table 3.2: Controlled and manipulated variables for sensitivity

analyses

Controlled variables

Crude density (ºAPI) 16 - 19 Boiling Range (˚C) Naphtha 32 - 204 Diesel 176 - 337 AGO 287-443 Specific gravity Naphtha 0.79 Diesel 0.845 Naphtha RVP (kPa),(psia) 28-69, 4-10

Manipulated variables range

Ratio of crude/diluent Naphtha 3.64-2.5 Lighter Crude 3.64-1.43 Number of trays (Main column) Atmospheric Tower 20-26 Vacuum Tower 6-10 Number of trays (in each Side columns)

4-10

Steam injected lb/bbl fed 10-50

The final configuration of the distillation columns was selected after the results

of sensitivity analyses were evaluated.

Some other variables design and operating specifications for the atmospheric

and vacuum distillation units are presented in Table 3.3.

Table 3.3: Modelling details, design considerations

Distillation Unit Specifications ADU VDU

Model in Aspen Plus Petrofrac Petrofrac Number of side products 3 2 Number of pumparound 2 2 Preheater temperature ˚C 287 Furnace temperature ˚C 349-398 388-454 Steam (temperature ˚C and pressure bar)

204.4 / 4.14

204.4/ 4.14

51

Figure 3.4: General crude distillation unit flowsheet

52

The properties evaluated in the different distillation products (Naphtha, diesel

and atmospheric gasoil) appear in Table 3.2. The effect of different diluents

selected and the ratio crude/diluent over these distillation products properties

was an important point studied to set up the necessary amount of diluent to

dilute the heavy crude oil. Additionally, the subsequent processing of these

distillation products depends on their properties.

Maximising the diesel recovery from the crude distillation unit is also important

because diesel is the main blending component in the syncrude.

The effect of the type of diluent on the main properties of the distillation

products can be seen in Figure 3.5.

Figure 3.5 shows the effect of crude/ diluent ratio, on the specific gravity (sg) of

the straight run naphtha and diesel when naphtha and light crude are used as

diluents. The result shows that a crude/naphtha ratio of 3.08 can meet the

specific gravity specifications of 0.79 (47 ºAPI) for straight run naphtha and 0.87

(31.4 ºAPI) for straight run diesel.

0,76

0,78

0,8

0,82

0,84

0,86

0,88

0,9

0,92

0,94

0,96

0,98

1

2,5 2,7 2,9 3,1 3,3 3,5 3,7

Sp

ec

ific

gra

vit

y

crude/diluent ratio

NAPHTHA (LIGHT CRUDE)

NAPHTHA (NAPHTHA)

DIESEL (LIGHT CRUDE)

DIESEL (NAPHTHA)

Figure 3.5: Effect of diluent flow rate on the specific gravity (sg) of the

straight run naphtha and diesel using naphtha and light crude as diluents

53

End-point values of approximately 160 ˚C for naphtha, 326.6 ˚C for diesel and

443 ˚C for atmospheric gas oil can be obtained with a crude/naphtha ratio of

3.08. In this case, for naphtha product the Reid vapour pressure is about 58.6

kPa. The cetane number for diesel is 36.5. These naphtha and diesel straight

run products do not meet the quality properties (RON > 90 and Cetane

number > 50) to be considered as transportation fuels (Gary et al., 2007). For

this reason, they are blending components for the synthetic crude oil.

The effect of crude/lighter crude ratio (31.4 ˚API) between the range calculated

in the dilution section (crude/diluent ratio= 3.64 - 1.43 on the specific gravity for

the main products: naphtha and diesel, can also be observed on Figure 3.5

It can be seen that approximately 1.72 crude/diluent ratio when using lighter

crude oil as diluent are necessary to reach the specification of specific gravity of

0.79 for straight run naphtha, which is used as diluent, and 0.884 for straight

run diesel.

Other important properties are also calculated with a crude/light crude oil ratio

1.72; for naphtha the Reid vapour pressure is 65.5 kPa; the end-point for diesel

is about 360˚C. However, a crude/ light crude oil ratio 1.43 and 1.49 necessary

to reach the specifications in the specific gravity and end point. Cetane number

for straight run diesel was 45. These products do not meet the RON and cetane

number to be used as transportation fuel.

Comparing simulation results for naphtha and lighter crude oil as diluents, one

may observe:

• With a crude/naphtha ratio of approximately 3.08, the density of the

crude fed is 17.5 °API and the properties in the straight run products

(density, end boiling point, diesel cetane number and naphtha RVP) can

be maintained as per specification.

• With crude/lighter crude oil ratio of approximately 1.49 the density in the

crude fed is 18.7 °API and the properties in the straight run products

(density, end boiling point, diesel cetane number and naphtha RVP) can

be maintained as per specification.

54

Therefore, the amount of lighter crude oil as diluent is approximately twice in

comparison with the amount of naphtha diluent.

Table 3.4 shows the results for mass balance in the unit, using naphtha as

diluent. The heavy crude/naphtha ratio processed was 3.08.

Table 3.4: Mass balance of dilution process and crude distillation unit of

the upgrader

HEAVY CRUDE

OIL NAPHTHA DILUENT

DILUTED CRUDE OIL LIGHT GAS

Temperature (˚C) 12.8 12.8 260 51.3 Pressure (kPa) 410 410 410 100 Mass Flow (t/h) 729.9 172.9 902.8 11.7

Volume Flow (bbl/d) 100,000 30,018.7 130,018.7 1981.1

Density (ºAPI) 9.8 47.5 17 50.2

NAPHTHA DIESEL

ATMOSPHERIC GASOIL

ATMOSPHERIC RESIDUE

Temperature (˚C) 51.3 204.6 317.3 369 Pressure (kPa) 100 150 160 170 Mass Flow (t/h) 193 75.4 69.2 554.1

Volume Flow (bbl/d) 33,141.3

11,685.4 10,000 73,210.9

Density (ºAPI) 45.7 28.3 17.6 4.7

OFF GAS

LIGHT VACUUM GASOIL

HEAVY VACUUM GASOIL RESIDUE

Temperature (˚C) 65.6 124.3 312.4 386.8 Pressure (kPa) 10 10 10 10 Mass Flow (t/h) 4.1 28.1 128.4 397.6

Volume Flow (bbl/d) 10.9 4174.6 18000 51025.4

Density (ºAPI) 33.8 13 0.8

STEAM TO

CDU STEAM TO

STRIPPER 2 STEAM TO

STRIPPER 3 STEAM TO VDU

Temperature (˚C) 204.4 204.4 204.4 204.4 Pressure (kPa) 410 410 410 410

The naphtha produced in this unit is sent to the dilution process and the rest is

blended into the syncrude. The middle distillates like diesel and light gas oils

55

were sent to HTU processes for further treatment, while the heavy gas oils were

sent to HCU. The residue from atmospheric distillation unit was sent to vacuum

distillation unit.

Products from vacuum distillation gas oils were sent to HCU and residue to

DCU.

3.2.4 Simulation of HTU, HCU and DCU for the upgrader

Once the best type and crude/diluent ratio for the distillation unit are determined

the flow rates and compositions of the main products provide inputs for

simulating the downstream processes HTU, HCU, DCU using correlations from

Gary and Handwerk (2007) and HPI Consultants correlations (Baird,1987).

Once the products from distillation unit are within the given ranges specified in

Table 3.2; HPI Consultants correlations were applied to estimate the yields and

properties of the hydrotreater (HTU), hydrocracker (HCU), and delayed coking

unit (DCU) of the upgrader. In this work, the HPI Consultants correlations were

implemented using Microsoft Excel.

The main characteristics and assumptions for the HPI correlations are

presented in Appendix C

The specific feedstock to upgrader units (hydrotreating, delayed coking and

hydrocracking), using HPI Consultants correlations are presented in Table 3.5.

This table shows the values of the main properties: volume average boiling

point (VABP), specific gravity (sg), sulphur wt% (S), cetane number, Watson

characterization factor (Kw) , Reid vapour pressure (RVP) and research octane

number (RON) of the feed to process units.

Table 3.5: Feedstock for simulation of upgrader units

Unit VABP (˚C) Kw sg S, wt% N, wt%

HTU 183.7 11.83 0.83 0.37 0.07

HCU 373.9 11 0.97 0.91 1.06

DCU 586.1 11.48 1.07 1.37 0.68

56

Table 3.6 shows the values of the main products properties: volume average

boiling point (VABP), specific gravity (sg), sulphur wt% (S), cetane number,

Watson characterization factor (Kw) , Reid vapour pressure (RVP) and research

octane number (RON) for HTU, DCU and HCU unit.

Table 3.6: Upgrader internal unit products

HCU PRODUCTS

LIGHT NAPTHA (C5 -82.2) ˚C)

HEAVY NAPHTHA (82.2-173 ˚C)

DIESEL (173-343 ˚C)

VABP (˚C) 54.8 133.55 259.11 Kw 12.45 11.67 11.59 sg 0.67 0.77 0.85

RVP (kPa) 89.6 5.52 S, wt% 0.00 0.00 0.02 N, wt% Cetane number

52.00

RON 87.80 74.90

DCU PRODUCTS

NAPHTHA LIGHT GASOIL

HEAVY GASOIL

COKE

VABP (˚C)

254.4 387.7

Kw 11.00 sg 0.76 0.84 1.03 1.20

RVP (kPa)

S, wt% 0.19 0.62 1.12 2.80 N, wt% 0.01 0.16 0.43 2.73 Cetane number

RON 66

HTU PRODUCTS

NAPHTHA DIESEL

VABP (˚C) 130.55 257.2 Kw 11.70 11.87 sg 0.77 0.83

RVP (kPa) 8.96 S, wt% 0.00 0.00 N, wt% 0.00 0.03 Cetane number 54.00

RON 55

57

The product from upgrader (syncrude) presented in Table 3.7 was made up of

naphtha and diesel from the hydrocracker and the hydrotreater units. The mass

flow rate and the ASTM-D86 distillation data of each naphtha and diesel were

blended to calculate the flow and ASTM-D86 curve for syncrude. The blending

process to get the syncrude was carried out in Aspen Plus, using the flow rate

and ASTM distillation data listed in Table 3.7.

Table 3.7: Syncrude blending components and ASTM distillation data

SYNCRUDE

STREAMS FROM

HYDROTREATER

STREAMS FROM HYDROCRACKER

NAPHTHA DIESEL LIGHT

NAPTHA

HEAVY NAPTHA

DIESEL

Flow rate (kt/d) 0.0014 5.97 0.41 1.49 5.69

Syncrude

flow rate (kt/d)

13.56

ASTM D-86 DISTILLATION DATA ( % Distillate Vs Temperature, ºC)

DIST , % 0 10 30 50 70 90 End point

Diesel hydrotreating unit

Naphtha 2.11 19.67 28.33 48.83 64.06 88.33 162.72

Diesel 192 216.11 228.78 260.61 288.5 315.22 342.67

Hydrocracking unit

Light naphtha 32.2 40.56 47.22 51.67 59.44 71.11 82.22

Heavy naphtha 96.1 109.44 124.44 132.22 141.11 157.78 173.89

Diesel 177.2 201.11 233.89 257.22 283.33 317.78 343.33

SYNCRUDE 52.2 144.96 227.80 258.17 286.43 319.3 346.78

Table 3.8 presents the main properties (density, sulphur content, RVP) for the

syncrude and recycling naphtha obtained from the simulation in comparison

with syncrude and naphtha property constraints. It can be observed that all the

properties obtained from simulation meet the limit property requirements

(maximum RVP, density and sulphur content) for both products. From the

58

blending of all the products, the synthetic crude is obtained with improved

characteristics in comparison to those of the heavy crude oil fed, that is, lower

density (higher API gravity), viscosity, amount of sulphur, nitrogen, metals and

Conradson carbon as expected. The density of the heavy crude oil fed to the

upgrader increases from 10.1 to 35.5 °API and sulphur content from 5.48 to

0.1% approximately. The API gravity obtained (35.5) was higher than the

obtained in the upgraders already in operation (PDVSA, 2013).

Table 3.8: Upgrader main products properties

Properties constraints Simulation results

Naphtha Syncrude Naphtha Syncrude

Density (°API) ( min) 47 31 46 35.5

RVP (psia) (max) 10.15 8.3

Sulphur content, (wt%) ( max)

0.001 0.001

According to the literature review presented, the proposed upgrader is a

medium to high API gravity process scheme (PDVSA, 2013).

The utilities and hydrogen demand for the upgrader appear in Table 3.9. The

negative sign indicates net consumption, while the positive sign indicates net

production.

Table 3.9: Upgrader utilities and hydrogen demand

Utilities and hydrogen demand

Fuel (kt/d)

LP steam (kt/d)

Cooling Water ((m

3/d).10

-3)

Power (MWh/d)

Chemicals (kt/d)

H2

(kt/d)

CDU -2.16 -0.83 -104.47 -134.93

HTU 0.00 0.00 0.00 -0.01 -0.02

-0.01

HCU -0.0012 -0.0026 -0.028 -0.0016

-0.33

DCU 0.00116 -0.0516 -0.0216 -4.92

TOTAL

-2.17

-0.88

-104.50

-139.88

-0.03

-0.34

59

The simulated flow rate in kt/d of the different streams for the production of

syncrude can be seen in Figure 3.3 in section 3.2.2. The products obtained are

syncrude (13.57 kt/d), naphtha for recycling (3.76 kt/d) and coke (1.64 kt/d). The

naphtha from the distillation unit without further treatment is used as diluent of

the heavy crude oil. The flow rate of the internal streams can be also seen in

Figure 3.3.

3.3 Upgrader economic evaluation

As mentioned in chapter two an economic analysis plays an important role in

providing quantitative evaluation of the economic worthiness of a particular

project. In this section the economic analyses for the upgrader is calculated via

discounted cash flow analysis (DCF), followed by a net present value (NPV)

analysis which uses discounted cash flows. Additionally, the profitability of

investments was also evaluated with the net present value index (NPVI). The

cost estimation in this study is based on a pre-design cost estimate that

includes some information to be considered in the followings categories: order-

of-magnitude, study, and preliminary estimates.

3.4 Capital cost analyses

The equation proposed by Kaiser and Gary (2007) as explained before in

Chapter 2, was used to calculate the cost in this study. The parameters required

for this correlation were obtained from different sources (Kaiser and Gary, 2007;

Meyer, 2004; Klara and Wimer, 2007).

The costs of equipments were obtained for different years and levelised to 2014

by using cost indexes. The cost indexes were taken from Chemical Engineering

Plant Cost Index (CEPCI) which is published monthly in Chemical Engineering

(Economic indicators, 2014).

The summary of upgrader capital cost for different syncrude productions are

presented in Table 3.10. The cost is a result of the sum of the individual cost for

different unit (CDU, HCU, HTU, DCU) and the utilities facilities like hydrogen

and cooling water plants and contingence cost. The results show that the total

cost increase as a consequence of heavy crude oil flow rate increase because

60

the product flow rate increase proportionally to heavy crude oil increase but not

the cost function which increase approximately 57 % from 102000 to 204000

and 27% from 204000 to 306000 bbl of heavy crude oil fed.

Table 3.10: Summary of upgrader capital cost

Syncrude production (bbl)

100000 200000 300000

Heavy crude oil fed

102150 204300 306500

Process Units

Feed to unit

kbbl/d Cost

(MM$)

Feed to unit

kbbl/d

Cost (MM$)

Feed to unit kbbl/d

Cost (MM$)

ADU 132.82 110.40

265.64 157.22 398.46 193.33

VDU 74.83 77.87 149.66 109.59 224.48 133.84

HTU 44.07 84.89 88.15 126.54 132.22 159.83

HCU 50.08 476.23 100.16 781.18 150.23 1043.47

DCU 52.12 305.64 104.24 479.77 156.36 624.57

SUBTOTAL 1055.03 1654.30 2155.05 Utilities Facilities

Hydrogen Production Unit 67.22

101.81

129.80

Cooling Water System 19.92

39.85

59.77

Storage 332.05 664.10 996.15

Offsites 221.13 369.01 501.11

SUBTOTAL 1695.35 2829.07 3841.88

Contingence 254.30 424.36 576.28

TOTAL 1949.66 3253.43 4418.16

3.5 Operating cost analyses

The operating costs were calculated using the process unit correlations from

Gary et al. (2007) and Meyer (2007) as indicated in Chapter 2. The cost for

utilities, such as steam, electricity, process and cooling water, compressed air,

natural gas and fuel oil, varies widely depending on the amount of consumption,

plant location and source (Petters et al., 2003). The prices for utilities in this

61

work were taken from different sources (Ulrich and Vasudevan, 2006; Gary and

Handwerk, 2007; Kraiser et al.,2007; EIA, 2012).

Table 3.11 reports the summary of operating cost for upgrader for different

syncrude productions. Because of the correlations for utilities cost estimation is

a function of feed flowrate cost is proportional to the heavy oil feed.

Table 3.11: Summary of upgrader operating cost

Syncrude production (bbl) 100000 200000 300000 Heavy crude oil fed 102150 204300 306500

Cost

Utilities (MM$/year) (MM$/year (MM$/year)

Fuel (Gcal/d) 269.71 539.43 809.14

LP Steam (ton/d) 53.60 107.20 160.81

HP Steam (ton/d) 39.11 78.23 117.34

Cooling Water (1000 m3/d) 13.75 27.50 41,25

Power (MWh/d) 19.90 39.80 59.69

Catalysts & Royalties ($/d) 6.86 13.72 20.58

Insurance 9.75 16.27 22.09

Maintenance 97.48 162.67 220.91

Plant Staff & Operators Salary 22.88 45.77 68.65

TOTAL 533.05 1030.57 1520.45

3.6 Discounted cash flow analyses


3.12. A plant life time of 25 years and 3 years construction time are typical in

most projects reported (IFC, 2014). The CO2 emission tax was taking from

Argus energy (Argus, 2014), A total of 8000 operating hours per year was

assumed (Ng et al., 2010). An international interest rate of 3% was assumed.

62

Table 3.12: Assumptions for economic analyses

Parameter Value

Plant Life 25 years


CAPEX at -2,-1 and 0 20%, 45%, 35%

Operating Percentage 90.41% (8000 h)


Interest Rate 3%

State + Federal Tax 60% of Income

3.6.1 Economic analyses of the upgrader

Table 3.13 summarizes the economic analysis for an upgrader with different

syncrude productions. It can be observed that all the values for the NPV are

positive and NPVI greater than 1 it means that all the schemes are profitable.

The NPV comparison shows that the upgrader producing 300000 bbl/d of

syncrude has the higher NPV of 8.320,23. These results indicate that the

upgrader is economically attractive for processing heavy crude oil. The NPVI

also indicated that the upgrader for producing 300000 bbl/d of syncrude is more

profitable.

Table 3.13: Economic evaluation for the upgrader

Syncrude flowrate (bbl)

100000 200000 300000

Heavy crude oil fed 102150 204300 306500

CAPEX (MM$) 1574.88 2558.82 3415.55

Gross Income (MM$/year) 3518.36 7036.72 10555.08

OPEX (MM$/year) 3177.84 6323,18 9461.52

Taxes (MM$/year) 112.82 250.45 393.44

Cash flow (MM$) 227.70 463.08 700.12

Net Present Value (MM$) 2180.06 5163.66 8320.23

Net Present Value index 1.25 1.83 2.20

Operating costs include the cost of raw material and capital costs include the

cost of the land and the working capital. The detailed calculations of the

economic evaluation for all cases are presented in Appendix C.

63

3.6.2 Sensitivity analyses of the economic indicators

The selected variables for sensitivity analysis study were the price for main

products (syncrude and coke) and raw material (heavy crude oil).

Figure 3.6 shows the effect of increasing the syncrude production over the

gross income and operating cost and taxes. It can be noticed how the increase

in production of syncrude generates more operating cost because of the more

needed of utilities and obviously more income.

Figure 3.6: Effect of syncrude production over the upgrader income and

operating cost

The Net present values and net present values index in Figure 3.7 shows

positives values for both NPV and the NPVI, and the increase of these

economic indicators with the syncrude flowrates, this means that the treatment

of heavy crude oil in the upgrader process is economically attractive and

increase with syncrude flowrates.

64

Figure 3.7: Effect of syncrude production over the upgrader NPV and NPVI

In Figure 3.8 it can be seen how the heavy crude oil and syncrude prices have

changed over the last five years. The fluctuations were from 63 to 108 for

syncrude and 40 to 78 for heavy crude oil.

Figure 3.8: Syncrude and heavy oil prices

65

Figure 3.9 presents the sensitivity analysis of NPVI with the syncrude and

heavy crude oil prices for the last five years. It can be observed that the NPVI is

quite sensible to the feed and product price and the worst scenario is when the

crude oil prices fall to a value of 63 in 2009, because of the economic recession.

Although of crude oil price fluctuation from 63 to 108 approximately and heavy

oil prices calculated from 40 to 78 it can be concluded that heavy oil upgrading

process is a good solution to process heavy crude oil.

Figure 3.9: Sensitivity analysis of NPVI with crude oil price

3.7 Chapter Summary

This chapter have presented the upgrader study for different syncrude

production. Simulation evaluates the performance of these proposed

processing schemes, rigorous simulation models in Aspen Plus 2006.5 were

developed for the upgrader units. The other upgrader units (Hydrocracking,

hydrotreating, delayed cooking and so on) were simulated using the petroleum

refinery correlations. The upgrader simulation firstly included a study of the

dilution process of the heavy crude oil with different diluents, to select the best

diluents to meet the quality of the blending crude oil fed to the upgrader

distillation unit.

66

The simulation results showed that naphtha recycled from the upgrader

distillation unit is the best option as diluent for the heavy crude oil. The

properties of the syncrude were met.

In the proposed schemes some heavy ends such as coke is produced, which

can be sent to other processes to produce more valuable products or to sell.

This chapter has presented the economic analyses to evaluate the proposed

processing schemes for upgrading heavy crude oil. The net present value and

net present value index have been used as methods for the economic

evaluation. The results have shown positive values for the NPV and the NPVI,

which indicates that the processing scheme studied, is profitable. Additionally

the sensitivity analyses of the NPV and NPVI with the variation of crude oil

prices during the last five years has indicated that processing heavy crude oil in

the upgraders facilities could be the solution for heavy crude oil.

67

Chapter 4 Refining processes

4.1 Introduction

In this chapter a typical refinery with processing capacity of 300000 barrel per

day is fed with light crude oil and synthetic crude oil (syncrude) both imported

from market place. The comparison between these two scenarios is made

technically and economically based in net present value for different feed

flowrates and different price of feed and products and the net present value

index to give an idea about the profitability of the project.

4.2 Simulation of refinery processes

Simulations were performed in Aspen Plus and Excel environment. The refinery

under consideration is a typical refinery which processes 300000 bbl/d of light

crude oil with a density of 31.5 °API. Figure 4.1 presents a general flowsheet for

this process (Sadhukhan et al., 2002).

The technologies used in this refinery are the ADU and VDU, reforming (REF),

isomerization (ISO), hydrodesulfurization (VGOHDS), hydrotreating for naphtha,

kerosene and light gas oil (NHT, KHT and GHT), fluid catalytic cracking (FCC),

visbreaking (VBU) and Delayed coking (DCU).

Gasoline with a high octane number and diesel are the main products from this

refinery. Cracked naphtha, naphtha from VGOHDS and DCU, Reformate,

isomerate and naphtha from cracking and GHT are blended to produce gasoline

with an octane number higher than 90. Diesel is obtained by blending distillates

from VGOHDS and GHT. Other products are kerosene from KHT and heavy

residues, including light and heavy cycle oils and slurry from the FCC unit and

visbreaking products and coke from DCU. Some heavy refinery products can be

considered for blending with the heavy ends from others processes to be used

as a feedstock for a gasification unit or can be sent to other plants for further

processing, whereas the main products can be sold to the market.

68

Figure 4.1: General flowsheet selected for refinery

Yields and utility consumption and generation of the various refinery processes

units (NHT, KHT, GHT, ISO, REF, VGO HDS, DCU, FFC and VBU), were

calculated from the HPI Consultants correlations (Baird, 1987) and from the

literature (Gary et al., 2007), respectively.

