PTQ+Q2+2011

special features

thermal combustion

separation

refininggas processingpetrochemicals

petroleum technology quarterly

ptqQ2 2011

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©2011. The entire content of this publication is protected by copyright full details of which are available from the publishers. All rights reserved. No part of this publication may be reproduced, stored in a retrieval system or transmitted in any form or by any means – electronic, mechanical, photocopying, recording or otherwise – without the prior permission of the copyright owner.The opinions and views expressed by the authors in this publication are not necessarily those of the editor or publisher and while every care has been taken in the preparation of all material included in Petroleum Technology Quarterly the publisher cannot be held responsible for any statements, opinions or views or for any inaccuracies.

3 Decision time ChrisCunningham

5 Processing Trends

15 ptq&a 29 Optimising safety relief and flare systems AlbanSirven,JulienGrosclaudeandGuillaumeFenolTechnip France JeremySaadaInvensys Operations Management

41 Energy performance monitoring RobertChares,HervéClosonandHuguesStefanskiBelsim Jean-ClaudeNoisierSIR

47 SO2 emission control for resid combustion

RickBirnbaumCansolv Technologies

51 Bulk separation of gas-liquid mixtures GiuseppeMosca,PierreSchaefferandBartGriepsmaSulzer Chemtech HarryKooijman Shell Global Solutions International

57 Diverting low-sulphur heavy stocks for fuel oil production RajeevKumar,ChithraV,PeddyVCRaoandNVChoudary Bharat Petroleum Corporation Ltd, India

65 Reducing carbon footprint TanmayTaraphdarTechnip KT India

75 A promoter for selective H2S removal: part II

GeraldVorberg,RalfNotzandTorstenKatz BASF SE WielandWacheandClausSchunk Bayernoil Raffineriegesellschaft

87 Main fractionator revamp JohnPayneandDanDarbyFoster Wheeler

97 Small-scale gas to liquids AndrewHolwellOxford Catalysts Group

103 Simulation of a visbreaking unit SRezaSeifMohaddecy,SepehrSadighi,OmidGhabuliandMahdiRashidzadeh Research Institute of Petroleum Industry

111 Modelling for ULSD optimisation KlasDahlgren Apex Optimisation/Dynaproc AnRigden Chevron HenrikTerndrup Apex Optimisation

119 Technology in Action

120 Industry News

126 New Products

Repsolrefinery,Tarragona,Spain Photo: Repsol

Q2 (Apr, May, Jun) 2011www.eptq.com

ptqYLRETRAUQYGOLONHCET MUELORTEP

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The European Union has arguably been the global leader in biodiesel production and use, with overall

biodiesel production increasing from 1.9 million tonnes in 2004 to nearly 10.3 million tonnes in 2007. Biodiesel production in the US has also increased dramatically in the past few years from 2 million gallons in 2000 to approximately 450 million gallons in 2007. According to the National Biodiesel Board, 171 companies own biodiesel manufacturing plants and are actively marketing biodiesel.1. The global biodiesel market is estimated to reach 37 billion gallons by 2016, with an average annual growth rate of 42%. Europe will continue to be the major biodiesel market for the next decade, followed closely by the US market.

Although high energy prices, increasing global demand, drought and other factors are the primary drivers for higher food prices, food competitive feedstocks have long been and will continue to be a major concern for the development of biofu-els. To compete, the industry has responded by developing methods to increase process efficiency, utilise or upgrade by-products and operate with lower quality lipids as feedstocks.

Feedstocks

Biodiesel refers to a diesel-equivalent fuel consisting of short-chain alkyl (methyl or ethyl) esters, made by the transesterification of triglycerides, commonly known as vegetable oils or animal fats. The most common form uses methanol, the cheapest alcohol available, to produce methyl esters. The molecules in biodiesel are pri-marily fatty acid methyl esters (FAME), usually created by trans-esterification between fats and metha-nol. Currently, biodiesel is produced from various vegetable and plant oils. First-generation food-based feedstocks are straight vegetable oils such as soybean oil and animal fats such as tallow, lard, yellow grease, chicken fat and the by-products of the production of Omega-3 fatty acids from fish oil. Soybean oil and rapeseeds oil are the common source for biodiesel produc-tion in the US and Europe in quanti-ties that can produce enough biodie-sel to be used in a commercial market with currently applicable

PTQ Q2 2011 3

Editor Chris Cunningham [email protected]

Production EditorRachel [email protected]

Graphics EditorRob Fris [email protected]

Editorial tel +44 844 5888 773fax +44 844 5888 667

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ISSN 1362-363X

Petroleum Technology Quarterly (USPS 0014-781) is published quarterly plus annual Catalysis edition by Crambeth Allen Publishing Ltd and is distributed in the USA by SPP, 75 Aberdeen Rd, Emigsville, PA 17318. Periodicals postage paid at Emigsville PA.Postmaster: send address changes to Petroleum Technology Quarterly c/o POBox 437, Emigsville, PA 17318-0437Back numbers available from the Publisher at $30 per copy inc postage.

Vol 16 No 3

Q2 (Apr, May, Jun) 2011

Decision time

W ith production in surplus, western majors moving out, diesel production in deficit and more regulations to cope with, Europe’s refining industry has some hefty decisions to make on investments.

However, the prospect of new owners from Asia appears likely to impose new rules on the decision-making process.

In its most recent report on the future of local refining, Blueprint for an inte-grated European energy network, the European Commission argued that European refiners can avert a significant and growing gap in trade should they go ahead with sufficient increases in the capacity and complexity of their sites to increase the rate of diesel production, at the expense of gasoline output.

And the trade gap does not only apply to automotive fuels; the EU was a net importer of gas oil and jet fuel to the tune of more than 30 million t/y in the latter years of the last decade. According to the refining industry’s own esti-mate, some 20 large-scale hydrocrackers would be needed to offset the short-fall, at a cost likely to approach €10 billion.

A new fleet of 20 hydrocrackers still could not restore the balance. The latest round of reductions in the sulphur content of marine fuel oil in compliance with the International Maritime Organisation’s Marpol Annex VI regulations will require another 15 million t/y of gas oil of the appropriate quality, which could mean another 10 big hydrocrackers by 2015, with a further requirement by 2020 as more global restrictions on marine fuel sulphur come into force.

With margins under pressure, local crude production in decline, and the backwash of economic downturn still taking its toll, the prospects for large- scale investment in new units and major revamps are not encouraging. Indeed, European refiners have cut back previous estimates of spending on improve-ments to the system over the next eight years by more than half.

All of this assumes that the refining industry is playing catch-up with changes in established trade patterns and, as ever, the steady march of fuel quality regulations. As the EC’s blueprint points out, there has been little incen-tive to invest in new European hydrocrackers to produce more diesel when the economic downturn has slowed growth in demand for middle distillates and the US has continued to provide a ready market for excess gasoline.

But changes may be afoot that are far more fundamental to the future shape of European refining. As PTQ went to press, Essar Energy and Shell were approaching a final decision on Essar’s purchase of the Stanlow refinery, the UK’s second largest. While the Indian-based major has expressed its admira-tion for Stanlow as a crude processing site, it has also been forthright about the refinery as an import terminal for production from its own refining fleet. The downturn has had less of an impact on some economies, most obviously India and China. Unlike their counterparts in Europe, Indian and Chinese refiners have not restrained their spending on new production and have, in fact, out-stripped growth in domestic demand with new capacity, nor has a limited local market done anything to restrain throughput.

As western oil majors move out of European refining and their Asian coun-terparts move in with their surplus tonnes for export, there is a consensus that Europe will become more import-dependent. When Asian refiners’ domestic markets catch up with their local production and the need to ship products through Europe subsides, the balance of trade may be radically different.

CHRIS CUNNINGHAM

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Gasoline and diesel imbalances in the Atlantic Basin Part 1: market outlookEric Benazzi, Marketing Director, Axens [email protected] European refining industry is coping with declin-ing domestic demand while the imbalance between product supply and market demand persists, in partic-ular the deficit in diesel supply and the excess in gasoline production.

This article presents the prospects for the evolution of this situation in the coming years. A second article in a coming issue of PTQ will address refinery technol-ogy and solutions to rebalance output.

Current market situation: origin of the imbalanceThe European imbalances are not recent; the European gasoline surplus has existed since the early 2000s, and in 2009 it exceeded 0.75 Mbdoe (million barrels per day of oil equivalent), while the diesel deficit reached about 0.5 Mbdoe (see Figure 1).

On the other side of the Atlantic, the US market pres-ents a recurring gasoline deficit of around 0.7 Mbdoe and, since 2008, diesel has been exported.

What are the main causes of this situation?In 2009, 144 refineries were operating in the US with an average crude oil distillation capacity by refinery of just above 120 000 b/d and a Nelson complexity index of 10.2. In Europe (EU-27), 110 refineries were listed with an average capacity of around 127 000 b/d and a Nelson complexity index of 7.3, lower than the US value.

At first glance, US and European refinery production seems to correspond well to demand. US production is around 18 million b/d and is mainly oriented toward gasoline (49% of production), while European refinery production is mainly dominated by middle distillates (42%) with a total production of about 14 million b/d.

However, is refinery production sufficiently tailored to demand? The following paragraph will further examine this question.

Figure 2 compares US refinery throughput volume with US demand for fuels. First, it can be seen that there is no fundamental inadequacy between the struc-ture of US production when compared with demand given as a percentage. The main problem stems from the insufficient gasoline throughput, about 7.7 Mbdoe, compared to US demand, which was set at 9.1 Mbdoe in 2009. As a result, the US needs to import gasoline.

The European situation is rather different. The European refineries do not produce enough diesel —their current output is 5.2 Mbdoe whereas 5.8 Mbdoe is required — and they produce too much gasoline — 3 Mbdoe, for a demand that reaches, with difficulty,

Reduce gasoline cutpoint

www.eptq.com PTQ Q2 2011 5

Processing Trends

0.6

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4.1 Mbdoe (incl. bio, 0.04 MBdoe)

1.5 Mbdoe

Figure 1 Gasoline and diesel surplus/deficit Source: JODI

Figure 2 US refinery throughput and demand Source: DoE, JODI, Axens

proc trends copy 8.indd 1 10/3/11 12:21:00

2020 (see Figure 4). Demand for diesel will continue to dominate the European fuels market with more than 30% of the market share.

During that time, the demand for gasoline will continue to represent the principal on-road fuel in the US with nearly 44% of US demand by 2020. Conversely, in Europe, gasoline will only represent 12.4% of the refined product demand including biofuels (see Figure 4).

As previously noted, on-road diesel demand should continue to grow on both sides of the Atlantic. Over the period 2009–2020, the incremental demand should represent about 0.4 Mbdoe for the US and 0.6 Mbdoe for the EU-27. Nevertheless, it is important to note that, in the US, a significant part of the incremental demand for on-road diesel for the next 10 years corresponds to a return to its highest level of 2007.

The second striking feature is that both US and European gasoline demand are expected to decrease, with European demand declining at a faster rate. In addition, if we consider the increasing amount of etha-nol that will be incorporated into the US gasoline pool, it is uncertain that excess European gasoline will continue to find an outlet on the US market.

This demand for on-road fuels forecast for the US is based on the fact that US passenger car sales through to 2020 will remain in the majority for gasoline cars, with the implementation of new CAFE programmes aimed at reducing car engine fuel consumption.

For Europe, our demand forecasts are based on a reference scenario for which, by 2020, diesel passenger

6 PTQ Q2 2011 www.eptq.com

2.3 Mbdoe. As a result, it can be said that the European imbalance is structural.

Consequently, Europe is importing diesel and export-ing gasoline to the US. How will the situation evolve through 2020?

Different parameters could affect the situation. The first is the change in product demand forecast. The potential impact of a change in automobile motorisation should also be taken into account. Other parameters include the effects of the level of biofuels incorporated into the on-road fuels pool and the reduction in refining capacity that could occur in the coming years.

Gasoline and diesel demand forecast: a 2020 outlookFor both Europe and the US, the demand for gasoline is forecast to decrease, whereas demand for on-road diesel should continue to increase between 2009 and

Figure 3 Diesel and gasoline imports across the Atlantic, 2009

eo

db

M

US

Market structure AAGR, 09-20

Figure 4 Forecast demand for US and European fuels

eo

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EU-27

Market structure AAGR, 09-20

Mbdoe 2009


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cars would represent 54% of sales, gaso-line passenger cars 36% and hybrid-gasoline 10%. With such a scenario, demand for on-road diesel will continue to increase and will represent just above 70% of on-road fuels consumption in the EU in 2020. At that time, the on-road diesel-to-gasoline ratio will reach a value of 2.5, which can be compared to the current value of 1.7 and to 0.3 for the US.

Could we expect an inversion of the trend in European demand by 2020? At this point the question is: can this scenario be challenged, especially the European situation and assumptions? For that, it is necessary to examine automo-bile motorisation further.

Analysing in more detail the European market for on-road diesel, it can be observed that on-road diesel consumption is mainly attributable to commercial vehicles, including freight trucks, light trucks and buses. In 2009, they were responsible for 58% of diesel demand, whereas passenger cars represented 42% (see Figure 5). According to the reference scenario, the demand for on-road diesel should reach 4.1 Mbdoe by 2020 (see Figure 5).

In order to envisage another trend, we have gener-ated an alternative scenario called Gasoline Plus. In this scenario, by 2020, the share of diesel passenger cars sales would decrease to 30% instead of 54%, while gasoline passenger cars sales would increase to repre-sent 60%. This is a complete reversal of the current trend. Nevertheless, even in this case, demand for on-road diesel would be at best stabilised compared to 2009 and would reach 3.7 Mbdoe by 2020, mainly attributable to sales of commercial vehicles (see Figure 5). Demand for gasoline would by 2020 still be lower than the current level. As a consequence, even a drastic change in passenger cars sales will not be enough to

significantly rebalance the demand for on-road fuels before 2020 in Europe. The impact would begin to be significant only after 2025.

These scenarios have established the forecast demand picture and its limits. What will be the effect of the incorporation of biofuels and the reduction in refining capacity on the market balance around the Atlantic Basin? To consider these aspects, the 2020 reference scenario has been completed, with assumptions about the level of incorporation of biofuels and reduction in refining capacity.

Concerning biofuels, our assumption is that regulatory levels will be reached, but later than the expected sched-ule. By 2020, we estimate that biofuels content will reach 8% (energy basis) in Europe, well below the regulatory level, which has been fixed at 10%. For the US, our assumption is 9.5% instead of 10.7%; this corresponds to 30 billion gallons of renewable fuels by 2020, according to the RFS2 rule. These values should be compared with the current biofuels content of on-road fuels for the US and Europe, which are about 4% (0.44 Mbdoe, mainly ethanol) and 4.3% (0.25 Mbdoe) respectively.

The execution of functional safety assessment and validation must take place on every safety project and must be carried out by functional safety experts

CV60%

PC40%CV

55%

PC45%

CV58%

PC42%

Reference scenario

Gasoline Plus scenario

Figure 5 Breakdown of demand for on-road diesel

US EU-27

Mbdoe MbdoeBiofuels in on-road fuels (energy basis)

Figure 6 Gasoline imbalance simulations


In addition, our reference scenario incorporates a reduc-tion in refi ning capacity, both in Europe and in the US. This could reach 10% of existing capacity in Europe, in addition to the closures already carried out in 2009–2010. For the US, we anticipate a 5% reduction in capacity, a lower value than in Europe because US refi nery closures have already been implemented and projects are planned to process heavy bitumen blends.

2020 fuels balance around the Atlantic BasinFigure 6 shows the results of our simulations regarding the gasoline balance in 2020 and the impact of various parameters, such as the incorporation of biofuels at different levels and refi ning capacity reduction. The purple dotted line represents the current situation, a gasoline defi cit, of about 0.6 Mbdoe in the US and a surplus of around 0.75 Mbdoe in Europe.

In Figure 6, the lowest blue bar, Prod. stable, illus-trates what could be the change in the gasoline balance in 2020 if refi ning capacity remains stable, while the biofuels content increases according to our reference scenario: to 8% in the EU and to 9.5% in the US. It can be seen that, due to the trends in gasoline demand in both zones, the US gasoline defi cit could move into small surplus, while in Europe the surplus could be multiplied by two in 2020.

The next blue bar in Figure 6, Reference scenario, represents a situation if, contrary to the previous case, refi ning capacities were reduced by 10% in the EU and by 5% in the US. It should be noted that the gasoline surplus in the EU would be only slightly reduced, whereas the US could once more experience a gasoline defi cit. This is our reference scenario.

The third blue bar, Prod. highly reduced, shows what could be the impact of higher refi ning capacity reduc-tions (20% in the EU and 15% in the US). In this case, the US could come largely into defi cit, whereas the European gasoline surplus would be reduced but would remain at a higher level than today.

The last case in Figure 6, Biofuels policy levels, examines what could be the situation if regulatory biofuels content levels were reached. The combination of those levels with a 5% reduction in refi ning capacity in the US and 10% in Europe would suppress the US gasoline defi cit, although in the EU-27 the gasoline surplus would rise.

To summarise, the main observations of this analysis are that, in the EU, the gasoline excess is structural and will be infl ated by the incorporation of biofuels

US EU-27

Mbdoe MbdoeBiofuels in on-road fuels (energy basis)

Figure 7 Diesel imbalance simulations



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Uhde

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(ethanol) into the gasoline pool and by the decline in gasoline demand. Even sizeable cuts in refinery capac-ity will not be enough to eliminate the gasoline surplus.

The US market presents a better fit between supply and demand. Gasoline deficit will be mechanically flat-tened by a decrease in gasoline demand and biofuels (mainly ethanol) incorporation.

Figure 7 shows the results of our simulations concerning the diesel balance in 2020 and the impact of various parameters, such as the incorporation of biofu-els at different levels and refining capacity reduction. The purple dotted line represents the current situation, a diesel surplus in the US of around 0.4 Mbdoe and a deficit in Europe of about 0.5 Mbdoe.

In Figure 7, the lowest purple bar, Prod. stable, illus-trates what could be the change in diesel balance in 2020 if the refining capacity remains stable while taking into account increases in biofuels content according to our reference scenario, to 8% in the EU and to 9.5% in the US. This shows that, due to the respective trends in diesel demand in both zones, the US diesel surplus could move into a deficit of around 0.5 Mbdoe, while Europe would continue to lack diesel.

The next bar, Reference scenario, represents what the situation would be if, conversely, refining capacities were reduced by 10% in the EU and by 5% in the US. It is notable that the diesel deficit would deepen,

especially in the EU, where it could reach 0.8 Mbdoe. This is our reference scenario.

The third bar, Prod. highly reduced, shows what the impact could be of larger capacity reductions (20% in the EU and 15% in the US). Due to the structural mismatch between EU refinery throughput and demand, such a scenario would significantly increase the EU diesel deficit, which could reach 1.4 Mbdoe in 2020. Although a reduction in refining capacity would be beneficial to the EU gasoline balance, the same cannot be said for its effects on the diesel balance. The US diesel balance would also appear to have deterio-rated badly, with a deficit peaking at 1 Mbdoe.