The simulation of the refinery crude oil distillation unit was built in Aspen Plus

2006.5, using the same procedure discussed in section 3.2.2 (Chapter 3). The

main properties of the light crude oil and syncrude appear in Table A3 in

Appendix A. The data for crude oil have been taken from Aspen plus 2006.5

library (Aspen Plus, 2007). The ASTM D-86 assay data are shown in Table A4

in Appendix A. Other specification including density (ºAPI) and properties

curves were included in the simulation. The syncrude to feed the refinery was

taken from upgrader product (Chapter 3).

The ASTM D-86 curves for both feedstock are showed in Figure 4.2, it can be

noticed that the syncrude has lighter fractions than regular light crude oil fed to

69

refinery. Its behaviour was as expected because the syncrude came from an

upgrading process (see Chapter 3) where its properties were improved. This

syncrude was made from naphtha and diesel from hydrotreating and

hydrocracking processes as shows Table 3.7; it does not has heavy fractions

because the bottom of barrel from atmospheric and vacuum distillation columns

and heavy end like coke were taken out in the upgrading process .

0

100

200

300

400

500

600

700

800

0 20 40 60 80 100

Te

mp

era

ture

(°° °°C

)

Liquid volume % distilled

Light crude oil

Syncrude

Figure 4.2: ASTM D-86 curves for different feedstock

The simulation general flowsheet of the crude oil distillation unit is presented in

Figure 4.3. The main streams results for refinery fed with light crude oil and

refinery fed with syncrude are presented in this Figure. It can be observed that

when conventional light crude oil is used more lighter components are obtained

from preflash column in comparison with those obtained when syncrude is fed;

980 t/d for light gases and 4102.6 t/d of light naphtha for conventional light

crude and traces for light gases and 2940.3 t/d for light naphtha are obtained

when syncrude is fed. The diesel and atmospheric gasoil production from

syncrude are greater than those obtained using conventional light crude oil,

11515.5 t/d and 16473.4 t/d in comparison with 5410 t/d and 4293 t/d

respectively. The bottom of barrel is greater for conventional oil (9604.1t/d), only

220.5 t/d using syncrude as feed are produced. These results are explained

70

because the syncrude is the product from heavy oil processing, made from

mainly middle distillates as diesel and atmospheric gasoil, the light fraction like

light gases and naphtha are separated from heavy crude oil and the bottom of

barrel from syncrude distillation is processed and the naphtha and light gasoil

are taken as part of syncrude. The syncrude is comparable with a refinery

product while the light conventional oil is taken directly from reservoir.

71

Crude: light crude oil API: 31.5

CRUDE LIGHT GASES

LIGHT NAPHTHA CDU-FEED HNAPHTHA KEROSENE DIESEL

ATMOSPHERIC GASOIL

LIGHT GASOIL

HEAVY GASOIL OFF-GAS

ATMOSPHERIC RESIDUE RESIDUE

Temperature ˚C 93.3 76.7 76.7 230.1 75.7 202.6 281.8 342.3 251.4 364.2 65.6 376.9 395.6

Pressure bar 4.14 2.74 2.74 3.08 1.08 1.46 1.55 1.61 0.83 0.88 0.80 1.70 0.93

Mass Flow t/d 41318.8 980.8 4102.6 36289.9 3585.4 3886.8 5410 4293.9 7196.5 2309 222.5 19114.5 9604.1

Liq Vol bbl/d 300,000.00 10249.7 36428.1 253322.2 30369.1 29989.2 39724.4 29999.9 48722.4 15000 36.7 123239.5 59480.5

Density (ºAPI) 31.5 113.9 68.1 25.2 58.9 41.8 33.4 25.4 20.5 14.3 41.1 13.2 7.5 Crude: Syncrude

API: 35.5

CRUDE LIGHT GASES

LIGHT NAPHTHA CDU-FEED HNAPHTHA KEROSENE DIESEL

ATMOSPHERIC GASOIL

LIGHT GASOIL

HEAVY GASOIL OFF-GAS

ATMOSPHERIC RESIDUE RESIDUE

Temperature ˚C 93.3 76.7 76.7 230.1 77.6 224.9 293.7 376 98.8 263.4 96.1 471.1 458.4

Pressure bar 4.14 2.74 2.74 3.08 1.08 1.46 1.55 1.61 0.83 0.88 0.80 1.70 0.93

Mass Flow t/d 40269.0 -- 2940.3 37340.8 3585.4 2731.0 11515.5 16473.4 1304.3 130.5 221.7 1659.5 220.5

Liq Vol bbl/d 300000 -- 24428.5 275571.5 39829.7 20819.3 85186.2 118326.3 9199.8 850 29.3 11410 1330.9

Density (ºAPI) 35.5 -- 55.1 34 48.6 39.5 34.4 29.6 26.7 14.6 33.7 22.7 3.8

Figure 4.3: Refinery distillation unit simulation flowsheet

72

As showed in the general scheme for refinery process showed in Figure 4.2, the

different slates from distillation unit were sent to other refining processes units,

including isomerization, hydrotreating, reforming, delayed coking, fluid catalytic

cracking and visbreaking units. The petroleum refining process correlations

(Blair, 1987) were used to calculate the yield and properties for different units.

The results from the simulation of refinery units fed with both conventional light

crude oil and syncrude are showed in tables 4.1 to 4.7, respectively. 77% of

vacuum distillation unit bottom of barrel is sent to DCU and the remained 23 %

to VBU. Table 4.1, shows the yield and main properties for refinery DCU for two

cases, case a presents the results for DCU when refinery is fed with

conventional light crude oil and case b presents the results when refinery is fed

with syncrude imported from an upgrader. In both cases, the DCU was fed with

vacuum residue. The main products are naphtha (C5/204,4 ºC), light gasoil

(204,4/315,6 ºC) and heavy gasoil (315,6 ºC+) and coke as residue.

As can be seen in Table 4.1, the products flowrate from DCU when light crude

oil is fed to refinery are bigger than those when the refinery is fed with syncrude

from upgrader. It is because of the lower bottom of barrel produced in distillation

unit when syncrude is fed to refinery unit. Naphtha produced in both cases need

to be sent to HTU to meet the required quality to be considered as a blending

component for final refinery products.

LGO and VGO for two cases are sent to VGOHDS to improve their properties to

blend into refinery final products. The coke production from refinery DCU for

case a (1.36 kt/d) and case b (0.03 kt/d) are lower than the upgrader production

from upgrader (4.87 kt/d).

Part of the distillation bottom of barrel is sent to refinery VBU. Table 4.2 shows

the main results for refinery VBU for the same cases explained before (light

crude and syncrude feedstock). It can be noticed that flowrate when syncrude is

use as a feedstock is lower than the one when light crude is used, for the same

reason explained for DCU. The feed and products density when syncrude is

used is higher in comparison with the use of light crude as feedstock, because

of the heavy crude.

73

Naphtha produced in VBU is sent to further processing to NHT, distillates to

VGOHDS and tar is disposal as heavy end.

Table 4.1: Refinery delayed coking unit (DCU) simulation results a) Case a

Feed NAPH LGO HGO COKE

API 7.50 55.00 36.00 16.66 - sg 1.02 0.76 0.85 0.96 1.2

S, wt% 3.74 0.52 1.68 3.06 7.63

RON

66 - - -

Flowrate:

kbbl/d 45.65 9.33 14.60 14.78 -

(kt/d) 7.39 1.12 1.96 2.24 1.36

b) Case b

Feed NAPH LGO HGO COKE

API 3.8 55.00 36.00 14.36 - sg 1.046 0.759 0.845 0.970 -

S, wt% 3.74 0.52 1.68 3.06 7.63

RON

66 - - -

Flowrate:

kbbl/d 1.02 0.21 0.34 0.33 -

(kt/d) 0.17 0.03 0.05 0.05 0.03

Case a :Refinery fed with light to medium crude; Case b : Refinery fed with syncrude

Table 4.2: Refinery visbreaking unit (VBU) simulation results a) Case a

Feed NAPH Distillate TAR

API 7.5 54.7 36.0 4.7

sg 1.02 0.76 0.85 0.96

RON 67 - -

Flowrate:

kbbl/d 13.70 1.50 1.45 10.77

(kt/d) 2.22 0.18 0.19 1.78 b) Case b

Feed NAPH Distillate TAR

API 3.8 49.7 31.6 0.9

sg 1.046 0.781 0.868 1.069

RON 67 - -

Flowrate:

kbbl/d 1.02 0.21 0.34 0.33

(kt/d) 0.05 0.0041 0.0045 0.04


74

Heavy gasoil from DCU and VBU along with heavy gas oil from vacuum

distillation unit are sent to vacuum gasoil hydrotreating unit VGOHDS. Table 4.3

depicts the result from this unit.

Table 4.3: Refinery vacuum gasoil Hydrodesulfurization unit (VGOHDS) simulation results

a) Case a

Feed NAPH Distillate HGO

API 21.7 58.10 38.78 25.74

sg 0.93 0.75 0.83 0.90

S, wt% 2.43 0.01 0.27

RON 65

Flowrate:

kbbl/d 44.370 1.70 3.65 39.97

(kt/d) 6.52 0.20 0.48 5.72

Cetane number 52.6 b) Case b

Feed NAPH Distillate HGO

API 18.8 54.6 35.6 22.7

sg 0.95 3.1 7.4 87.75

S, wt% 2.43

0.01 0.27

RON 65

Flowrate: kbbl/d 1.51 0.058 0.125 1.366

(kt/d) 0.23 0.01 0.02 0.20

Cetane number 52.7


Naphtha, middle distillates and heavy gasoil are the main products from this

process with low sulphur content and but still without the requirements to be

final products, for this reason this products are sent to other processes like

reforming unit (REF).

Light and heavy gasoil from atmospheric distillation unit, VBU and DCU are

treated in middle distillate hydrotreating unit (DHT). Table 4.4 shows the main

products for this process, light naphtha and distillate with lower sulphur content

than the feedstock. Naphtha from this process is sent to reforming unit (REF).

75

Table 4.4: Refinery distillate hydrotreating unit (DHU) simulation results

a) Case a b) Case b

Feed NAPH

Distillate Feed NAPH

Distillate

API 22.3 43.8 24.4 29.8 52.4 32.0 sg 0.92 0.87 0.908 0.88 0.769 0.865 S (% w) 1.75 - 0.137 1.53 - 0.120 RON Flowrate: kbbl/d 78.56 1.52 76.19 127.5 2.47 123.90 (kt/d) 11.49 0.20 11.00 17.78 0.30 17.05 Cetane number

45.71 53.7 45 53.7


Kerosene from distillation unit was sent to KHT unit to low the sulphur content.

The results for this process are presented in Table 4.5.

Table 4.5: Refinery kerosene hydrotreating unit (KHU) simulation results

a) Case a b) Case b

Feed NAPH

Kerosene Feed NAPH

Kerosene

API 36.9 48.1 38.2 35.9 47.1 37.4 sg 0.84 0.79 0.83 0.845 0.79 0.84 S (% w) 0.70 0.054 RON 65 0.045 65 Flowrate:

kbbl/d

69.58 0.67 68.932

106 1.018 105.00 (kt/d) 9.30 0.08 9.14 14.25 0.13 13.99 Cetane number

54.4

54


The product from this unit is kerosene which can be sold to the market. The

small amount of naphtha produced is sent to NHT to be further processing.

Heavy Naphtha from distillation unit, DCU and VBU are sent to NHT to get

product with low sulphur content as can be seen in Table 4.6. Naphtha product

from this unit met the sulphur requirement but the octane number (RON) do not

76

meet the quality requirements to be considered as transportation fuel (> 90), for

this reason this naphtha is sent to reforming process (REF), the results for this

process is showed in Table 4.7, where can be observed that naphtha product

met the octane number required (> 90). This product is considered as blending

component to get the final gasoline.

Table 4.6: Refinery naphtha hydrotreating unit (NHU) simulation results a) Case a b) Case b

Feed

Naphtha Feed

Naphtha

API 64.9 66.7 55.4 57.1

sg 0.72 0.71 0.757 0.750

S (% w) 0.134 0.0001 0.006 0.0001

RON 65.27 64.7 65 66.4

Olefins 6.15 0.9 0.257 0.039

Flowrate:

kbbl/d 47.20 47.54 24.68 24.92

(kt/d) 5.40 5.40 2.97 2.97


Table 4.7: Refinery reforming unit (REF) simulation results a) Case a b) Case b

Feed

(NHT-NAPH) REFORMATE Feed

(NHT-NAPH) REFORMATE

API 66.7 45.6 57.1 37.0

sg 0.71 0.80 0.750 0.840

RON 65 100 66.4 100

Flowrate:

kbbl/d 47.54 37.82 24.91 19.82

(kt/d) 5.40 4.80 2.97 2.65


Table 4.8 presents the results for simulation of FCC unit. In FCC, the heavy

gasoil from VGOHDS is converted into gasoline light and heavy cycle oil and

slurry. Naphtha produced has a high octane number (91) and can be

considered as the main component in the gasoline. The flowrate is lower when

syncrude is used as feedstock because most of the middle distillate production

77

is in the range of diesel. For both cases the amount of heavy ends production is

lower in comparison with the upgrader. LCO and HCO are final products and

slurry is taking out as heavy ends.

Table 4.8: Refinery FCC unit simulation results

a ) Case a

Feed LNAPH HNAPH LCO HCO SLURRY

API 25.7 71.1 50.5 20.0 17.4 - sg 0.900 0.698 0.778 0.934 0.950 -

S, wt% 0.27 0.03 0.03 0.40 0.65 3.24

N, wt%

RON

91 91

Flowrate:

kbbl/d 39.97 8.49 14.62 5.54 2.86 -

(kt/d) 5.72 0.94 1.81 0.82 0.43 0.27

b ) Case b

Feed LNAPH HNAPH LCO HCO SLURRY

API 22.7 71.1 50.5 17.6 11.5 - sg 0.918 0.698 0.778 0.949 0.990 - S, wt% 0.27 0.03 0.03 0.40 0.64 2.92

N, wt%

RON

92 92

Flowrate:

kbbl/d 1.36 0.30 0.51 0.19 0.1 - (kt/d) 0.20 0.03 0.06 0.03 0.02 0.01


Light naphtha directly from crude distillation unit is processed in the

isomerization unit to increase the octane number. Table 4.9 presents the results

for this process where an isomerate is produced with an octane number of 82.

78

Table 4.9: Refinery isomerization unit (ISO) simulation results

c) Case a d) Case b

Feed

(C5-NAPH) ISOMERATE Feed

(C5-NAPH) ISOMERATE

API 58.9 53.2 49.00 53.2

sg 0.74 0.77 0.78 0.77

RON 65 82 65 82

Flowrate:

kbbl/d 30.35 28.9 39.83 39.98

(kt/d) 3.59 3.52 4.96 4.87


The products for refinery are high octane number gasoline, diesel, kerosene

and heavy products like LCO, HCO and coke. The properties for the main

products gasoline, diesel and kerosene, are showed in Table 4.10. The high

octane number gasoline is made of naphtha from hydrotreater unit,

isomerization unit an FCC unit. Diesel is made of DHT distillate and VGOHDS

distillate, and Kerosene from KHT distillate. When light crude oil is fed to a

refinery, gasoline is produced as a main product containing a octane number of

92, 49.85 kPa in RVP and 0.78 in specific gravity to meet the requirements, an

octane number higher than 90 met the requirement as can be seen in Table

4.10. Kerosene also met the requirements in the most important properties

while Diesel met the main property, that is cetane number but the specific

gravity did not meet, probably because a heavy gasoil is part of the feedstock.

The amounts of product are comparable: 11.47 for gasoline, 11.63 for diesel

and 9.22 for kerosene, respectively.

79

Table 4.10: Refinery main products, utilities and hydrogen demand. Light

crude oil feedstock.


Gasoline

Diesel

Kerosene

Gasoline

Diesel

Kerosene

RON, min >90 92.35

RVP, (kPa) 60-65 49.85

S, (%) 0.010 0.03 0.010 0.023

sg, max 0.78 0.845 0.84 0.78 0.91 0.83

Cetane number (min)

45

Cetane number

53.65

Mass flow (kt/d)

11.47 11.63 9.22

Heavy products

Mass flow (kt/d)

Light cycle oil (LCO)= 0.82

Heavy cycle oil (HCO)= 0.43

Slurry= 0.27

Visbreaker products (VB)=1.78

Coke= 1.36

Total utilities and hydrogen balance

Fuel

(Gcal/d)

High pressure

steam (kt/d)

Low pressure

steam (kt/d)

Cooling

Water (m3/d)

10-3

Power

(MWh/d)

H2

(kt/d)

TOTAL -24814 -2.47 -2.62 -880.66 -1574.18 -0.15

Table 4.11, shows the results for the main products, utilities and hydrogen

requirements for a refinery using syncrude as feedstock. Gasoline met the

octane number but the specific gravity did not meet the specification; it is a little

higher in comparison with the specification and with the refinery fed with light

crude oil. Kerosene met the quality requirements and diesel met the cetane

80

number but did not meet the specific gravity, probably because of the quality of

the cut sent to this unit (heavy gasoil).

It also can be observed in both cases that all the demand values are negative,

this mean that the utilities and hydrogen requirements are greater than its

generation in the refinery. Therefore it is necessary to retrofit or to import

utilities and hydrogen.

Table 4.11: Refinery main products, utilities and hydrogen demand.

Syncrude feedstock


Gasoline Diesel Kerosene Gasoline Diesel Kerosene

RON, min >90 90

RVP, (kPa) 60-65 71.5

S, (%) 0.010 0.03 0.010 0.03

sg, max 0.78 0.845 0.84 0.83 0.87 0.84

Cetane number (min) 45 53.7

Mass flow (kt/d) 7.92 17.25 14.18

Heavy products

Mass flow (kt/d)

Light cycle oil (LCO)= 0.03

Heavy cycle oil (HCO)= 0.02

Slurry= 0.01

Visbreaker products (VB)=1.041

Coke= 0.03

Total utilities and hydrogen balance

Fuel

(Gcal/d)

High pressu

re steam

(kt/d)

Low pressure

steam (kt/d)

Boiling feed water

(m3/d)

10-3

Cooling

Water

(m3/d)

10-3

Power

(MWh/d)

Chemicals

(kt/d)

H2

(kt/d)

TOTAL -19631 -1.635 -1.275 0 -756.29 -1190 0 -0.165

4.3 Refinery economic evaluation

The economic analysis for refinery fed with light crude oil and syncrude are

presented below. The analysis will be presented in two parts; first part is the

results for the simulation of refinery for different feedstock (light crude oil and

syncrude), and the second part is the sensitivity analyses of economic

indicators.

81

4.4 Capital cost analyses

The methodology for capital cost analyses was the same as described in

chapter 3 for upgrader.

Table 4.12 and Table 4.13 show the summary of refinery capital cost for

different refinery feedstock (light crude oil and syncrude) with different flowrates.

The cost is a result of the sum of the individual cost for different refinery units

and the utilities facilities like hydrogen and cooling water plants and contingence

cost. For both feedstock, as expected, the cost increase with the increase in

feed flow rate. The results show that the total cost increase from 729 to 1423

MM$ while the feed flowrate range between 100 to 300 Kbbl/d. When syncrude

is fed to refinery the cost range between 373.48 to 688.47 MM$ respectively.

Table 4.12: Summary of refinery capital cost. Light crude oil

Light crude oil flowrate (bbl)

100000 200000 300000

Process Units Feed

kbbl/d Cost

(MM$) Feed

kbbl/d Cost

(MM$) Feed

kbbl/d Cost

(MM$)

ADU 100.00 95.52 200.00 136.03 300.00 167.28

VDU 40.98 57.87 81.95 81.44 122.93 99.46

ISO 10.12 12.79 20.23 18.92 30.35 23.79

REF 15.85 43.01 31.70 63.94 47.54 80.63

NHT 12.13 20.44 24.26 29.40 36.40 36.35

KHT 23.19 58.65 46.39 87.43 69.58 110.43

DHT 9.98 36.08 19.96 53.79 29.94 67.94

VGHDS 14.79 90.58 29.58 161.48 44.37 226.45

FCC 13.32 137.44 26.65 195.72 39.97 240.68

VBU 4.56 19.88 9.13 33.22 13.69 44.86

DCU 15.22 156.83 30.43 245.07 45.65 318.19

SUBTOTAL 729.10 1159.09 1423.77 Hydrogen Production Unit 6.05

9.17

11.69


30.00

45.00

Storage 250.00 500.00 750.00

Offsites 150.02 246.84 333.41

SUBTOTAL 1150.17 1892.44 2556.18

Contingence 172.53 283.87 383.43

TOTAL 1322.70 2176.31 2939.60

82

It is important to note that when syncrude is used as a feed, a drastically fall in

the cost of approximately more than 50% is because of the nature of syncrude.

Syncrude come from a refining process and it has less bottom of barrel than a

light crude oil which has not been treated previous to the distillation unit.

Practically the bottom of barrel processes when used syncrude is not necessary.

Table 4.13: Summary of refinery capital cost. Syncrude

Syncrude flowrate (bbl)

100000 200000 300000

Process Units

Feed to unit

kbbl/d Cost

(MM$)

Feed to unit

kbbl/d

Cost (MM$)

Feed to unit kbbl/d

Cost (MM$)

ADU 100.00 95.52 200.00 136.03 300.00 167.28

VDU 3.80 17.93 7.61 25.23 11.41 30.81

ISO 13.28 14.91 26.55 22.06 39.83 27.74

REF 8.30 29.72 16.61 44.18 24.91 55.71

NHT 8.14 16.59 16.29 23.85 24.43 29.50

KHT 35.34 74.74 70.67 111.42 106.01 140.73

DHT 39.44 79.63 78.88 118.70 118.33 149.93

VGHDS 0.51 5.42 1.01 9.67 1.52 13.56

FCC 0.46 24.56 0.91 34.98 1.37 43.01

VBU 0.10 1.19 0.20 1.99 0.31 2.68

DCU 0.34 13.56 0.68 21.19 1.02 27.51

SUBTOTAL

373.78 549.30 688.47 Utilities Facilities


24.55

31.30


30.00

45.00

Storage 250.00 500.00 750.00

Offsites 98.25 165.58 227.22

SUBTOTAL 753.23 1269.43 1741.99

Contingence 112.98 190.41 261.30

TOTAL 866.22 1459.84 2003.28

83

4.5 Operating cost analyses

The prices for utilities in this work were taken from different sources (Ulrich and

Vasudevan, 2006; Gary et al, 2007; Kraiser et al.,2007; EIA, 2012). As

explained before in chapter 3.

Table 4.14 and Table 4.15 report the summary of operating cost for refinery

using light crude oil and syncrude as feedstok. The correlations for utilities cost

estimation are a function of feed flowrate cost for this reason it is proportional to

the feedstock. Additionally, the operating cost are the same for both feedstock

because the operating cost function depend on the feedstock flowrate.

Table 4.14: Summary of upgrader operating cost. Light crude oil

Light crude oil flowrate (bbl) 100000 200000 300000 Cost


Fuel (Gcal/d) 114.09 228.17 342.26

Boiler Feed Water (m3/d) 0.00 0.00 0.00

LP Steam (ton/d) 14.89 29.79 44.68

HP Steam (ton/d) 18.89 37.78 56.67

Cooling Water (1000 m3/d) 10.84 21.68 32.52

Hydrogen (ton/d) 0.00 0.00 0.00

Power (MWh/d) 8.66 17.32 25.97

Chemicals ($/d) 0.00 0.00 0.00


Insurance 6.61 10.88 14.70

Maintenance 66.13 108.82 146.98


TOTAL

265.50

505.20

739.94

84

Table 4.15: Summary of refinery operating cost. Syncrude

Syncrude flowrate (bbl) 100000 200000 300000 Cost


Fuel (Gcal/d) 111.59 223.17 334.76


LP Steam (ton/d) 8.40 16.80 25.19

HP Steam (ton/d) 12.49 24.98 37.46

Cooling Water (1000 m3/d) 9.31 18.62 27.93

Hydrogen (ton/d) 0.00 0.00 0.00

Power (MWh/d) 6.55 13.10 19.64

Chemicals ($/d) 0.00 0.00 0.00


Insurance 4.33 7.30 10.02

Maintenance 43.31 72.99 100.16


TOTAL 219.66 424.32 626.22

It is clear that the main contribution to the total operating cost are the fuel gas

and maintenance.

4.6 Discounted cash flow analyses


3.12. (Chapter 3).

4.6.1 Economic analyses of the refinery

Tables 4.14 and 4.17, summarizes the economic analysis for refinery with light

crude oil and syncrude for different feed flow rate. It can be observed that all the

values for the NPV and NPVI for both feedstock are greater than zero; it means

that all the schemes are profitable. The NPV comparison shows that the refinery

processing 300000 bbl/d of syncrude has the higher NPV of 15910.24 MM$,

while the equivalent using light crude oil as a feedstock is 4326.82 MM$ These

results indicate that the refinery is economically attractive for processing with

both feedstock,it is, light crude oil and syncrude. The NPVI also indicated that

the refinery for processing 300000 bbl/d of syncrude is the most profitable case.

85

Table 4.16: Economic evaluation for the refinery feed with light crude oil

Light crude oil flowrate (bbl)

100000 200000 300000

CAPEX (MM$) 1322.70 2176.31 2939.60


OPEX (MM$/year) 265.50 505.20 739.94

Taxes (MM$/year) 125.18 279.92 440.34

Cash flow (MM$) 149.35 294.94 439.81



Table 4.17: Economic evaluation for the refinery feed with syncrude

syncrude flowrate (bbl)

100000 200000 300000

CAPEX (MM$) 866.22 1459.84 2003.28


OPEX (MM$/year) 219.66 424.32 626.22

Taxes (MM$/year) 461.34 939.86 1421.53

Cash flow (MM$) 349.61 697.05 1044.08



Operating costs include the cost of raw material and capital costs include the

cost of the land and the working capital. The detailed calculations of the

economic evaluation for all cases are presented in Appendix C.

4.6.2 Sensitivity analyses of the economic indicators

A comparison between different sensitivity analyses for refinery fed with light

crude oil and syncrude is presented.

Figure 4.4 and 4.5 show the effect of increasing the light crude oil and syncrude

flowrate over the gross income and operating cost and taxes. It can be noticed

in Figure 4.4, how the increase in flowrate generates more operating cost

86

because of the more needed of utilities and obviously more income. The same

behaviour is observed in Figure 4.5 when the refinery is fed with syncrude


operating cost


operating cost

Comparing the results for both figures it can be noticed that when syncrude is

used more income is generated and more taxes probably because of the CO2

emission.

87

The Net present values and net present values index for both refinery

feedstocks are presented in figures 4.6 and 4.7. Figure 4.6 shows positives

values for both NPV and the NPVI, and the increase of these economic

indicators with the light crude oil flowrates, this means that the refinery fed with

light crude oil is economically attractive and increase with light crude oil

flowrates.

Figure 4.6: Effect of light crude flowrate over the refinery NPV and NPVI

Figure 4.7 shows the refinery VPN and VPNI with syncrude feedstock and

similarly to the previous case positives values for both NPV and the NPVI, and

the increase of these economic indicators with the syncrude flowrates, indicated

that this process is economically attractive and increase with syncrude flowrates.

88

Figure 4.7: Effect of syncrude flowrate over the refinery NPV and NPVI

Comparing both feedstocks, as expected, for refinery using syncrude as

feedstock the values for NPV (5106.14, 10.483.25 and 15910.24) and NPVI

( 5.33, 6.50 and 7.19) were higher than VPN obtained using light crude oil as

feedstock (1101.48, 2669.26 and 4326.82) and NPVI (0.75,1.11,1.33).

4.7 Chapter summary

In this chapter the refinery was evaluated for two feedstock, light crude oil and

syncrude. This evaluation allowed comparing the production, quality and the

economic indicator for both cases.

To evaluate the performance of these proposed processing schemes, rigorous

simulation models in Aspen Plus 2006.5 were developed for the refinery

distillation units. The other refinery units (Hydrocracking, hydrotreating, delayed

cooking and so on), were simulated using the petroleum refinery correlations.

The results have shown how profitable is the refinery for both feedstock. All the

values for the NPV and NPVI for both feedstock are greater than zero; it means

that all the schemes are profitable. The NPV comparison shows that the refinery

89

processing 300000 bbl/d of syncrude has the higher NPV of 15910.24 MM$,

while the equivalent using light crude oil as a feedstock is 4326.82 MM$. The

refinery is economically attractive for processing with both feedstock, it is, light

crude oil and syncrude. Refinery for processing 300000 bbl/d of syncrude is the

most profitable case with NPVI of 7.19.

90

Chapter 5 Integration of an Upgrader with a Petroleum Refinery

5.1 Introduction

Nowadays, the upgraders are exploited as the upstream facilities to existing

refineries. The upgraders are designed to produce synthetic crude oil (syncrude)

to be sold to a refinery for further processing and the refinery produces the final

transportation fuels (PDVSA, 2013 ).

As mentioned in chapter 2, upgraders and refineries have been largely treated

separately, for this reason in this chapter different integration opportunities

between upgraders and refineries are studied.

Two groups of proposed processing schemes for the integration between the

upgrader and the refinery were studied:

The first group includes two proposed processing schemes for the integration of

upgraders and refineries where the upgrader is considered as heavy crude oil

pre-treatment to be fed to the refinery. The first proposed scheme the syncrude

produced from the upgrader is sent to the refinery and in the second proposed

scheme some intermediate products from refinery which are not main blending

components in the final transportation fuels (diesel, gasoline and kerosene), are

processed in the upgrader and refinery and upgrader share utilities and

hydrogen demand.

The second group of proposed schemes includes two particular cases for the

integration between upgrader and refinery. Conventional upgraders produce as

final product light and sweet synthetic crude. These new schemes for

processing heavy crude oils for the production of final transportation fuels such

as gasoline and diesel instead of syncrude are presented as a particular case

for integration.

An economic evaluation based on NPV and NPVI for the different proposed

schemes is presented at the end of the chapter.

91

The integration of an upgrader with a refinery was performed using simulation

studies. Simulations were built in Aspen Plus and Excel environment.

The chapter includes the explanation for different schemes, the economic

evaluation for each cases, the comparison and selection of the best schemes.

Finally the chapter summary is presented.

5.2 Strategies for integration of an upgrader to a refinery

To evaluate the integration between the upgrader and the refinery two groups of

proposed processing schemes were studied. Following the explanation for each

group is presented in detail. The first group includes two cases:

5.2.1 Proposed processing scheme 1

The integration between an upgrader to a refinery for this case, as depicted in

Figure 5.1, has been studied. The synthetic crude oil (syncrude) produced in the

upgrader is dispatched to a refinery. The upgrader operates so that the

syncrude meets the quality constraints (density, sulphur, etc.) of the crude oils

that the refinery processes (see Table 3.8).

The upgrader considered for this proposed scheme has been studied earlier in

Chapter 3 and the refinery fed with the syncrude produced in this upgrader was

discussed in Chapter 4.

UPGRADER REFINERY

Refinery products

Heavy ends

Dilutedcrude

Diluent

DILUTION PROCESS

Heavy crudeproduced

Recycled diluent

Syncrudeto refinery

Heavy ends

Figure 5.1: Proposed processing scheme 1

92

Upgrader main products properties, utilities and hydrogen demand for this case

were shown in tables 3.8 and 3.9, in Chapter 3. The syncrude produced in the

upgrader was directly sent to the refinery and the results for the refinery: main

products, utilities and hydrogen demand were shown in table 4.11 in Chapter 4.

The advantage for this scheme in comparison with the two sites separately is

that refinery buys directly heavy crude oil instead light crude oils. The heavy

crude oil is cheaper and more abundant than light crude oil. The final decision

for this scheme in terms of economics is discussed at the end of this chapter.


Proposed processing schemes 2 as can be seen in Figure 5.2., the upgrader

receives intermediate products from refinery for further processing. As

discussed in chapter 4, refinery fed with syncrude has the advantage that no so

much bottom of barrel is produced, it is due to the feedstock is previously

treated in the upgrader. It could be a good opportunity for integration between

refinery and upgraders. The refinery bottom of barrel is sent to upgrader DCU

for further processing. Additionally refinery intermediate products like FCC

heavy products are sent to upgrader DCU for further processing. The refinery

will receive the synthetic crude from the upgrader. The total flow rate to the

refinery is 300,000 bbl/d. The effect of incorporating theses products on the

upgrader units and the quality for upgrader final synthetic crude oil to be sent to

refinery was evaluated. Additionally the final products from refinery were

evaluated and compared with those produced when refinery is fed with

synthetic crude without integration of intermediate products, utilities and

hydrogen, and with the refinery fed with light crude oil.

In this proposed processing scheme the demand for utilities and hydrogen was

evaluated taking account the total site (upgrader and refinery). To evaluate the

integration between upgrader and refinery the simulation was carried out under

the consideration mentioned in Chapters 3 and 4.

93

UPGRADER REFINERY

Refinery products

Heavy ends

Utilities and H2

Dilutedcrude

Diluent

DILUTION PROCESS

Heavy crudeproduced

Recycled diluent

Syncrudeto refinery

FCC Heavy products Heavy ends


The upgrader simulation was done including some intermediate products from

the refinery. Table 5.1 lists the flow rates and properties for the intermediate

products from refinery (LCO and HCO from FCC unit), which were considered

for further processing in upgrader units. The selection of the upgrader process

unit to send this refinery streams was done according to the intermediate

product density: the FCC light and heavy cycle oils are light and heavy gasoil

respectively, for this reason they were sent to upgrader hydrocracking unit and

the refinery vacuum distillation unit bottom of barrel which is a heavy residue

stream was fed to upgrader delayed coking unit (DCU).

Table 5.1: Refinery products to be processed in upgrader.

Properties of refinery products to be processed in upgrader

FCC LCO FCC HCO Vacuum

Residue

Mass flow (kt/d) 0.02 0.01 0.22

sg 0.97 1.06 3.8

S (wt%) 0.40 0.62 3.7

The mass balance in the upgrader and the utilities and hydrogen demand for

upgrader before and after the integration can be observed in Table 5.2. As can

be observed in Table 5.2, for the same syncrude production, less heavy crude

oil flow rate has to be fed after integration of the upgrader to a refinery due to

the incorporation of the refinery streams. The properties for the products in the

affected units (HTU, DCU and HCU), were not significantly affected because

the products from refinery have low flow rates. The mass flow rate fed to the

94

hydrocracking unit slightly decreased by 0.05 kt/d while the feedstock to DCU

and HTU increased by 0.11 kt/d and 0.02 kt/d, respectively.

The total consumption of utilities in the upgrader after integration is presented in

Table 5.2. It can be seen that total utilities consumption decreased, it probably

happened because the feed flow rate to distillation unit, the unit that has the

highest utilities requirement, is less in the integration scheme for this reason

less utilities are needed.

The quality of the upgrader products after integration with the refinery, along

with ASTM D-86 distillation data for new syncrude is shown in Table 5.3. The

addition of refinery intermediate products to the upgrader improve the quality of

the syncrude, it produce a slight decrease in its density from 34.4 ºAPI to 40.1

ºAPI after the integration. As can be observe in the syncrude distillation curve

the composition of syncrude is mostly intermediate products.

Table 5.2: Upgrader units balance. Proposed processing scheme 2

Updated upgrader mass balance (kt/d) in affected unit Hydrocracking unit (HCU)

Input LIGHT GASOLINE

NAPHTHA

DIESEL FUEL

Before integration 23.12 1.35 4.19 17.41 After integration 23.07 1.35 4.18 17.38

Delayed coking unit (DCU)

Input NAPHTHA LGO HGO COKE

Before integration 26.53 4.04 7.05 8.06 4.87 After integration 26.68 4.06 7.08 8.09 4.89

Hydrotreating unit (HTU)

Input NAPHTHA DIESEL

Before integration 17.46 0.05 17.41 After integration 17.48 0.05 17.43

Utilities and hydrogen demand after integration

Fuel

(Gcal/d) LP

steam (kt/d) LP steam

(kt/d)

Cooling Water

(m3/d).10-3 Power

(MWh/d)

H2

(kt/d) Before

integration -50500.12 -8143.55 -5122.69 -1116.82 3617.75

-0.991

After

integration -50412.90 -8142.33 -5106.34 -1114.51 3614.08

-0.990

95

Table 5.3: Upgrader main products after integration with refinery.

Proposed processing scheme 2

Properties of main product

Syncrude

40.1

1.1

39.35

Density (ºAPI)

S (%)

Mass flow rate (kt/d)

SYNCRUDE COMPOSITION , wt% STREAMS FROM DHT STREAMS FROM HCU

NAPHTHA DIESEL LIGHT NAPHTHA

HEAVY NAHPTHA

DIESEL

0.001 0.431 0.033 0.104 0.430 ASTM D-86 DISTILLATION data % Distillate Vs ºC

DIST , % 0 10 30 50 70 90 End point Temperature 52.67 147.93 230.45 259.47 287.22 319.57 346.73

The feed to the crude distillation unit in the refinery was the syncrude. All the

syncrude produced in the upgrader was sent to the refinery. The refinery

capacity is 300000bbl/d. Table 5.4 shows the properties of the refinery main

products (gasoline, kerosene and diesel), utilities and hydrogen demand after

integration. It can be noticed that the distribution of products is different from the

refinery base case; in this case less gasoline and more diesel and kerosene

flow rates are produced, these results are as expected because the syncrude

has more intermediate product than light crude oil.

The main quality specifications of the products were modified in comparison

with the refinery fed light crude, but the products still meet most specifications,

only the reid vapour pressure for gasoline is over the range but it can be adjust

using some additives. Additionally specific gravity for diesel is high as observed

in the based case.

96

Table 5.4: Refinery main products, utilities and hydrogen balance after

integration with upgrader. Proposed processing schemes 2

Properties of main products

Gasoline Kerosene Diesel

RON 91.3

RVP (kPa) 69.6

sg 0.79 0.83 0.91

Mass flow (kt/d) 7.85 14.12 17.06

Cetane number 53.65

Total utilities and hydrogen balance after integration

Fuel (Gcal/d)

HP steam (kt/d)

LP steam (kt/d)

Cooling water

(m3/d)10-3

Power (MWh/d)

H2 (kt/d)

-19442.7 -1626.7 -1251.62 -762.6 -1180.6 -0.164

The utilities consumption is considerably higher than those obtained using light

crude oil as a feed because for the integration some unit are not needed. The

consumption of hydrogen is higher it goes from 0.151 kt/d for light crude as a

feed to 0.164 kt/d for integration, it happen because more flow rate is sent DHT

unit and it increases the hydrogen consumption.

The second group of proposed schemes includes two particular cases for the

integration between upgrader and refinery considering the reduction in the unit

number. Two schemes for processing heavy crude oils to produce final products

are proposed, both produce transportation fuels and some heavy ends. The

proposed processing schemes are the following:

1. Production of transportation fuels (Diesel and Gasoline) using solvent

de-asphalted for residue processing and including reforming (REF),

diesel hydrotreating (DHT), and hydrocracking units (HCU) (see

Figure 5.3).

2. Production of transportation fuels (Diesel and Gasoline), using

delayed coking for residue processing and including reforming (REF),

97

diesel hydrotreating (DHT) and hydrocracking units (HCU) (see

Figure 5.4).

The procedure for the simulations of the two proposed processing schemes for

heavy crude oils is the same described in Chapter 3 for upgrader and Chapter 4

for refinery, using Aspen Plus 2006.5 (Aspen Plus, 2007) and the petroleum

refinery correlations (Baird, 1997).

The same volume for feed flow rate used for all the integration schemes is

considered (306500 bbl/d of heavy crude oil).


Figure 5.3 presents the first proposed processing scheme for processing heavy

crude oils to produce gasoline and diesel, and coke as heavy end. This scheme

includes atmospheric and vacuum distillation units (ADU and VDU), two primary

upgrading processes (delayed coking as residue processor and hydrocracking

unit (HCU)), and two secondary upgrading processes (reforming (REF) and

diesel hydrotreating units (DHT)).

Heavy oil is diluted with naphtha recycled from the atmospheric distillation unit

to produce the diluted heavy crude oil to be fed to the distillation unit. The

required amount of straight run naphtha for diluting heavy crude oil is recycled

without further treatment while the remaining is sent to the hydrotreating unit,

along with naphtha and light gasoil from the delayed coking unit, and the middle

distillates. Straight run gasoil, light and heavy vacuum gasoil are sent to the

hydrocracking unit from which light and heavy naphtha and diesel are produced.

The naphtha product from the hydrotreating unit is treated then in the reformer

unit to produce reformates which is a blending component in the gasoline. The

naphtha from hydrocracking unit is also sent as blending component in the

gasoline. The diesel product is made from diesel fuels produced in the

hydrocracking and hydrotreating units. In this scheme coke is produced in the

delayed coking unit as heavy end.

98

ADU

REF

HCU

DCU

BLEND

ING

SR naphtha

SR diesel

Light vacuum gasoil

SR gas oil

VacuumResidue

Diesel

Diesel

Heavy naphtha

Light naphtha

Coke

VDU

DIESEL

Naphtha Light gasoil

Heavy gasoil

RESERVOIR

naphtha For Recycling

Fresh naphtha

Light gasoil f rom DCU

Naphtha f rom DCU

Heavy vacuum gasoil Heavy oil

DilutedHeavy oil

HTU

Reformate

GASOLINE

COKE

Naphtha



This proposed processing scheme, shown in Figure 5.4, has a process for

processing heavy crude oil to produce transportation fuels (gasoline and diesel)

and asphalt as heavy ends. This scheme includes atmospheric and vacuum

distillation units (ADU and VDU), two primary upgrading processes (solvent

deasphalting unit (SDA) as residue processor and hydrocracking unit (HCU)),

and two secondary upgrading processes (reforming (REF) and diesel

hydrotreating units (DHT)).

In this case, the vacuum residue is sent to the solvent deasphalting unit to

produce a deasphalted oil (DAO) and asphalt as heavy ends. The deasphalted

oil is sent to the hydrocracking unit. The rest of the units are similar to those

described in the previous cases.

99

ADU

REF

HCU

SDA

BLEND

ING

SR naphtha

SR diesel

Light vacuum gasoil

SR gas oil

Vacuumresidue

Diesel

Diesel

Heavy naphtha

Light naphtha

VDU

DIESEL

DAO

RESERVOIR

naphtha For Recycling

Fresh naphtha

Heavy vacuum gasoil Heavy oil

DilutedHeavy oil

HTU

Reformate

GASOLINE

ASPHALT

Naphtha


The results for the simulation of both proposed processing schemes are shown

in Table 5.5, 5.6 and 5.7. In Table 5.5, the comparison of the mass flows (kt/d)

of the main products (gasoline and diesel), end products such as coke and

asphalt for both proposed schemes are presented.

Table 5.5: Mass flows of the products (kt/d) for proposed schemes 3 and 4

Gasoline Diesel Coke Asphalt

Proposed scheme 3

4.5

34.8

4.87

-

Proposed scheme 4

6.0 30.1 - 9.17

Comparing both schemes, it can be observed in Table 5.5 that when the solvent

desphalting unit is used the production of gasoline increases from 4.5 to 6.0 kt/d

100

and the diesel production decreases from 34.8 to 30.1 kt/d, respectively. The

increase of naphtha production is due to the greater amount of naphtha

produced in the hydrocracking unit, as a consequence of sending all DAO

production (17.4 kt/d) from the solvent deasphalting unit to the hydrocracking

unit. In proposed scheme 3 coke is a final product. It produces 4.87 kt/d while in

proposed scheme 4 heavier ends is produce, approximately 9.2 kt/d of asphalt

as a final product.

Table 5.6: Gasoline and diesel properties for proposed processing

schemes 3 and 4.

Gasoline Diesel

Proposed Schemes 3

RVP (kPa) 35.65

RON >95

Cetane number 52.0

Specific gravity 0.75 0.84

Gasoline Diesel

Proposed Schemes 4

RVP (kPa) 35.51

RON >95

Cetane number 52.4

Specific gravity 0.75 0.85

Properties of gasoline and diesel products are shown in Table 5.6. In all he

processing schemes, gasoline and diesel met the quality requirements needed

to be sent to the market as transportation fuels: RON, density and RVP for

gasoline, cetane number and density for diesel.

Utilities and hydrogen demand are shown in Table 5.7 for all the schemes. The

negative sign indicates net consumption, while the positive sign indicates net

production.

101

Table 5.7: Utilities and hydrogen demand for proposed schemes 3 and 4

Utilities and hydrogen demand

Fuel (Gcal/d)

LP steam (kt/d)

Cooling Water

((m3/d).10-3) Power

(MWh/d)

HP steam (kt/d)

H2

(kt/d)

Proposed scheme 3 -50610.1 -8785.6 -1154.8 -3713.8 -4799.1

-0.991

Proposed scheme 4

-46110.9 -8446.3 -1161.4 -3530.6 -4590.6 -0.930

As can be seen in Table 5.7, for both schemes 1 and 2, it is necessary to import

all the utilities including the hydrogen. Scheme 3 has more requirements in

utilities than schemes 4, it is because schemes two includes DCU which

requires more utilities than SDA.

The proposed schemes produce less gasoline but more diesel in comparison

with the refinery based case, because these schemes do not have FCC unit.

The economic analysis at the end of this chapter presents a comparison

between them.

5.3 Economic analysis of the integration between upgrader and refinery

The economic analysis was applied to the two groups of different proposed

processing schemes evaluated in this chapter, for the integration of an upgrader

to a refinery. In this point, two cases in each group have been considered:

proposed scheme 1: heavy cruel oil is fed to an upgrader and the syncrude

produced in this unit is directly sent to a refinery, proposed scheme 2: various

streams and units utilities were integrated between the upgrader and the

refinery. The second group of proposed processing schemes are: the scheme

to produce transportation fuels using DCU as bottom of barrel processor and

adding reformer to produce a naphtha with the properties of gasoline and the

last one the proposed scheme to produce transportation fuels using SDA as

102

bottom of barrel processor and adding reformer to produce a naphtha with the

properties of gasoline.

Table 5.8 summarizes the economic analysis for proposed processing schemes

1. For proposed processing scheme 1 the results show that all the values for

the NPV are positive and NPVI greater than 1; it means that all the schemes are

profitable. The NPV comparison shows that the upgrader NPV is

7226.7MM$ and 15,910.2 MM$ for refinery, both, without integration and

23136.9 MM$ for the upgrader - refinery integration. These results indicate that

the upgrader - refinery integration is economically attractive for processing

heavy crude oil. However, the NPVI indicates that refinery alone is more

profitable than the upgrader alone and the upgrader – refinery integration, due

to the CAPEX and OPEX for the last two scenarios are higher.

Table 5.8: Economic evaluation for proposed processing scheme 1

Upgrader Refinery Upgrader -

refinery integration

CAPEX (MM$) 4418.16 2,213.6 6421.45


OPEX (MM$/year) 1520.45 626.22 2146.67

Federal taxes (MM$/year) 759 1419.3 2178.3 CO2 emission taxes (MM$/year) 24.3 2.3 26.5

Cash flow (MM$) 702.6 1044.9 1746.7



Similarly to proposed processing scheme 1, Table 5.9 summarizes the

economic analysis proposed processing scheme 2. It can be observed that all

the values for the NPV are positive and NPVI greater than 1; this means that all

the schemes are profitable. The NPV comparison shows that the upgrader NPV

is 7486.2 MM$, the refinery NPV is 15994.0 MM$, both, without integration and

for the upgrader - refinery integration the NPV is 23535.4 MM$. These results

103

indicate that the upgrader - refinery integration is economically attractive for

processing heavy crude oil. On the other hand, the NPVI indicates that the

refinery alone is more profitable than the upgrader and the upgrader – refinery

integration, for the same reasons explained for processing schemes 1.

Comparing the two proposed processing schemes the upgrader and refinery

integration sending the syncrude to refinery and refinery sending some

intermediate and bottom product for further processing in the upgrader is the

more economically attractive scheme. Additionally, since light crude oil reserves

are decreasing worldwide, for the future upgrading of heavy crude oil and the

integration of upgrader and refinery units could be the only option to accomplish

the requirements for transportation fuels, which will grow in the next thirty years

since there is no other energy source comparable to crude oils, as seen in

Chapter 1 (OPEC, 2013; EIA, 2013).