The final case examines what the situation could be if regulatory levels for biofuels content were reached. The combination of those levels with a 5% capacity reduction in the US and a 10% reduction in the EU would help to reduce diesel deficits both in the EU and the US. But they would remain at high levels.

As a result, on-road diesel demand is expected to grow both in the EU and the US. Both zones will expe-rience on-road diesel deficits by 2020, but US refineries are more flexible; they have the ability to implement cut-point adjustments to adapt their production to these changes. The EU’s on-road diesel deficit is likely to be higher than that in the US, essentially due to an increase in demand from a situation today that is already in deficit.

Incorporating more biodiesel will contribute to a reduction in the deficit of on-road diesel, but will not be sufficient to eliminate it completely.

ConclusionThe prospect for the European market in the coming years is to remain imbalanced; the gap between supply and demand is structural and will not be impacted by new passenger car sales trends before 2020. The trends that will influence the European market’s evolution are: • Biofuels incorporation will worsen the gasoline surplus while reducing but not eliminating the diesel deficit• Capacity reduction cannot address both the gasoline surplus and the diesel deficit; the gasoline surplus will reduce as the diesel balance deteriorates.

The situation in the US is less strained because of a better fit between supply and demand. Furthermore, the more flexible US refineries will find it easier to adapt to changes in demand than their European coun-terparts. Probable trends to be monitored are:• The incorporation of biofuels and improvements in fuel efficiency will reduce the gasoline deficit to very low levels, except for the scenario where capacity is reduced by 15%• The US market will pass from a diesel surplus to a diesel deficit.

It is notable that whatever scenario is envisaged, mismatches in supply and demand will persist. Technology and catalysts will play a major role on both sides of the Atlantic in adapting refinery production to market demand. The technical solutions to address these challenges will be presented in a second article.




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fueling hiSfuture

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Q We are looking to switch from a high zeolite-to-matrix (Z/M) catalyst in our FCC unit to a lower Z/M formulation for processing resid. Which operating parameters will need special attention following the switch?

A Alan English, Senior Staff Consultant, KBC Advanced Technologies Inc, [email protected] time a catalyst formulation change is made, unit operation and yield should be monitored to assure the expected benefits are realised. Key variables such as gas yield, regenerator temperature, product yields and gasoline octanes should be plotted versus the percent-age of inventory changed out. The percentage of inventory changed out can be calculated from the fresh catalyst addition rate and unit inventory, assuming the inventory is well mixed (hence, losses and withdrawals will be a mix of old and new catalyst). The percentage of change-out can also be monitored by tracking a chemical component contained in different concentra-tions in the two catalysts. Si, Al and Re are often used for this if they differ sufficiently between the two cata-lysts. Other chemical species can be used but may require additional testing. Keep in mind that the magni-tude of most changes, such as yield selectivity shifts, will be observed in direct proportion to the percentage of change-out, but some effects, such as gasoline octane and LPG olefinicity, may not be observed until more than 80% of the bed has been changed. Of course, those parameters normally monitored, such as MAT, surface area, unit cell size, bulk density, APS and size distribu-tion, should also be plotted over time to assure no unexpected changes are occurring.

Changes specific to lowering the Z/M ratio will be a lower surface area, lower zeolite content, improved bottoms cracking selectivity, increased gas selectivity and increased regenerator temperature. The Z/M ratio should be optimised for the specific intended operation, since changing this ratio will have both beneficial and undesirable effects for resid cracking. For example, a higher matrix activity will increase the delta coke and regenerator temperature, but will also improve bottoms cracking selectivity. Changing the amount of matrix in the formulation may also affect metals tolerance. Increased vanadium tolerance with a lower nickel toler-ance is possible. Other changes to the formulation, such as the addition of metals traps, should be considered.

Of course, introducing resid will cause additional changes to the catalyst and unit as a whole. This will make monitoring catalyst effects difficult, since some changes caused by the resid will be larger than the effect the catalyst has on the same parameter. The refinery

P

may want to consider allowing the catalyst change-out to proceed to 50% or 70% before changing feed so the effects of each can be monitored independently. Catalyst metals levels, corrosion coupons and sour water quali-ties should also be closely monitored.

A Phillip Niccum, Director of FCC Technology, KBR, [email protected] lower Z/M surface area may improve the coke selectivity to help with coke burning constraints and heat balance when processing the residue. However, the catalytic sites of the matrix are more accessible to the large residue molecules characteristic of residue feedstocks. The dilemma with residue processing is that while low matrix surface area catalysts may have good coke selectivity, the bottoms cracking capability is compromised. The ideal situation when processing residue is to use a variable-duty, dense-phase catalyst cooler and perhaps partial CO combustion to control the heat balance so you have the freedom to choose catalysts with good bottoms upgrading ability when processing residue.

While we are generally referring to matrix surface area when discussing the Z/M ratio, it is recognised that the catalytic character of the matrix in terms of activity and selectivity is also a function of the matrix composition and manufacture.

With respect to operating parameters needing special attention, since less coke is likely to deposit on the matrix of a low Z/M catalyst, you may find coke deposits in the reactor will increase while processing residue using low matrix surface area catalysts. The use of closed reactor cyclone systems and steam purg-ing of the reactor void spaces will minimise reactor coke deposits.

A Carel Pouwels, Global FCC SpeScialist Resid, Albemarle, [email protected], we recommend monitoring the quality of your equilibrium catalyst and closely following the catalyst activity, surface area, micro-pore volume (MiPV) and mesoporous surface area (MSA). And since you are processing (more) resid, you should keep a close watch on contaminant metals, such as Ni, V, Na, Fe and Ca. These analyses help you to manage catalyst addition, which is an important operational parameter.

An important operating parameter to watch is the regenerator temperature, which is often limiting the unit operation. Of course, also monitor the catalyst cooler duty, as it has a strong influence on the regener-ator temperature.


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Another parameter important to monitor is the hydrocarbon-to-coke (H/C) ratio of the coke. A limita-tion in the stripper will lead to more entrained hydrocarbons and commonly a higher H/C, which will have a detrimental effect on the regenerator temperature. A lower Z/M ratio usually alters the cata-lyst architecture, which, in turn, affects the strippability of the catalyst. The stripping steam rate might need to be adjusted accordingly. Coke selectivity, strippability, metals tolerance and bottoms conversion power are thus the main concerns to be properly addressed by the catalyst design.

Albemarle’s low Z/M catalysts Upgrader and Upgrader R+ can process the heaviest residues, and their high accessibility leads to better strippability and thus a lower H/C ratio. Also, the balanced amount of different matrices and zeolite leads to an optimal balance in metals resistance, coke selectivity and bottoms cracking.

A Michel Melin, Director Technical Service EMEA, Grace Davison Refining Technologies, [email protected], moving to a lower Z/M ratio when processing more resid is not necessarily the right option. The reason is that a lower Z/M ratio will generally increase the unit delta coke and therefore the regenerator temperature. As the feed concarbon will go up, as well as the equilibrium catalyst metals content, there may not be much margin left for an increase in delta coke

from a catalyst reformulation. Grace Davison usually recommends selecting a low delta coke catalyst, allow-ing a high catalyst-to-oil ratio to crack the additional heavy compounds coming in with the resid. The oper-ating parameter that needs most attention when processing resid are the feed nozzles. In addition, the feed preheat and nozzle steam rates have to be adjusted accordingly.

A Stefano Riva, Regional Technical Service Manager EMEA, BASF Catalysts, [email protected] from a high to a lower Z/M catalyst can be done in different ways: increasing the matrix surface area alone, reducing the zeolite surface area alone, doing both, changing the matrix and zeolite technology and so on, so it really depends on how the lower Z/M ratio is achieved. Typically, a change in the Z/M ratio is accompanied by a change in operating conditions (ie, from gasoline to distillate modes) or feed quality (for instance, the need to enhance the bottoms cracking of heavier feeds when conversion is constrained), so all in all it depends on unit-specific objectives, constraints and the nature of the catalyst change. In general, more matrix enhances the bottoms cracking (ie, giving higher distillate to slurry ratios), sometimes at the expense (depending on the matrix technology) of coke selectiv-ity. Therefore, unless the unit severity is reduced at the same time, the regenerator may face a constraint, such as air availability or high temperature. BASF’s Prox-SMZ technology platform has commercially demonstrated the highest e-cat MSA in the market, in excess of 100 m2/g, and the lowest Na, below 0.1 wt%, to enhance catalyst stability, reduce hydrogen transfer and improve LCO quality, as well as enhanced coke selectivity compared to other high MSA technologies. A lower Z/M ratio has also been recognised to enhance product olefinicity at constant conversion, but if the conversion is reduced the overall LPG production may actually be lower and require the addition of ZSM-5 to maintain the downstream units, such as alky, MTBE, polymerisation, at full capacity. Depending on how much the zeolite surface area is reduced, if at all, a lower unit conversion at constant fresh catalyst make-up may also be observed, so unless this was specifically desired for distillate operation, particular attention should also be kept on the e-cat activity monitoring.

Q Our refinery wants to make the most of opportunities in the propylene market. What strategies are there for maximising propylene from the FCC unit?

A Alan English, Senior Staff Consultant, KBC Advanced Technologies Inc, [email protected] production can be increased on an existing FCC unit by raising the riser outlet temperature (ROT), adding ZSM-5 additive, reducing catalyst rare earth metals, allowing nickel content to increase and raising the feed rate. Of course, any of these changes will also increase wet gas production and have other effects on the unit. For instance, if ZSM-5 or ROT are increased,



inlet and steam production will give you an indication of the energy performance.

When the fouling factor starts to increase, you can think about online cleaning the loop. The slurry loop can be cleaned online in as little as 24–36 hours. Simultaneously, the main fractionator can be online cleaned together with the entire slurry loop.

As far as coke on catalyst is concerned, while a certain amount of coke will act as fuel to sustain the reaction, excess coke will reduce catalyst performance and yield. We have developed a new chemistry that will improve yield and reduce coke on catalyst.

The assumption is that you are not considering a modi-fication to your unit to improve its energy performance, but to optimise the existing unit. The actual variables to be monitored will depend upon your unit design and constraints. General areas to be monitored are:• One key focus area is to minimise the processing of cold feed from tankage and maximise hot feed from the upstream unit. While it may be necessary to process some amount of cold feed to keep the lines to and from the tank farm from setting up, the amount of cold feed should be minimised. Feed preheat from the main fractionator pumparound/product streams should be maximised to decrease the firing of the feed heater

• Often the single largest energy consumer in an FCC unit is the main air blower. Operating the converter at the lowest pressure consistent with maxi-mum superficial velocities, cyclone velocities and wet gas compressor constraints will reduce the energy consumption for the main air blower and typically improve yields of higher-value products. If the regen-erator operates in full combustion, the excess O2 should be reduced as low as possible, consistent with constraints on CO emissions and any afterburn issues. If the regenerator operates in partial combustion, the CO production should be maximised, consistent with the limitations of the CO boiler, coke on regenerated catalyst and unit heat balance. The excess O2 from the CO boiler should be minimised, consistent with the constraints on CO emissions. Stripping steam should be minimised while maintaining adequate removal of hydrocarbons from the spent catalyst. Atomising steam to the feed nozzles should be minimised, but be careful not to sacrifice yields in the pursuit of reducing utilities• Optimise pumparounds to provide maximum feed preheat and reboiler heating for the gas plant and minimise any heat that goes to air coolers or cooling water exchangers. Minimise steam to the side strippers, consistent with flash point specifications. Remember, even if low-value, low-pressure steam is used for stripping, it will generate sour water that must be treated• Control the wet gas compressor to eliminate/minimise any spillback. For pumps with minimum flow control, ensure the control is set prop-erly to eliminate any spillback when not necessary for pump protection. Minimise reflux ratios on towers, consistent with product specifications. Minimise heat input to the primary stripper, consistent with adequate stripping to minimise recycle from the stripper to the high-pressure separator. Minimise amine circulation, but ensure limits on rich amine loading are not exceeded. Minimise any off-spec product that will require reprocessing• In the winter months, control cooling water temperatures so as not to overcool reflux streams, pumparound returns, and so on. Ensure a good steam trap maintenance programme is followed. Optimise the usage of steam turbines/electric motors on pumps to properly balance the steam system.

When reviewing these different areas, always keep in mind the potential for maintenance or capital projects to further improve the energy performance of the unit.

Column performance can make a large contribution to energy costs in an FCC unit; distillation is an energy-intensive process. While there are many options to improve energy usage during the design phase, there are also several things that refiners can watch for and monitor on their operating units to ensure their columns are optimised.

For all towers in general, use of high-efficiency trays

MD, Coral MD and Amber MD are designed to achieve

High accessibility to maximise the diffusion of oil feed molecules to the active sites for maximum bottoms

The high accessibility enhances the diffusion of -

quently minimises the conversion of LCO molecules to

High matrix-to-zeolite ratio to minimise cracking of

Tailored zeolite activity and selectivity for optimum

Maximum resistance to the deleterious effect of feed contaminants like nitrogen, carbon residue and metals.

-ages our catalyst technology and provides extreme bottoms conversion power in an additive form. This is

-tions caused by sudden or short-term yield degradation due to opportunity feedstocks being sent to the FCCU.

Octane loss incurred across FCC gasoline hydrotreaters is primarily a result of olefins saturation. The goal of this operation is to achieve sufficient removal of

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the refinery might have to reduce its feed rate to control wet gas production, so a net increase in propyl-ene yield might not be observed. The refiner must also consider the difference between producing more propylene in the reactor and recovering more propyl-ene as a saleable product. Increases in the ROT or catalyst nickel, in particular, will increase the yield of non-condensable gases, which will lower propylene recovery in the absorber. The best strategy for a given unit will therefore be dependent on the particular constrains for that unit. One effective strategy often used on units constrained by wet gas capacity is to add ZSM-5 to increase the yield and olefinicity of LPG, while reducing the ROT to compensate for the increase in the wet gas rate. Lowering the ROT will also lower the C2- gas and LPG yield. Since ZSM-5 has little effect on C2 gases, the net result will be increased LPG and olefinicity at the same wet gas rate, resulting in more propylene production and perhaps improved recovery.

Before embarking on a propylene maximisation oper-ation, the refiner should also consider if their downstream equipment, especially their sulphur removal equipment, is capable of handling the addi-tional load. Several vendors are now offering mechanical designs targeted at increased propylene production. Any of these options would require a major unit revamp.

A Carel Pouwels, Global FCC Specialist Resid, Albemarle, Carel [email protected] downstream of the FCC unit often deter-mines how much LPG and C3= can be produced. Debottlenecking of this downstream equipment is required to participate in the market of high propylene producers.

Once the FCCU is viable for high propylene produc-tion, the operation is typically optimised through a maximum reactor temperature and high catalyst-to-oil ratio. One often has limited flexibility to increase steam partial pressure in the riser. A refiner can also consider the use of external streams with a high propensity for propylene such as naphthas with a high content of gasoline olefins.

The choice of catalyst, on the other hand, can make a substantial difference. Albemarle has a wealth of expe-rience with maximum propylene catalysts, ranging from very light feeds to the heaviest residues. Our AFX catalyst is a good choice for any grassroots maximum propylene FCC unit. One of its key features is the high accessibility of the catalyst particles, which enables cracking products to leave the catalyst quickly and reduces any unwanted secondary cracking such as hydrogen transfer. Hydrogen transfer reactions are one the most important competing reactions, which should be minimised for maximum propylene. Also, in these maximum propylene applications, the FCC unit often operates at or near the so-called propylene plateau, where the addition of any additional ZSM-5 additive does not yield any extra propylene. The use of AFX increases the propylene plateau and maximises the propylene level.

When installing a grassroots FCC unit for maximum propylene, an array of technology options is available. Several processes are offered, whereby the common variables are: • High severity through a high reactor temperature and high catalyst-to-oil ratio • Low hydrocarbon pressure • Optionally, one can install a dedicated riser to crack streams rich in olefins to further enhance propylene production.

A Romain Roux, Deputy Product Line Manager Catalytic & Thermal Cracking, Axens Technology Department, [email protected] The demand for higher propylene is driven in the FCC unit by the differential of price between gasoline and propylene. To be economical, the production of addi-tional propylene has to minimise the production of fuel gas. Indeed, the cost of producing propylene from the FCC unit is additional wet gas compressor capacity and propylene/propane splitter capacity. The propyl-ene/propane capacity is mainly driven by the propylene yield, and the propylene/propane ratio has a limited impact. The wet gas compressor capacity is driven by the yields of dry gas, LPG and C5. As a consequence, to increase the propylene yield inside the wet gas compressor, you have to either reduce the dry gas production in the FCC riser or the C5 yield by increasing cooling upstream of the wet gas compressor. The first option is generally used.

The most effective way to increase propylene produc-tion while minimising dry gas is to use ZSM-5-based additives. ZSM-5 is a zeolite widely used for this purpose and provides gasoline cracking towards LPG for almost no additional dry gas production.

Another classical way is to increase the riser outlet temperature. This solution significantly increases the dry gas yield. So to reduce the additional dry gas, refiners usually install a riser separator at the riser outlet; for instance, Axens/Shaw RS. This technology usually reduces the dry gas yield by 15–30%. Axens has recently licensed a new technology called FlexEne, enabling selective cracking of the C4 and C5 olefins produced by the FCC unit into propylene. Instead of directly recycling these cuts to the high-severity riser — for instance, our PetroRiser — we have improved our technology by converting the C4 and C5 olefins into C8 to C10 olefins through an oligomerisation unit. These longer olefins are recycled to the existing riser and crack selectively toward propylene. The selectiv-ity factor of additional propylene (mol)/additional dry gas (mol) is three times higher with FlexEne than with the direct recycle of C4 or C5 cuts in a separate riser. This new technology also offers the advantage of flexibility. As the propylene price is fluctuating, FCC users are willing to switch from maximum gaso-line or LCO to high propylene quickly. FlexEne scheme enables instant valourising of the produced oligomer into gasoline or even kerosene almost instantly. Flexibility of operation and selectivity is the target of FlexEne.


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A Michel Melin, Director Technical Service EMEA, Grace Davison Refining Technologies, [email protected] a refinery wants to simply increase its propylene yield in the FCC unit, the most common and conve-nient method is to use a ZSM-5 additive — a shape-selective zeolite from the Pentasil family — which will crack gasoline-range olefins into propylene and other light olefins. However, a tailored integral catalyst system will provide even higher propylene yields. This is because, in addition to containing Pentasil, an integral catalyst system is designed to minimise the H-transfer reactions of gasoline-range olefins, which compete with the cracking reactions that produce propylene. Increasing operational severity either by using a higher riser outlet temperature or higher catalyst-to-oil ratio will also increase your propylene yield. In addition, processing feeds with higher paraffinicity will result in an increased propyl-ene yield due to the higher conversion and reduced H-transfer associated with the paraffinic feed.

A Stuart Foskett, Regional Technical Service Manager Asia/Pacific, BASF Catalysts, [email protected] There are many publications that address FCC propyl-ene maximisation; for example, J McLean and G Smith, Maximizing propylene in the FCC unit — beyond conven-tional ZSM-5 additives, NPRA Annual Meeting AM-05-61. A short summary is given below.

Most FCC units can make 1% or 2% higher propylene

yield simply by using ZSM-5 additive. There will be many constraints in the FCC gas plant that need to be considered, especially wet gas compressor capacity, and loadings in the primary absorber, de-ethaniser and debutaniser columns, to name but a few. To truly maxi-mise propylene to around 15–20 vol%, the propylene yield requires a more specialised catalyst solution. First, the activity dilution effect of high levels of ZSM-5 additive usage must be overcome. BASF has addressed this by developing a high-activity catalyst featuring built-in ZSM-5 functionality, Maximum Propylene Solution (MPS). MPS is currently in use in several refineries around the world processing a variety of feedstocks. Separate addition of ZSM-5 additive may still be used with the catalyst having built-in ZSM-5 functionality, but typically a level of 5–10% additive is the most that is required.

The second important aspect to recognise for maxi-mum propylene mode is that the base catalyst properties define the maximum achievable propylene yield, not the amount of ZSM-5 crystal in the unit. With increasing concentrations of ZSM-5, at a certain point the gasoline olefin feedstock for ZSM-5 cracking is depleted, and a maximum plateau propylene yield is reached. Changes to the base catalyst properties, espe-cially to minimise hydrogen transfer reactions, can preserve more olefins in the gasoline range and signifi-cantly raise this ceiling for propylene yield. For this reason, maximum propylene catalysts typically have low rare earth. Higher fresh surface area is needed to compensate for the lower rare earth to maintain satis-factory activity and catalyst addition rates. In this regard, catalyst technologies offering the highest fresh surface area have a distinct advantage of achieving a more optimised balance of activity (catalyst consump-tion) and propylene yield.

Process conditions also significantly affect propylene yield, in particular unit pressure. As hydrocarbon partial pressure increases, the kinetics of hydrogen transfer reactions is more favoured over ZSM-5 cracking, leading to less preservation of gasoline olefins for cracking to propylene. Increasing the reactor pressure from 30 psi to 45 psi can reduce the propylene yield by as much as 4 vol%. Increasing riser steam will reduce the hydrocarbon partial pressure and should be considered for revamps to maximum propylene operation. A higher riser outlet temperature also increases the propylene yield and in most units is generally limited to around 1005°F (541°C) to keep the dry gas yield within an economically acceptable level.

A Jill Brown Burns, Senior Refinery Design and Applications Engineer, Sulzer Chemtech USA, [email protected] adjusting reactor or catalyst conditions in the FCC unit can make a large impact on propylene yields, a large amount of that valuable C3= product can end up as fuel gas without proper optimisation of the GasCon columns. Optimisation begins with under-standing how the absorber/stripper columns operate and how propylene is recovered.


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The absorber and stripper (de-ethaniser) columns are the primary vessels in the FCC GasCon for propylene recovery. The downgrade of propylene from product/alky feed to fuel gas can be very significant. FCC unit engineers should focus on optimising several key areas of the absorber/stripper operation for maximum propylene recovery as follows: stripping, water removal and intercooler usage.

One common problem in this area of the FCC unit is overstripping. Often conservative targets are set to minimise H2S in the debutaniser overhead; this results in recycling LPG back up the column, increasing the amount of absorption required by the lean oil. If the column operation cannot respond with increased lean oil or increased cooling, this may result in propylene product lost to the offgas. The overstripping effect may become even more pronounced in units that operate separate absorber and stripper columns, with the high-pressure separator operating as the receiver. Overstripped gases are recycled back to the high- pressure coolers upstream of the separator. If they condense, they are re-fed to the stripper. If they do not condense, they are fed to the absorber, absorbed and end up right back in the high-pressure separator. The recycled traffic consumes hydraulic capacity in the column and cooling capacity in the condensers. Overstripping can be a hugely wasteful loop from both energy efficiency and hydraulics perspectives.

Another common problem in the absorbing/strip-ping section is poor water removal. Any water present, whether from free water carryover or water precipitat-ing from solution, will create excess traffic in the top of the stripping section through the middle of the absorb-ing section. Water travels down the column to the point where it is hot enough to vapourise, usually around the middle of the stripping section. The water vapour then travels back up the column to the point where it is cool enough to condense. This loop takes up valuable tray capacity. Water can also initiate a Ross-type foaming just before it begins to condense or drop out of solution and two distinct liquid phases become present. Whether due to hydraulics or foam-ing, the presence of water can impact tray efficiencies and contact time, adversely affecting the absorption of propylene from the offgas stream.

Finally, the absorption process generates heat, and many columns use side intercoolers or pumparounds to limit the column temperature rise, which improves recovery of LPG from the offgas. While refiners under-stand the effects of pumparounds in a fractionator, often the importance of their role in the absorber is not as well understood and the systems are often neglected. Monitoring fouling in the intercooler exchanger (usually exchanged with cooling water, a notoriously problematic system for many refineries) and maintaining cooling water flow rates are a few on-stream optimisations the FCC engineer can look at. From a design perspective, draws, returns and transi-tion zones can be tricky and warrant careful study during the design phases.

Q Just how much energy can I expect to save by optimis-ing plant metrics such as pumparound rates and reflux ratios rather than committing to large-scale capex?

A Allan Rudman, Regional Practice Coordinator EMEA, KBC Process Technology Ltd, [email protected] KBC’s experience, it is possible to achieve energy savings between 2% and 5% by optimising plant metrics. However, due to the strong link between energy and yield, this type of optimisation typically opens up the potential to improve yield and/or throughput and invariably, due to the value of yield improvement compared to energy saving, the former tends to prevail.

For example, a European refinery with a 400 t/h CDU was achieving a furnace inlet temperature of 242°C. The CDU was reviewed during an energy assessment of the site as part of a recent KBC energy review. KBC observed that the naphtha/kero gap was large and recommended a reduction by 5°C. Also, 84% of the total heat was removed by an overhead condenser and top pumpar-ound, so it was recommended to move the heat down the column to improve preheat train performance.

As a result, KBC modelled the column and preheat train in Petro-SIM, increased the medium pumparound and bottom pumparound heat removal up to the pump’s limit, and the furnace inlet increased by 10°C, saving 2.2 Gcal/h worth €0.5 million/y. The metrics to be monitored are the naphtha/kero gap and MPA/BPA flow rate.

Q We would like to integrate biomass conversion (for second-generation biofuels production) in our refinery, but is a large-scale, reliable supply of biomass a realistic proposition?

A Rick Manner, Principal Consultant, KBC Advanced Technologies Inc, [email protected] supply of biomass for a second-generation plant is usually not feasible. It is typically very difficult to grow, harvest and transport more than 1000–2000 b/d worth of feedstock for any biofuels process. In special circumstances, such as a refinery in the middle of a rainforest, it may be possible to gather as much as 5000 b/d worth of feedstock.

Second-generation process implies the feedstock should be a non-food crop. Relieving this limitation by accepting vegetable and waste oils from the food processing industry may allow a larger plant, perhaps 5000 b/d or more, at some locations. Some biofuels producers would consider this large scale, but most refineries probably would not.

We do not believe any large-scale second-generation biofuels processes currently exist. Most biological processes such as cellulosic ethanol are expensive and have not yet been demonstrated on a large scale. Fischer-Tropsch processes have been demonstrated commercially, but the investment required is very high.

There are some projects in the conceptual stage for biomass collection in Europe based on gathering and generating the intermediate oils on a distributed basis


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and then refining them on a central basis. Whether the technology currently exists to make this feasible without displacing food crops, or is in development, is to be seen.

A Edgar Steenwinkel, Global Business Director Alternative Fuel Technologies, Albemarle, [email protected] are several ways of looking at this. In the first place, one must consider that extensive transportation of biomass to a processing facility will be at the cost of the benefits of using renewable feedstocks. Therefore, it is suggested to consider locally produced biomass as favourable for intake into a facility.

A second option is to somehow densify the energy locally, near where the biomass is available, and then transport this intermediate to a refinery. Several options are available for this, such as thermal flash pyrolysis or catalytic pyrolysis. These processes are being developed and should be relatively flexible in the type of biomass they take in. They would “yield” a green crude that needs subsequent processing in a refinery. Note that these green crudes still contain oxygen as a chemically bound impurity. Albemarle is looking at all these options to retain the carbon effi-ciency of these processes, while still resulting in manageable intermediates and products.

Q If we want to switch frequently between high and low delta coke operation in the FCC unit, do we have to change

the catalyst every time, or will a multipurpose catalyst give an acceptable balance of operational performance without changeovers?

A Alan English, Senior Staff Consultant, KBC Advanced Technologies Inc, [email protected], we assume the questioner intends to alternate between operations with high coke-producing feed, such as contaminated gas oils, coker gas oils or resid, and low coke-producing feeds, such as clean gas oils. Typically, the catalyst formulation would be optimised for a specific type of feed, but occasionally opportunity feeds may be available for a relatively short period of time. Since a catalyst change-out typically requires three or more months to accomplish, the refinery may need to process these feeds with a less than ideal cata-lyst formulation.

As with all FCC questions, the best answer will vary with the specific limits and capabilities of the unit considered. Choosing a middle-of-the-road catalyst for all anticipated feeds will provide some degree of flexibil-ity, but limit the extreme in either direction. A unit with flexibility in one particular direction may therefore want to optimise the catalyst in the opposite direction. For example, a unit with fired feed preheating capacity might be better served by selecting a low delta coke catalyst and low feed preheat to enhance resid cracking. It could then use the same low delta coke catalyst and high preheat temperature during periods with clean gas oil. On the other hand, a unit with catalyst coolers could tolerate a high delta coke catalyst even with some resid, while assuring the unit can heat balance with clean gas oils by shutting down the catalyst coolers. Of course, in any of these scenarios, once the unit has reached the limits of its flexibility to process opportunity feeds, the refiner will need to consider if the opportunity exists for sufficient time to justify a catalyst change, which would allow further optimisation.

If frequent feed changes are inevitable, a better alternative might be to speed up the transition from high to low delta coke catalyst. Several vendors offer automated catalyst addition devices, which are inte-grated with fairly large hoppers. This offers the refiner the opportunity to economically manage multiple cata-lyst additives. In this case, the refiner could select a low delta coke catalyst as the primary catalyst, but have a separate hopper and additions system for a high delta coke additive. During periods of low resid feed, the high delta coke additive would be injected without the ship-ping and inventory delays inherent with a change of the primary catalyst. Once a resid cracking opportunity again occurs, the additive injection would be terminated and the amount of resid in the feed ramped up as the inventory returns to its base delta coke behaviour.

A Phillip Niccum, Director of FCC Technology, KBR, [email protected] you plan to frequently switch between high delta coke and low delta coke operations, switching catalysts will be impractical due to the time required to change the composition of the catalyst inventory. This is



evaluation service, which identifies the optimum coat-ing solution for delayed coker heater fouling.

The key factor in mitigating fouling is having a good fired heater design. Some of the critical design factors include heat flux, cold oil velocity, tube spacing and external surface-to-volume ratio. KBR’s long experience in fired heater design and in handling difficult coker feedstocks insures the design will maximise the run length for any given feedstock. The next most important factor is good desalting of the crude oil upstream of the coker. Salts and solids left in the crude end up in the coker and promote fouling of the heater tubes.

With a good fired heater design and proper desalting of the coker feed, there is little else that can be done. Several chemical vendors offer antifoulants that have shown success in reducing fouling. Whether they are cost effective or not is pretty much a case-by-case trial.

The degradation of refractory in catalyst transfer lines can be due to a number of mechanisms depending on the design and service.

Today, most catalyst transfer lines in FCC units consist of carbon steel shells containing a vibration cast, erosion-resistant refractory lining. Assuming the refractory materials are chosen and installed properly, and the design and installation of the support anchors are adequate, the more prevalent mechanism of degra-dation occurs in reactor service such as the feed riser, where coke impregnation can cause surface spalling of thin flakes of refractory that in time accumulate to substantially thin the lining. On the regenerator side of the unit, the coking mechanism is absent, so the refrac-tory tends to provide longer life.

If the velocity of the catalyst and vapours is above the recommended design velocity, erosion can also be an issue. Catalyst transfer lines that carry catalyst at high velocity include the feed riser and sometimes spent catalyst transfer lines that carry catalyst in a dilute phase back into an elevated regenerator. Generally, erosion is proportional to the velocity raised to at least the third power, so designing for a reason-able velocity can mitigate the erosion in these lines.

Standpipes are the least likely catalyst transfer lines to suffer from refractory damage because the velocity is low and generally the lines are of smaller diameter; the smaller diameter increases the key-holing effect that can help to hold together even fractured linings with broken anchors.

An older technology, dual-layer refractory, utilises a dense abrasion-resistant refractory on the inner surface, backed by a thicker layer of lower-density, better-insulating refractory. These lines can suffer from a destructive mechanism not seen in the monolithic vibration cast linings. In these dual-layer systems, the abrasion-resistant refractory is typically held in place by a stainless steel mesh commonly referred to as hex-metal. Since the hex-metal runs at a very high temperature relative to the carbon steel shell, it can fail from the effect of its thermal expansion, causing it to buckle against the restraining force of the colder outer shell and insulating lining. In recent decades, the FCC industry has moved away from the use of dual-layer linings, but some are still in service.

Finally, in reactor riser service, hex-metal welded directly to a hot shell can be wrenched loose over time by coke growing between the hot shell and the hex-metal/refractory system. This mechanism is exacerbated when the hex-steel is placed on the outside of a line or on a flat plate, as might be done to protect an internal riser from an external source of potential erosion. In these cases, the aforementioned key-holing effect is effectively absent, allowing the refractory to be more easily forced from the surface by the coke.

In summary, to mitigate problems seen in the reli-ability of refractory linings in catalyst transfer lines, it is important to identify the mechanism of degradation and also to check the quality of the design, materials and installation. Once a mechanism is identified as the likely cause of the degradation, the solutions will become apparent. For instance, if the problem is high velocity, install a larger transfer line or reduce the vapour rates. If the problem is coking, decide if the

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especially true for units with large catalyst inventories. A more expedient way to effect the change in the cata-lysts coke-making propensity is to simply change the fresh catalyst addition rate as needed to adjust the activity. If the regenerator needs to heat balance at a higher temperature, you can increase activity and, if you want to reduce the regenerator temperature, you can reduce activity. For this, the middle-of-the-road catalyst may be best.

The delta coke of the FCC operation can also be adjusted by changing the feed dispersion and stripping steam rates. For example, while processing a severely hydroprocessed VGO feedstock, you may need to increase the catalyst microactivity to 75% or more, while running the feed and stripping steam rates at the minimums recommended by the FCC licensor. If you then switch to a higher carbon residue content gas oil or good-quality atmospheric residue, you may need to lower the activity into the low 60s and increase the dispersion and stripping steam rates to their recom-mended maximums just to prevent the regenerator from overheating.

It may also be useful to consider keeping a smaller inventory of very high-activity, high matrix-activity catalyst, which would be used as an additive when there is a need to heat the regenerator. In this scenario, the base catalyst would more optimally be one with a minimal coke-making character.

A Bob Skocpol, Senior Technical Consultant FCC Catalysts, Albemarle, [email protected] assume when you say high and low delta coke operation that you mean large changes in FCC feed quality. For example, we have customers changing between significantly more or significantly less resid in their FCC feed. The best answer probably depends on how frequently the changes happen and how expected or unexpected they are.

For slower and more predictable feed changes, such as a heavier feed in the winter, you could deliver and exchange the catalyst in time to meet the seasonal needs. For example, on the heavier feed, you might apply a coke-selective catalyst to increase conversion and reduce coke. But for more rapid changes, you need other approaches. For fast feed changes, we recom-mend the use of our Bottoms Cracking and Metals Tolerant (BCMT) additive. BCMT enhances bottoms conversion, metals tolerance and catalyst accessibility when more refractory and contaminated (residual) feedstocks are routed to the FCC unit.

For rapid and unpredictable changes in feed quality, there are some functionalities that can be built into the catalyst, which help with the bad feeds and do no harm with the good feeds. Stable and metals-tolerant zeolites, nickel passivation matrices and vanadium-tolerant technology are available from Albemarle to cope with sudden shifts to poor feed quality. Also, the use of Albemarle’s high accessibility technology within our catalyst families provides yet another means to be prepared catalytically when low feed quality is encountered.

In Albemarle’s experience, each FCC unit has different goals, limits and issues, so each solution needs to be customised. In many cases, we have developed, together with the customer, operating strategies to handle shifts between high and low delta coke operation. An example is a customer with a very clean hydrotreated FCC feed who adds vacuum resid when the economics and avail-ability are positive. The opportunities often appear with little or no warning. In one particular case, our analysis showed that the spike in metals led to longer-term loss of catalyst activity, which then hurt the operation on the better feed for a while. By comparing what is practical for the refinery with our dynamic activity models, we developed a strategy of adding a certain percentage of extra tonnes/day of catalyst during the resid operation and continuing for the proper number of days after-wards. This effectively eliminated the negative effect of the metals on the catalyst with the better feed.

A Romain Roux, Deputy Product Line Manager Catalytic & Thermal Cracking, Axens Technology Department, [email protected] Changing the catalyst is certainly a good option. Nevertheless, it may not fit with the production plan. Here, the best solution is to control the delta coke separately. Recycle of HCO and slurry in the riser is one solution if the heat balance is in deficit. Recycling heavy naphtha with Axens/Shaw mixed temperature control (MTC) injectors is another solution if the heat balance is in excess. Finally, torch oil is a practical solu-tion for feed with a very low vanadium content.

A Michel Melin, Director Technical Service EMEA, Grace Davison Refining Technologies, [email protected] is unlikely that the catalyst formulation can be adjusted frequently enough to follow a change in unit delta coke. It is advisable rather to adjust the e-cat activ-ity by adding more or less catalyst. A move to a lower delta coke operation will require an increase in e-cat activity. In addition, the possibility to add a (cheaper) coke precursor feed when required is usually a good option. Adding a bottoms-cracking additive when being in a low delta coke mode is not recommended, because building a concentration of this additive in the unit over a few weeks will lead to a drop in e-cat activity, with a subsequent drop in unit conversion.

A Stefano Riva, Regional Technical Service Manager EMEA, BASF Catalysts, [email protected] the FCC unit allows it, the easiest, fastest and most profitable way to change delta coke is to change the catalyst addition rate and therefore the e-cat activity. If this is not possible, you should define the predominant operating mode, the reason for the desired delta coke change and what “frequently” really means. For exam-ple, is the feed source changing from hydrotreated to non-hydrotreated? The frequency of the change should be compared to the unit-specific combined catalyst changeover time. The term “combined” means the sum of all the times required to process the change: the supplier’s manufacturing time, the delivery time, the


Q&A copy 6.indd 7 10/3/11 12:38:41

fresh catalyst hopper inventory replacement and, finally, the unit-specific time needed to reach at least 70–80% changeover. BASF’s manufacturing capability enables delivery of a new formulation almost on a shipment-by-shipment basis, so the first piece of the timeline is significantly reduced. However, if the overall time is still too much compared to the frequency of the change, the base case catalyst should be optimised to match the most profitable yearly scenario, taking into consider-ation the duration of each mode, and the benefit and limitations encountered when running each mode with a catalyst not specifically designed for that mode. Once the catalyst has been optimised, the BASF Co-Catalyst solution should also be considered, as it offers a fast way to follow the rapidly changing FCC objectives. Co-Catalysts combine all the positive features of additives in terms of flexibility, together with the ability to affect the FCC selectivities.