Table 5.9: Economic evaluation for proposed processing schemes 2

Upgrader Refinery Upgrader -

refinery integration

CAPEX (MM$) 4405.22 1941.3 6346.55


OPEX (MM$/year) 1517.68 620.2 2137.93

Federal taxes (MM$/year) 781.0 1425.1 2,210.8

CO2 emission taxes (MM$/year) 24.3 2.3 26.5

Cash flow (MM$) 716.6 1044.9 1764.7



The results for second group of processing schemes are showed in tables 5.10.

Two proposed processing schemes for heavy crude oil to produce

transportation fuels instead of syncrude are compared economically in Table

5.10.

104

Table 5.10: Proposed processing schemes 3 and 4

Case Proposed scheme 3 Proposed scheme 4

Units for Heavy Process

CAPEX (MM$) 4509.64 4215.13

Gross Income (MM$/year) 13420.2 12194.2

OPEX (MM$/year) 1502.07 1616.29

Federal taxes (MM$/year) 2223.9 1432.5 CO2 emission taxes (MM$/year) 20.9 18.7

Cash flow (MM$) 1687.2 1147.0

Net Present Value (MM$) 24267.6 15196.2

Net Present Value index 4.87 3.26

As can be observed, all the proposed processing schemes for heavy crude oil

have positive values of NPV and NPVI, which indicate that all of them are

profitable. The proposed schemes for processing heavy crude oil, using delayed

coking and hydrocracking units as primary upgrading technologies has a higher

NPV (24267.6 MM$) than those using solvent deasphalting units (15196.2

MM$). The NPVI are 4.87 and 3.26 respectively. This makes proposed scheme

3 the best option to select; it has the higher NPV and NPVI for the integration

options. Still the refinery alone is more profitable but as explained before the

integration feeding directly heavy crude oil probably will be the solution for

produce transportation fuels in the future.

The detailed calculations of the economic evaluation for all cases are presented

in Appendix C.

5.4 Chapter summary

In this chapter different proposed processing schemes for the integration of an

upgrader to a refinery have been explored. Little information related to this topic

has been reported in open literature. Most work so far has studied the upgrader

and the refinery individually. However, the growing need to process the heavy

crude oil requires that different proposed processing schemes are studied for

105

processing heavy crude oils to satisfy the future requirements in transportation

fuels.

To evaluate the performance of these proposed processing schemes, rigorous

simulation models in Aspen Plus 2006.5 were developed for the upgrader and

the refinery distillation units. The other refinery units (Hydrocracking,

hydrotreating, delayed cooking and so on) were simulated using the petroleum

refinery correlations (Baird, 1997).

The properties of the refinery products were met when a syncrude was

processed in the refinery; additionally the integration allowed disposing heavy

products of the FCC unit and the bottom of barrels from refinery vacuum

distillation unit in the upgrader installations. The number of process unit in the

refinery diminished after integration.

In all proposed processing schemes for the integration of an upgrader to a

refinery some heavy ends such as coke and asphalt are produced, which can

be sent to other processes to produce more valuable products.

Conventional upgraders produce as a final product a light and sweet synthetic

crude oil called syncrude which is sent to a refinery for further processing. New

schemes for processing heavy crude oils for the production of final

transportation fuels such as gasoline and diesel instead of syncrude were

presented as a particular case for integration. Gasoline and diesel with the

quality properties required for transportation fuels were obtained from all the

proposed schemes. In scheme 4, which uses solvent deasphalting as residue

processor, the amount of heavy ends produced was greater than scheme 3

which uses delayed coking.

The net present value and net present value index have been used as methods

for the economic comparison of the processing schemes. All the processing

schemes have a positive NPV and NPVI, which indicates that all the processing

schemes studied are profitable. The NPV analysis shows that the integration of

the upgrader and refinery in all cases studied is attractive economically.

106

The best processing scheme includes distillation, delayed coking, hydrocracking,

hydrotreating and reforming units to produce gasoline, diesel, and the heavy

end (coke).

The refinery alone is more profitable but the schemes for integration where the

refinery feed is directly heavy crude oil probably will be the solution for produce

transportation fuels in the future.

The proposed schemes for processing heavy crude oil which includes delayed

coking and hydrocracking units as primary upgrading technologies (proposed

scheme 3 ) has a higher NPV (24267.6 MM$) and NPVI (4.87). This makes this

scheme the best option for the integration. It is important to mention that

delayed coking is the primary upgrading technology used for upgrading heavy

oils.

107

Chapter 6 Integration between upgrader, refinery and integrated gasification combined cycle (IGCC)

6.1 Introduction

The objective in this chapter is the evaluation of different opportunities for

integration between upgraders, refineries and integrated gasification combined

cycle plant (IGCC) to process the heavy fractions and to produce high value

products such as hydrogen, power and steam to satisfy the requirements for the

total site.

In this chapter IGCC plants receiving different feedstocks from refinery,

upgrader, refinery- upgrader integration, were compared economically. It is

important to mention that the capital costs for the IGCC unit are really high;

therefore only schemes with high level integration could be attractive for

processing heavy crude oil. Previous work showed that if upgrader-refinery-

IGCC integration is well exploited, this integration will be profitable, considering

the transformation of heavy cuts into more valuable light cuts, and power and

hydrogen savings (Sadhukhan, 2002).

This chapter include three parts: the first part addresses the simulation of the

IGCC process using Aspen plus 2006.5 (Aspen plus 2008), all the details about

the model and the validation of the model. The second part is the evaluation for

the different integration schemes: upgrader and IGCC; refinery and IGCC; and

upgrader, refinery and IGCC. For all cases sensitivity analyses for power or

hydrogen production and decarbonisation are analysed. Additionally, the

economic analysis includes firstly the sensitivity analysis that shows the

changes in the economic indicator with the fraction to power, the

decarbonisation grade and the carbon tax for all integration schemes are

presented and finally the comparison of all the schemes with a fixed values for

fraction to power, decarbonisation grade and carbon tax.

6.2 Integrated gasification combined cycle

Figure 6.1 presents the flowsheet for gasification combined cycle for

simultaneous production of power and H2. The feedstock, which can be coal or

108

heavy ends from upgrader or refinery processes, along with oxygen and steam

are fed to gasifier reactor (GASIFIER). The reactor reaches a temperature of

1,300 ºC and the main exothermic reactions, partial oxidation, steam reforming

and water gas shift, happen. Synthetic gas or syngas containing mainly

hydrogen, hydrogen sulphide and carbon monoxide (H2 + H2S + CO) is

produced. The hot syngas is sent to syngas cooler (COOLER) where

superheated steam is generated, which can be sent later to the steam turbine to

generate power to be used in the site. Ashes are removed from the syngas, in

the cyclone (CYCLONE) and 99.9 % of the sulphur is removed from the gas in

the hydrogen sulphide removal unit (CLAUS). In this case, the clean syngas,

composed mainly of CO and H2, can be divided in two streams; one stream is

sent to the high and low temperature water gas shift converter reactors

(HTWGSR and LTWGSR) for the conversion of CO into H2 and CO2. The other

part of clean syngas can be mixed with refinery light gases, light liquid fuels, off

gases from hydrogen network and sent to a gas turbine (GAS TURBINE) for the

production of power.

GASIFIER

HPSteam

Feed

Oxygen

COOLER

HTWGSR/LTWGSR(CO+H2O�CO2+H2)

H2SEP

Syngas:H2COH2S

CLAUSS

H2S

CO2 *

H2

Clean SyngasTo

Export

Gas Turbine

Power

Natural Gas

Air

To ExportUpgrade/refineryr

Utility Network

HRSG

Steam

Steam

Steam

Turbine

Power

To ExportUpgrader/refinery

Utility Network

Steam

BFW

CO2Steam

CCS

To Export

To Hydrotreatin/hydrocracking

CO2 *CO2

Upgrader/refineryUtility Network

Figure 6.1: General flowsheet for IGCC process.

The product from the low temperature water gas shift reactor (LTWGSR) is sent

to a carbon capture and storage unit (CCS); this unit is represented for a CO2

109

separation column and CO2 compressor as appear in detail elsewhere (Ng et al.;

2010). The hydrogen rich stream is recovered via membrane or pressure swing

adsorption process (H2SEP), obtaining H2 at 99% to be used in

hydroprocessing.

In this IGCC plant, saturated and superheated steam is generated in syngas

coolers and heat recovery steam generators (HRSG) using gas turbine

exhausts. Steam can be used in the site for power production or can be

exported.

6.2.1 Simulation framework for IGCC

The simulation and analysis of IGCC plants is complicated due to the large

number of units involved, and the interaction between them require an

extensive computation and optimisation effort. In this work IGCC simulation was

based on rigorous steady-state simulation using the commercial software Aspen

Plus 2006.5 (Aspen Plus, 2008).

In this study, oxygen blown entrained-flow gasifier was selected. This type of

gasifier is the most flexible and appropriate for large IGCC plants (Ng et al.;

2010). It is the one used in GE (formerly Texaco) and E-Gas. Texaco and Shell

entrained flow gasifiers are used in approximately 75% of the gasification plants

all over the world (Emun et al., 2010). Both, Polk central IGCC plant and

Wabash IGCC plant in Indiana in United States use this type of gasifier to

process 2.2 kt/d of coal or petroleum coke.

The Figure 6.2 shows a simulation flowsheet in Aspen Plus. The specifications

for IGCC are based on Shell and Texaco gasification processes (Shell, 2000;

Higman and Van Der Burgt, 2003; Zhent et al., 2005).

6.2.1.1 Components, stream class and physical properties

Various types of components are involved in gasification feedstock such as

conventional, conventional solid (carbon) and non-conventional, e.g. ashes. For

conventional and non-conventional solids, the particle side distribution (PSD)

110

was an input to the simulation. When these types of components are present,

the option MCINCPSD for stream class has to be selected.

The activity coefficient model for the gasification reactor is the non-random two

liquid (NRTL). The properties of solid components were taking from the work of

Ng et al., (2010).

111

FEED BFW O2 6.00 7.00 STEAM GAS SYNGAS FUELGAS H2PRODUC AIR BFW2 BFW3 M GAS2 CO2+ H2

Temperature C 25.00 25.00 25.00 600.00 300.00 750.00 1300.00 450.00 450.00 450.00 25.00 25.00 25.00 673.80 100.00 30.00 250.00

Pressure bar 1.00 1.00 1.00 50.00 55.00 50.00 60.00 60.00 60.00 60.00 1.00 1.00 1.00 2.00 2.00 80.00 1.00 Mass Flow tonne/day 2800.00 3464.77 2038.40 2038.40 2800.00 3464.77 8303.16 8271.76 2481.53 5790.23 16914.65 1989.96 211.38 26592.61 26592.61 5755.46 448.49

Figure 6.2: Simulation flowsheet for coal IGCC unit

STEAM

6

7

GAS

COLDGASGAS1

ASH

SYNGAS

H2PRODUC

FUELGAS

G-TURBIN

M

POWER

W

GAS2

BFW 4

POWER1

W

9

10

L

11

CO2-L

POWER3

W

H2S

QSTEAM

Q

GAS3

CO2

H2+

VAP

H2

4

STEAM2

STEAM3

BFW

O2

8FEED

AIR

POWER2

W

CO2+

BFW 2

BFW 3

GASIFIER

CYCLONE

S2

G-TURBIN

S-TURBIN

COMBUST

CO2-COMP

SEP2

CLAUS

COOLER

SEP2

CCS

B5

SEP2

H2SEP

HTWGSR

LTWGSR

B1

B2

B3

P1

AIRCOMP

B9

B10

B11B12

B13

HRSG

WATER

12

N2

B14

SCSTEAM

6.2.1.2 Unit operation models

For the simulation of the gasification unit the model RGIBBS was chosen

according with previous works (Ng et al., 2010; Emun et al., 2010). RGibbs

models rigorous reaction and/or multiphase equilibrium based on Gibbs free

energy minimization (Aspen Plus, 2008). In this reactor the following products

were specified: H2O, N2, O2, C, CO, CO2, H2S, CH4, COS, CL2, SO2, SO3, NO,

NO2. RGibbs uses Gibbs free energy minimization with phase splitting to

calculate equilibrium. RGibbs does not require specifying the reaction

stoichiometry. RGibbs is recommended to model reactors with:

• Single phase (vapour or liquid) chemical equilibrium

• Phase equilibrium (an optional vapour and any number of liquid phases)

with no chemical reactions

• Phase and/or chemical equilibrium with solid solution phases

• Simultaneous phase and chemical equilibrium

RGibbs can calculate the chemical equilibria between any number of

conventional solid components and the fluid phases. RGibbs also allows

restricted equilibrium specifications for systems that do not reach complete

equilibrium (Aspen Plus, 2008).

Main reactions:

Gasification reactions:

Equations 1, 2 and 3 depict the main gasification reactions (Higman and Van

Der Burgt, 2003):

C + 1/2 O2 → CO (1)

C + H2O → CO + H2 (2)

CO + H2O → CO2 + H2 (3)

113

For the simulation of the HTWGSR and LTWGSR the model REquil was chosen

according to previous work (Ng et al., 2010), and for the combustion reactor

RStoic was selected. REquil is used to model reactors when:

• Reaction stoichiometry is known and

• Some or all reactions reach chemical equilibrium

REquil calculates simultaneous phase and chemical equilibrium; it also allows

restricted chemical equilibrium specifications for reactions that do not reach

equilibrium. The model can perform rigorous one- and two-phase equilibrium

reactors on stoichiometric approach, and it is used to model a reactor when:

• Reaction kinetics are unknown or unimportant and

• Stoichiometry and the molar extent or conversion is known for each

reaction

RStoic can model reactions occurring simultaneously or sequentially, in addition,

that can perform product selectivity and heat of reaction calculations (Aspen

Plus, 2008).

Combustion reactions:

CO + 1/2 O2 → CO2 (4)

H2 + 1/2 O2 → H2O (5)

CH4 + 2 O2 → 2 H2O + CO2 (6)

Reaction in HTWGSR and LTWGSR

CO + H2O → CO2 + H2 (6)

114

The rest of unit operations models to simulate the IGCC in Aspen Plus and their

specifications are presented in Table 6.1. The values of the specification

processes correspond to typical IGCC facilities (Shell, 2008; Coca, 2003).

6.2.1.3 Process specifications

The syngas final composition depends on the conditions of pressure,

temperature and also of gasifying agents (air, oxygen and steam).

At the gasification temperature (1300 ˚C) , the oxygen consumption is high, but

the challenge for gasification operation is to keep the oxygen close to the

minimum required, the less oxygen that is used, the more steam is needed and

the cost of the process will be less because oxygen is more expensive than

steam (Higman and Van Der Burgt, 2003). In this study stoichiometric amount

of oxygen for total combustion was calculated, and 1/3 of this value was used

as gasifying agent as appear in literature (Higman and Van Der Burgt, 2003).

FORTRAN blocks in Aspen Plus (Aspen Plus, 2008) were used to calculate:

• Amount of air to gas turbine from stoichiometric for the combustion

reactions

• Amounts of steam to HTWGSR and LTWGSR using the stoichiometric

amounts for the water gas shift reactions

• The mass balance in the separators from the split fractions of hydrogen

sulphide, carbon dioxide and hydrogen

• The total amount of steam to the steam turbine

115

Table 6.1: Unit Operation models for IGCC. Base case

Process in the IGCC Unit Operation in Aspen Plus

Model in Aspen Plus

Process specifications

GASIFIER Reactor RGibbs P= 55 bar T=1300 ˚C

Specification of possible products (see list above)

COOLER Heat Exchangers Heater P= 60 bar T=450 ˚C CYCLONE Mixer/Splitter SSplit Split fraction of gas=1

CLAUS Separator Sep2 Split fraction of H2S in

outlet stream=0.99

S2 SSplit

Split fraction of the stream going to the combined heat and power (CHP)=0.7

GAS TURBINE

Reactor (COMBUST)

RStoic P= 55 bar T=1260 ˚C Combustion reactions

(see above) pressure changers

(G-TURBIN) Compr (turbine)

Discharge pressure=2 bar

HRSG Heat Exchangers HeatX (Shortcut design model)

Outlet temperature: 100 ºC

Water inlet temperature:

S-TURBIN pressure changers Compr (turbine)

3 stages turbine T=100 ˚C


HTWGSR Reactor REquil P= 55 bar T=450 ˚C see reactions above)

LTWGSR Reactor REquil P= 55 bar T=250 ˚C

(see reactions above)

B3 Heater P= 55 bar T=250 ˚C

reactions

CCS

Separator (CCS)

Sep2

Mass flow Split fraction of CO2 in

outlet stream=0.99

pressure changers (CO2-COMP)

Compr Compressor


H2SEP Separator

Sep2

Mass flow Split fraction of H2 in outlet stream=0.999

6.2.2 Validation of IGCC model

Two previous works were used to validate the model built in ASPEN PLUS

2006.5 (Ng et al., 2010; Jung and Kim, 2007). In both works the scheme of

IGCC has been proposed to produce power. In this work the IGCC scheme is

for the production of hydrogen and power; however, the data of the work of Ng

et al. (2010) and Jung and Kim (2007), are appropriate for validation of the

gasifier unit and thermodynamic methods. The Gibbs free energy minimization

116

method to calculate the composition of the gasifier gas output was used in

these works.

In the first simulation based in the work of Ng et al. (2010), the thermodynamic

model NRTL for the simulation of IGCC was validated. In this simulation, coal

data from EIA were used (EIA, 2013)

The second simulation based on the work of Jung and Kim (2007) used Illinois

N˚ 6 coal as feedstock with different amount and different temperature and

pressure condition. In addition, for Illinois N˚ 6 coal operational data are also

presented and were used to compare with the syngas composition obtained in

the simulation. Table 6.2 shows the composition for Illinois N˚ 6 and IEA coal,

taken as a reference in this work

Table 6.2: Coal composition

Composition wt%

C H S N ASH O

Illinois N˚ 6 71.72 5.06 2.08 1.41 10.91 8.08 Coal

(IEA 2003) 82.5 5.60 1.10 1.77 9.00

Table 6.3 shows the comparison for syngas composition from previous works

and after simulating IGCC plant with Aspen plus 2006.5. As can be seen in

Table 6.3, there is an acceptable agreement between the CO and H2

composition of the two previous works and the simulation results. As a

conclusion the simulation model in Aspen plus is appropriated to evaluate the

behaviour of the IGCC plant with different feedstock and conditions.

117

Table 6.3: Validation of IGCC. Syngas composition

Syngas (Gas in Figure 4.10) composition, %mol

IEA Coal Illinois N˚ 6 coal

Simulation

Previous work

(Ng et al. (2010),

Simulation Previous

work (IEA, 2003)

Operational data

CO 29.27 28.6 63.85 63.8 61

CO2 9.65 10.5 2.51 1.5 1

H2 30.72 29.0 30.46 28.6 35

N2 1.15 1.2 2.14 3.4 1

CH4 0.02 1 1.1 1

COS + H2S 0.01 0.98 1.2 1.5

H2O 29.18 30.5

6.3 Schemes for integration between upgrader, refinery and IGCC

After validation of IGCC model, some proposed processing scheme were

studied to analyse different schemes for producing hydrogen and utilities with

low emission and high efficiency that satisfy the requirements of all parties on

the site.

Integration opportunities between upgraders and refineries with IGCC plants

were studied. Different options for interchanging streams between the refinery,

the upgrader and the IGCC unit can be generated. Both, refinery and upgrader

can send heavy ends to the IGCC and receive from it utilities and hydrogen for

hydrotreating processes.

6.3.1 IGCC Feedstocks

Table 6.4 lists the composition of different feedstocks corresponding to the

integration schemes considered in this work: coke from the integration

upgrader-IGCC, coke and visbreaking product from the refinery-IGCC

integration scheme, coke , visbreaking product or asphalt from the integration

upgrader-refinery-IGCC. The composition analysis listed in this table for

different feedstocks was calculated from the correlation used and compared

with literature (Higman and Van Der Burgt, 2003).

118

Table 6.4 Feedstocks to IGCC

Composition, wt%

C H S N Volatiles Ash O H2O

Upgrader coke 82.10 3.11 2.80 2.73 11.16 0.26 6.52

Refinery Coke 82.21 3.11 5.50 1.90 11.16 0.26 6.52

Refinery visbreaking

product 85.27 10.08 4.0 0.30 0.15 0.20

Table 6.5, 6.6 and 6.7 show the feedstocks to gasifier for the different schemes

for integration. Table 6.5 shows the integration upgrader and IGCC, Table 6.6

shows the integration refinery-IGCC with both feedstocks studied they are light

crude oil and syncrude and finally Tables 6.7 to 6.10 show the feedstock for

integration upgrader- refinery and IGCC for all proposed processing schemes

evaluated in chapter 5. The gasifier agent (O2) and the steam necessary for

gasification for different feedstock are also presented in theses tables.

Table 6.5 Feedstock amounts to gasifier. Upgrader-IGCC

Feedstock to gasifier

Syncrude production bbl/d 300

Requirements (kt/d)

Coke Oxidant (95% O2)100 Steam

1.62 2.09 0.39

Table 6.6 shows clearly that the feedstock to gasifier from the refinery being fed

with syncrude is less because practically not bottom of barrel is produced with

this feedstock.

Table 6.6 Feedstock amounts to gasifier. Refinery-IGCC


Refinery fed with

Light crude Refinery fed

with Syncrude

(kt/d)

Coke 0.92 0.03

VB product 1.20 0.04

Oxidant (95% O2) 2.46 0.08

Steam 0.51 0.02

119

Table 6.7 presents the feedstocks to gasifier for proposed processing schemes

5 and 6. Proposed processing scheme 5 corresponds to proposed processing

schemes 1 integrated with IGCC, for this scheme coke from refinery and

upgrader and visbreaking product is sent to gasifier. Proposed processing

scheme 6 corresponds to proposed processing schemes 2 integrated with

IGCC. As can be observed in this table, for this scheme the feedstock to IGCC

is only coke from upgrader because all bottom of barrel from refinery is

integrated with upgrader DCU for further processing.

Table 6.7 Feedstock amounts to gasifier.

Proposed processing scheme 5 and 6


Syncrude processed bbl/d: 300

Proposed scheme 5

(proposed scheme 1+IGCC)

Proposed scheme 6


(kt/d)

Upgrader Coke 4.87 4.89

Refinery Coke 0.03 -

Refinery VB product 0.04 -

Oxidant (95% O2) 5.67 5.61

Steam 1.19 1.17

Table 6.8 presents the feedstocks to gasifier for proposed processing schemes

7 and 8. Proposed processing scheme 7 corresponds to proposed processing

schemes 3 integrated with IGCC and proposed processing scheme 8

corresponds to proposed processing schemes 4 integrated with IGCC. As can

be observed in this table, the proposed scheme 8 which use SDA as a residue

processor produce more heavy ends.

120

Table 6.8 Feedstock amounts to gasifier.

Proposed processing scheme 7 and 8


Heavy crude oil processed bbl/d: 306.5

Proposed scheme 7

(proposed

scheme 3+IGCC)

Proposed scheme 8


Coke (kt/d) 4.87 -

Asphalt (kt/d) - 9.17

Oxidant (95% O2) 5.59 9.58

Steam (kt/d) 1.20 2.20

The IGCC unit was simulated for all these feedstocks. Table 6.9 shows the

results for the syngas composition after simulation of the IGCC with all different

feedstocks.

Table 6.9 Syngas composition for different Feedstock

Syngas composition, % wt using different feedstock

Upgrader-IGCC

(coke) Refinery-IGCC (Refinery

fed with Light crude ) Refinery-IGCC (Refinery

fed with Syncrude)

CO 63.06 41.08 37.74

CO2 22.83 4.79 6.86

H2 5.01 39.61 35.56

H2O 9.10 13.32 18.62

Proposed scheme 5

Proposed scheme 6

Proposed scheme 7

Proposed scheme 8

CO 77.20 77.17 77.17 72.73

CO2 7.88 7.88 7.88 7.42

H2 5.29 5.31 5.31 9.69

H2O 9.63 9.64 9.64 10.95

Once the syngas is produced it can be sent to combined heat and power and/or

hydrogen production.