Q We are getting close to our wastewater discharge limit. How can we increase our water recycling rate?

A Gert-Jan Fien, Senior Consultant, KBC Process Technology, [email protected] recycling in refineries is a complex issue. Without knowing the local situation regarding current effluent characteristics and constraints, only very generic strategies can be proposed. KBC suggests a four-step approach:• Reduce water flows at the point of use. This should have a notable impact on the total wastewater flow. Obviously, there are many ways to affect such reduc-tions and, as a start, a comprehensive water balance (inside and outside battery limits) should be constructed. Software tools are available to assist with constructing balances from the typically patchy water data available. Actual consumptions should then be compared to realistic targets for gap closure• Reduce water contamination. This will offload the wastewater treatment plant (WWTP) and give it a better chance to produce recyclable water; for example, as cooling water make-up. Again, this requires a close look at many types of operation, to judge if and how water contamination, with oil, caustic, amines and ammonia, can be minimised• Examine the routings of water and wastewater streams. Can some effluent streams be reused without having to go to the WWTP? Can streams bypass parts of the WWTP; for instance, stop sending non-oily water through the API?• Optimise the WWTP, to improve the quality of (part of) the effluent. From the water balance and reuse targets (determined in step 1), it can be judged how much treated wastewater of sufficient quality could real-istically be recycled. The WWTP should then be improved and perhaps expanded to generate this water.

Over recent years, KBC has put much effort into the tools and methodologies needed for this approach and achieved good results. Potential water savings of 25–40% are regularly identified in this manner.

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Optimising safety relief and flare systems

Over the past decades, the refining industry has contin-uously moved towards

higher levels of crude conversion and more stringent product specifi-cations. When they are combined with aromatics production in partic-ular, refinery schemes have become more complex. Additionally, name-plate capacity limits for grassroots plants have been gradually pushed upwards.

At the same time, along with dieselisation of the European vehi-cle fleet and the strengthening of product specifications, the refining scheme of European refineries has become increasingly complex. The addition of new units to an existing refining scheme affects flare

Understanding the behaviour of refinery units by dynamic modelling of emergencies enables the prediction of realistic relief loads

AlbAn Sirven, JUlien GrOSclAUde and GUillAUme FenOl Technip FranceJeremy SAAdA Invensys Operations Management

systems, leading to a revamp of the existing flare network or to the addition of a new flare network and the consequences of altering the network.

Flare systems are primarily sized with regard to common failure modes. General electrical power fail-ure (GEPF) results in a simultaneous loss of condensation for most proc-ess systems, and all corresponding individual relief loads are summed up to determine the required capac-ity of the flare systems.

In the context of both an increased number of interconnected process units and higher processing capaci-ties, flare systems approach critical sizes when industry-standard calcu-lation methods for the determination

of individual relief loads are applied. To overcome the related issues in terms of mechanical and structural design, supply and constructability, as well as to satisfy the requirements of refinery turna-round and scheduled maintenance, the configuration of the relief disposal system for grassroots designs should consist of several flare systems, the largest with a main header diameter as wide as 100in or more and a flare stack as high as 200m.

The cost of providing adequate protection systems for any refining complex is substantial. At this point, understanding and model-ling the dynamic behaviour of refinery units in emergency


Semi-dynamic approachConventional approach Dynamic approach

Figure 1 Comparison of three methods for sizing flare systems

invensys.indd 1 10/3/11 12:52:15

situations becomes necessary to assess the required capacity of flare systems more accurately than conventional calculations methods, which tend to produce conservative results.

Modelling refinery equipment behaviour in emergency conditionsThe sizing of refinery flare systems requires prediction of the behaviour


of equipment (or refinery subsys-tems). Calculation methods used for the determination of relief loads have, since the beginning of the refining industry, required static and semi-dynamic calculations.

Conventional approachThe conventional approach is based on the use of process data shown in the unit heat and mass balance,

corresponding to steady-state oper-ating conditions.

Semi-dynamic methodsSemi-dynamic calculations require the enhanced use of static modelling tools such as SimSci-Esscor Pro/II to better model upset scenarios. This method will correct the results of static methods, accounting for basic equipment design data or the change of key fluid properties between operating and relief conditions. This type of analysis also aims to evalu-ate the relief start/end time for critical equipment and systems. The result is less conservative relief loads than with static calculations, but the analysis cannot account for complex phenomena related to the transient responses of process systems to major upsets.

Dynamic methodsWith the increasing performance of calculation tools, dynamic process simulators are now able to work on desk computers with reasonable computation time. Process dynamic simulations are based on validated thermodynamic models as well as dynamic heat and material balance equations. They also consider the response times of control loops, thereby predicting more realistic relief scenarios. Figure 1 shows a comparison of the accumulated relief loads used for sizing a flare system, obtained by these different methods.

In addition to the reduction in relief loads, dynamic simulation offers the advantage of evaluating the impact of different system response times on accumulated relief loads.

However, dynamic modelling is time consuming compared to earlier calculation methods. This is due in particular to the amount and diver-sity of input data that must be defined for the description of any system; for instance: • Definition of process streams • Operating conditions based whenever applicable on licensor information• Detailed equipment design and geometry data based on vendor information• Instrumentation, automation and safeguarding data.

Qualitativeanalysis

Conventionalapproach

Review of the system

characteristics

Implementation of flare loads mitigation measures

Screening and identificationof large relief

loads

Semi-dynamicapproach for these large relief loads

Dynamicsimulation

of the system. New relief loads

Quantitativeanalysis

Does this system match

Technip criteria for dynamic simulation?

Sizing of flare systems

Dynamicsimulation

Figure 2 Dynamic simulation within flare study methodology

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Technip developed a methodol-ogy for targeting opportunities for dynamic simulations and set up a list of criteria to select candidate systems for the use of dynamic simulation. This methodology for a refinery flare study (GEPF scenario) is detailed in Figure 2.

Dynamic simulation softwareComputer simulations were made using the SimSci-Esscor Dynsim tool from Invensys as the dynamic process simulator. The program provides a series of capabilities that enable the modelling of rigorous transient processes and facilitate the development of dynamic simula-tions for applications from process studies through to operator training systems.

The equipment models are rigor-ous dynamic models. A library of unit operations, equipment types, control functions and other algo-rithms have been developed, enabling the specification of models for high-level concepts.

In dynamic relief load calcula-tions, the user can often come across unexpected operating condi-tions and effects. The provision of accurate thermodynamic predic-tions enables these effects to be modelled correctly. Ranges of rigor-ous thermodynamic methods enable the user to predict phase behaviour and so achieve accurate simulations of real processes.

Other features that make the soft-ware suitable for relief load analysis are the interactive run control and scenario handling options. These include the ability to change from one operating condition to another quickly and easily, speed up or slow down the execution of the simula-tion, and record/play back operating scenarios. These features enable the user to investigate how relief loads are impacted by different emergency situations, initial conditions and operating procedures. The software also includes a basic data historian and options for trending and storing process variables so that relief load scenarios can be viewed and recorded for future analysis.

Dynamic simulations were built using the software. One of the tech-nical challenges of applying process


dynamic simulations lies in the capacity of having a system descrip-tion capable of supporting the generation of a realistic transient response of process equipment and systems during upset and emer-gency scenarios.

Throughout the studies, several of the default modules available with the software were analysed not to have sufficient descriptive capabilities and functionalities so it was mandatory to incorporate user models. For example, the default furnace module has been improved by providing a reliable furnace model for the simulation of the furnace behaviour under process upsets.

Dynamic simulation resultsDynamic simulation of furnaceshutdownThe vapours generated in column reboiling furnaces are major contributors to flare relief loads if the column condensation is stopped (single condensation failure or common mode of failure such as GEPF). In the specific scenario of GEPF, the feed to the furnace is stopped if the reboiler feed pump is electric motor-driven or may be continuously pumped in the case of a steam turbine driver. Even when furnace firing is shut down (the normal scenario for fired heaters equipped with a forced-draft fan), heat transfer and vapour generation continue as a result of thermal iner-tia of the radiant box walls.

For these reasons, a reliable

representation of the furnace is necessary. The software’s default furnace module enables simulation of the furnace’s thermal inertia. However, this module is unable to model the response of the liquid inventory in furnace coils during upsets. Thus, Technip has developed in-house modelling modules and techniques for furnaces, based on Dynsim’s default furnace module, to have a proper and realistic represen-tation of the key parameters affecting the determination of relief rates.

There are no standard criteria regarding the residual heat duty transferred to the process in the case of trip of process feed and/or furnace firing. The industry consid-

ers values as high as 100% of normal duty if heater firing is maintained. This results in substantially conserv-ative relief flow rates to the flare.

The case study in this context evaluates the behaviour (residual duty transferred to process) of a reformate splitter furnace in three different configurations:• Charge stopped and firing maintained• Charge and firing stopped• Charge maintained and firing stopped.

Dynamic simulations have demon-strated that, if the process feed to the furnace is stopped, the residual duty transferred to the fluid drops immediately (see Figure 3). It then does not exceed a peak value of 40% of normal duty. Averaged over a few minutes, it does not exceed 30%

60

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Charge stopped, firing stopped

Charge maintained, firing stoppedCharge stopped, firing maintained

Figure 3 Furnace behaviour in three shutdown cases

invensys.indd 3 10/3/11 12:52:38

behaviour of crude distillation units (CDU) after process upsets, as well as the importance of the appropri-ate configuration of pressure safety valves (PSV) to reduce relief flow rates.

Actual PSV relief flow rates depend mainly on two factors:• Dynamic behaviour of the equip-ment itself • PSV configuration (number, size, set pressure stepping).

The case study in this context considers a crude distillation unit of 250 000 b/d capacity and compares the benefits of optimisa-tion using dynamic simulation against a conventional approach.

Results of a conventional approach The sizing of the relief installed capacity (orifice area) by means of a

conventional approach using heat and material balances leads to the following results:• Reflux failure: 860 t/h (sizing case)• GEPF: 200 t/h• PSV configuration: 12 PSV/T orifice/balanced bellows/set pres-sure = 3.5 barg• Installed orifice area: 2020 cm2.

Optimisation of installed orificearea using dynamic simulationThe conventional approach gives an over-sized installed orifice area. Dynamic simulation of the system under upset may help in optimising the required installed orifice area.

The installed orifice area has been optimised down to 930 cm2, corre-sponding to 9 PSV (Orifice R). Figure 4 shows that the relief flow rate in the event of reflux failure would be stable at around 414 t/h, compared to the 860 t/h initially envisaged. With regard to these two cases, dynamic simulation shows a reduction of relief loads of about 50%. It is also noticeable that the number of PSVs can be reduced, leading to a reduction in capital and maintenance costs.

Staggering of safety valve set pressuresSafety valve staggering is most of the time neglected in conventional calculations, but dynamic simula-tion shows that it is a key parameter in a correct estimation of flow rates. Even if the installed orifice area is reduced, as shown previously, dynamic simulation shows chatter-ing conditions for non-sizing scenarios such as GEPF.

Figure 5 shows this chattering situation in the event of GEPF, where peak relief flow rates reach about 400 t/hr, the average relief load being about 110 t/h. The stag-gering of the PSV avoids chattering situations and cancels high peak flows. Table 1 shows the selected configuration of the PSV complying with reflux failure and GEPF requirements.

Figure 6 shows that, by imple-menting this configuration, the relief flow rate would now be stable around 110 t/h in the event of GEPF, compared with 200 t/h initially envisaged with the


of normal duty. These values corre-spond to the fact that drastic changes in flow patterns, resulting from the interruption of process feed to the furnace coils, affect heat transfer coefficients. Vapour generation is significantly reduced, even if furnace firing were to be maintained. This reduction is more significant if both process feed and furnace firing are stopped.

Dynamic simulations have also demonstrated that the heat duty transferred to the process decreases more sharply if the furnace feed pump is stopped, compared to the furnace feed being maintained.

Crude distillation column dynamic simulationThe use of dynamic simulations helps in an understanding of the

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Figure 4 CDU behaviour under reflux failure after optimisation of the installed orifice area

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Figure 5 Chattering during GEPF

invensys.indd 4 10/3/11 12:52:48

conventional approach. In this case, a reduction of about 50% in relief rate is indicated.

Mitigation measure: design pressure selection with dynamic simulationCommon failure modes such as GEPF are generally the sizing scenario for the flare system (main flare header, KO drums, flare stack and tip, radiation calculations). Avoiding any relief situation during GEPF for the CDU column would contribute to reduced aggregate refinery flare loads. Dynamic simu-lation offers this possibility, for new designs, by optimising the column design pressure.

In the case under study, increas-ing this design pressure to 4.3 barg enables the relief flow rate to be cancelled during GEPF (see Figure 8). Table 2 shows the selected PSV configuration in the event of a set pressure increase, the sizing case corresponding to reflux failure.

This set pressure increase makes recondensation of hydrocarbon vap-ours more important, and reduces the flow rate of low pres-sure stripping steam to almost nil. Consequently, the relief flow rate observed during the reflux pump failure will be smaller (113 t/h versus 414 t/h, see Figure 7).

Dynamic simulation results for aCDU case studyDynamic simulation shows its value in optimising relief valve design and configuration for a CDU column through the follow-ing main results:• Installed orifice area and relief flow rates can be decreased by about 50% in any upset case by choosing the proper PSV configu-ration (number, size and staggering)• Relief flow rates and installed orifice area can be further signifi-cantly reduced by increasing the design pressure of the CDU column. In this case, relief flow rates during GEPF can be cancelled.

The potential benefits of the selected PSV configuration can be estimated as follows:• Reduction in investment and maintenance costs (fewer PSVs,


reduction in flare header size and flare height as a consequence of relief flow rates reduction)• Reduction in chattering inducing a reduction in pipe stress.

Potential savings for grassroots or upgraded refineriesGeneral overviewThe approach in this case is to use dynamic simulations to predict

Set pressure, barg PSV number/type Installed orifice areaConventional calculations 3.5 12 T 2020 cm2

Dynamic simulation optimisation 3.5 3 R

930 cm2 3.7 6 R

Selected PSV configuration

Table 1

Set pressure, barg PSV number/type Installed orifice areaConventional calculations 3.5 12 T 2020 cm2

Dynamic simulation optimisation 4.3 3 Q 210 cm2

Selected PSV configuration with new set pressure

Table 2gra

b,

eruss

erp

dae

hrev

O

rh/t

,et ar

wol f

feil

eR

Time, min

Figure 6 CDU behaviour in the event of GEPF with selected configuration

Time, min

Figure 7 CDU behaviour in the event of reflux failure with increased set pressure

invensys.indd 5 10/3/11 12:52:57


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emergency situations applied to grassroots or revamp projects. For grassroots refinery projects, the establishment as early as possible of a clear philosophy regarding the optimisation of flare systems is beneficial in terms of capital spend-ing. Flare header size, flare stack height and the dimensions of flare KO drums can be reduced significantly.

For refinery upgrading projects, the use of dynamic simulations becomes especially meaningful when assessing the capacity of existing flare systems to accommo-date additional relief loads from new and revamped units. The contribution of dynamic simulation can lead to important reductions in relief loads to the flare and to the conclusion that it is not necessary to implement additional flare network. When the objective is to connect relief valves previously discharging to atmosphere to the existing flare network, dynamic simulations produce realistic

predictions of the new relief flow rates and confirm whether or not the existing flare network is under-sized.

Case study for a grassroots refinery with an aromatic complexTechnip’s methodology for opti-mising flare network sizes was

applied to a grassroots refinery including an aromatic complex. Relief loads were drastically reduced (2700 t/h from 6500 t/h), as well as flare header size (76in from 130in) and flare stack height (180m from 280m). In this case, expected investment cost savings approached €20 million.

Figure 8 CDU behaviour in the event of GEPF with increased set pressure

From other units

From other units

GP

NHTNaphthasplitter

CDU

VDU

VB

FCC

11 t/h

265 t/h 275 t/h 275 t/h

90 t/h

89 t/h

28 t/h

185 t/h

180 t/h

From other units

HDS489 t/h

310 t/h 330 t/h

530+ 230 t/h

500 t/h

540 t/h

1300 t/h37º API0.8% S

From other units

ISOM diH

Gasoline

Kero DieselHeatingoil

FG

LPG

REF

ALKY

HDS

FO

Figure 9 Case study for refinery upgrading scheme (before revamp)


invensys.indd 6 10/3/11 12:53:06

Case study for a refinery upgradeThis study considers upgrading an existing 1300 t/h capacity European topping/reformer/catalytic cracker (TRC) refinery with new conversion units to achieve higher conversion and desulphurisation. Figure 9 shows the refinery’s block flow diagram before and after the revamp. New units included a hydrocracking complex (supple-mented by the required hydrogen plant), whereas existing CDU bottoms and VDU units needed to be revamped. This upgrading solu-tion is a good approach to processing heavier and higher-sulphur crude to meet the current European specifications.

Table 3 shows the relief flow rates to be considered: in the event of GEPF for the flare system design in the existing refinery configuration; for the upgraded refinery configu-ration without the implementation of dynamic simulation; and follow-ing the implementation of dynamic simulation.

Considering that the original refinery flare network is designed to handle 985 t/h in the event of GEPF, the use of conventional calculations after the refinery upgrade with a new hydrocracker unit leads to an increase in the flare network’s design flow rate to 1285 t/h (+27%). In this context, it would be necessary to consider the imple-mentation of an additional flare system or expanding the existing flare system, which are both costly solutions.

Existing simulation models were adapted in this case. Relief load contributors such as crude distilla-tion columns, vacuum distillation columns, reformate splitter and FCC main fractionators were simu-lated and their associated relief loads recalculated. In spite of the introduction of two new units (hydrocracker and hydrogen plant), aggregate relief loads to the flare in the event of GEPF were reduced by 8%, from 985 to 900 t/h, generating substantial investment cost savings

(around €15 million, or about 4% of the total investment).

Further benefits of dynamic simulationA process model used as a knowl-edgebase can enable savings in many areas. The use of a dynamic simulation rather than a steady-state representation opens up new opportunities. The flare study work reported here targets the capital cost of construction, with a particu-lar focus on the materials of construction. This provides a real multimillion-dollar return from the flare system. In the design phase of any project, the capital expenditure budget is paramount.

The benefits of simulation are not limited to design. Dynamic simula-tion can be used to investigate specific implementation tasks and to assess operability and trip settings ahead of the plant being built. More intensive process design, with wider feedstocks and narrow product qualities, is making

Figure 9 Case study for refinery upgrading scheme (after revamp)


invensys.indd 7 10/3/11 12:53:20

process plant harder to operate. Questions such as “Does the opera-tor have sufficient time to respond to a disturbance?” are increasingly important. A trip that takes out a unit in a refinery can cost $0.5 million. If the trip propagates to other units through the utility systems, the costs can be even greater.

In addition, the training of opera-tors so that they maintain a well-regulated process and know how to operate the units for optimum production, while respect-ing the alarm and trip aspects of the process, can regularly improve production by $10–100 000/day.

Every grassroots refinery goes through a commissioning phase, not least the control and trip systems. Checking out the configu-ration using a dynamic simulation provides a more rigorous and extensive review of the control system’s implementation and catches configuration, parameter and design errors before they are identified on plant. In general, the

General electrical power failure TRC refinery TRC refinery upgraded with HCK20% cut-off failing Conventional Conventional With dynamic method, t/h method, t/h simulation, t/hCDU/VDU/gas plant 220 280 190Gasoline units (NHT/Reformer/ISOM/Alky) 660 650 515HDS 50 50 50FCC complex (inc CatNaphtha HDS) 55 55 0Visbreaker 0 0 0HCK (inc H

2 plant) - 220 145

Total 985% 1255% 900% (+27%) (-8%)

Dynamic simulation results for upgraded refinery flare study

Table 3

greater the level of automation, the greater the saving.

SimSci-Esscor, DYNSIM and PRO/II are trademarks of Invensys plc, its subsidiaries or affiliates.

Alban Sirven is Refining Chief Engineer, Process and Technology division, at Technip France, Paris. He holds an engineer’s degree from École Centrale Paris, France. Email: [email protected] Julien Grosclaude is Lead Process Engineer, Refining, Process and Technology division, at Technip France, Paris. He holds an engineer’s

degree from École Centrale Paris and a MSc from IFP School, Paris. Email: [email protected] Guillaume Fenol is Process Engineer, Refining, Process and Technology divison, at Technip France, Paris. He holds an engineer’s degree from Lyon’s School of Chemistry and a MSc from IFP School, Paris.Email: [email protected] Jeremy Saada is a Client Sales Executive for Invensys Operations Management in charge of the SimSci-Esscor simulation software business in France. He holds a master’s in process engineering from École Nationale Supérieure des Industries Chimiques, France. Email: [email protected]



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Energy performance monitoring

For heavy energy consumers such as refineries, total energy expenses represent a consider-

able amount. On the other hand, complex processes require a certain minimal amount of energy to guar-antee the desired production output. The question that naturally arises is how to use the given energy in the most efficient way or, in other words, how to increase energy efficiency.

There are several approaches to tackling the vast challenge of opti-mally using energy. One of these is the implementation of an energy management system (EMS) for closely monitoring and increasing energy performance. An EMS can be interpreted in different ways. Belsim’s conception of an EMS is to prepare all process-related energy information so that it is reliable, centralised and easily readable, and can be readily used as support for intelligent business decisions.

This information includes:• Energy performance indicators (EPIs) of equipment, single units or the whole site• Energy balances to determine accurately the energy consumption of equipment and units• Monitoring of emissions (CO2, SO2, NOx).

It is measured on a short-term basis (hourly or daily) and is acces-sible via customised reporting tools. The results that are reported through an EMS are key to provid-ing an insight into process performance, which is the basis for changes in operating mode to improve economic performance.

However, the quality of business decisions depends directly on the

An integrated energy management system supports improvements to a refinery’s performance in energy consumption and emissions control

RobeRt ChARes, heRvé Closon and hugues steFAnski BelsimJeAn-ClAude noisieR SIR

quality of the data on which the decisions are based. Therefore, an EMS should have an additional component that increases the qual-ity of measurement data. In particular, EPIs are typically calcu-lated based on other measurements. If those measurements are already false, their errors will propagate and amplify. Consider, for example, an efficiency defined as the ratio between two quantities, A and B. If A and B are both erroneous, it is possible that the efficiency would be calculated as higher than 100%, which is clearly impossible.

Advanced data validation and reconciliation (DVR) is a technology that uses measurement data and process information to correct measurements as little as necessary so that all process constraints (such as material and energy balances or thermodynamic equilibria) are satis-fied, while taking into account the uncertainty of measurement.

Belsim’s EMSs always include a

DVR component. There are two major reasons for this inclusion:• The reported information is relia-ble and dedicated actions can be taken to improve energy performance• The impact of the corrective actions can be seen directly, since energy-related data is cleared of any noise.

In that sense, an EMS can be viewed as a tool that gives continu-ous decision support to tackle various energy-related challenges arising in complex systems, such as refineries, leading to corrective actions that have to be taken to amend them. These challenges could be, for example, energy imbalances (fuel, electricity) around some process units or the inexplica-bly high energy consumption of some piece of equipment. They are observations on the process level based on measurement data.

The results supplied by an EMS:help to truly understand the


ActionManagement

ProcessInsight into processes

Verify corrective actions

EMS

Challenge

Reporting

DVR

Figure 1 Belsim’s conception of an EMS

belsim.indd 1 10/3/11 13:12:23

process by providing transparent and reliable energy-related informa-tion that refl ects the actual condition of the refi nery; are presented in a transparent and easily readable manner; can be exported and used for further analysis; are quickly accessible and updated frequently to enable the right decision to be made at the right time.

Based on the results provided by an EMS, dedicated actions can be initiated at the management level — for example, sensors, equipment and control settings can be ques-tioned and causes of poor effi ciency can be detected — leading to a reproduction of better perform-ances. The impact of these actions can be verifi ed with the help of an EMS. In order to ensure a continu-ous improvement, there should be a frequent interaction between the EMS, situated at the process level, and the actions taken at the management level. This concept is illustrated in Figure 1.

Implementation of an EMS This conception of energy perform-ance monitoring has been successfully implemented at Société Ivoirienne de Raffi nage (SIR) in Abidjan, Ivory Coast. SIR refi nes crude oil to produce various prod-ucts for inland markets and for export. At the moment, about 3.8


million tonnes of crude oil are proc-essed in the refi nery each year. Within the framework of a complete revamp of the refi nery’s informa-tion system (Projet d’Intégration de l’Information — PII), Belsim imple-mented an EMS that meets the following requirements:• Implementation and confi gura-tion of a completely integrated solution that automatically calcu-lates energy balances and EPIs with a high precision• Accurate monitoring of air emis-sions for environmental reporting• Energy performance monitoring of equipment, several units and the whole site• Gives everybody in the refi nery access to the results.

The overall objective of the imple-mentation of the EMS was to improve the refi nery’s performance and to make it more competitive within the West African area. This objective was to be achieved by calculating reliable and accurate EPIs and mass and energy balances, leading to an improved monitoring of fuel consumption, steam usage, electricity consumption and fl are, and to a decrease in overall energy consumption.

The solution that was imple-mented at SIR is entirely integrated in PII via a third-party integration platform (m:pro) that is the connec-

tion point for the EMS and all of the other modules. It links different types of data and information coming from various systems in diverse forms and arriving at differ-ent frequencies. It includes reporting facilities that provide the requested information to the right person at the right time. The integration plat-form combines different applications, such as refi nery material balance, refi nery unit performance monitor-ing, energy management of utilities, with databases and provides a single user interface for engineers from various departments all over the refi nery.

The EMS automatically fetches measurement data from the data historian (DAHS), which retrieves input from the laboratory informa-tion management system (LIMS) and the distributed control system (DCS). The measurements can be modifi ed by manual inputs. The results of the EMS are stored on a Belsim SQL database, which can be accessed through a communication tool on PII by anyone and any appli-cation connected to the integration platform. The interconnection between the different components is illustrated in Figure 2.

The EMS is entirely integrated into PII, and all other applications can access results generated by the EMS (see Figure 2). It is also completely

Integrationplatform

Manualencoding

Oracle

DAHS

Massbalances

Performancemonitoring of

units

Production scheduling

...

LIMS

User interface DCS

Energymanagement

Figure 2 Integration of an EMS at SIR

belsim.indd 2 10/3/11 13:12:36

embedded in the daily business workflow of the end users via the integration platform to ensure a continuous improvement in terms of the management of energy usage. The end users review the results and present them during meetings where different business units are present.

Results achievedThe project yielded several benefits for SIR. Some of them are described below. More benefits are to be expected as soon as further correc-tive actions are undertaken by the management of the refinery.

Gains in time and efficiencyThe integration of the EMS into the daily workflow is the foundation of a sustainable improvement through increased awareness of energy consumption at the refinery. Running the application takes only a few minutes every day, which means a significant gain in time and efficiency: engineers can now spend time on analysing data and not waste time on collecting them.

The flow sheet-based reporting tools continuously provide reliable information on the use of energy at the refinery in a transparent manner. Previously, energy balances were established on a monthly basis — not frequent


enough to take action in due time.Since it is entirely integrated

through the integration platform, the whole refinery can gain from the information provided by the EMS. For example, the refinery production balances are established partially based on results provided by the EMS. It turned out that this helped to calculate the balances more quickly, with a 50% reduction in time.

Prevention of lossesSociété Multinationale de Bitumes (SMB) is a company that is located on the SIR site and strongly inte-grated with SIR in terms of energy consumption among other links. As

Figure 3 Prevented losses in fuel invoicing, January–June 2010

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a consequence of this energy inte-gration, SIR sells SMB fuel gas and fuel oil to be used for treating crude oil for the production of road bitu-men and by-products. It turned out that SMB’s consumption of fuel has been systematically underestimated by about 215 tonnes of fuel oil equivalent each month. As a result, the average price for fuel per tonne of crude was underestimated by almost 20%. This underestimation resulted in undercharging for fuel over recent years. With the help of the EMS, the true energy consump-tion of SMB has been determined with a higher precision, and losses have been prevented by invoicing the correct amount of energy


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right indicate the measured and reconciled values respectively. One can see that the electricity consump-tion is significantly underestimated and the production is slightly over-estimated, when considering measured values only.

Moreover, the information provided by the EMS helped to pinpoint erroneous electricity counters. The imbalance of 2 MW was partly due to the counters at the electricity process unit substa-tion, which had never been calibrated over the past 30 years. The calibration of these devices

helped to reduce the imbalance in measurements to below 1 MW.

Supported by the EMS, further actions will be taken by the refinery to close the remaining imbalance of the electricity measurements, in particular by checking the accuracy of electricity production. This will help to distribute energy correctly among the process units in order to accurately calculate their effective operating costs.

Gas flaringRefineries have flare systems for burning excess gas. The burning of some gas might be necessary because it could be difficult to transport it to other places in the refinery, or to release some pressure for security reasons. A large amount of gas being flared indicates non-optimal operation of the refinery.

The results of the EMS indicate that 2000 tonnes of gas is flared each month, corresponding to about 24 000 Gcal. Figure 4 illustrates the average amount of gas flared per hour on 1 March 2010.

This flaring of gas amounts to more than $1 million of losses per month. The EMS detected and iden-tified one flow meter to the flare that is systematically providing wrong data. Despite this erroneous flow meter, the EMS calculated the reconciled flow rate to that flare. A team was appointed by the refin-ery’s management to analyse the possibility of installing a compres-sor, with the aim of transforming part of the flare gas to fuel gas.

Emissions monitoringEmissions are now monitored on a more frequent basis, daily as opposed to monthly, on several levels — refinery-wide, on a unit level and for each item of equip-ment — and in a more rigorous manner. The emissions are calcu-lated based on the consumption of fuel and its sulphur content, which is measured four times per month. Previously, the monthly average of these values had been used to determine the sulphur content. The EMS always uses the latest values to precisely determine emissions of CO2, SO2 and NOx, which are then reported directly to the Ministry of


consumed by SMB. The losses that were prevented in the first six months of 2010 are shown in Figure 3.

Table 1 shows the average monthly invoiced amount based on true fuel consumption, as opposed to wrongly estimated fuel consumption using raw measure-ments. One can see that an average $60 000 of losses could be prevented per month. This finding alone ensures a return on investment of less than six months.

Electricity imbalance detectedAn imbalance in the electricity grid has been identified and quantified at 2 MW, representing about 10% of total electricity production. This imbalance represents an accumu-lated monthly value of about $235 000 based on western European pricing. It is therefore important to properly understand whether some equipment is consuming more elec-tricity than is expected or whether the difference is due to measure-ment errors. As the EMS is based on a data validation and reconcilia-tion technology, the electricity balance is now closed. Table 2a shows SIR’s electricity production on a particular day (14 October 2009) and Table 2b shows electricity consumption on the same day. In both tables, the two columns on the

600

700

500

400

300

200

100

HSK2 SMB DHC Total

Acid

fla

re p

roducti

on,

kg/h

0

MeasuredReconciled

Figure 4 Amount of acid gas flared

Table 1

Item Amount, $Jan-Jun 2010 average monthly invoice for SMB, using true energy consumption 321 000Jan-Jun 2010 average monthly invoice for SMB, using underestimated energy consumption 261 000Prevented losses due to energy consumption underestimation 60 000

Effects on fuel invoicing of true and underestimated energy consumption

Unit Measured ReconciledGTA5230 11.71 11.25GTA5240 11.28 11.09CIE 0.00 0.00Total 22.98 22.34

Electricity import/production, MW

Unit Measured ReconciledP10 3.72 3.91P11 2.08 2.14P12 3.75 3.94P13 8.78 9.82P14 2.45 2.53Total 20.78 22.34

Electricity import/consumption, MW

Table 2a

Table 2b

belsim.indd 4 11/3/11 14:43:50

Environment. The daily availability of these values has been acknowl-edged by ISO 14001 auditors.

Lessons learnedThe implementation of an EMS based on an advanced DVR tech-nology has clearly demonstrated its first benefits. It is completely inte-grated into the daily workflow at the refinery of SIR. It is actively used for providing continuously reliable energy-related data. Business decisions can now be based on a trustworthy set of infor-mation. SIR has acquired ownership and therefore full control of the application. Awareness in terms of energy consumption has been increased and losses can be prevented without installing new equipment or instrumentation.

The management of SIR has noticed a huge time saving in the collection of reliable information for both energy balances and environ-mental indicators as a result of the implementation of this project. The time saved is available for engineers to analyse in greater depth the


refinery’s current situation and to propose operating improvements. In other words, SIR has highlighted a noticeable improvement in work-ing productivity.

Actions taken by the management of SIR are steps towards the improved competitiveness of the refinery. Moreover, the effects of corrective actions can be detected more easily and quickly as a result of the removal of measurement noise by the DVR technology.

Success factorsThere were three important factors in the success of this project. First, there was a strong commitment to the project by the management of SIR and a positive attitude towards the management of change. Second, a dedicated, multi-disciplinary team was established, formed by staff from SIR and Belsim, working together towards the defined objec-tives. Third, the transfer of knowledge from Belsim to SIR was provided by training, on-site missions and continuous support during the project and following its

completion. These factors not only guaranteed that the project was successful for both parties, but also ensured continuous improvement after the project was completed.

Robert Chares is Marketing Solution Engineer at Belsim, Awans, Belgium, specialising in energy management and energy production. He has a master’s in mathematics from Chemnitz University of Technology, Germany, and a PhD in applied mathematics from University of Louvain, Belgium. Email: [email protected]é Closon is Director, Services and Technology, at Belsim. He has a master’s in chemical engineering from the University of Liège, Belgium, and in management from HEC, Liège. Email: [email protected] Noisier is Performance and Quality Manager at SIR in Abidjan, Ivory Coast. He has a master’s in engineering and a PhD in physics and mathematics from the French Petroleum Institute, part of Paris University. Email: [email protected] Stefanski is an Advanced Process Engineer at Belsim. He specialises in energy management, performance monitoring and production accounting for downstream applications and has a master’s in chemical engineering from the University of Liège, Belgium. Email: [email protected]


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SO2 emission control for resid combustion

Various environmental and market pressures may provide the refiner with an

incentive to consume modest quan-tities of high-sulphur residuum on site, to generate power or steam that is consumed locally.

Resids contain elevated amounts of sulphur compared to sweet liquid fuels or natural gas. SO2 scrubbing is likely to be required to support a refinery’s co-generation project that uses a heavy sour liquid fuel stream. Designers need to consider whether to install a non-regenerable SO2 scrubbing system that uses an alkaline agent such as caustic soda (NaOH), lime or lime-stone, or whether a regenerable system will be used that generates a pure stream of SO2 that can read-ily be converted to elemental sulphur or 98% sulphuric acid. Depending on the sulphur content of the fuel in question, one or the other type of system is preferred. This article describes the regenera-ble Cansolv SO2 system and compares it to non-regenerable systems that employ NaOH, lime or limestone as the sorbent.

Resid combustion and cogenerationThe refining industry continues to be pressured to process more diffi-cult crudes and, at the same time, respond to ever-tighter limitations on product sulphur content. Investment in bottom-of-the-barrel hydrocracking or coking systems is economical for large volumes of resid, but options are limited when smaller volumes of high-sulphur products must be accommodated. Co-generation projects may be able to serve as outlets for future quanti-

Regenerable and non-regenerable SO2 scrubbing systems for high-sulphur

residuum combustion are compared

Rick BiRnBaumCansolv Technologies

ties of high-sulphur refinery streams. These projects will require SO2 scrubbing to prevent the sulphur in the fuel mix from escap-ing to the atmosphere.

A refiner may wish to consider a co-generation project for three

reasons. First, co-generation projects can be designed to accommodate a wide range of fuels, from cycle oils to asphalt and coke. Second, on-site generation of power and steam reduces the refiner’s reliance on external, purchased balancing fuels. Finally, the versatile nature of the

co-generation system to accept vari-able fuel qualities increases the refinery planner’s ability to fill the refinery crude slate with a wider range of crudes and to maximise margins.

SO2 scrubbing

SO2 scrubbing can be effected by non-regenerable and regenerable means. Non-regenerable systems consume an alkaline agent such as sodium hydroxide, limestone or dry lime and generate a waste stream of sodium sulphate or gypsum (calcium sulphate). Sodium sulphate is disposed of in wastewater treatment systems, while gypsum is most often disposed of in a landfill site.

Regenerable systems use an alka-line agent, such as sodium sulphite or Cansolv DS, to capture SO2 and release it in pure form from a regenerator that is designed to split the reagent from the SO2. In the refinery, SO2 is converted to sulphur in the refinery’s sulphur recovery unit (SRU).

Non-regenerable systems capture one tonne of SO2 from flue gas and


1Flue gas

management

2Reagent

preparation

3SO2

absorption

4SO2

regeneration

5Waste

management

6Byproduct

management

Figure 1 Process subsystems for flue gas desulphurisation

co-generation projects can be designed to accommodate a wide range of fuels, from cycle oils to asphalt and coke

cansolv.indd 1 10/3/11 13:17:34

generate between two and three tonnes of dry equivalent waste, which is discharged as a dilute stream of liquid (sodium sulphate) or as a wet, hydrated solid (CaSO4). Regenerable systems, which convert SO2 to elemental sulphur in the refinery, remove a tonne of SO2 from flue gas and generate only half a tonne of high-value, marketable product.