A flowsheet for different alternatives for processing this syngas is represented in

Figure 6.3.

121

Figure 6.3: Alternative schemes for syngas processing (Hydrogen or CHP)

The clean syngas can be sent (as explained before) to hydrogen production that

includes water gas shift reactors, CO2 capture and hydrogen separation

(dashed line in Figure 6.3) or can be sent to combined heat and power that

includes gas and steam turbines for power generation and heat recovery steam

generator and CO2 capture (solid line). Different scenarios were evaluated for

the different integration schemes:

6.3.1 Maximum power and hydrogen production from IGCC

Sensitivity analyses were performed to evaluate the production of hydrogen and

power as a function of the amount of syngas sent to the hydrogen or combined

heat and power production (CHP) route for all the integration schemes.

a) Integration upgrader and IGCC

Figure 6.4 shows the results for sensitivity analyses for the upgrader feedstocks.

In this figure the amounts of hydrogen and power production in the IGCC,

respectively, are presented versus the fraction of syngas sent to power

production.

Gasification +syngas cooling

+ syngas cleaning

Hydrogen production + Carbon separation

Combined heat and power (CHP)

Syngas

?

Feedstock

122

Figure 6.4: Hydrogen and power production as a function of fraction to power. Feedstock: coke from upgrader

It can be observed in Figure 6.4 that when using coke from the upgrader as

feedstock the maximum production of hydrogen is 0.345 kt/d when all the

syngas is sent to the production of hydrogen.

Compared with the upgrader requirement (0.991 kt/d ) as can be seen in Table

6.10 there is a deficit, part of the amount of hydrogen has to be imported.

Table 6.10: Power and H2 requirements and production. Integration

upgrader and IGCC

Power and H2 requirements

Power (MW) H2 (kt/d)

150.73 0.991

IGCC maxima H2 and power production

Feedstock Power (MW)

(max) H2 (kt/d)

H2 (kt/d)

(max)

Power (MW)

Upgrader coke 270.4 to 0% CCS

263.8 to 100% CCS 0 0.35 0

The other limited operation condition is for maximum power production (Figure

6.4). In this figure two lines appear, the first one (solid green) is for production

without CCS and the second one (dashed purple line) is for production with

123

100% of CCS. Using the upgrading coke as feedstock the maximum power

generation is 270.4 MW without CCS and 263.8 MW for 100% of CCS. The

difference between them is the additional power required for CO2 compression.

The requirement of the upgrader is 150 MW (Table 6.10). There is an amount of

power that can be exported; with 55% approximately of the power produced the

upgrader requirement can be satisfied using upgrader coke as a feedstock.

b) Integration refinery and IGCC

Figure 6.5 and 6.6 show the results for sensitivity analyses for the refinery

fed with light crude and syncrude. In Figure 6.5 the amounts of hydrogen

and power production in the IGCC, respectively, are presented versus the

fraction of syngas sent to power production.

0.0

50.0

100.0

150.0

200.0

250.0

300.0

350.0

400.0

0.0

50.0

100.0

150.0

200.0

250.0

300.0

350.0

400.0

450.0

0.0 0.2 0.4 0.6 0.8 1.0P

ow

er

(MW

)

H2

(to

n/d

)

Fraction to Power

H2 Power without CCS Power with 100% CCS

Figure 6.5: Hydrogen and power production as a function of fraction to power. Feedstock: coke and VB product from refinery fed with light crude

oil

Figure 6.5 shows that when using coke and VB product from refinery fed with

light crude oil the maximum production of hydrogen is 0.42 kt/d, it correspond

to all the syngas sent to the production of hydrogen.

Compared with the refinery requirement (0.15 kt/d ) as can be seen in Table

6.11 this amount of hydrogen can satisfy the refinery hydrogen requirements

and the remaining can be exported.

124


6.5). In this figure two lines appear, the first one (solid green) is for production

without CCS and the second one (dashed purple line) is for production with

100% of CCS. Using the light crude oil as feedstock in refinery the maximum

power generation is 338.7 MW without CCS and 330.1 MW for 100% of CCS.

The narrow difference between them as explained for the previous case is the

additional power required for CO2 compression. The power requirement of the

refinery is 65.6 MW (Table 6.11). There is an amount of power (264.51 MW)

that can be exported.

Table 6.11: Power and H2 requirements and production. Integration

refinery and IGCC

Power and H2 requirements

Refinery (light crude oil) Refinery (syncrude crude oil)

Power (MW) H2 (kt/d) Power (MW) H2 (kt/d)

65.6 0.15 49.6 0.17



(max)

H2

(kt/d)

H2 (kt/d)

(max)

Power

(MW)

Coke from Refinery (fed with

light crude oil)

338.7 to 0% CCS

330.1 to 100% CCS 0 0.42 0

Coke from Refinery (fed with

syncrude crude oil)

11.5 to 0% CCS

11.3 to 100% CCS 0 0.015 0

Figure 6.6 shows the results for IGCC production when the feedstocks coke and

VB products come from refinery fed with syncrude. The results have shown that

when using syncrude as feedstock to refinery the production of heavy products

(coke and VB products) is really insignificant in comparison with the previous

case, for this reason the maximum production of hydrogen from IGCC is only

0.015 kt/d.

Compared with the refinery requirement (0.17 kt/d ) as can be seen in Table

6.11 this amount of hydrogen cannot satisfy the refinery hydrogen

requirements and has to be imported.

125


6.6). Using syncrude as feedstock the maximum power generation is 11.5 MW

without CCS (solid line) and 11.3 MW for 100% of CCS. The narrow difference

between them as explained for the previous case is the additional power

required for CO2 compression. The power requirement of the refinery is 49.6

MW (Table 6.11). There is a deficit, of power, 38.3 MW has to be imported.

The integration Refinery fed with syncrude and IGCC probably is not necessary

because not much bottom of barrel is produced.

0.0

2.0

4.0

6.0

8.0

10.0

12.0

14.0

0.0

2.0

4.0

6.0

8.0

10.0

12.0

14.0

16.0

0.0 0.2 0.4 0.6 0.8 1.0

Po

we

r (M

W)

H2

(to

n/

d)

Fraction to Power


Figure 6.6: Hydrogen and power production as a function of fraction to power. Feedstock: coke and VB product from refinery fed with syncrude

c) Integration upgrader- refinery and IGCC

In this case the heavy products from integration of upgraders and refinery

are sent to IGCC to produced H2 and power to supply the requirements from

upgrader and refinery. All proposed processing schemes studied are

presented.

Proposed processing scheme 5:

Figure 6.7 shows the results for proposed processing schemes 5, here the

feedstock considered were coke from upgrader and refinery and refinery VB

126

product as seen in Table 6.7. The maximum production of hydrogen was

1.12 kt/d when all the syngas is sent to the production of hydrogen.

Compared with the total site hydrogen requirement (1.16 kt/d) as appear in

Table 6.12, this amount of hydrogen cannot satisfy the hydrogen

requirements. Part of the hydrogen has to be imported.

0.0

100.0

200.0

300.0

400.0

500.0

600.0

700.0

800.0

900.0

1000.0

0.0

200.0

400.0

600.0

800.0

1000.0

1200.0

0.0 0.2 0.4 0.6 0.8 1.0

Po

we

r (M

W)

H2

(to

n/d

)

Fraction to Power


Figure 6.7: Hydrogen and power production as a function of fraction to power. Proposed processing scheme 5

The other limited operation condition is for maximum power production

(Figure 6.7). In this figure two lines appear, the first one (solid green) is for

production without CCS and the second one (dashed purple line) is for

production with 100% of CCS. Using refinery and upgrader heavy ends as

feedstock the maximum power generation is 876.5 MW without CCS and

852.3 MW for 100% of CCS. The power requirement of total site was 185.4

MW (Table 6.12). The power requirements for the total site can be met with

this integration; there is an amount of power that can be exported.

127

Table 6.12: Power and H2 requirements and production. Proposed

processing scheme 5 and 6

Power and H2 requirements for total site

Proposed processing scheme 5 Proposed processing scheme 6


185.4 1.16 199.7 1.15

IGCC maxima H2 and power production in IGCC


(max)

H2

(kt/d)

H2 (kt/d)

(max)

Power

(MW)

Coke from upgrader and coke and

VB product from refinery

(proposed processing scheme 5)

876.5 to 0% CCS

852.3 to 100% CCS 0 1.12 0

Coke from upgrader


877 to 0% CCS

857 to 100% CCS 0 1.13 0

The other limited operation condition is for maximum power production

(Figure 6.7). In this figure two lines appear, the first one (solid green) is for

production without CCS and the second one (dashed purple line) is for

production with 100% of CCS. Using refinery and upgrader heavy ends as

feedstock the maximum power generation is 876.5 MW without CCS and

852.3 MW for 100% of CCS. The power requirement of total site was 185.4

MW (Table 6.12). The power requirements for the total site can be met with

this integration; there is an amount of power that can be exported.


Figure 6.8 shows the results for proposed processing schemes 6, here the

feedstock considered was only coke from upgrader as seen in Table 6.7.

128

0.0

100.0

200.0

300.0

400.0

500.0

600.0

700.0

800.0

900.0

1000.0

0.0

200.0

400.0

600.0

800.0

1000.0

1200.0

0.0 0.2 0.4 0.6 0.8 1.0

Po

we

r (M

W)

H2

(to

n/d

)

Fraction to Power



As can be observed in Figure 6.8, the maximum production of hydrogen is

1.13 kt/d when all the syngas is sent to the production of hydrogen.

Compared with the total site requirements (1.15 kt/d) as can be seen in

Table 6.12, this amount of hydrogen is not enough to satisfy the total site

hydrogen requirements, it needs to be imported.

Maximum power production as can be seen in Figure 6.8 is 877 MW without

CCS (solid line) and 857 MW for 100% of CCS (dashed line). The power

requirement for the total site was 199.7 MW (Table 6.12). The power

requirements for the total site can be met with this integration scheme; there

is an important amount of power that can be exported.


Figure 6.9 shows the results for proposed processing schemes 7. In this

proposed scheme as presented in Table 6.8 the feedstock considered was

coke from DCU. The maximum production of hydrogen for this scheme is

1.10 kt/d, compared with the total site requirements (0.99 kt/d) as can be

seen in Table 6.13, this amount of hydrogen can satisfy the total site

hydrogen requirements and part of it can be exported.

129

0.0

100.0

200.0

300.0

400.0

500.0

600.0

700.0

800.0

900.0

1000.0

0.0

200.0

400.0

600.0

800.0

1000.0

1200.0

0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0

Po

we

r (M

W)

H2

(to

n/d

)

Fraction to Power



Maximum power production as can be seen in Figure 6.9 is 865 MW without

CCS (solid line) and 845 MW for 100% of CCS (dashed line). The power

requirement for the total site was 154.7 MW (Table 6.13). The power


is an important amount of power that can be exported.

Table 6.13: Power and H2 requirements and production. Proposed

processing scheme 7 and 8

Power and H2 requirements for total site

Proposed processing scheme 7 Proposed processing scheme 8


154.7 0.99 147.1 0.93



(max)

H2

(kt/d)

H2 (kt/d)

(max)

Power

(MW)

Coke


865 to 0% CCS

845 to 100% CCS 0 1.10 0

Asphalt


2033 to 0% CCS

2000 100% CCS 0 2.77 0

130


Figure 6.10 shows the results for proposed processing schemes 8. The

feedstock considered for this proposed scheme was asphalt from SDA (see

Table 6.8). The maximum production of hydrogen for this scheme was 2.77

kt/d, compared with the total site requirements (0.93 kt/d) as can be seen in

Table 6.13, this amount of hydrogen can satisfy the total site hydrogen

requirements and an important part of it can be exported.

0.0

500.0

1000.0

1500.0

2000.0

2500.0

0.0

500.0

1000.0

1500.0

2000.0

2500.0

3000.0

0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0

Po

we

r (M

W)

H2

(to

n/

d)

Fraction to Power



Maximum power production as can be seen in Figure 6.10 was 2033 MW

without CCS (solid line) and 2000 MW for 100% of CCS (dashed line). The

power requirement for the total site is 147.1 MW (Table 6.13). The power


is an important amount of power that can be exported. The production

hydrogen and power for this proposed scheme is the highest, however the

economic analysis will show the profitability for all proposed scheme.

In general the requirements in power are met with no more than 30 % of

syngas to the power production, only for upgrader and IGCC integration

schemes the power requirements are higher. The production of hydrogen

only can be satisfied in some cases.

The main advantages of the proposed schemes for the integration of

upgrader-refinery and IGCC is that more products are obtained for example

131

power and hydrogen that cannot be produce in the schemes proposed by

Aguilar et al. (Aguilar et al., 2012) .

6.3.2 Economic analysis for integration schemes

All integration schemes: upgrader and IGCC; refinery and IGCC, and upgrader -

refinery and IGCC were evaluated economically. The economic evaluation

includes the sensitivity analysis of these indicators with the fraction to power,

the decarbonisation grade and the carbon tax, and then with these values fixed

a comparison for all cases is presented.

a) Integration upgrader and IGCC

The net present value index as a function of fraction to power is presented in

Figure 6.11. The NPVI indicates that is better to produce hydrogen than power.

0.0

0.2

0.4

0.6

0.8

1.0

1.2

1.4

1.6

0.0 0.2 0.4 0.6 0.8 1.0

NP

V IN

DEX

Fraction to Power

Figure 6.11: NPVI vs. fraction to power. Upgrader-IGCC integration

Figure 6.12 shows the NPVI for different syncrude productions with and without

CCS. It can be observed that the NPVI increases with the syncrude production

for both schemes with and without CCS. As expected, the NPVI is lower for

100% CCS.

132

Figure 6.12: Effect of syncrude production on NPVI. Upgrader-IGCC

integration

It is important to see how important the CCS for the integration upgrader-IGCC

plants is, the Figure 6.13, shows the NPVI as a function of CO2 tax with and

without CCS. It can be observed that the NPVI is highly influence with the

increase in carbon tax. When carbon tax increases the profitability of the project

without CCS can be less than those using CCS, apart from that the

environmental problem is reduce using CCS facilities.

0.50

0.60

0.70

0.80

0.90

1.00

1.10

1.20

1.30

1.40

0 5 10 15 20 25 30 35 40

NP

V IN

DEX

CO2 Tax ($/ton)

100 % CCS 0% CCS

Figure 6.13: Effect of CO2 tax over the NPVI. Upgrader-IGCC integration

133

b) Integration refinery and IGCC

Figure 6.14 shows the net present value as a function of fraction to power with

and without carbon capture and storage (CCS) for refinery fed with light crude

and syncrude respectively. For both cases the NPVI indicates that is better to

produce hydrogen than power. The NPVI without CCS is higher than the NPVI

for 100 % CCS because additional investment is needed for CCS.

0.90

0.95

1.00

1.05

1.10

1.15

1.20

1.25

1.30

0.0 0.2 0.4 0.6 0.8 1.0

NP

V I

ND

EX

Fraction to Power

0% CCS

100% CCS

a) Feedstock: light crude oil

10.0

10.1

10.2

10.3

10.4

10.5

10.6

10.7

10.8

10.9

11.0

0.0 0.2 0.4 0.6 0.8 1.0

NP

V IN

DEX

Fraction to Power

0% CCS

100% CCS

b) Feedstock: syncrude

Figure 6.14 : NPVI vs. fraction to power. Refinery- IGCC integration

As can be observed in figure 6.14, the NPVI is bigger for refinery using

syncrude that using light crude oil, this results were as expected because de

syncrude receive previous treatment in the upgrader and has less light and

heavy fractions in comparison with the light crude.

Figure 6.15 shows the NPVI for different feedstock flow rates with and without

CCS for refinery fed with light crude and syncrude. It can be observed that the

134

NPVI increases with the feedstock for both light crude oil and syncrude for both

schemes, with and without CCS. As expected, the NPVI is lower for 100% CCS.

As can be observed in Figure 6.15b, using syncrude as feed not significant

differences with and without CCS is obtained because the production of heavy

ends for this integration is low.

0.4

0.5

0.6

0.7

0.8

0.9

1.0

1.1

1.2

100.0 150.0 200.0 250.0 300.0

NP

V IN

DE

X

Feedstock (kbbl/d)

0% CCS

100% CCS


7.0

7.5

8.0

8.5

9.0

9.5

10.0

10.5

100.0 150.0 200.0 250.0 300.0

NP

V I

ND

EX

Feedtock (kbbl/d)

0% CCS

100% CCS


Figure 6.15: Effect of feedstock flow rate on NPVI. Refinery- IGCC

integration.

Figures 6.16, shows the NPVI as a function of CO2 tax with and without CCS

for refinery fed with light crude (figure a) and syncrude (Figure b). As explained

before the NPVI can be affected with the increase in carbon tax. When carbon

tax increases the NPVI without CCS can be less than those using CCS when

light crude oil is processed in refinery. The low flow rate to IGCC when refinery

is fed with syncrude produces not significant differences between NPVI values

135

for the process with and without CCS. For all values in CO2 tax the NPVI values

with 100% CCS are lower than NPVI with 0% CCS.

0.70

0.80

0.90

1.00

1.10

1.20

1.30

0 5 10 15 20 25 30 35 40

NP

V I

ND

EX

CO2 Tax ($/ton)

100% CCS

0% CCS


10.00

10.10

10.20

10.30

10.40

10.50

10.60

10.70

10.80

10.90

11.00

0.00 5.00 10.00 15.00 20.00 25.00 30.00 35.00 40.00

NP

V I

ND

EX

CO2 Tax ($/ton)

100% CCS

0% CCS


Figure 6.16: Effect of CO2 tax over the NPVI. Refinery-IGCC integration

136

c) Integration upgrader, refinery and IGCC

- Effect of fraction to power on NPVI for proposed processing

scheme 5-8:

Figure 6.17 presents the results of net present value index as a function of

fraction to power with and without carbon capture and storage (CCS) for

proposed processing schemes 5 to 8. For all cases the NPVI indicates that is

better to produce hydrogen than power. The NPVI is higher without CCS until

approximately when 30% of the syngas production is sent to power production,

after that the NPVI are similar due to the investment needed and CO2 tax. For

proposed processing schemes 7 and 8 the same behaviour is obtained but for

these cases the NPVI with CCS is highest than 0% CCS for all fraction to power

values.

2.40

2.50

2.60

2.70

2.80

2.90

3.00

3.10

0.0 0.2 0.4 0.6 0.8 1.0

NP

V IN

DEX

Fraction to Power

0% CCS

100% CCS

a) Proposed processing scheme 5

2.40

2.50

2.60

2.70

2.80

2.90

3.00

3.10

0.0 0.2 0.4 0.6 0.8 1.0

NP

V I

ND

EX

Fraction to Power

0% CCS

100% CCS

b) Proposed processing scheme 6

0.90

1.40

1.90

2.40

2.90

3.40

3.90

4.40

0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0

NP

V I

ND

EX

Fraction to Power

0% CCS

100% CCS

c) Proposed processing scheme 7

0.90

1.10

1.30

1.50

1.70

1.90

2.10

2.30

2.50

2.70

0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0

NP

V I

ND

EX

Fraction to Power

0% CCS

100% CCS

d) Proposed processing scheme 8

Figure 6.17 : NPVI vs. fraction to power

137

- Effect of feedstock flow rate on NPVI for proposed processing

scheme 5-8:

Figure 6.18 shows the NPVI for different feedstock flow rates for all proposed

processing schemes. It can be observed that the NPVI increases with the

feedstock for all cases. For proposed schemes 7 and 8 the NVPI had the same

behaviour founded for most of cases, that it higher NPVI for 100% of CCS while

for proposed processing scheme 5 and 6 no differences between The NPVI for

0% CCS and 100% NVPI was found.

1.40

1.60

1.80

2.00

2.20

2.40

2.60

100 150 200 250 300

NP

V I

ND

EX

Feed flow rate (kbbl/d)

0% CCS

100% CCS


1.4

1.6

1.8

2.0

2.2

2.4

2.6

100.0 150.0 200.0 250.0 300.0

NP

V I

ND

EX


0% CCS

100% CCS


0.4

0.9

1.4

1.9

2.4

2.9

3.4

100.0 200.0 300.0

NP

V I

ND

EX


0% CCS

100% CCS


0.4

0.6

0.8

1.0

1.2

1.4

1.6

1.8

100.0 200.0 300.0

NP

V IN

DEX


0% CCS

100% CCS


Figure 6.18: Effect of feedstock flow rate on NPVI.

- Effect of CO2 on NPVI for proposed processing scheme 5-8:

Figure 6.19 shows the NPVI as a function of CO2 tax. It can be observed that

the integration is profitable for both cases, but it is important to know that when

carbon tax increases the NPVI without CCS can be less than those using CCS.

138

As can be observed proposed processing scheme 7 is the most profitable for

this CO2 tax variation. It is important to mention that the curves for proposed

processing schemes have approximately the same NPVI values for 13 $/ton in

CO2 tax. All schemes were calculated taking as based case 13 $/ton for CO2

tax, for this reason not differences can be observed in NPVI for proposed

processing schemes 5 and 6 (Figure 6.18).

1.50

1.70

1.90

2.10

2.30

2.50

2.70

2.90

0 5 10 15 20 25 30 35 40

NP

V IN

DE

X

CO2 Tax ($/ton)

100 % CCS

0% CCS


1.50

1.70

1.90

2.10

2.30

2.50

2.70

2.90

0 5 10 15 20 25 30 35 40

NP

V IN

DE

X

CO2 Tax ($/ton)

100% CCS

0% CCS


2.50

2.60

2.70

2.80

2.90

3.00

3.10

3.20

3.30

3.40

3.50

0 5 10 15 20 25 30 35 40

NP

V I

ND

EX

CO2 Tax ($/ton)

100% CCS

0% CCS


1.00

1.10

1.20

1.30

1.40

1.50

1.60

1.70

1.80

1.90

2.00

0 5 10 15 20 25 30 35 40

NP

V IN

DE

X

CO2 Tax ($/ton)

100% CCS

0% CCS


Figure 6.19: Effect of CO2 tax over the NPVI.

The comparison between all proposed schemes is presented in Table 6.14. For

this comparison the fraction to power as explained before was fixed in 0.3.

Instead that CCS implies higher CAPEX the legislation in emission are stricter

nowadays, for this reason the decarbonisation grade was fixed in 100%.

139

Table 6.14: Economic evaluation for upgrader – refinery and IGCC

integration

CAPEX (MM$)

Gross Income

(MM$/year)

OPEX (MM$/year)

Cash flow

(MM$)

NPV (MM$)

NPVI

Upgrader+IGCC 6311.4 10217.2 1200 751 6614.9 1.05

Refinery+IGCC. (Feedstock : light crude oil)

4263.3 11991.6 787.9 528 4827 1.13

Refinery+IGCC. (Feedstock : Syncrude)

2257.6 13988.3 622.64 1498.6 23779 10.53

Proposed scheme 5

9174.2 14246.5 2308.12 1782.3 21626.4 2.36

Proposed scheme 6

8897.4 14012.8 1784.3 1963.5 25064.5 2.82

Proposed scheme 7

7140.4 13159.2 1179.06 1812.3 24234.7 3.39

Proposed scheme 8

8494.8 12502.3 1409.4 1498.3 17376.9 2.05

Table 6.14 shows the results for the main economic parameters (CAPEX, gross

income, OPEX, cash flow, NPV and NPVI) for all cases studied. The positive

values for the NPV and NPVI indicate the profitability of all the cases. It can be

observed that the higher NPV (25064.5) was for proposed processing scheme 6

(Integration upgrader, refinery and IGCC that includes the integration of some

intermediate products between upgrader and refinery) while the highest NPVI

(10.53) was for the scheme in which the refinery fed with syncrude is integrated

with IGCC; as can be observed in the same figure it has lowest CAPEX

because in this case the bottom of barrel is really low, thus the investment in

DCU is reduce.