When external SO2 from a regen-erable SO2 scrubbing system is fed to the refinery SRU, it displaces combustion air and can be used alone or in combination with other strategies, to debottleneck or reduce the cost of the SRU, which is often under pressure to process greater quantities of H2S from elsewhere in the refinery.

SO2 scrubbing systems are composed of up to six process blocks (see Figure 1). The lime and limestone non-regenerable systems


include a reagent preparation area (block 2) and a gypsum filtration area (required for blocks 5 and 6). Caustic requires no reagent prepa-ration or byproduct management block, since NaOH is sourced as a bulk liquid and sodium sulphate is discharged to wastewater-treating systems as a dilute solution of sodium sulphate. The non- regenerable systems, by definition, do not have a regeneration block.

Regenerable systems have little or no requirement for a reagent prepa-ration area, but dedicate significant resources and space to solvent regen-eration. While both regenerable and non-regenerable systems can produce a marketable byproduct, it is more common for the caustic, lime and limestone non-regenerable systems to direct their byproducts to waste, as markets for sodium sulphate or gypsum are limited. This contrasts with the regenerable systems that

direct SO2 into sulphuric acid or elemental sulphur markets.

Table 1 compares the process areas required for several common non-regenerable systems and the Cansolv SO2 scrubbing system.

Caustic systems have lower capi-tal costs than other non-regenerable systems. Reagent is purchased as a concentrated liquid or dry solid and wastes are directed to the refinery wastewater treatment system. NaOH prices are quite volatile, because it is a co-product in the manufacture of chlorine. When chlorine is in high demand, NaOH pricing tends to drop, whereas when chlorine is in low demand NaOH prices tend to rise. Chlorine markets are sensitive to world demand for chemical products such as vinyl chloride monomer, used to manufacture polyvinyl chloride plastics.

Sodium carbonate and sodium bicarbonate are less expensive alter-natives to caustic and can be substituted for caustic, but their cost is higher than for limestone or lime, and they require investment in reagent preparation and manage-ment systems.

Capital costs for limestone- and lime-based systems show an advan-tage over regenerable systems. Although investment is required for reagent preparation and byproduct or waste management, by defini-tion, no investment is required for solvent regeneration systems.

Regenerable systems have a higher capital cost than any of the non-regenerable systems, since solvent regeneration and byproduct conversion systems are required.

Table 1

Figure 2 Comparative capital and operating costs: Cansolv vs non-regenerable scrubbing systems

Cansolv Caustic Limestone Lime spray dry

Capital costOp. cost 4.6%SOp. cost 2.6%S

Area Cansolv NaOH Limestone Lime spray dryReagent prep – – Grinding and storage Storage and slakin

Pre-scrubbing Required upstream of absorber Incorporated in absorber Incorporated in absorber Incorporated in absorber

Absorber type Packed Multiple spray Multiple spray Spray atomisation

Regeneration Steam strip – – –

Waste management Minor blowdown from Na2SO

4 to wastewater CaSO

4, dry filter cake to landfill CaSO

4, dry filter cake to landfill

pre-scrubber and amine or to possible sale purification system

Byproduct recovery Sulphur plant; sulphuric acid; – Gypsum – liquid SO

2

Process subsystem requirements: regenerable and non-regenerable scrubbing systems

Table 1

cansolv.indd 2 10/3/11 13:17:43

Instead of consuming non-regener-able reagents, the regenerable systems consume relatively large quantities of steam, electricity and cooling water to split SO2 from the reagent.

SummaryIn summary, site-specific economics and the amount of SO2 that must be captured from a given flue gas stream dictate whether or not a regenerable system is favoured over a non-regenerable system.

Optimum conditions for regener-able systems apply when:• Fuel sulphur content is high• Energy costs are low• Access to alkaline reagents is limited• Access to wastewater or waste landfill systems is limited.

Optimum conditions for non-regenerable systems apply when:• Fuel sulphur content is low• Energy costs are high• Free access to alkaline reagents is available• Free access to wastewater and solid waste disposal sites is available.

Figure 2 shows high-level capital and operating cost comparisons for four SO2 flue gas desulphurisation (FGD) technologies. The chart illustrates that the capital cost for the Cansolv system is higher than for the non-regenerable systems, but that its operating cost is lower. A second case comparing the costs of scrubbing 2.6% sulphur fuel is shown alongside the base case. The capital and operating costs for Cansolv are relatively constant, but the operat-ing costs of the non-regenerable systems, which are dominated by reagent costs, change in direct proportion to the amount of SO2 that must be captured from the flue gas.

Net Present Value (NPV) calcula-tions using a discount rate of 10% show that for the 4.6% sulphur case, non-regenerable technologies cost more than the regenerable technol-ogy. The comparison is shown in Table 2. Table 2 confirms that for reduced sulphur feeds, the NPV of the regenerable process still prevails over caustic and limestone, but


becomes competitive with lime spray dry systems.

The cost basis for the four cases is highly dependent on costing assumptions used for the analysis. The competitive NPV for the

2.6% sulphur case between Cansolv and the non-regenerable processes requires that the costs basis be clearly defined, while the conclusions for the 4.6% sulphur case are more clear cut. Table 3 shows the economic analysis assumptions used for these comparative cases.

ConclusionsFor regenerable systems, steam represents 40% of the operating cost. For caustic scrubbed systems, reagent makes up nearly 60% of the operating cost of the system. In limestone and lime systems, the reagent costs represent 25–35% of the operating cost, respectively.

Cansolv SO2 scrubbing systems can be economically justified over non-regenerable systems for cases that involve the combustion of high-sulphur resids.

Reagent availability, the client’s ability to dispose of waste byprod-ucts from the non-regenerable system and materials-handling problems associated with solid reagents and byproduct can drive the decision to use a regenerable technology for sour fuel fired co-generation applications.

Rick Birnbaum is Licensing Manager, Oil and Gas, for Cansolv Technologies Inc’s SO

2

and CO2 scrubbing technology businesses in

Montreal. He graduated from McGill University in Montreal, Quebec.

Table 5.

Technology NPV compared to Cansolv at 4.6% NPV compared to Cansolv at 2.6% S in fuel; discounted at 10% S in fuel; discounted at 10%Cansolv 100 100Caustic 130 107Limestone 117 113Lime spray dry 110 99

NPV of 4.6% and 2.6% sulphur in fuel case: regenerable vs non-regenerable

Table 2

Boiler sizing basis Flue gas basis 4.6% sulphur case Flue gas basis 2.6% sulphur caseFuel feed rate, t/h 35 35Sulphur in fuel, wt% 4.6 2.6Boiler power production, MWe 160 160SO

2 capture capacity, t/y 28 000 16 644

SO2 content of flue gas, vppm 2400 1400

Utility cost basis Cooling water, $/m3 0.02 Steam, $/ton 10.00 Electricity, $/kW 0.085 DI water, $/m3 1.80Utility water, $/m3 1.00ChemicalsNaOH, $/t 300Limestone, $/t -100% NaOH 30Lime, $/t 100Waste, $/t - dry basis Na

2SO

4;CaSO

4 20

Byproduct credit - elemental sulphur, $/t 60

Operating cost basis for 4.6% S and 2.6% S, 160 MW boiler SO2 scrubber

Table 3

Regenerable systems have a higher capital cost than the non-regenerable systems, but their operating cost can be a lot lower

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Bulk separation of gas-liquid mixtures

T he operating conditions of mixed phases and the requirements for separation

efficiency may vary widely. Therefore, special care should be taken in selecting the most appro-priate device to match the specific duty. For application to bulk gas-liquid separation, where generally not more than 95% of the liquid must be removed from the gas stream, the Schoepentoeter is a proprietary feed inlet vane device used for introducing gas-liquid mixtures into distillation columns or gas-liquid separators. It has two main functions: • To separate the liquid from the gas • To distribute the vapour in the gas compartment of the column.

The device accomplishes these objectives by slicing up the mixed-phase feed into a series of flat jets by means of properly distributed and oriented vanes. The jets dissi-pate a large part of the kinetic energy due to the vanes so that the vapour enters the gas compartment of the column in a smooth and uniform manner. The vanes also provide the mixed-phase feed with centrifugal acceleration to promote and/or enhance the separation of the liquid from the vapour — other-wise possible only by gravitational force.

For any given duty, the device allows for a considerably smaller feed entry section in the vessel, thus reducing the total column’s height and costs.

Process design parameters The main design parameters for a Schoepentoeter are the sizing of the

Effective gas-liquid separation is increasingly important to produce high-quality products from feedstocks of decreasing quality

GiusEPPE Mosca, PiErrE schaEffEr and BarT GriEPsMa Sulzer Chemtech harry KooijMan Shell Global Solutions International

feed inlet nozzle, the flow parame-ter and the column load factor. These factors are important in predicting its efficiency. The sizing of the feed inlet nozzle of a vessel equipped with a Schoepentoeter should be based on the maximum flow rate, including the design margin. The internal nozzle diame-ter can be taken to be equal to that of the upstream feed piping to the vessel, provided that the maximum momentum criterion is satisfied. In some applications where the gas density is very low — for instance, in refinery vacuum towers — the velocity of the gas at the feed inlet nozzle should be somewhat lower than the critical velocity of the gas (the speed of sound of the gas mixture) to prevent choking or damage due to vibrations. The flow parameter is used to characterise the type of gas-liquid mixture enter-ing the vessel or the relative importance of the liquid load approaching the feed inlet device. It is proportional to the ratio of the liquid mass flow to gas mass flow.

Additionally, the performance of the device — in particular, the separation efficiency — is greatly affected by the column load factor, also known as the capacity factor. This factor is proportional to the volume flow of the gas to the cross- section of the tower.

separation efficiency The separation efficiency of a feed inlet device for gas-liquid mixtures is normally defined by the ratio of the liquid flow rate separated from the gas stream and the liquid flow rate originally contained in the mixed-phase stream.

For a Schoepentoeter, the separa-tion efficiency can be expressed as a function of the nozzle’s and column’s diameters, the column load factor, the flow parameter and the ratio of the surface tension of the liquid compared with the surface tension of water.

Mechanical design parameters The device should be designed to comply with and satisfy the follow-ing mechanical requirements and criteria: • A maximum operating load over the feed inlet nozzle of 15 000 Pa • Withstand its own weight plus the weight of the fluid at process conditions • Downward or upward deflection under operating loads not exceed-ing 1% of the nozzle diameter or 15 mm, whichever is larger• Tilt of the Schoepentoeter not exceeding 1% of the column diame-ter or 15mm, whichever is smaller• Thermal expansion during normal operation and transient conditions — for instance start-up and shutdown — should be also considered• For mega-sized Schoepentoeter devices, those with a nozzle diame-ter exceeding 3m and length exceeding 9m, additional detailed mechanical strength calculations and vibration calculations should be performed.

There are cases — refinery vacuum tower revamps or flare system knockout drums — in which the Schoepentoeter is subject to loads even heavier than those mentioned above. Therefore, some additional measures should be taken to avoid vane tips being bent


sgs.indd 1 10/3/11 13:25:13

or broken, by for instance using thicker material or employing stiff-ening strips at the back of long, unsupported vane tips.

Established performance In most cases, the conventional Schoepentoeter has been proven to match and even exceed expected performance. There are only a few applications, such as refinery vacuum towers, where the separa-tion efficiency was measured to be lower than expected. Those meas-urements may have occurred because the liquid, separated by the vane, is not conveyed. Rather, it leaves the vane in the shape of a thin curtain, which, on its way to the bottom section of the tower, is subject to the upward momentum of the ascending vapour. A portion of the separated liquid (entrain-ment) may be carried to the feed entry zone of the tower. Therefore, the resultant separation efficiency can be lower than expected, espe-cially under severe operating conditions that are commonly


encountered in refinery vacuum towers, such as when the inlet nozzle momentum is above 7000–8000 Pa or the column load factor is above 0.09 m/s.

Research and development Extensive research and develop-ment work was completed in the form of experimental tests and computational fluid dynamics anal-ysis at the Sulzer Chemtech pilot plant in Winterthur and at the Shell Technology Center in Amsterdam.

The aim of the study was to opti-mise the separation efficiency without compromising the hydrau-lic capacity, particularly the pressure drop through the feed nozzle and the Schoepentoeter itself. The idea was to design a feature that would collect the separated liquid in a way to coun-terbalance the upward momentum of the ascending vapour. Several types of advanced vanes were tested. The goal was achieved by modifying the back end of the vane from a straight and flat vertical

plate to a sloped and curling plate — the so-called catching rim.

The catching rim collects the separated liquid and conveys it into a rivulet heavy enough to win the upward momentum of the ascend-ing vapour and reach the bottom section of the tower, thus minimis-ing the entrainment. The tests were performed at different capacity factors and flow parameters (see Figure 1).

At low column load factors, no major difference was measured; both the conventional and new Plus devices performed sufficiently. At higher capacity factors, typically encountered in several industrial columns, the separation efficiency of the Plus version was consistently higher than the conventional one; the entrainment was even less than one-third for values typically encountered in several industrial columns. The improvement was achieved without any significant increase in pressure drop (see Figure 2).

A new correlation for the predic-tion of the entrainment was developed by analytical regression of the experimental data, which considers the effect of the new vanes. A new tool has been engi-neered to manufacture the catching rim.

Computational fluid dynamic study Within the last decade, computa-tional fluid dynamics (CFD) has reached such maturity that it is now considered an indispensable analy-sis and design tool in a wide range of industrial applications, including feed entry sections of distillation towers or gas-liquid separators. Therefore, a CFD study was performed to check the efficiency of the feed inlet device in terms of vapour distribution. For this scope, the flash zone of a refinery vacuum tower was modelled and analysed with both the devices. The follow-ing operating conditions were set: • A feed inlet nozzle momentum of 6370 Pa • A column load factor of 0.097 m/s • A collector tray with a 30% open area above the Schoepentoeter• A combined bed of Mellapak

Entr

ain

ment

Column load factor

Schoepentoeter ConventionalSchoepentoeter Plus

Figure 1 Tests were performed at different column load factors and flow parameters. At higher column load factors, the entrainment of the Schoepentoeter Plus is even less than one-third of the conventional one

Figure 2 The improvement is achieved without any significant increase in pressure drop

Pre

ssure

dro

p

(Column load factor)2

Max two-column widthBlue background around brown.Always translate into UK English (except images).Leading cap on �rst word only.

Schoepentoeter ConventionalSchoepentoeter Plus

sgs.indd 2 10/3/11 13:23:56

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125X and Mellagrid 64X structured packings above the collector tray.

The vertical vapour velocities over the horizontal plane were checked at different tower eleva-tions, particularly underneath the combined bed of Mellagrid and Mellapak in the wash section. There is no significant difference between the two devices: the vapour distri-bution efficiency is good for both the distributors.

Fields of application In general, the Schoepentoeter Plus could be used in all applications suitable for a conventional device, such as separation in oil and gas upstream units or distillation in oil and gas downstream plants. This article focuses mainly on the second application. There are cases where there is no need for the


Schoepentoeter Plus, such as when the inlet device is used for a single-phase stream and no significant benefit in distribution efficiency would be achieved. The higher cost of the Plus version makes the conventional one more attractive.

The best candidates for installa-tion of the Plus device are vacuum towers, crude distillation main frac-tionators and hydrocracking main fractionators in oil refineries, where the separation efficiency of vapour from liquid plays a significant role in the performance of the units.

Case study: vacuum tower revamp The column is located at a major European refinery. The main duty of the tower is the recovery of light and heavy vacuum gas oil (LVGO and HVGO) from the long residue coming from the primary

distillation of the crude oil. The feed, preheated to 400–420°C and partially vapourised, accesses the flash zone of the column through the feed inlet device, which performs a bulk separation of the liquid from the vapour as well as a vapour distribution in the gas compartment of the column.

The liquid drops down to the stripping section for ultimate recovery of the light hydrocarbons and is finally drawn off as short residue from the bottom of the tower. The vapour phase is fraction-ated into LVGO and HVGO in the upper sections. The LVGO is drawn off at the top section of the tower.

The HVGO is generally the first useful side cut above the flash zone. A pumparound provides the column with the duty necessary to condense the right amount of vapours coming from the wash section. A portion of the condensate — the wash oil — is pumped back to the bed below to control the quality of the drawn-off product. Among other factors, such as feed composition, wash section configu-ration and operating parameters, the quality of the HVGO may also be affected by the separation effi-ciency of the feed inlet device.

Concerns at the existing tower The flash zone of this column was originally equipped with a conven-tional Schoepentoeter. Since the separation efficiency was lower than expected, the liquid carryover to the wash section (entrainment), which was made up of the heaviest hydrocarbons and should have followed the short residue at the bottom of the tower, was higher than expected. As a consequence, the slop wax flow rate (generally an unwanted product) was consist-ently higher than foreseen.

In an attempt to maximise the yield of HVGO while minimising the production of slop wax, the wash oil was substantially reduced, even below the minimum, causing a deterioration in the wash bed’s performance: • Poor-quality HVGO: a high Conradson carbon residue (CCR) and metals content with a negative impact on the downstream fluid

Figure 3 Coke in the wash bed leads to a higher pressure drop and lower recovery of distillates, resulting in shorter plant run length and unexpected shutdown

sgs.indd 3 10/3/11 13:24:05

refinery. All the other tower inter-nals were retained. The tower has recently been started up.

Increased separation efficiency The Schoepentoeter Plus is a tool with which to improve the bulk separation efficiency of gas-liquid mixtures. The main fields of appli-cation are the refinery towers in vacuum distillation units, crude distillation units and hydrocracking plants. The best fit is the revamp of vessels equipped with radial feed inlet devices. In new columns, the higher cost of the Plus version may make the conventional device more attractive, provided that the performance requirements are not excessively high.

Schoepentoeter is a mark of Shell. Mellagrid and Mellapak are marks of Sulzer Chemtech.

Giuseppe Mosca is the Global Refinery Technology Manager of Sulzer Chemtech in Milan, Italy. He leads the design of mass transfer components for distillation towers, absorbers and strippers, and is

involved in process simulation, revamping proposals, troubleshooting, commissioning of tower internals and start-up assistance of fractionation equipment. He holds BS and MS degrees in chemical engineering from the University La Sapienza in Rome.Email: [email protected] Pierre Schaeffer is an R&D Engineer in the Laboratory for Mass Transfer Technology at Sulzer Chemtech in Winterthur, Switzerland. As a specialist in experimental methodology, he contributes to the development of separators and fractionation trays. Email: [email protected] Griepsma is a Senior Mechanical Specialist with Sulzer Chemtech in Winterthur. He is currently responsible for global support in mechanical engineering, global technical sales support for mechanical revamping proposals, and for mechanical design optimisation and product improvement. Email: [email protected] Harry Kooijman is Senior Separations Equipment Consultant for Shell Global Solutions International and is a Subject Material Expert for transport properties (diffusion). He has a MS degree in chemical engineering from Delft University of Technology, Netherlands, and a PhD in distillation from Clarkson University, New York.


catalytic cracking (FCC) unit. These results led to lower liquid yields and a higher catalyst make-up rate than expected• Coking up of the wash bed: higher pressure drop and less recovery of distillates. This effect led to a shorter plant run length and unexpected shutdown, and thus a reduced plant utilisation factor and increased maintenance costs (see Figure 3).

Tower modifications After an in-depth investigation and detailed analysis of the tower performance, Sulzer decided to replace the existing conventional Schoepentoeter with the Plus version. The wash bed was replaced due to coke formation, and the exist-ing combination of Mellagrid and Mellapak was retained. In addition, the two pumparound beds were replaced with the same type of packing within the scheduled main-tenance programme of the unit during the overall turnaround of the


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Diverting low-sulphur heavy stocks for fuel oil production

Burgeoning use of natural gas, including liquefied/compressed natural gas (LNG/

CNG), as a cheaper fuel option has made low-sulphur heavy stocks (LSHS) a product in surplus since 2008–2009. In fuel-consuming indus-tries such as power production, liquid fuels like naphtha and LSHS have faced significant competition from an increased use of natural gas.1 However, fuel oils have maintained a better market as a result of steady local demand and export markets, reflecting their ease of transportation.

LSHS is normally produced by blending two straight-run refinery streams; namely, low-sulphur vacuum residues (LS-VR) and clari-fied oil (CLO), with LS-VR as more than 90 wt% of the blend (see Figure 1). This is one of the simplest options for utilising residues. It has been practised for years and no processing units are required. In a period of low demand, large quan-tities of LSHS occupy storage tanks, which puts pressure on the capacity of low-sulphur-processing crude distillation units for the continuous production of LS-VR. In these circumstances, it is essential to evacuate LSHS to sustain crude throughput and add value.

Studies enabled a refinery to divert low-value product stocks into higher-value products and preserve crude throughput levels

Rajeev KumaR, ChithRa v, Peddy v C Rao and N v ChoudaRy Bharat Petroleum Corporation Ltd, India

In the current refining climate, refinery configurations have been improved to cater for developments in the processing of low-cost oppor-tunity crude oils, including high-sulphur feeds, high TAN crudes and heavy oils.2 Additionally, residue upgrading techniques and hydrotreating options are applied to refinery bottoms to achieve

higher margins. However, many refineries still do not have the best hardware configurations to match their processing needs, or they may suffer space constraints for residue upgrading processes, such as delayed coking, visbreaking and solvent deasphalting.3 Older crude

distillation units (CDU) were designed to process only low-sulphur crude oils producing LS-VR streams. This also leads to the production of quantities of LSHS, which fills the storage tanks in a surplus market. At one of BPCL’s refineries, in Mumbai, there are three CDUs with a total capacity of 12 million t/y. The two older units (CDU-I and CDU-II) have a combined capacity of 6 million t/y and they process only low-sulphur crude oils. Thus, the continuous operation of these units leads to LSHS production. When LSHS is in low demand, it fills many storage tanks. In this scenario, the evacua-tion of LSHS is essential to sustain the crude throughput of these older units. The LSHS product specifica-tions are shown in Table 1.4

Conversely, CDU-III has a capac-ity of 6 million t/y and processes high-sulphur crude oils, generating


LShS Fuel oil LSFoProperties Normal Coastal LowS 180cStgrade 380cStgrade 180cStgrade 380cStgradeDensity@15°C,gm/cm3 ReportedSulphur,wt% Max1.0 Max1.5 Max0.45 Max4.0 Max4.0 Max4.0 Max4.0Pourpoint,°C Max66 Max45 Max63 Max27 Max27 Max27 Max27Kin.viscosity@50°C,cSt Max500 - Max180 Max380 Max180 Max180 Max380Flashpoint,°C Min76 Min66 Min66 Min66 Min66 Min66 Min66

Specifications for LShS, fuel oil and LSFo

table 1

Blending

CLO

LS-VRLSHS

Figure 1 BlendingschemeforLSHS

in fuel-consuming industries, liquid fuels like naphtha and LShS have faced significant competition from an increased use of natural gas

bpcl.indd 1 10/3/11 13:31:58

high-sulphur vacuum residue (HS-VR). These streams are used in the production of fuel oil. Fuel oil is normally produced by blending HS-VR, LS-VR and kerosene/high-speed diesel (HSD) to meet product specifications for flash point, kine-matic viscosity, pour point and so on. Normally, two grades of fuel oil are produced for different applica-tions; namely, 180 cSt and 380 cSt. The product specifications are shown in Table 1. This is, again, a simple residue utilisation scheme for which no complex processes or operations are needed. The existing correlation for the fuel oil blending scheme has always had some qual-ity giveaways while meeting all the specifications. Both of the viscosity grades for fuel oil either fail to meet specifications or have quality givea-


ways. In order to address these issues, optimisation of the fuel oil production scheme and the devel-opment of new correlation models has been taken up at BPCL’s Corporate R&D Centre. A blending scheme for fuel oil is shown in Figure 2.

In this article, a modelling approach is proposed, exploring

various possibilities for dealing with surplus LSHS. Absorption of LSHS into fuel oil and/or making a new low-sulphur fuel oil (LSFO) product were ideas that evolved into immediate solutions. New correlations have been developed

for the prediction of critical specifi-cations for LSHS, fuel oil and LSFO products. The model uses Aspen+ and crude oil management tools for simulations.

The immediate need to deal with LSHS by diverting it into the fuel oil pool was the first choice for development and implementation. Experiments were carried out at BPCL (R&D) with refinery samples to generate a database for estimat-ing parameters for the model. The model has been developed and implemented at the refinery and proved helpful in minimising qual-ity giveaways.

Following implementation of the model at the Mumbai refinery, the work received a Judges Special Award in a corporate competition for new ideas.

Methodology for diverting LSHS into fuel oil5

During the processing of low-sulphur crude oils in refineries, volumes of LS-VR are produced in excess, hence more LSHS is produced than is required. During 2008–2009, when demand for LSHS suddenly fell away, production moved into surplus and tank stor-age of the product for long periods became difficult. Various possibili-ties have been explored in order to avoid its production to sustain crude throughput — for instance, producing LSFO and absorbing LSHS into fuel oil — to achieve continuous operation of the low-sulphur crude processing units.

For maximising the absorption of surplus LSHS into fuel oil for both product grades, a methodology has been developed on the basis that LSHS is simply a blend of two components, LS-VR and CLO, which are common constituents of fuel oil blends (see Figures 1 and 2). The concept of diverting LSHS into fuel oil is shown in Figure 3.

Production of LSFO5 In view of stringent environmental regulations and particular concern about air quality in metropolitan cities, there is growing customer demand for LSFO. The scheme for producing LSFO is similar to the fuel oil blending scheme (see Figure

Feed streamsFeed properties HS-VR LS-VR Kero CLO LSHSDensity @ 15°C, gm/cm3 0.9953 0.9605 0.7979 0.9916 0.9621Specific gravity 0.9959 0.9611 0.7985 0.9922 0.9627° API 10.58 15.73 45.71 11.11 15.48Sulphur, wt% 4.567 0.552 0.07 0.754 0.5621Pour point, °C 34 50 <-60 34 45Kin. viscosity @ 50°C, cSt 493 880 10 226 1.0583 31.39 3848Flash point, °C 352 344 46.2 72.8 -

Characterisation of refinery streams and LSHS

Table 2

BlendingCLO

LS-VRHS-VRKERO

FO

Figure 2 Blending scheme for fuel oil

CLOLS-VR

HS-VRKERO

FO

LSHS

Figure 3 Blending scheme for diversion of LSHS into fuel oil

New correlations have been developed for the prediction of critical specifications for LSHS, fuel oil and LSFO products

bpcl.indd 2 10/3/11 13:32:07

2), but the specification for LSFO is slightly different from fuel oil with regard to sulphur composition (see Table 1). As with fuel oil, two grades of LSFO, 180 cSt and 380 cSt, are required. To meet the LSFO require-ment, an alternative scheme was developed involving surplus LSHS. LSFO is an upgraded version of the fuel oil product but with the sulphur specification reduced to 2.0 wt%. The other specification parameters are similar. Since sulphur content is the cut-off property for LSFO prod-uct, the composition can be simulated and fixed for sulphur-limiting streams. Optimisation of the other streams can then be carried out to meet the other specifications.

In the present case, the possibility of diverting LSHS into LSFO prod-uct was explored. Since the LSFO product was new, optimisation of the blend composition was carried out as part of the study. This proved helpful in producing LSFO with minimum quality giveaways.

Experimental detailsSamples were collected from the refinery for experimental investiga-tion and to generate data for modelling and simulation. The streams, obtained from storage tanks, were HS-VR, LS-VR, kero, CLO and LSHS. These streams were characterised for their physical properties (see Table 2). The data were used in developing the corre-lation models for predicting the properties of the final product streams in order to meet product specifications.

Modelling and simulationProduction of the fuel oil blend is done using the streams shown in Figure 2. A simple correlation was developed for predicting the critical properties of fuel oil blends, such as flash point, kinematic viscosity and pour point. The correlation applies to fractions with a wide range of boiling points, 140–800°C. The model used Aspen+ and Crude Manager database tools for simula-tion, and the input data required for these streams included boiling distribution-distillation profile data (ASTM D86 or SimDist), kinematic viscosity @ 50°C, pour point and


95

100

90

85

80

75

70

65 70 75 80 85 90 95 100

Sim

ula

tion,

ºC

Experimental, ºC

65R2 = 0.9506

Figure 4 Validation of flash point model

350

400

300

250

200

150

100

50 100 150 200 250 300 350 400

Sim

ula

tion

, cS

t

Experimental, cSt

50R2 = 0.9741

Figure 5 Validation of viscosity point model

Sim

ula

tion

, ºC

Experimental, ºC

R2 = 0.9812

–30 –20 –10 0 10 20 30 40 50

30

50

40

20

10

0

–10

–20

–30

Figure 6 Validation of pour point model

bpcl.indd 3 10/3/11 13:32:18

The proposed models are given as:

Flash point, °C = 0.1813*[MABP(°C)] - 0.12 (1)

Kinematic viscosity, cSt @ 50°C = Exp (k1*Exp

[((MABP(°C)/323.15)+0.8453)*(SG)k2])+k3 (2)

For KV : k1,2,3

(1.0185, 4.4419, 3.89)

Pour point, °C = [(X1P

1)+ (X

2P

2)+ (X

3P

3)+

(1.9975*X4P

4)] - 3.77 (3)

(where Xi are compositions and P

i are pour

point values)

Sulphur is a linear property (4)

The new correlation models open up several opportunities for the production of various grades of product. The new blending schemes made for fuel oil and LSFO are shown in Tables 3 and 4. A blend-ing scheme for using LSHS in the production of LSFO has also been established (see Table 5). The new scheme achieves minimal quality giveaways.

Diversion of LSHS into fuel oil:implementation A modelling approach has been employed to explore the possibilities for diverting surplus LSHS into fuel oil. The kinematic viscosity, flash point and pour point values for LSHS are very high compared to those of fuel oil, hence it is clear that the absorption of LSHS into fuel oil can have a major effect on viscosity and pour point values only. However, the sulphur content is very low for LSHS compared to fuel oil. This implies that the absorption of LSHS into fuel oil would have a diluting effect on sulphur content. Therefore, no concern arises in meet-ing the sulphur specifications for a fuel oil blend.

There are two constraints (kine-matic viscosity and pour point) in the diversion of LSHS into fuel oil to achieve a final blend. Experiments were carried out with the refinery samples to establish a profile of variations in these two properties. Various blends were prepared with increasing composition of LSHS in the existing fuel oil product. The property variation profiles are shown in Figures 7 and 8.

The results showed that up to 10–15 wt% and 25–30 wt% of LSHS can be diverted into the existing fuel oil scheme to produce 180 cSt and 380 cSt grades respectively, while meeting product specifica-tions. The model showed better accuracy with the experimental results with R2 ~0.98. Sulphur content is linear and so it can be one of the check points to restrict the percentage of high-sulphur components for producing other products such as LSFO to specifica-tion. The simulation results for maximisation of LSHS into fuel oil streams are shown in Table 6.

Economics In the first instance, around 5000 tonnes of LSHS was diverted into fuel oil in three batches. Two batches were used for 180 cSt-grade fuel oil and one batch was used for 380 cSt-grade fuel oil. By diverting LSHS to fuel oil, crude throughput was sustained and, hence, the derived benefit was quantified using PIMS tools as $110 million.


specific gravity (60/60°F), to esti-mate the characteristics of fuel oil blend compositions.

The input data required for esti-mating the properties of a fuel oil blend that meets product specifica-tions are mainly mean average boiling point, specific gravity and pour point. These data were used in the correlation model for predicting product properties. The sulphur specification can be obtained by optimisation of the stream compositions. The correla-tion models for the prediction of product properties are shown in Equations 1–4.

Experiments were carried out with various compositions of the streams for validation of the model predictions. The experimental results were in-line with the simu-lation results for flash point, kinematic viscosity @ 50°C and pour point (see Figures 4–6). The model showed better accuracy in the prediction of heavy petroleum fraction properties with R2 >0.95.

Table 3

FP, °C KV, cSt PP, °C S, wt%FO compositions, wt% Sim Exp Sim Exp Sim Exp ExpLS-VR HS-VR CLO Kero12 55 10 23 76.42 75 181 175 -8.13 -7 2.6626.88 44.48 10.06 18.58 80.96 80 375.09 370.12 4.83 3 2.27

Blending scheme for fuel oil

FP, °C KV, cSt PP, °C S, wt%FO compositions, wt% Sim Exp Sim Exp Sim Exp ExpLS-VR HS-VR CLO Kero40 20 20 20 77.03 76 195.19 186.28 4.66 9 1.340 27 20 13 87.24 90 378.41 372.09 15.85 18 1.61

Blending scheme for LSFO

Table 4

FP, °C KV, cSt PP, °C S, wt%FO compositions, wt% Sim Exp Sim Exp Sim Exp ExpLS-VR HS-VR CLO Kero20 40 10 30 67.21 68 184.36 176.66 -14.84 -12 2.0430 40 10 20 79.01 80 391.45 382 2.58 6 2.09

Blending scheme for utilisation of existing LSHS for LSFO

Table 5

bpcl.indd 4 10/3/11 13:32:33

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On implementation of the model, quality giveaway was minimised. Also, during the production of export grades of fuel oil (380 cSt), benefits of $0.20 million/month were derived. SummaryIn a surplus resid scenario, a model-ling approach for the diversion of low-value product streams into a


high-value product pool, and/or their use in for making new prod-ucts, presented an immediate solution. By applying this approach, surplus products in stock with a temporarily low value can be utilised to achieve higher margins. The new correlation models devel-oped in this study can be useful for minimising quality giveaway in heavy products.

AcknowledgementThe authors express their sincere thanks to K V Seshadri, ED (MR/R&D) of BPCL for his continuous mentoring on research activities and encouragement. Also, special thanks to Supriya Sapre and P V Ravitej for their constant support in implementing the model.

References1 ICRA rating feature report on Rating Methodology for Downstream Oil Companies, Oct 2009 (www.icra.in).2 Kumar Rajeev, Thorat T S, Chithra V, Rathore V, Peddy V C Rao, Choudary N V, Processing opportunity crude oils — catalytic process for high-acid crudes, Hydrocarbon World, 4, 2, 2009, 64–68.3 Rana M S, Sa´mano V, Ancheyta J, Diaz J A I, A review of recent advances on process technologies for upgrading of heavy oils and residua, Fuel, 86, 2007, 1216–1231.4 A report on optimization of cutter reduction for fuel oil and LSFO, study carried out for Mumbai Refinery (India) at Corporate R&D Centre, BPCL, Greater Noida, CRDC-MR/LSHS-FO-001/2009. 5 A report on development of new schemes for utilization of LSHS, study carried out for Mumbai Refinery (India) at Corporate R&D Centre, BPCL, Greater Noida, CRDC-MR/LSHS-FO-002/2009.

Rajeev Kumar is a Deputy Manager (R&D) with Bharat Petroleum Corporation Ltd, India, specialising in the development of new processes in crude areas, crude compatibility, resid upgrading, biodiesel and biolubricant processes. He holds a master’s in chemical engineering from Indian Institute of Technology, Kanpur, India. Email: [email protected] V is a Senior Research Scientist (R&D) with Bharat Petroleum Corporation Ltd, specialising in crude evaluations, fuel characterisations and new product development. She holds a master’s in chemistry. Email: [email protected] V C Rao is a Senior Manager (R&D) with Bharat Petroleum Corporation Ltd, and has 24 years’ experience in the petroleum, biofuels and petrochemicals industry. He holds a doctorate in chemistry from Indian Institute of Technology, Bombay, India. Email: [email protected] V Choudary is a Chief Manager (R&D) with Bharat Petroleum Corporation Ltd. He has 26 years’ research experience in petroleum refining, catalysis, adsorption and thermodynamics and holds a doctorate in chemistry. Email: [email protected]

Table 6

LSHS, wt% KV @ 50°C, cSt PP, °C SUL, wt% Fuel oil10 154.95 -4.06 2.78 180 cSt (max)14 177.74 -6.00 2.68 30 307.70 7.23 2.28 380 cSt (max)36 378.00 10.61 2.13

Maximisation of LSHS into fuel oil for 180 cSt and 380 cSt grades

1260

1980180016201440

1080900720540360180

3240

39604140

378036003420

306028802700252023402160

0 20 40 60 80 100

kV,

cS

t at

50ºC

LSHS, wt%

0

y = 109.96e0.0343x

R2 = 0.9852

Expon. (KV)KV

180 cSt380 cSt

Figure 7 Variation in kinematic viscosity by diversion of LSHS into fuel oil

0 20 40 60 80 10010 30 50 70 90

PP,

ºC

LSHS, wt%

–10–505

101520253035404550

y = 0.5642x – 9.7006

R2 = 0.9828

180 cSt

380 cSt

Linear (PP, ºC)PP, ºC

Figure 8 Variation in pour point by diversion of LSHS into fuel oil

bpcl.indd 5 10/3/11 13:32:44

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Optimizing CO2 Capture, Dehydration and Compression Facilities

PROCESS INSIGHT

The removal of CO2 by liquid absorbents is widely implemented in the fi eld of gas processing, chemical production, and coal gasifi cation. Many power plants are looking at post-combustion CO2 recovery to meet environmental regulations and to produce CO2 for enhanced oil recovery applications. The fi gure below illustrates actual data of fuel consumption in 2005 and an estimate of energy demand for various fuels from 2010 to 2030. The world energy demand will likely increase at rates of 10–15% every 10 years. This increase could raise the CO2 emissions by about 50% by 2030 as compared with the current level of CO2 emissions. The industrial countries (North America, Western Europe and OECD Pacifi c) contribute to this jump in emissions by 70% compared to the rest of the world, and more than 60% of these emissions will come from power generation and industrial sectors.