Instead the scheme of refinery fed with syncrude and IGCC integration got the

highest NPVI, it has the limitation that the feedstock depends of other parties

(upgraders) to process the heavy crude oil for this reason the proposed

processing schemes in this work are those directly fed with heavy crude oil.

Thus proposed processing schemes 6 and 7 are the best option, the proposed

7 has lower investment as a consequence higher NPVI.

140

Table 6.15 summarises the results for NPV and NPVI for all the schemes

studied in this work for upgrader, refinery, upgrader-refinery integration and

upgrader-refinery-IGCC integration.

Table 6.15: Economic evaluation comparison for all schemes studied

Upgrader Refinery

(Feedstock : light crude oil)

Refinery (Feedstock : Syncrude)

Proposed scheme 1

Proposed scheme 2

NPV (MM$)

8320.2 4326.8 1590.2 23136.9 23535.4

. NPVI 2.20 1.33 7.19 3.26 3.36

Upgrader+

IGCC

Refinery+ IGCC(Feedstock :

light crude oil)

Refinery+IGCC.(Feedstock :

Syncrude)

Proposed scheme 5

Proposed scheme 6

NPV (MM$)

6614.9 4827.4 23779.7 21626.4 25064.5

. NPVI 1.05 1.13 10.53 2.36 2.82

Proposed scheme 3

Proposed scheme 4

Proposed scheme 7

Proposed scheme 8

NPV (MM$)

24267.6 15196.2 24234.7 17376.9

NPVI 4.87 3.26 3.39 2.05

The results have shown that all schemes studied are profitable since all NPV

are positives and all NPVI are greater than cero.

The best scheme without integration is the refinery treating as feedstock

syncrude as explained in previous chapter this crude has been treated

previously in the upgrader plant.

In all the integration proposed processing schemes between upgrader and

refinery, it can be observed that proposed processing scheme 2 which include

the upgrader and refinery integration presented the highest NPV, but proposed

processing scheme 3 (which is a new schemes for production of transportation

fuels that includes DCU as residue processor) got the highest NPVI.

Table 6.15 shows that for most of schemes of integration with IGCC the NVP

values are higher in comparison with the schemes without IGCC, but the NPVI

for all cases diminished due to the high investment that IGCC requires. Only for

the scheme for integration of refinery and IGCC where the refinery is fed with

141

syncrude the NPVI was the highest because the heavy fractions in this scheme

are low.

All proposed processing schemes were more economically attractive in

comparison with previous works (Aguilar et al, 2012); additionally previous work

did not consider the NPVI as economic indicator.

The integration schemes with IGCC could allow getting it off the bottom of barrel

and at the same time to meet with the environmental requirements for CO2

emission; for these reasons the proposed schemes 6 and 7 are chosen as the

best option to process the heavy crude oil.

6.4 Chapter summary

In this chapter different proposed processing schemes for the integration of an

upgrader to a refinery and IGCC have been compared technical and

economically.

This chapter included three parts: the first part addressed the simulation of the

IGCC process using Aspen plus 2006.5 (Aspen plus 2008). The second part

was the evaluation for the different integration schemes: upgrader and IGCC;

refinery and IGCC; and upgrader, refinery and IGCC and in the third part

sensitivity analyses for all cases were carried out. Sensitivity analysis that show

the changes in the economic indicator with the fraction to power, the

decarbonisation grade and the carbon tax for all integration schemes were

presented. Then a comparison between all schemes that includes fixed values

for fraction to power, the decarbonisation grade and the carbon was evaluated.

Finally a comparison table between al schemes studied in this work was

discussed.

All the schemes studied are economically attractive. In spite of the scheme of

refinery fed with syncrude and its integration with IGCC got the highest NPVI, it

has the limitation that the feedstock depends of other parties (upgraders) to

process the heavy crude oil. The proposed processing schemes 5 to 8 in this

chapter are the best options because the feedstock is directly heavy crude oil.

Proposed processing schemes 6 and 7 are chosen as the best option to

142

process heavy crude oil, the proposed 7 has lower investment as a

consequence higher NPVI.

The results for this chapter have demonstrated as previous work (Sadhukhan,

2002) that the integration of IGCC to refineries is economically attractive.



emission, for these reasons the proposed schemes 6 and 7 are chosen as the

best option to process the heavy crude oil.

143

Chapter 7 Conclusions and future work

7.1 Conclusions

In this work various approaches for processing heavy crude oils have been

presented: an upgrader to produce syncrude, the integration of an upgrader to a

refinery, and the integration of an upgrader and a refinery with an IGCC unit for

processing the heavy ends.

Naphtha recycled from the upgrader distillation unit is the best option as diluent

for the heavy crude oils. When naphtha is used as diluents, the diluted crude oil

density can be kept in the required value for the feed to the upgrader distillation

columns with a crude/diluent ratio of approximately 3.

The properties of gasoline and naphtha products were met when syncrude was

processed in the refinery. Additionally, the integration of the upgrader to a

refinery allows both the treating heavy streams of the refinery and their

transformation to products of higher qualities. The integration of the IGCC unit

to the upgrader and the refinery permits a complete elimination of the heavy

residues produced in these units, thus producing power and hydrogen to meet

the upgrader and refinery requirements.

All the proposed processing schemes have a positive NPV and NPVI, which

indicates that all the processing schemes studied are profitable. The NPV

analysis showed that the integration of the upgrader – refinery and IGCC unit in

all cases studied is attractive economically, since light crude oil worldwide

reserves are decreasing.

In spite of the scheme of refinery fed with syncrude and its integration with

IGCC got the highest NPVI, it has the limitation that the feedstock depends of

other parties (upgraders) to process the heavy crude oil.

144



emission.

The proposed schemes 6 and 7 that are fed directly with heavy crude oil are

chosen as the best option to process the heavy crude oil. Proposed 7 has lower

investment as a consequence higher NPVI.

The simulation models for the dilution process, distillation and IGCC units

developed in Aspen Plus 2006.5, along with the petroleum refining processes

correlations developed by Gary and HPI Consultants, are adequate for the

evaluation of the integration of upgrader, refinery and IGCC.

7.2 Future work

Future research work on heavy crude oil processing can address the following

issues:

1. Detailed study of different upgrader configurations and a comparison with

upgraders in operation.

This work has a limitation that a typical upgrader configuration was selected

based on the upgraders installed in the countries with the biggest reserve of

heavy crude oil in the world, but future work should explores between

different upgrader configuration and compare economically all the schemes.

2. Detailed study of different configurations for processing heavy crude oil

to produce directly transportation fuels, instead of syncrude.

This work has presented eight different proposed processing schemes to

integrate upgrader and refinery to produce directly transportation fuels but

still a future work in which more proposed schemes and the optimisation of

these schemes are needed.

3. Integration of upgrader with other industries, for example petrochemistry

industry for producing methanol, ammonia, plastic, and so on.

145

The integration of upgrader or refinery fed with heavy crude oil with other

industries could represent a challenge for the petroleum industry in the

future.

4. Detailed simulation of different refinery units, compared with calculations

of petroleum refining correlations.

5. Simulation of all the models for the integration of upgrader-refinery-IGCC

using one software application, for example, ASPEN PLUS, MATLAB,

MATHCAD, among others.

In the present work has been studied the simulation of upgrader/refinery

unit using a combination between commercial simulation programs and

petroleum refining correlations, because the commercial simulation

programs at the moment when this thesis was done did not include the

simulation of refinery units.

Nowadays some commercial programs include this unit, for this reason as

future work could be simulated all cases proposed in this work using the

same program and take advantage of the other tools that these programs

have.

6. Detailed energy integration study of different refinery units.

Detailed simulation models can be developed for each unit individually and

then be incorporated into the different simulations programs.

7. To continue the study of the proposed processing schemes chosen in

this work to optimise them.

146

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151

Appendix A: Properties of Crude Oils

Table A1. Properties of Venezuelan Boscán Crude Oil

Boscán crude 10.1 ˚API

True boiling point curve (liquid volume basis)

Percent distilled Temperature (F)

0.55 200

1.87 300

2.75 350

4.03 400

7.99 500

10.65 550

17.11 650

Gravity and aniline curves Mid percent

distilled API gravity Aniline point

3.39 112.69

4.26 113.07

6.01 34.00 113.83

9.32 28.62

13.88 25.54 117.30

25.00 130.00

35.00 140.00

45.00 150.00

55.00 160.00

58.55 6.10

65.00 170.00

75.00 180.00

85.00 190.00 95.00 200.00

152

Table A2. Properties of crude used as diluents

True Boiling Point curve (liquid volume basis)

Lighter crude 34.1 ˚API Naphtha 47.0 ˚API

Percent distilled

Temperature (F) Percent distilled

Temperature (F)

6.8 130 5 68.14

10 180 10 94.88

30 418 30 114.01

50 650 50 135.25

62 800 70 164.17

70 903 90 213.62

76 1000 95 311.65

Lighter crude 34.1 ˚API

Mid percent distilled

API gravity Light

component Volume Fraction

5 90 C1 0.001

10 68 C2 0.0015

15 59.7 C3 0.009

20 52 IC4 0.004

30 42 NC4 0.016

40 35 IC5 0.012

45 32

50 28.5

60 23

70 18

80 13.5

153

Table A3. Modeling details, design considerations

Distillation Unit Specifications Atmospheric

tower Vacuum

tower Model in Aspen

Plus Petrofrac Petrofrac

Number of trays for for main

column

20-26 6-10

Number of side products

3 2

Number of trays for for side column

4-10

Number of pumparound

2 2

Preheater temperature˚C

287

Furnace temperature˚C

349-398 388-454

Steam (temperature ˚C

and pressure bar)

204.4 / 4.14

204.4/ 4.14

Steam injected lb/bbl fed

10-50 10-50

154

Appendix B: Economic Evaluation for Upgrader and Refinery Integration

Cost Analysis Results

Table B.1. Upgrader: Summary of Unit Costs

Summary of Unit Costs Feed to Unit Cost (MM$)

Process Units kbbl/d ºAPI kt/d at 2013

Atmospheric Destillation (ADU) 398.46 17.31 60.24 193.33

Vacuum Destillation (VDU) 224.49 5.02 36.99 133.84

Hydrotreater (HTU) 132.22 38.84 17.46 159.83

Hydrocracker (HCU) 150.24 14.72 23.12 1 043.47

Delayed Coking (DCU) 156.36 1.11 26.53 624.57

SUBTOTAL 2 155.05

Utilities Facilities



Storage 996.15

Offsites 501.11

SUBTOTAL 3 841.88

Contingence 576.28

TOTAL 4 418.16

Table B.2. Upgrader: Utility Requirements for Units

Utilities CDU/VDU DCU HCU HTU HPU TOTAL

Fuel (Gcal/d) 6.716.90 5.519.43 10 102.65 3 333.87 24 827.28 50 500.12

Boiler Feed Water (m3/d) 0.00 0.00 0.00 0.00 0.00 0.00

LP Steam (ton/d) 0.00 1 545.16 6 239.01 359.38 0.00 8 143.55

HP Steam (ton/d) 2.803.25 0.00 0.00 0.00 2.319.44 5 122.69

Cooling Water (1000 m3/d) 359.06 41.96 316.77 152.05 246.98 1 116.82

Power (MWh/d) 425.96 146.18 2 454.06 264.44 327.10 3 617.75

Hydrogen (ton/d) 0.00 0.00 965.45 25.76 0.00 991.21

Chem. and Catalysts ($/d) 0.00 0.00 48 829.71 2.644.42 10 903.36 62 377.49

Pooled Components CDU/VDU DCU HCU HTU HPU TOTAL

Hydrogen (ton/d) 0.00 0.00 0.00 0.00 991.21 991.21

Gas LHV (ton/d) 0.00 1 232.95 358.26 3.40 0.00 1 594.61

155

Table B.3.Upgrader: Summary of Operation Costs

Summary of Operation Costs

Cost

Utilities TOTAL Cost ($/units) (MM$/year)

Fuel (Gcal/d) 34 811.23 70.44 809.14

LP Steam (ton/d) 8 143.55 59.84 160.81

HP Steam (ton/d) 5 122.69 69.41 117.34

Cooling Water (1000 m3/d) 1 116.82 111.91 41.25

Power (MWh/d) 3 617.75 50.00 59.69

Catalysts & Royalties ($/d) 62 377.49 - 20.58

Insurance 22.09

Maintenance 220.91

Plant Staff & Operators Salary 68.65

TOTAL 1 520.45

Table B.4. Upgrader: Products and Raw Materials

Upgrader Products and Raw Materials Cost

Unit/d Cost ($/units) (MM$/year)

Raw Materials

Crude Oil 9.8 ºAPI (bbl) 306 464 78.35 7 923.79

H2 Prod. Gas Feed (MSCF) 97 538 2.25 72.42

TOTAL 7 996.21

Products

Syncrude 35.5 ºAPI (bbl) 300 000 110.00 10 890.00

Coke (ton) 4 873 70.00 112.56

TOTAL 11 002.56

Table B.5. Upgrader: Costs and Revenues

Cost and Revenues Cost

(MM$/year)

Gross Income 11 002.6

TOTAL 11 002.6

Production Costs

Raw Materials 7 996.2

Operation Costs 1 520.5

Depreciation 220.9

TOTAL 9 737.6

Income before tax 1 265.0

Less tax (60%) 759.0

CO2 Emission Tax 24.3

Net Income 481.7

CASHFLOW 702.6

156

Table B.6. Upgrader: Assumptions Used and NPV Analysis

Assumptions and NPV Analysis

Parameter Value

Plant Life 25 years


CAPEX at -2,-1 and 0 20%, 45%, 35%

Operating Percentage 90.41%


Interest Rate 3%

Total Invesments (MM$)

Construction Cost 1 476.4

Land Cost 7.4

Working Capital 147.6

TOTAL 1 631.4

Calculated NPV 3 794.3

NVP INDEX 2.3

Table B.7. Summary of upgrader for different syncrude production

Summary of Upgrader CASE A CASE B CASE C

Syncrude production (kbbl/d) 100.00 200.00 300.00

Heavy crude oil fed (kbbl/d) 102.15 204.31 306.46

CAPEX (MM$) 1 574.88 2 558.82 3 415.55

Gross Income (MM$/year) 3 518.36 7 036.72 10 555.08

OPEX (MM$/year) 3 177.84 6 323.18 9 461.52

Taxes (MM$/year) 112.82 250.45 393.44

Cash flow (MM$) 227.70 463.08 700.12

Net Present Value (MM$) 2 180.06 5 163.66 8 320.23


157

Table B.8. Utilities for Refinery: A) Light crude oil.

Utilities CDU/VDU ISO DCU VBU NHT REF KHT DHT VGOHDS FCC HPU TOTAL

Fuel (Gcal/d) 4.708,79 1.449,60 1.611,28 2.963,89 1.190,15 3.596,26 2.631,68 2.971,26 2.237,53 1.007,88 446,13 24.814,45

Boiler Feed Water (m3/d) 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00

LP Steam (ton/d) 0,00 0,00 430,30 320,16 128,29 646,10 252,16 284,70 201,00 0,00 0,00 2.262,72

HP Steam (ton/d) 1.903,18 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 529,08 41,68 2.473,94

Cooling Water (1000 m3/d) 244,07 93,07 12,25 13,88 54,28 72,90 106,69 120,46 85,04 73,58 4,44 880,66

Power (MWh/d) 306,88 33,38 40,71 6,38 94,40 142,63 208,74 235,68 266,22 233,27 5,88 1.574,18

Hydrogen (ton/d) 0,00 3,21 0,00 0,00 15,95 0,00 44,35 16,79 70,87 0,00 0,00 151,17

Chemicals ($/d) 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00

Catalysts & Royalties ($/d) 0,00 1.517,37 0,00 0,00 944,03 950,85 2.087,45 2.356,81 3.105,92 15.988,95 195,93 27.147,29

Pooled Components CDU/VDU ISO DCU VBU NHT REF KHT DHT VGOHDS FCC HPU TOTAL

Hydrogen (ton/d) 0,00 0,00 0,00 0,00 0,00 133,36 0,00 0,00 0,00 0,00 17,81 151,17

Power (MWh/d) 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00

Gas LHV (ton/d) 0,00 71,71 343,38 63,17 18,29 142,20 65,08 149,38 24,12 148,18 0,00 1.025,51

158

Table B.9. Refinery: A) Light crude oil. Summary of unit costs



Atmospheric Destillation (ADU) 300,00 31,52 41,40 167,28

Vacuum Destillation (VDU) 122,93 13,18 19,11 99,46

Isomerization (ISO) 30,35 58,92 3,59 23,79

Reforming (REF) 47,54 66,70 5,40 80,63

Naphta Hydrotreater (NHT) 36,40 64,91 5,41 36,35

Kerosene Hydrotreater (KHT) 69,58 36,88 9,30 110,43

Gas Oil Hydrotreater (DHT) 29,94 22,31 11,49 67,94

Gas Oil Hydrodesulph (VGHDS) 44,37 21,68 6,52 226,45

Fluid Catalitic Cracking (FCC) 39,97 25,74 5,72 240,68

Visbreaking Unit (VBU) 13,69 7,50 2,22 44,86

Delayed Coking (DCU) 45,65 7,50 7,39 318,19

SUBTOTAL 1.416,08


Hydrogen Production Unit 11,69

Cooling Water System 45,00

Storage 750,00

Offsites 333,41

SUBTOTAL 2.556,18

Contingence 383,43

TOTAL 2.939,60

Table B.10. Refinery: A) Light crude oil: summary of utilities costs

Summary of Operation Costs Cost


Fuel (Gcal/d) 14.724,75 70,44 342,26

Boiler Feed Water (m3/d) 0,00 5,00 0,00

LP Steam (ton/d) 2.262,72 59,84 44,68

HP Steam (ton/d) 2.473,94 69,41 56,67

Cooling Water (1000 m3/d) 880,66 111,91 32,52

Hydrogen (ton/d) 0,00 2.150,00 0,00

Power (MWh/d) 1.574,18 50,00 25,97

Chemicals ($/d) 0,00 - 0,00

Catalysts & Royalties ($/d) 27.147,29 - 8,96

Insurance 14,70

Maintenance 146,98

Plant Staff & Operators Salary 67,20

TOTAL 739,94

159

Table B.11. Refinery: A) Light crude oil. Products and Raw Materials

Refinery Products and Raw Materials Cost

Unit/d Cost

($/units)

(MM$/year)

Raw Materials

Lagomedio Crude Oil 31.5 ºAPI 300.000 102,75 10.172,25

H2 Prod. Gas Feed (MSCF) 1.753 2,25 1,30

TOTAL 10.173,55

Products

Gasoline 95 99.199 128,43 4.204,22

Diesel 87.675 133,77 3.870,34

Kerosene 69.598 161,66 3.712,88

Coke ton/d ($/ton) 269 70,00 6,21

TOTAL 11.793,64

Table B.12. Refinery: A) Light crude oil. Cost and Revenues


(MM$/year)

Gross Income 11.793,6

TOTAL 11.793,6

Production Costs

Raw Materials 10.173,6

Operation Costs 739,9

Depreciation 147,0

TOTAL 11.060,5

Income before tax 733,2

Less tax (60%) 439,9

CO2 Emission Tax 0,4

Net Income 292,8

CASHFLOW 439,8

160

Table B.13. Refinery: A) Light crude oil. Assumptions and NPV Analysis


Parameter Value

Plant Life 25 years


CAPEX at -2,-1 and 0 20%, 45%, 35%

Operating Percentage (%) 90,41

CO2 Emission Tax ($/ton) 13

Interest Rate 3%


Construction Cost 2.939,6

Land Cost 14,7

Working Capital 294,0

TOTAL 3.248,3

Calculated NPV 4.326,8

NPV INDEX 1,33

161

Table B.14. Summary of Refinery: A) Light crude oil

Summary of Unit Costs CASE A CASE B CASE C

Lagomedio crude oil fed (kbbl/d) 100,00 200,00 300,00

Process Units Cost at 2013 (MM$)

Atmospheric Destillation (ADU) 95,52 136,03 167,28

Vacuum Destillation (VDU) 57,87 81,44 99,46

Isomerization (ISO) 12,79 18,92 23,79

Reforming (REF) 43,01 63,94 80,63

Naphta Hydrotreater (NHT) 20,44 29,40 36,35

Kerosene Hydrotreater (KHT) 58,65 87,43 110,43

Gas Oil Hydrotreater (DHT) 36,08 53,79 67,94

Gas Oil Hydrodesulph (VGHDS) 90,58 161,48 226,45

Fluid Catalitic Cracking (FCC) 137,44 195,72 240,68

Visbreaking Unit (VBU) 19,88 33,22 44,86

Delayed Coking (DCU) 156,83 245,07 318,19

SUBTOTAL 729,10 1.159,09 1.416,08


Hydrogen Production Unit 6,05 9,17 11,69

Cooling Water System 15,00 30,00 45,00

Storage 250,00 500,00 750,00

Offsites 150,02 246,84 333,41

SUBTOTAL 1.150,17 1.892,44 2.556,18

Contingence 172,53 283,87 383,43

TOTAL 1.322,70 2.176,31 2.939,60



Utilities Cost at 2013 (MM$)

Fuel (Gcal/d) 114,09 228,17 342,26


LP Steam (ton/d) 14,89 29,79 44,68

HP Steam (ton/d) 18,89 37,78 56,67

Cooling Water (1000 m3/d) 10,84 21,68 32,52

Hydrogen (ton/d) 0,00 0,00 0,00

Power (MWh/d) 8,66 17,32 25,97

Chemicals ($/d) 0,00 0,00 0,00

Catalysts & Royalties ($/d) 2,99 5,97 8,96

Insurance 6,61 10,88 14,70

Maintenance 66,13 108,82 146,98

Plant Staff & Operators Salary 22,40 44,80 67,20

TOTAL 265,50 505,20 739,94

162

Table B.14. Summary of Refinery: A) Light crude oil. (Continuation)



CAPEX (MM$) 1.322,70 2.176,31 2.939,60

Gross Income (MM$/year) 3.931,21 7.862,43 11.793,64

OPEX (MM$/year) 265,50 505,20 739,94

Taxes (MM$/year) 125,18 279,92 440,34

Cash flow (MM$) 149,35 294,94 439,81

Net Present Value (MM$) 1.101,48 2.669,26 4.326,82

Net Present Value index 0,75 1,11 1,33

163

Table B.15. Utilities for Refinery: B) Syncrude.

Utilities CDU/VDU ISO DCU VBU NHT REF KHT DHT VGOHDS FCC HPU TOTAL

Fuel (Gcal/d) 3.865,80 1.902,53 36,00 68,05 622,16 1.884,46 4.009,27 4.821,92 76,48 34,44 2.309,81 19.630,91 Boiler Feed Water (m3/d) 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00

LP Steam (ton/d) 0,00 0,00 9,88 7,35 67,07 338,56 384,16 462,03 6,87 0,00 0,00 1.275,92

HP Steam (ton/d) 1.401,34 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 18,43 215,79 1.635,57

Cooling Water (1000 m3/d) 180,50 122,14 0,27 0,32 28,38 38,20 162,54 195,49 2,91 2,56 22,98 756,29

Power (MWh/d) 273,42 43,81 0,93 0,15 49,35 74,74 318,02 382,47 9,10 8,13 30,43 1.190,55

Hydrogen (ton/d) 0,00 4,21 0,00 0,00 8,34 0,00 71,31 79,33 2,44 0,00 0,00 165,64

Chemicals ($/d) 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00

Catalysts & Royalties ($/d) 0,00 1.991,48 0,00 0,00 493,50 498,25 3.180,15 3.824,74 106,16 546,31 1.014,40 11.654,98

Pooled Components CDU/VDU ISO DCU VBU NHT REF KHT DHT VGOHDS FCC HPU TOTAL

Hydrogen (ton/d) 0,00 0,00 0,00 0,00 0,00 73,42 0,00 0,00 0,00 0,00 92,22 165,64

Power (MWh/d) 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00 0,00

Gas LHV (ton/d) 0,00 99,27 7,88 1,45 7,54 78,29 99,73 231,11 0,83 5,34 0,00 531,43

Table B.16. Refinery: B) Syncrude. Summary of unit costs






Reforming (REF) 24,91 57,15 2,97 55,71








SUBTOTAL 688,47




Storage 750,00

Offsites 227,22

SUBTOTAL 1.741,99

Contingence 261,30

TOTAL 2.003,28 Table B.17. Refinery: B) Syncrude: summary of utilities costs


Utilities TOTAL Cost

($/units) (MM$/year)

Fuel (Gcal/d) 14.402,31 70,44 334,76


LP Steam (ton/d) 1.275,92 59,84 25,19

HP Steam (ton/d) 1.635,57 69,41 37,46

Cooling Water (1000 m3/d) 756,29 111,91 27,93

Hydrogen (ton/d) 0,00 2.150,00 0,00

Power (MWh/d) 1.190,55 50,00 19,64

Chemicals ($/d) 0,00 - 0,00


Insurance 10,02

Maintenance 100,16


TOTAL 626,22

165

Table B.18. Refinery: B) Syncrude. Products and Raw Materials


Unit/d Cost


Raw Materials

syncrude 300.000 110,00 10.890,00


TOTAL 10.896,74

Products

Gasoline 95 64.559 128,43 2.736,15

Diesel 126.774 133,77 5.596,32

Kerosene 106.019 161,66 5.655,86

Coke ton/d ($/ton) 10 70,00 0,24

TOTAL 13.988,57

Table B.19. Refinery: B) Syncrude. Cost and Revenues


(MM$/year)


TOTAL 13.988,6

Production Costs



Depreciation 100,2

TOTAL 11.623,1

Income before tax 2.365,4

Less tax (60%) 1.419,3


Net Income 943,9

CASHFLOW 1.044,1

166

Table B.20. Refinery: B) Syncrude. Assumptions and NPV Analysis


Parameter Value

Plant Life 25 years


CAPEX at -2,-1 and 0 20%, 45%, 35%



Interest Rate 3%



Land Cost 10,0


TOTAL 2.213,6


NPV INDEX 7,19

167

Table B.21. Summary of Refinery: B) Syncrude


Syncrude oil fed (kbbl/d) 100,00 200,00 300,00

Process Units Cost at 2013 (MM$)

Atmospheric Destillation (ADU) 95,52 136,03 167,28

Vacuum Destillation (VDU) 17,93 25,23 30,81

Isomerization (ISO) 14,91 22,06 27,74

Reforming (REF) 29,72 44,18 55,71

Naphta Hydrotreater (NHT) 16,59 23,85 29,50

Kerosene Hydrotreater (KHT) 74,74 111,42 140,73

Gas Oil Hydrotreater (DHT) 79,63 118,70 149,93

Gas Oil Hydrodesulph (VGHDS) 5,42 9,67 13,56

Fluid Catalitic Cracking (FCC) 24,56 34,98 43,01

Visbreaking Unit (VBU) 1,19 1,99 2,68

Delayed Coking (DCU) 13,56 21,19 27,51

SUBTOTAL 373,78 549,30 688,47


Hydrogen Production Unit 16,21 24,55 31,30

Cooling Water System 15,00 30,00 45,00

Storage 250,00 500,00 750,00

Offsites 98,25 165,58 227,22

SUBTOTAL 753,23 1.269,43 1.741,99

Contingence 112,98 190,41 261,30

TOTAL 866,22 1.459,84 2.003,28



Utilities Cost at 2013 (MM$)

Fuel (Gcal/d) 111,59 223,17 334,76


LP Steam (ton/d) 8,40 16,80 25,19

HP Steam (ton/d) 12,49 24,98 37,46

Cooling Water (1000 m3/d) 9,31 18,62 27,93

Hydrogen (ton/d) 0,00 0,00 0,00

Power (MWh/d) 6,55 13,10 19,64

Chemicals ($/d) 0,00 0,00 0,00

Catalysts & Royalties ($/d) 1,28 2,56 3,85

Insurance 4,33 7,30 10,02

Maintenance 43,31 72,99 100,16

Plant Staff & Operators Salary 22,40 44,80 67,20

TOTAL 219,66 424,32 626,22

168

Table B.21. Summary of Refinery: B) Syncrude. (Continuation)



CAPEX (MM$) 866,22 1.459,84 2.003,28

Gross Income (MM$/year) 4.662,86 9.325,71 13.988,57

OPEX (MM$/year) 219,66 424,32 626,22

Taxes (MM$/year) 461,34 939,86 1.421,53

Cash flow (MM$) 349,61 697,05 1.044,08

Net Present Value (MM$) 5.106,14 10.483,25 15.910,24

Net Present Value index 5,33 6,50 7,19

169

Table B.22. Upgrader-refinery integration Proposed processing scheme 1: summary of unit costs


Process Units kbbl/d ºAPI kt/d at 2013 Atmospheric Destillation (UP-ADU) 398.46 17.31 60.24 193.33

Vacuum Destillation (UP-VDU) 224.49 5.02 36.99 133.84

Hydrotreater (UP-HTU) 132.22 38.84 17.46 159.83

Hydrocracker (UP-HCU) 150.24 14.72 23.12 1,043.47

Delayed Coking (UP-DCU) 156.36 1.11 26.53 624.57 Atmospheric Destillation (R-ADU) 300.00 35.56 40.40 167.28

Vacuum Destillation (R-VDU) 11.41 23.18 1.66 30.81

Isomerization (R-ISO) 39.83 49.03 4.96 27.74

Reforming (R-REF) 24.91 57.15 2.97 55.71

Naphta Hydrotreater (R-NHT) 24.43 55.39 2.97 29.50

Kerosene Hydrotreater (R-KHT) 106.01 35.90 14.25 140.73

Gas Oil Hydrotreater (R-DHT) 118.33 29.84 17.78 149.93 Gas Oil Hydrodesulph (R-VGHDS) 1.52 18.75 0.23 13.56

Fluid Catalitic Cracking (R-FCC) 1.37 22.69 0.20 43.01

Visbreaking Unit (R-VBU) 0.31 3.78 0.05 2.68

Delayed Coking (R-DCU) 1.02 0.00 0.17 27.51

SUBTOTAL 2,843.52

Utilities Facilities - Upgrader



Storage - Upgrader 996.15

Offsites - Upgrader 501.11

Utilities Facilities - Refinery



Storage - Refinery 750.00

Offsites - Refinery 227.22

SUBTOTAL 5,583.87

Contingence - Upgrader 576.28

Contingence - Refinery 261.30

TOTAL 6,421.45

170

Table B.23. Upgrader-refinery integration Proposed processing scheme

1: summary of utility requirements

Utilities UPGRADER REFINERY TOTAL

Fuel (Gcal/d) 50,500.12 19,630.91 70,131.03


LP Steam (ton/d) 8,143.55 1,275.92 9,419.47

HP Steam (ton/d) 5,122.69 1,635.57 6,758.26

Cooling Water (1000 m3/d) 1,116.82 756.29 1,873.11

Power (MWh/d) 3,617.75 1,190.55 4,808.30

Hydrogen (ton/d) 991.21 165.64 1,156.85

Chemicals ($/d) 0.00 0.00 0.00

Catalysts & Royalties ($/d) 62,377.49 11,654.98 74,032.47

Pooled Components UPGRADER REFINERY TOTAL

Hydrogen (ton/d) 991.21 165.64 1,156.85

Power (MWh/d) 0.00 0.00 0.00

Gas LHV (ton/d) 1,594.61 531.43 2,126.04


1: summary of operation costs



($/units)

(MM$/year)

Fuel (Gcal/d) 49,213.54 70.44 1,143.90


LP Steam (ton/d) 9,419.47 59.84 186.00

HP Steam (ton/d) 6,758.26 69.41 154.80

Cooling Water (1000 m3/d) 1,873.11 111.91 69.18

Hydrogen (ton/d) 0.00 2,150.00 0.00

Power (MWh/d) 4,808.30 50.00 79.34

Chemicals ($/d) 0.00 - 0.00

Catalysts & Royalties ($/d) 74,032.47 - 24.43

Insurance 32.11

Maintenance 321.07


TOTAL 2,146.67

171

Table B.25. Upgrader-refinery integration Proposed processing scheme 1: products and raw materials


Unit/d Cost

($/units)

(MM$/year)

Raw Materials

Crude Oil 9.8 ºAPI (bbl) 306,464 78.35 7,923.79

H2 Prod. Gas Feed (MSCF) 106,612 2.25 79.16

TOTAL 8,002.95

Refinery Products

Gasoline 95 64,559 128.43 2,736.15

Diesel 126,774 133.77 5,596.32

Kerosene 106,019 161.66 5,655.86

Coke ton/d ($/ton) 4,883 70.00 112.80

TOTAL 14,101.13 Table B.26. Upgrader-refinery integration Proposed processing scheme 1: costs and revenues


(MM$/year)

Gross Income 14,101.1

TOTAL 14,101.1

Production Costs

Raw Materials 8,002.9

Operation Costs 2,146.7

Depreciation 321.1

TOTAL 10,470.7

Income before tax 3,630.4

Less tax (60%) 2,178.3


Net Income 1,425.6

CASHFLOW 1,746.7

172


1: assumptions and NPV analysis


Parameter Value

Plant Life 25 years


CAPEX at -2,-1 and 0 20%, 45%, 35%

Operating Percentage (%) 90.41


Interest Rate 3%


Construction Cost 6,421.4

Land Cost 32.1


TOTAL 7,095.7

Calculated NPV 23,136.9

NPV INDEX 3.26

173

Table B.28. Upgrader-refinery integration Proposed processing scheme 2: summary of unit costs


Process Units kbbl/d ºAPI kt/d at 2013 Atmospheric Destillation (UP-ADU) 396.42 17.31 59.93 192.83 Vacuum Destillation (UP-VDU) 223.33 5.02 36.80 133.50



Delayed Coking (UP-DCU) 155.56 1.11 26.39 622.48 Atmospheric Destillation (R-ADU) 300.00 40.02 39.35 167.28 Vacuum Destillation (R-VDU) 11.41 23.18 1.66 30.81


Reforming (R-REF) 24.67 57.16 2.94 55.40 Naphta Hydrotreater (R-NHT) 24.43 55.40 2.94 29.50 Kerosene Hydrotreater (R-KHT) 106.01 35.90 14.25 140.73 Gas Oil Hydrotreater (R-DHT) 118.33 29.84 17.78 149.93 Gas Oil Hydrodesulph (R-VGHDS) 0.85 14.53 0.13 8.34 Fluid Catalitic Cracking (R-FCC) 0.76 18.64 0.11 31.98


SUBTOTAL 2,792.50

Utilities Facilities - Upgrader Hydrogen Production Unit 129.69




Utilities Facilities - Refinery Hydrogen Production Unit 31.21




SUBTOTAL 5,518.74



TOTAL 6,346.55

174


2: summary of utility requirements

Utilities UPGRADER REFINERY TOTAL

Fuel (Gcal/d) 50,388.44 19,442.73 69,831.17


LP Steam (ton/d) 8,140.00 1,251.62 9,391.63

HP Steam (ton/d) 5,104.93 1,626.77 6,731.70

Cooling Water (1000 m3/d) 1,114.02 752.56 1,866.58

Power (MWh/d) 3,611.96 1,180.62 4,792.58

Hydrogen (ton/d) 989.77 164.50 1,154.26

Chemicals ($/d) 0.00 0.00 0.00

Catalysts & Royalties ($/d) 62,281.64 11,353.02 73,634.66

Pooled Components UPGRADER REFINERY TOTAL

Hydrogen (ton/d) 989.77 164.50 1,154.26

Power (MWh/d) 0.00 0.00 0.00

Gas LHV (ton/d) 1,587.62 518.83 2,106.45





($/units)

(MM$/year)

Fuel (Gcal/d) 49,106.48 70.44 1,141.41


LP Steam (ton/d) 9,391.63 59.84 185.45

HP Steam (ton/d) 6,731.70 69.41 154.20

Cooling Water (1000 m3/d) 1,866.58 111.91 68.93

Hydrogen (ton/d) 0.00 2,150.00 0.00

Power (MWh/d) 4,792.58 50.00 79.08

Chemicals ($/d) 0.00 - 0.00


Insurance 31.73

Maintenance 317.33


TOTAL 2,137.93

175



Unit/d Cost

($/units)

(MM$/year)

Raw Materials



TOTAL 7,962.18

Refinery Products

Gasoline 95 63,974 128.43 2,711.36

Diesel 127,353 133.77 5,621.90

Kerosene 106,019 161.66 5,655.86

Coke ton/d ($/ton) 4,893 70.00 113.03



(MM$/year)


TOTAL 14,102.1

Production Costs



Depreciation 317.3

TOTAL 10,417.4


Less tax (60%) 2,210.8


Net Income 1,447.3

CASHFLOW 1,764.7

176




Parameter Value

Plant Life 25 years


CAPEX at -2,-1 and 0 20%, 45%, 35%



Interest Rate 3%



Land Cost 31.7


TOTAL 7,012.9


NPV INDEX 3.36 Table B.34. Upgrader-refinery integration Proposed processing scheme

3: summary of unit costs






Hydrocracker (HCU) 150.24 14.72 23.12 1,043.47

Reformer (REF) 47.24 58.45 5.60 80.34


SUBTOTAL 2,235.39




Storage 996.15

Offsites 511.49

SUBTOTAL 3,921.42

Contingence 588.21

TOTAL 4,509.64

177

Table B.35. Upgrader-refinery integration Proposed processing scheme 3: summary of utility requirements

Utilities TOTAL

Fuel (Gcal/d) 50,610.13

Boiler Feed Water (m3/d) 0.00

LP Steam (ton/d) 8,785.60

HP Steam (ton/d) 4,799.10

Cooling Water (1000 m3/d) 1,154.81

Power (MWh/d) 3,713.84

Hydrogen (ton/d) 991.21

Chem. and Catalysts ($/d) 60,856.36

Pooled Components TOTAL


Gas LHV (ton/d) 1,742.06





($/units)

(MM$/year)

Fuel (Gcal/d) 33,470.50 70.44 777.97

LP Steam (ton/d) 8,785.60 59.84 173.48

HP Steam (ton/d) 4,799.10 69.41 109.93

Cooling Water (1000 m3/d) 1,154.81 111.91 42.65

Power (MWh/d) 3,713.84 50.00 61.28


Insurance 22.55

Maintenance 225.48


TOTAL 1,502.07

178



Unit/d Cost

($/units)

(MM$/year)

Raw Materials



TOTAL 7,986.11

Products

Gasoline (bbl) 41,916 128.43 1,776.47

Diesel (bbl) 261,216 133.77 11,531.15

Coke (ton) 4,873 70.00 112.56



(MM$/year)


TOTAL 13,420.2

Production Costs



Depreciation 225.5

TOTAL 9,713.7


Less tax (60%) 2,223.9


Net Income 1,461.7

CASHFLOW 1,687.2

179




Parameter Value

Plant Life 25 years


CAPEX at -2,-1 and 0 20%, 45%, 35%



Interest Rate 3%



Land Cost 22.5


TOTAL 4,983.2


NPV INDEX 4.87 Table B.40. Upgrader-refinery integration Proposed processing scheme

4: summary of unit costs







Reformer (REF) 56.81 58.27 6.74 89.28 Solvent Deasphalting Unit (SDA) 156.36 1.11 26.53 208.73

SUBTOTAL 2,020.29




Storage 996.15

Offsites 478.09

SUBTOTAL 3,665.33

Contingence 549.80

TOTAL 4,215.13

180

Table B.41. Upgrader-refinery integration Proposed processing scheme 4: summary of utility requirements

Utilities TOTAL

Fuel (Gcal/d) 46,110.94

Boiler Feed Water (m3/d) 0.00

LP Steam (ton/d) 8,446.32

HP Steam (ton/d) 4,590.63

Cooling Water (1000 m3/d) 1,161.42

Power (MWh/d) 3,530.57


Chem. and Catalysts ($/d) 63,925.50

Pooled Components TOTAL


Gas LHV (ton/d) 656.60





($/units)

(MM$/year)

Fuel (Gcal/d) 39,650.82 70.44 921.63

LP Steam (ton/d) 8,446.32 59.84 166.78

HP Steam (ton/d) 4,590.63 69.41 105.15

Cooling Water (1000 m3/d) 1,161.42 111.91 42.89

Power (MWh/d) 3,530.57 50.00 58.25


Insurance 21.08

Maintenance 210.76


TOTAL 1,616.29

181

Table B43. Upgrader-refinery integration Proposed processing scheme 4: products and raw materials


Unit/d Cost

($/units)

(MM$/year)

Raw Materials



TOTAL 7,979.60

Products

Gasoline (bbl) 50,410 128.43 2,136.46

Diesel (bbl) 224,138 133.77 9,894.36

Asphalt (ton) 9,166 54.01 163.38



(MM$/year)


TOTAL 12,194.2

Production Costs



Depreciation 210.8

TOTAL 9,806.6


Less tax (60%) 1,432.5


Net Income 936.3

CASHFLOW 1,147.0

182




Parameter Value

Plant Life 25 years


CAPEX at -2,-1 and 0 20%, 45%, 35%



Interest Rate 3%



Land Cost 21.1


TOTAL 4,657.7


NPV INDEX 3.26

Table B.46. Upgrader-IGCC integration: summary of unit costs






Hydrocracker (HCU) 150.25 14.72 23.12 1043.5


Integ, Gas Comb, Cycle (IGCC) 4,87 1010,9

SUBTOTAL 3611.5




Storage 996.2

Offsites 239.9

SUBTOTAL 4966.7

Contingence 745.0

TOTAL 5711.7

183

Table B.47. Upgrader-IGCC integration: summary of operation costs




Fuel (Gcal/d) 16,687.81 70.44 387.88

Boiler Feed Water (m3/d) 1,308.45 5.00 2.16

LP Steam (ton/d) 10,940.52 59.84 216.04

HP Steam (ton/d) 3,429.68 69.41 78.56

Cooling Water (1000 m3/d) 1,525.50 111.91 56.34

Hydrogen (ton/d) 0.00 2,150.00 0.00

Power (MWh/d) 0.00 50.00 0.00

Chemicals ($/d) 1,543.87 - 0.51


Insurance 28.56

Maintenance 285.58


TOTAL 1,200.00 Table B.48. Upgrader-IGCC integration: products and raw materials


Unit/d Cost


Raw Materials



TOTAL 7,943.90

Products

Syncrude 35.5 ºAPI (bbl) 300,021 102.75 10,172.96

Hydrogen (ton/d) 0 1350.00 0.00

Power (MWh/d) 2,680 50.00 44.22

Coke (ton) 0 70.00 0.00

TOTAL 10,217.18

184

Table B.49. Upgrader-IGCC integration: costs and revenues


(MM$/year)


TOTAL 10,217.2

Production Costs



Depreciation 285.6

TOTAL 9,429.5

Income before tax 787.7

Less tax (60%) 315.1


Net Income 466.1

CASHFLOW 751.6 Table B.50. Upgrader-IGCC Integration: assumptions and NPV analysis


Parameter Value

Plant Life 25 years


CAPEX at -2,-1 and 0 20%, 45%, 35%



Interest Rate 3%



Land Cost 28.6


TOTAL 6,311.4


NPV INDEX 1.05

185

Table B.51. Refinery A) with IGCC integration: summary of unit costs





Isomerization (ISO) 29.21 51.53 3.59 23.28

Reforming (REF) 46.34 61.17 5.41 79.46

Naphta Hydrotreater (NHT) 38.65 59.35 5.42 37.52

Kerosene Hydrotreater (KHT) 60.55 38.23 8.03 101.93

Gas Oil Hydrotreater (DHT) 94.36 24.51 13.61 131.61

Gas Oil Hydrodesulph (VGHDS) 47.19 19.18 7.05 238.39

Fluid Catalitic Cracking (FCC) 42.52 23.21 6.18 248.38

Visbreaking Unit (VBU) 9.21 6.55 1.50 33.43


Integ, Gas Comb, Cycle (IGCC) - - 2.12 740.74

SUBTOTAL 2,122.32




Storage 750.00

Offsites 437.60

SUBTOTAL 3,354.92

Contingence 503.24

TOTAL 3,858.16 Table B.52. Refinery A) with IGCC integration: summary of operation costs



($/units)

(MM$/year)

Fuel (Gcal/d) 13,601.45 70.44 316.15


LP Steam (ton/d) 3,282.83 59.84 64.82

HP Steam (ton/d) 2,226.64 69.41 51.00

Cooling Water (1000 m3/d) 1,110.85 111.91 41.02

Hydrogen (ton/d) 0.00 2,150.00 0.00

Power (MWh/d) 0.00 50.00 0.00

Chemicals ($/d) 687.28 - 0.23


Insurance 19.29

Maintenance 192.91


TOTAL 787.88

186

Table B.53. Refinery A) with IGCC integration: products and raw materials


Unit/d Cost

($/units)

(MM$/year)

Raw Materials

Lagomedio Crude Oil 31.5 ºAPI 300,000 102.75 10,172.25

H2 Prod. Gas Feed (MSCF) 0 2.25 0.00

TOTAL 10,172.25

Products

Gasoline 95 98,646 128.43 4,180.79

Diesel 100,915 133.77 4,454.78

Kerosene 60,656 161.66 3,235.85

Hydrogen (ton) 243 1350.00 108.18

Power (MWh) 729 50.00 12.02

Coke (ton) 0 70.00 0.00

TOTAL 11,991.62

Table B.54. Refinery A) with IGCC integration: costs and revenues


(MM$/year)


TOTAL 11,991.6

Production Costs


Operation Costs 787.9

Depreciation 192.9

TOTAL 11,153.0

Income before tax 838.6

Less tax (60%) 503.2


Net Income 335.4

CASHFLOW 528.3

187

Table B.55. Refinery A) with IGCC integration: assumptions and NPV analysis


Parameter Value

Plant Life 25 years


CAPEX at -2,-1 and 0 20%, 45%, 35%



Interest Rate 3%



Land Cost 19.3


TOTAL 4,263.3


NPV INDEX 1.13

188

Table B.56. Refinery B) with IGCC integration: summary of unit costs






Reforming (REF) 24,91 57,15 2,97 55,71








Integ, Gas Comb, Cycle (IGCC) - - 0,07 27,73

SUBTOTAL 716,21




Storage 750,00

Offsites 231,38

SUBTOTAL 1.773,88

Contingence 266,08

TOTAL 2.039,96 Table B.57. Refinery B) with IGCC integration: summary of operation costs



Fuel (Gcal/d) 14.403,09 70,44 334,78


LP Steam (ton/d) 1.275,92 59,84 25,19

HP Steam (ton/d) 1.635,64 69,41 37,47

Cooling Water (1000 m3/d) 756,29 111,91 27,93

Hydrogen (ton/d) 0,00 2.150,00 0,00

Power (MWh/d) 915,60 50,00 15,11

Chemicals ($/d) 0,00 - 0,00


Insurance 10,20

Maintenance 102,00


TOTAL 624,63

189

Table B.58. Refinery B) with IGCC integration: products and raw materials

Products and Raw Materials Cost

Unit/d Cost ($/units) (MM$/year)

Raw Materials

Synthetic Crude Oil 35.5 ºAPI 300.000 110,00 10.890,00


TOTAL 10.896,74

Products

Gasoline 95 64.559 128,43 2.736,15

Diesel 126.774 133,77 5.596,32

Kerosene 106.019 161,66 5.655,86

Hydrogen (ton) 0 1350,00 0,00

Power (MWh) 0 50,00 0,00

Coke ton/d ($/ton) 0 70,00 0,00

TOTAL 13.988,33

Table B.59. Refinery B) with IGCC integration: costs and revenues


(MM$/year)


TOTAL 13.988,3

Production Costs



Depreciation 102,0

TOTAL 11.623,4

Income before tax 2.365,0

Less tax (60%) 946,0


Net Income 1.393,7

CASHFLOW 1.495,7

190

Table B.60. Refinery B) with IGCC integration: assumptions and NPV

analysis


Parameter Value

Plant Life 25 years


CAPEX at -2,-1 and 0 20%, 45%, 35%



Interest Rate 3%



Land Cost 10,2


TOTAL 2.254,2


NPV INDEX 10,53

191

Table B.61. Upgrader-Refinery-IGCC integration proposed processing scheme 5: summary of utilities

Utilities UPGRADER REFINERY IGCC TOTAL

Fuel (Gcal/d) 50,500.06 19,630.88 0.78 70,131.71

Boiler Feed Water (m3/d) 0.00 0.00 2,388.61 2,388.61

LP Steam (ton/d) 8,143.54 1,275.92 2,838.42 12,257.88

HP Steam (ton/d) 5,122.68 1,635.56 0.07 6,758.32

Cooling Water (1000 m3/d) 1,116.82 756.29 492.72 2,365.83

Power (MWh/d) 3,617.74 1,190.55 0.00 4,808.29

Hydrogen (ton/d) 991.21 165.64 0.00 1,156.85

Chemicals ($/d) 0.00 0.00 1,566.83 1,566.83

Catalysts & Royalties ($/d) 62,377.41 11,654.97 35,676.76 109,709.13

Pooled Components UPGRADER REFINERY IGCC TOTAL

Hydrogen (ton/d) 991.21 165.64 0.00 1,156.85

Power (MWh/d) 0.00 0.00 20,455.25 20,455.25

Gas LHV (ton/d) 1,594.61 531.43 0.00 2,126.04

192

Table B.62. Upgrader-Refinery-IGCC integration Upgrader-Refinery-IGCC integration proposed processing scheme 5: summary of unit costs



Atmospheric Destillation (UP-ADU) 398.46 17.31 60.24 193.33




Delayed Coking (UP-DCU) 156.36 1.11 26.53 624.57

Atmospheric Destillation (R-ADU) 300.00 35.56 40.40 167.28



Reforming (R-REF) 24.91 57.15 2.97 55.71



Gas Oil Hydrotreater (R-DHT) 118.33 29.84 17.78 149.93

Gas Oil Hydrodesulph (R-VGHDS) 1.52 18.75 0.23 13.56



Delayed Coking (R-DCU) 1.02 0.00 0.17 27.51

Integ, Gas Comb, Cycle (IGCC) - - 4.94 1,422.29

SUBTOTAL 4,265.80











SUBTOTAL 7,219.49



TOTAL 8,302.41

193

Table B.63. Upgrader-Refinery-IGCC integration Upgrader-Refinery-IGCC integration proposed processing scheme 5: summary of operation costs



($/units)

(MM$/year)

Fuel (Gcal/d) 49,214.25 70.44 1,143.92


LP Steam (ton/d) 12,257.88 59.84 242.05

HP Steam (ton/d) 6,758.32 69.41 154.81

Cooling Water (1000 m3/d) 2,365.83 111.91 87.37

Hydrogen (ton/d) 0.00 2,150.00 0.00

Power (MWh/d) 0.00 50.00 0.00

Chemicals ($/d) 1,566.83 - 0.52


Insurance 41.51

Maintenance 415.12


TOTAL 2,308.12 Table B.64. Upgrader-Refinery-IGCC integration Upgrader-Refinery-IGCC integration proposed processing scheme 5: products and raw materials


Unit/d Cost

($/units)

(MM$/year)

Raw Materials



TOTAL 8,002.94

Products

Gasoline 95 64,559 128.43 2,736.14

Diesel 126,774 133.77 5,596.31

Kerosene 106,019 161.66 5,655.86

Hydrogen (ton) 0 1350.00 0.00

Power (MWh) 15,647 50.00 258.17

Coke ton/d ($/ton) 0 70.00 0.00

TOTAL 14,246.48

194

Table B.65. Upgrader-Refinery-IGCC integration Upgrader-Refinery-IGCC

integration proposed processing scheme 5: costs and revenues


(MM$/year)


TOTAL 14,246.5

Production Costs



Depreciation 415.1

TOTAL 10,726.2


Less tax (60%) 2,112.2


Net Income 1,367.2

CASHFLOW 1,782.3


integration proposed processing scheme 5: assumptions and NPV analysis


Parameter Value

Plant Life 25 years


CAPEX at -2,-1 and 0 20%, 45%, 35%


CO2 Emission Tax ($/ton) 13.00

Interest Rate 3%



Land Cost 41.5


TOTAL 9,174.2


NPV INDEX 2.36

195

Table B.67. Upgrader-Refinery-IGCC integration proposed processing scheme 6: summary of utilities

Utilities UPGRADER REFINERY IGCC TOTAL

Fuel (Gcal/d) 30,844.68 19,442.87 0.00 50,287.55

Boiler Feed Water (m3/d) 0.00 0.00 1,110.41 1,110.41

LP Steam (ton/d) 8,140.00 1,251.64 2,807.83 12,199.48

HP Steam (ton/d) 3,279.09 1,626.77 0.00 4,905.86

Cooling Water (1000 m3/d) 919.60 752.56 596.01 2,268.18

Power (MWh/d) 3,354.47 1,180.62 0.00 4,535.09

Hydrogen (ton/d) 989.77 164.50 0.00 1,154.26

Chemicals ($/d) 0.00 0.00 1,550.19 1,550.19

Catalysts & Royalties ($/d) 53,698.64 11,353.02 35,303.57 100,355.23

Pooled Components UPGRADER REFINERY IGCC TOTAL

Hydrogen (ton/d) 209.49 164.50 780.27 1,154.26

Power (MWh/d) 0.00 0.00 5,970.62 5,970.62

Gas LHV (ton/d) 1,587.62 518.83 0.00 2,106.45

196




Atmospheric Destillation (UP-ADU) 396.42 17.31 59.93 192.83




Delayed Coking (UP-DCU) 155.56 1.11 26.39 622.48

Atmospheric Destillation (R-ADU) 300.00 35.56 40.40 167.28



Reforming (R-REF) 24.67 57.16 2.94 55.40



Gas Oil Hydrotreater (R-DHT) 118.33 29.84 17.78 149.93

Gas Oil Hydrodesulph (R-VGHDS) 0.85 14.53 0.13 8.34




SUBTOTAL 4,160.56











SUBTOTAL 7,001.70



TOTAL 8,051.96

197




($/units)

(MM$/year)

Fuel (Gcal/d) 29,562.83 70.44 687.15


LP Steam (ton/d) 12,199.48 59.84 240.89

HP Steam (ton/d) 4,905.86 69.41 112.37

Cooling Water (1000 m3/d) 2,268.18 111.91 83.77

Hydrogen (ton/d) 0.00 2,150.00 0.00

Power (MWh/d) 0.00 50.00 0.00

Chemicals ($/d) 1,550.19 - 0.51


Insurance 40.26

Maintenance 402.60




Unit/d Cost

($/units)

(MM$/year)

Raw Materials



TOTAL 7,962.18

Products

Gasoline 95 63,975 128.43 2,711.36

Diesel 127,353 133.77 5,621.90

Kerosene 106,019 161.66 5,655.86

Hydrogen (ton) 0 1350.00 0.00

Power (MWh) 15,480 50.00 255.43

Coke ton/d ($/ton) 0 70.00 0.00

TOTAL 14,244.55

198

Table B.71. Upgrader-Refinery-IGCC integration Upgrader-Refinery-IGCC integration proposed processing scheme 6: costs and revenues


(MM$/year)


TOTAL 14,012.8

Production Costs



Depreciation 402.6

TOTAL 10,092.1


Less tax (60%) 2,352.4


Net Income 1,560.9

CASHFLOW 1,963.5 Table B.72. Upgrader-Refinery-IGCC integration Upgrader-Refinery-IGCC



Parameter Value

Plant Life 25 years


CAPEX at -2,-1 and 0 20%, 45%, 35%



Interest Rate 3%



Land Cost 40.3


TOTAL 8,897.4


NPV INDEX 2.82

199








Reformer (REF) 47.24 58.45 5.60 80.34



SUBTOTAL 3,802.36




Storage 996.15

Offsites 732.92

SUBTOTAL 5,619.04

Contingence 842.86

TOTAL 6,461.90 Table B.74. Upgrader-Refinery-IGCC integration Upgrader-Refinery-IGCC integration proposed processing scheme 7: summary of operation costs




Fuel (Gcal/d) 14,007.73 70.44 325.59

LP Steam (ton/d) 11,581.79 59.84 228.70

HP Steam (ton/d) 2,980.83 69.41 68.28

Cooling Water (1000 m3/d) 1,554.74 111.91 57.42

Power (MWh/d) 0.00 50.00 0.00


Insurance 32.31

Maintenance 323.09


TOTAL 1,179.06

200

Table B.75. Upgrader-Refinery-IGCC integration Upgrader-Refinery-IGCC integration proposed processing scheme 7: products and raw materials


Unit/d Cost


Raw Materials



TOTAL 7,929.34

Products

Gasoline (bbl) 37,446 128.43 1,587.04

Diesel (bbl) 261,216 133.77 11,531.15

Hydrogen (ton) 0 1350.00 0.00

Power (MWh) 2,488 50.00 41.06

Coke (ton) 0 70.00 0.00

TOTAL 13,159.25 Table B.76. Upgrader-Refinery-IGCC integration Upgrader-Refinery-IGCC

integration proposed processing scheme 7: costs and revenues


(MM$/year)


TOTAL 13,159.2

Production Costs



Depreciation 323.1

TOTAL 9,431.5


Less tax (60%) 2,236.7


Net Income 1,489.2

CASHFLOW 1,812.3

201

Table B.77. Upgrader-Refinery-IGCC integration Upgrader-Refinery-IGCC integration proposed processing scheme 7: assumptions and NPV

analysis


Parameter Value

Plant Life 25 years


CAPEX at -2,-1 and 0 20%, 45%, 35%



Interest Rate 3%



Land Cost 32.3


TOTAL 7,140.4


NPV INDEX 3.39 Table C.78. Upgrader-Refinery-IGCC integration Upgrader-Refinery-IGCC integration proposed processing scheme 8: summary of unit costs







Reformer (REF) 56.81 58.27 6.74 89.28 Solvent Deasphalting Unit (SDA) 156.36 1.11 26.53 208.73


SUBTOTAL 4,757.02




Storage 996.15

Offsites 871.94

SUBTOTAL 6,684.87

Contingence 1,002.73

TOTAL 7,687.60

202




($/units)

(MM$/year)

Fuel (Gcal/d) 20,518.69 70.44 476.93

LP Steam (ton/d) 13,232.50 59.84 261.29

HP Steam (ton/d) 2,803.25 69.41 64.21

Cooling Water (1000 m3/d) 2,040.31 111.91 75.35

Power (MWh/d) 0.00 50.00 0.00


Insurance 38.44

Maintenance 384.38




Unit/d Cost


Raw Materials


H2 Prod. Gas Feed (MSCF) 0 2.25 0.00

TOTAL 7,923.79

Products

Gasoline (bbl) 45,034 128.43 1,908.64

Diesel (bbl) 224,138 133.77 9,894.36

Hydrogen (ton) 1,175 1350.00 523.56

Power (MWh) 10,652 50.00 175.76

Asphalt (ton) 0 54.01 0.00

TOTAL 12,502.33

203

Table B.81. Upgrader-Refinery-IGCC integration Upgrader-Refinery-IGCC integration proposed processing scheme 8: costs and revenues


(MM$/year)


TOTAL 12,502.3

Production Costs



Depreciation 384.4

TOTAL 9,717.6


Less tax (60%) 1,670.9


Net Income 1,113.9

CASHFLOW 1,498.3




Parameter Value

Plant Life 25 years


CAPEX at -2,-1 and 0 20%, 45%, 35%



Interest Rate 3%



Land Cost 38.4


TOTAL 8,494.8


NPV INDEX 2.05

204

Appendix C: Petroleum refining correlations

For simulating the refining processes HTU, HCU, DCU the correlations from

Gary and Handwerk’s Handbook (2007) and Baird (1987) were used in this

work. Once the products from distillation unit are within the given ranges

specified, the HPI correlation can be applied.

The petroleum refining process correlations are empirical in nature, having

been derived from a number of different sources. These expressions are

adequate to estimate the yields and properties that a refining unit can achieve

in a commercial operation (Baird 1987).

Petroleum Refining Process Correlations were published in 1987 and

nowadays they are used to developed lot of important refinery project

(NEXIDEA™ System, 2010)

The performance of refining process depends upon a number of variables.

They are feedstock quality, operating severity, conversion level, process

conditions, catalyst and physical characteristics of the units, which are

apparently captured within the above correlations. The main assumptions and

correlations for the units are described, the complete information for all units

can be found elsewhere in the work of Baird (Baird 1987).

a) Hydrocracking correlation:

For hydrocracking correlation the product distribution is a function of

operating severity and feedstock quality and uses the light gasoline (C5/

82.22 ºC) yield as a measure of the conversion during the process. The

other product yields are correlated with the gasoline yield. The light

205

gasoline is calculated from the feed gravity and characterization factor. In

this work the type of operation selected was maximum diesel fuel

production.

The hydrocracking correlation can be used for gas oils boiling range

between 315.5 ºC and 515.5 ºC and a characterization factor ranging from

11.0 to 12.2.

Yield correlations

Feed properties required : API = Density (API gravity)

Sf = Sulfur content , Nf = Nitrogen content, Kf = Watson factor

SGf = Specific gravity, PPf = Pour point, VABPf = Volume average boiling

point

1-Hydrogen Consumption, WT % H = 21.14 - 1.76 * (Kf) + 0.003 * VABPf

+ 0.0625 * (Sf) + 0.214 * (Nf)

2-Hydrogen Sulfide Yield, WT % H2S = 1.0625 * (Sf)

3-Ammonia Yield , WT% NH3 = 1.214 * (Nf)

4-Hydrocarbon Yield, WT % HC = 100 + H - H2S - NH3

5-Light Gasoline Yield, WT % LG = (HC / 100) * (0.15 * (APIf) + 2.4 *

(Kf) - 22.89)

6-Refinery Fuel Gas (C3 and Lighter) Yield, WT %

RG = 0.094 * (LG) + 1

C4 LPG Yield, WT %

C4LPG = 0.59 * (LG)

7-Naphtha Yield , WT% HN = 3.1 * (LG)

8-Diesel Fuel Yield , WT% DF = HC - RG - C4LPG - LG - HN

206

Properties of products correlations

LIGHT GASOLINE (C5/180 F):

Characterization Factor Kg = 0.25 * (Kf) + 9.7

Specific Gravity SGg = 8.3897 / Kg

Research Octane Number (Clear) RONCL = -5.2 * (Kf) + 145

RON3 = 0.57 * (RONCL) + 49.5

Motor Octane Number (Clear) MONCL = 0.8 * (RONCL) + 13.6

Reid Vapour Pressure (PSIA) RVPg = 13

Characterization Factor Kn = 0.25 * (Kf) + 8.92

Specific Gravity SGna = (9.014 / Kn)

Research Octane Number (Clear) RONCLn = -19.5 * (Kf) + 289.4

RON3n = 0.76 * (RONCLn) + 33

Motor Octane Number (Clear) MONCLn = 0.91 * (RONCLn) + 5.4

Reid Vapour Pressure (PSIA) RVPn = 0.8

Sulfure content, WT% Sn = 0.001

DIESEL FUEL(345/650 F):

Volumetric Average Boiling Point (F) VABPd = 498.4

Characterization Factor Kd = 0.425 * (Kf) + 6.91

Specific Gravity SGd = (9.8582 / Kd)

Aniline Point (F) APd = -375.4 + (439.1 / SGd)

'Sulfure content, WT% Sd = 0.02 * (Sf)

207

b) Hydotreating Correlations:

Hydrogen consumption:

For Hydrotreating correlations the hydrogen consumed is a function of

quantity of sulphur, nitrogen and oxygen removed, the degree of olefins

and aromatics saturation, the metal content in the feedstock and the

molecular characteristic of the feedstock. The following assumptions were

made (Baird 1987).

• The hydrogen consumption is highly influenced by sulphur removal. The

consumption depends upon the feed constituents. A linear correlation

between hydrogen consumption and sulphur removal, based on the type of

chargestock (virgin or cracked) is assumed.

• Linear relationship for nitrogen removal in spite of the fact that the quantity

varies considerably with the molecular type of feedstock.

• The oxygen removal is taken as a constant

• The main objective is the desulfurization of the feedstock and thus, the

hydrogen consumed for aromatics saturation is insignificant.

• The hydrogen consumed for olefin saturation is accounted for explicitly

only in the naphtha hydrotreating correlation. In middle distillate and gas oil

correlation, the higher hydrogen consumption for olefin saturation in

cracked stocks is reflected in the different coefficients for bromine number

and the proportion of FCC stocks in the feed.

208

• To consider a metal, especially vanadium and nickel in the feedstock a

metal adjustments factor is applied to the hydrogen consumption equation.

• A non- chemical constant hydrogen consumption are assumed :

• Hydrogen lost by solution : 20 scf/bbl (for naphtha hydrotreating) and

35 scf/bbl (for gas oil hydrotreating)

• Hydrogen lost by Leaks: 5 scf/bbl (for naphtha hydrotreating) and 5

scf/bbl (for gas oil hydrotreating)

• The yield of hydrogen sulphide is proportional to the sulphur removed

which is fixed for each type of hydrotreating unit.

• The ammonia yield is proportional to the nitrogen removed and is fixed for

each type of hydrotreating unit.

• The water yield is not considered. The gas (C1 through C4’s) yield and

composition in naphtha hydrotreating is a function of the sulphur and nitrogen

removed. For middle distillate and heavy gas oil hydrotreaters the gas yield is

proportional to the hydrogen consumption. The gas composition was

considered constant.

• For naphtha hydrotreater the naphtha yield is the difference between the

feed (hydrocarbon plus hydrogen) and the light products (hydrogen sulphide,

ammonia, and gas). For middle distillate and gas oil hydrotreaters the naphtha

yield is a function of the sulphur removed and the molecular weight or boiling

range of the feedstock. The naphtha yield correlation is based on the degree

209

of hydrogenation as measured by the hydrogen consumption. For residue

hydrodesulfurizer correlation a adjustment factor for feed gravity is included.

• In naphtha hydrotreater correlations the sulphur and nitrogen content of

the naphtha product is fixed by the degree of desulfurization and

denitrogenation and adjustment is made to the product gravity and octane

number based on the amount of olefin and aromatic saturation. The volatility

characteristics of the product and products are considered the same.

• For middle distillate, gas oil and residue hydrotreaters the correlation

assumed that the properties for naphtha are constant for all cases

• In gas oil hydrodesulfurization correlations the middle distillate yield is

calculated directly with the sulphur removed and in residue

hydrodesulfurization the sulphur removed and distillate yield is made indirectly

by naphtha yield as the correlating patrameter.

• Since the gravity of hydrotreated middle distillate varies with the amount of

cracking which occurs during the processing, the amount of hydrogen

consumed is used as correlating parameter.

• The amount of sulphur and nitrogen is fixed by the degree of

denitrogenation and desulfurization. The flow properties like pour point, freeze

point and viscosity do not change a lot as a result of hydrotreatment .

• The properties of the middle distillate produced in gas oil and residue

hydrodesulfirization are assumed to be the same for all operations.

210

The heavy gas oil yields are calculated by difference to complete the unit

material balance.

• The sulphur and nitrogen in the product are fixed by the operating severity

and the product gravity is correlated with the amount of sulphur removed. The

properties of the heavy gas oil produced in gas oil and residue

hydrodesulfirization are assumed to be the same for all operations.

Yield correlations

Feed properties required : API = Density (API gravity)

Sf = Sulfur content , Nf = Nitrogen content, Kf = Watson factor

SGf = Specific gravity, PPf = Pour point, VABPf = Volume average boiling

point,

1- Hydrogen Consumption,(scf/bbl) Ho = 120 * (Sf) + 159 * (Nf) +

8 * (SGf * Bf) + 25

2- Hydrogen Consumption,WT% H = (1 / 658.29) * (Ho / SGf)

3- Hydrogen Sulfide Yield, WT % H2S = 1.01 * (Sf)

4- Ammonia Yield , WT% NH3 = 0.669 * (Nf)

5- Refinery Gas Yield, WT % RG = 0.132 * (H)

6- Naphtha Yield , WT% HN = 0.2 * (H)

7- Middle Distillate Yield , WT% MD = 100 + H - H2S - NH3 -

RG - HN

211


NAPHTHA:

Volumetric Average Boiling Point (F) VABPn = 267

Characterization Factor Kn = 11.7

Specific Gravity SGna = 0.7732

Research Octane Number (Clear) RONCLn = 55

Sulfure content, WT% Sn = 0.0085 * (Sf)

Nitrogen content, WT% Nn = 0.07 * (Nf)

MIDDLE DISTILLATE:

Specific Gravity SGd = SGf - 0.026 * H

Characterization Factor Kd = Kf * (SGf - SGd) + Kf

Volumetric Average Boiling Point (F)

VABPd = ((Kd) * (SGd)) ^ 3 - 460

Pour Point (F) PPd = PPf + (40 * (Bf / (PPf + 100)))

Freeze Point (F)

FPd = FPf + (40 * (Bf) / (FPf + 100))

Sulfure content, WT% Sd = (5 * (Sf) - (HN) * (Sn)) / MD

Nitrogen content, WT% Nd = (45 * Nf - (HN) * (Nn)) / MD

APId APId = (141.5 / SGd) - 131.5

Cetane Index

CI = -420.34 + 0.016 * (APId) ^ 2 + 0.192 * (APId) * (2.85) + 65.01 * (2.85)

^ 2 - 0.0001809 * (2.85) ^ 2

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c) Delayed coking correlation

The yield correlations use the conradson carbon residue of the feedstock as

the correlation parameter to calculate the yields for delayed coking units

(Baird 1987).

The correlations are based on the following conditions:

a. Coke drum pressure between 35 to 45 psig

b. Feedstock is a straight run residue

c. Coker temperature between 432.2 and 440.5 ◦C

d. Recycle of 10% volume of fresh feed

The constraints for the quality of final products and feedstocks, hardware

maximum capacities, utility and product supply and demand are considered

for the whole flowsheet design. The consumption and production of utilities,

consumption of chemicals and catalysts are also predicted, these correlations

are usually linear with respect to the feed flow (Baird 1987). The correlations

are the following:

Yield correlations

1. Hydrogen Sulfide Yield, WT % H2S = 0.25 * (Sf)

2. Refinery Gas Yield, WT % RG = 3.5 + 0.1 * (CCRf)

3. C3/C4 LPG Yield, WT % LPG = 4.3 + 0.044 * (CCRf)

4. Naphtha (C5/400 F) Yield , WT% NAP = 11.38 + 0.335 * (CCRf)

5. Total Gas oil (400/950 F) Yield , WT% TGO = 100 - RG - LPG - NAP -

213

COKE

6. Weight Ratio Light Gas oil to Total Gas Oil Yield

RATIO = 0.38 + 0.011 * (CCRf) - 3# * (10 ^ (-4)) * (CCRf) ^ 2

7. Light Gas oil (400/650 F) Yield , WT% LGO = RATIO * TGO

8. Heavy Gas Oil( 650/950 F) Yield , WT% HGO = TGO - LGO

9. Coke Yield , WT% COKE = 1.6 * (CCRf)


NAPHTHA (C5/400 F):

Sulfure content Sn = 0.14 * (Sf)

Nitrogen content Nn = 0.01 * (Nf)

LIGHT GAS OIL(400/650 F):

Sulfure content Slgo = 0.45 * (Sf)

Nitrogen content Nlgo = 0.24 * (Nf)

HEAVY GAS OIL(650/950 F):

Specific Gravity SGhgo = 0.58 * SGf + 0.4053

Sulfure content Shgo = 0.82 * (Sf)

Nitrogen content Nhgo = 0.63 * (Nf)

Bromine Number BNhgo = 283 - 270 * (SGhgo)

Aniline Point, F APhgo = (637.6 / SGhgo) - 530

214

COKE, WT%:

Sulfure content

SCOKE = (100# * (Sf) - 23.52 * (Sf) - 0.14 * (Sf) * (NAP) - 0.45 * (Sf) * (LGO) -

0.82 * (Sf) * (HGO)) / COKE

Nitrogen content

NCOKE = (100# * (Nf) - 0.01 * (Nf) * (NAP) - 0.24 * (Nf) * (LGO) - 0.63 * (Nf) *

(HGO)) / COKE