Despite the strong recommendations from certain governments, there are very few actual investments in CO2 capture facilities geared toward reducing greenhouse gas emissions mainly because of the high cost of CO2recovery from fl ue gas. CO2 capture costs can be minimized, however, by designing an energy effi cient gas absorption process. Based on the fi ndings of recent conceptual engineering studies, HTC Purenergy estimated the production cost to be US$ 49/ton CO2 (US$ 54/ tonne CO2) for 90% CO2 recovery of 4 mole% CO2 content in the fl ue gas of NGCC power plants. A separate study showed the cost for 90% CO2 recovery of 12 mole% CO2 from a coal fi red power plant to be US$ 30/ton CO2 (US$ 33/tonne CO2). The cost of CO2 recovery from coal power plant fl ue gas is substantially less than that of NGCC power plant fl ue gas due to the higher CO2 content in the feed. The energy effi ciency of a CO2 capture plant depends primarily on the performance of the solvent and optimization of the plant. In traditional fl ue gas plant designs, MEA was the primary solvent and was limited to 20 wt% to minimize equipment corrosion. Recent developments in controlling corrosion and degradation has allowed an increase in the solvent concentration to about 30 wt% thus decreasing the required circulation and subsequent steam demand. A recent DOE study shows the steam consumption for an existing CO2 plant using 18 wt% MEA (Kerr McGee Process) is 3.45 lb of steam per lb of CO2 for amine regeneration. A modern process that uses 30 wt% MEA is expected to use 1.67 lb of steam per lb of CO2 for amine regeneration. The HTC formulated solvent is a proprietary blend of amines and has a lower steam usage than the conventional MEA solvent. Based on the material and energy balances for the plant designed in the recent study, the reboiler steam consumption is estimated at about 1.47 lb steam/lb CO2 using the proposed

formulated solvent without implementing any split fl ow confi gurations. This is much less than the reported steam usage for the MEA solvent. The design of a facility to capture 90% of the CO2 from the fl ue gas of a coal fi red power plant is based on the specifi ed fl ue gas conditions, CO2 product specifi cations, and constraints. Using the ProMax® process simulation software from Bryan Research & Engineering, CO2 capture units can be designed and optimized for the required CO2 recovery using a variety of amine solvents. The following fi gure represents a simplifi ed process fl ow diagram for the proposed CO2 Capture Plant.

The table below presents the main fi ndings for CO2 capture from the coal fi red power plant and the NGCC power plant, each designed to produce about 3307 ton per day (3,000 TPD metric). To produce the same capacity of CO2, only one train with smaller column diameters is required in the case of the coal power plant and two trains with larger column diameters are required in the NGCC Power Plant case. This is mainly due to processing a larger fl ue gas with lower CO2 content in the NGCC power plant. Consequently, a substantial reduction in the capital and production cost was reported for the coal fi red power plant CO2 recovery facility.

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Reducing carbon footprint

Conversion of crude oil into valuable fuel products prima-rily consists of the age-old

distillation process, which is very energy intensive. On top of that, environmental regulations necessi-tate the production of ultra-clean fuels that require more hydrogen for hydrotreatment, which leads to more energy usage and higher CO2 emissions. At the same time, refin-ers face the demand of curtailing CO2 emissions to reduce their carbon footprint. These conflicting demands have put immense pres-sure on refiners to sustain their business with a healthy gross refin-ery margin (GRM) while minimising CO2 emissions. An integrated programme, looking into refinery process units and the whole refin-ery operation, can be adopted to address this complex issue. This programme is focused on reducing the carbon footprint without carbon capture and storage, thus avoiding the need for government legislation to make the project viable.

There are several techniques, including energy audit and the use of energy-efficient equipment and advanced process controls, which are adopted to reduce the carbon footprint in small steps. However, this article discusses methods that address energy conservation in major CO2-emitting units and equipment for significant reduc-tions in CO2 emissions.

Refinery carbon balanceCarbon atoms in crude oil essen-tially end up in CO2 when fuel products from a refinery are consumed by end users. The refin-ery business chain starts from crude

An integrated programme of process integration techniques lowers CO2

emissions levels in refineries through energy savings

TAnmAy TARAphdARTechnip KT India

transportation, followed by refining operations and finally product transportation. Figure 1 shows the typical contribution to CO2 emis-sions at different stages of the refinery business chain, assuming that the refinery is located away from its crude source(s) and that refinery products are consumed within the region where the refin-ery is located.

It is evident that refining opera-tions are the major contributor to CO2 emissions in the entire refinery business chain. Refinery operations

essentially involve the removal of impurities from crude oil and the rearrangement of hydrocarbon molecules to produce fuels of the desired quality. A simple carbon balance for a refinery is shown in Figure 2.

The objective of the carbon balance is to establish the number of carbon atoms lost through CO2 emissions and the development of a mechanism to reduce this loss of carbon atoms. CO2 is emitted primarily from process heaters, util-ity generation systems (mainly steam and power), hydrogen gener-ation units and fluid catalytic cracking (FCC) unit regenerators. Typical contributions towards CO2 emissions from these different sources in a refinery are presented in Figure 3, assuming that the refin-ery has an FCC unit as one of its secondary processing units.

Figure 3 shows that about 50% of emissions are from process heaters, while utility generation systems and hydrogen plant contribute about 30% and 16% respectively. FCC is an endothermic process that requires a lot of energy to perform cracking reactions. This energy is primarily supplied by burning coke


Refining,60%

Prod. transport,

10%

Crudetransport,

30%

Figure 1 Typical CO2 contribution at

different stages of the refinery business chain

Carbon atoms in crude

Carbon atoms in suppl. fuel+feed (if any)

Carbon atoms in products

Carbon atoms in emitted CO2 – defines the carbon footprint of the refinery

Refinery

Figure 2 Overall refinery carbon balance

TECHNIP.indd 1 10/3/11 13:58:58


associated with the catalyst, which leaves little scope for a reduction in CO2 emissions. However, there are greater opportunities for CO2 reduc-tion in three other areas that contribute significantly to emis-sions. Hence, it is important to pay attention to these areas to reduce the refinery’s carbon footprint.

Integrated programmeThe integrated programme starts with linear programming (LP) modelling of all refinery operations to establish the carbon footprint. Figure 4 depicts the masterplan showing different stages of the

integrated programme. LP model-ling should be rigorous to take care of various operating scenarios (different crude cases, product slates and so on) associated with a different set of constraints. It should also consider any future expansion plan for the refinery. It is important to finalise base cases after analysis of all these scenarios.

After establishing base cases, the next step is to look deep into the process units to apply various proc-ess optimisation techniques, such as column targeting and unit pinch analysis, to get as close to the ther-modynamic minimum in terms of energy usage. These process opti-misation techniques ensure a minimum hot and cold utility requirement in process units, thereby reducing the duty of proc-ess heaters, utility boilers and gas turbines. Typically, a process unit is considered to be a three-layered system. The reaction and separation system forms the innermost layer, the heat transfer system forms the middle layer, while the utility system forms the outermost layer. Separation systems, involving distil-lation columns mainly, are optimised using column targeting. Heat transfer and utility systems are optimised using unit pinch analysis and furnace efficiency improvement methods.

Similarly, a hydrogen generation unit, a major contributor to CO2 emissions, is optimised by applying hydrogen network management and state-of-the-art technologies to minimise CO2 emissions. The schemes and strategies that emerge by applying the various techniques and technologies mentioned need careful economic evaluation before their implementation.

The next step is to apply total site integration techniques to optimise the site’s steam power and fuel network to further reduce energy consumption vis-à-vis CO2 emis-sions. Finally, LP modelling of process units, utilities and the H2 network helps in the reduction of the refinery’s carbon footprint.

Scoping analysisThe purpose of scoping analysis is to identify areas where improvements

Process heaters,50%

FCC-regen, 4%

Utilities,30%

Hydrogen plant,16%

Figure 3 Typical CO2 emissions profile of

a refinery

Refinerymodelling

Scopinganalysis

Road map forimplementation

Establishbenefits –

reduction in carbon

footprint

Input

Programme

Unit pinchFurnace

efficiencyColumn

targeting

Refinery data & information

Future expansion

planConstraints

Output

Refinerycarbonbalance

H2management

Total site integration

Projects identifed for improvement

Figure 4 Masterplan of integrated programme for reduction of a refinery’s carbon footprint

TECHNIP.indd 2 10/3/11 13:59:14


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with respect to CO2 emissions are possible and to quantify the extent of emissions reduction. The tools used are column targeting, unit pinch analysis, furnace efficiency improvement methods, hydrogen network management and finally total site integration. Each of these tools and their potential benefits are discussed here with case studies.

Column targetingTypically, a distillation column consumes about 30% of all the energy used in a refinery. Optimisation of a distillation column leads to significant savings in energy. Column targeting is a powerful tool to optimise the design and operation of a distillation column. A tray-by-tray column enthalpy profile is generated from the results of a converged column simulation. This enthalpy profile is known as the column grand composite curve (CGCC). The pinch point of the CGCC is located at the column feed. The CGCC indicates at what temperature heat needs to be supplied and rejected up and down the column. Not all heat needs to be provided at reboil temperature. Some can also be supplied at lower temperatures. Likewise, not all heat needs to be removed at the condensing temper-ature. Partial heat removal at higher temperatures may be appropriate. During the design of the distillation column, once column pressure, the number of trays and feed tray loca-tion are decided, the following can be optimised using CGCC:• Feed conditioning CGCC indicates the possibility of feed preheating or feed precooling. It is also possible to determine the extent of feed preheating or precooling• Reflux ratio CGCC helps to optimise the reflux ratio, while maintaining the product specifications• Side reboiler/condenser The possi-bility of placing a side reboiler/ condenser is clearly shown by CGCC.

A case study illustrates the appli-cation of column targeting and its potential benefits.

A study was performed for a LPG column separating LPG components


0 1000500 1500 2000 2500 3000 3500 4000 4500 5000

CGCC for LPG column

Reduction in Reboiler duty

Scope for feed preheating

Feed

Condenser duty

Figure 5 CGCC for a LPG column


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at which heat must be supplied and can be removed. The heat exchanger network and utility requirement of the process unit is optimised by analysing the composite curves and the GCC.

Furnace efficiency improvement programmeRefinery furnaces contribute to more than 50% of the total CO2 emissions. Hence, it is important to design and operate the furnaces for optimum performance to minimise emissions. Several steps are involved to reduce CO2 emissions from an operating furnace. The first is to establish the operating efficiency of the furnace by simulation using operating data. The second is to identify bottlenecks by studying key components of the furnace such as burners, air preheater, induced draft (ID)/forced draft (FD) fans and refractory. The final step is to develop solutions to minimise energy consumption vis-à-vis CO2 emissions.

Refinery hydrogen managementThe hydrogen generation unit is a significant contributor to refinery CO2 emissions. It contributes in two ways while converting hydrocarbon feed into hydrogen through steam reforming. One is related to the conversion of the carbon content of hydrocarbon feeds to CO2 and the other to fuel firing for supplying the necessary heat for the highly endo-thermic steam reforming process.

Process CO2 emissions from steam reforming can be lowered only by reducing the size of the hydrogen generation unit. This can be achieved by developing a hydrogen balance model across the refinery,

identifying the constraints and flex-ibility of hydrogen usage, and hydrogen pinch analysis to identify possible alternatives for hydrogen reuse from refinery off-gas (ROG).

It should be noted that a careful techno-economic evaluation is required before implementing any project for hydrogen recovery from ROG. The reason is that, on the one hand, it reduces the size of the hydrogen generation unit and thus CO2 emissions and, on the other hand, it degrades the quality of the ROG in terms of calorific value and reduces the opportunity to burn hydrogen to lower CO2 emissions. However, for a larger hydrogen contributor, a dedicated recovery system is justified to reduce overall CO2 emissions.

Three different options are availa-ble to reduce CO2 generated from fuel firing in a hydrogen generation unit. These are combustion air preheating, pre-reforming and post-reforming. Air preheating against flue gas is widely practised to improve the thermal efficiency of the reformer and, hence, reduce CO2 emissions. Air preheat levels of up to 500°C result in reductions of 12–15% in the total CO2 emissions from the hydrogen plant.

Pre-reforming is an adiabatic low-temperature reforming step that helps to lower CO2 emissions from the hydrogen generation unit by reducing the reformer firing duty. Typical reductions with pre- reforming are in the range 5–10%. Pre-reforming combined with air preheating can provide a 20% reduction in CO2 emissions from the hydrogen generation unit.

Post-reforming involves utilisation

from natural gas liquids (NGL). The LPG column was designed with the reboiler using medium-pressure (MP) steam. Column targeting of this column revealed that it is possi-ble to reduce about 50% of the reboiler duty by feed preheating with low-pressure (LP) steam. Figure 5 shows a CGCC of the LPG column.

Since the site had an excess of LP steam, this modification helped to increase LP steam consumption and reduce MP steam consumption, thereby increasing power cogeneration potential on the site.

Once the separation system of the process units is optimised, it is necessary to optimise the heat trans-fer system and utility system using the unit pinch analysis concept.

Unit pinch analysisThe basic principle of unit pinch analysis is to find minimum heating and cooling requirements in a proc-ess unit with a number of hot and cold streams. This obviously depends on the minimum approach temperature differential between the hot and cold streams. Considering the heating and cooling duties of all the streams for different temperature intervals, it is possible to generate the hot and cold composite curves by plotting the enthalpy of the cumulative hot and cold streams against the temperature levels. The composite curves enable calculation of the hot and cold utility targets and an understanding of the driving force potential and location of the heat recovery “pinch”.

The utility targets depend on the value of the minimum approach temperature, DTmin. A small DTmin brings the curves closer together, reducing the hot and cold utility demand and giving lower operat-ing costs. However, this is at the expense of a larger heat exchange area and, hence, higher capital costs. The optimum choice of DTmin depends on the trade-off between capital and energy.

The ideal interface between the process and utility system can be represented by the grand composite curve (GCC). The GCC is derived directly from the composite curve data and indicates the temperature


Units Hot utility, MMkcal/hr Target from unit Actual Savings % Savings pinch analysis consumption potential potentialAVU I/II 29 X 2 33.5 X 2 4.5 X 2 13.4CRU 8.7 9.4 0.7 7.4LRU 2.4 2.6 0.2 7.5Coker B 16.9 17.9 1 5.6SDU 2.9 3.7 0.8 21.5AVU III 43.4 43.8 0.4 1Coker A 27.5 28.1 0.6 2NMPU 7.8 8.7 0.9 10

Results of unit pinch analysis

Table 1

TECHNIP.indd 4 10/3/11 13:59:41


of the high-level heat of the reformed gas used to reform part of the feed in a heat-exchange reformer, thus reducing reformer duty and CO2 emissions. Technip has a proprietary enhanced heat transfer reformer (EHTR) that provides a reduction in reformer firing of up to 25%.

Total site integrationThe total site concept goes beyond the boundaries of individual proc-ess units by taking into account

auxiliary systems for steam and power generation, the cooling water system, other utility facilities, offsite facilities and all other infrastruc-tural support systems. The objective is to integrate all hot and cold utili-ties, along with the power system, in an optimal manner. A GCC for a single process can be split into two at the pinch. Above the pinch, the process requires heat input and is therefore a heat sink. Below the pinch, the process has excess heat and hence is a heat source.

A site source sink profile (SSSP) is constructed by combining all of the site’s heat sources and sinks. To construct the SSSP, the enthalpy versus temperature data above the pinch are extracted from each GCC. The data are then combined for all the units to give one site sink curve. The site source curve is similarly obtained by summation of the sections of the individual GCCs that fall below the individual process pinches. In the site profile, the source and the sink curves do not overlap, thereby improving the integration between processes through the utility system. The site profiles are used to set targets for the total site.

From the source profile, it is possi-ble to identify and evaluate the quantity and quality of steam that can be generated. The requirement for steam can also be established from the sink profile. After matching these two streams, one can look into the possibility of co-generation and the net requirement of high-pressure steam and the corresponding fuel requirement.

A case study for a refinery demonstrates the potential for CO2 emissions reduction by applying unit pinch analysis, a furnace efficiency improvement programme and total site integration.

Case studyStudy basis and methodologyA study was performed on a 140 000 b/d refinery. The refinery has three trains of atmospheric and vacuum distillation units (AVU-I, II and III), a catalytic reforming unit (CRU), an LPG recovery unit (LRU), a NMP extraction unit (NMPU), a solvent dewaxing unit (SDU) and two trains of coker units (Coker A and Coker B). The CRU consists of three sections: the naphtha splitter section, the naphtha hydrotreat-ment section and the reforming section. Six new process units — a diesel hydrotreater (DHDT), a hydrogen generation unit (HGU), an amine regeneration unit (ARU), a sour water stripper (SWS), a sulphur recovery unit (SRU) and a FCC unit — had just been commis-sioned at the time of the study.

Unit pinch analysis was carried out for all nine process units. A

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fi lters, while bringing outstanding longevity.

SSX and HCX Solve The Following Common Problems:Plugging of heat exchangers, piping, towers

Poor hydrocarbon separation from

Deactivation of Claus Catalyst in the SRU

SSX and HCX processes work by removing solids to submicron size particles, while also removing hydrocarbons


TECHNIP.indd 5 10/3/11 13:59:52

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furnace improvement study was conducted for the AVU furnaces, following unit pinch analysis of the units. The six new process units as well as the utility and offsite facili-ties are considered “black boxes”; that is, fixed utility sources or sinks. A GCC for each of the nine process units was generated during unit pinch analysis. The SSSP was then constructed from the GCCs obtained from unit pinch analysis and the black boxes. Different energy-saving schemes were generated during the unit pinch analysis, furnace effi-ciency improvement programme and total site integration. Site-wide steam power balances were gener-ated for each of these cases. A cost benefit analysis was carried out for each energy-saving scheme and an optimal road map was developed for the implementation of the vari-ous energy-saving schemes.

Results of unit pinch analysisThe objective of the unit pinch analysis was to identify the poten-tial for heat recovery in the individual process units, and the generation of a GCC for utility opti-misation and total site integration. The results of the unit pinch analy-sis are given in Table 1.

Total hot utility consumption in these units is 181.2 MMKcal/h and the total savings potential from unit pinch analysis is 13.6 MMkcal/h. In addition, AVU-III has MP steam generation potential equivalent to 6 MMKcal/h. This results in a total savings potential of about 19.6 MMkcal/h, which is 10.8% of total consumption. However, it is not possible to achieve the entire poten-tial savings given the constraints of the existing units. Some of the units for which major energy savings could be achieved are discussed below.

AVU I/II are two identical units with the same capacity. Analysis of the composite curve shows that cross pinch heat transfer equivalent to 4.5 MMkcal/h is occurring in the existing heat exchanger network. About 4.3 MMKcal/h of energy is achievable through modification of the existing heat exchanger network. This will improve the crude preheat temperature, leading

to less firing in the crude furnace and a reduction in CO2 emissions.

Analysis of the existing heat exchanger network and the compos-ite curves of Coker B show that cross pinch heat transfer equivalent to about 1.0 MMkcal/h is taking place. The existing heat exchanger network has been modified to remove this cross pinch heat transfer. This modi-fication saves hot utility equivalent to 0.7 MMkcal/h by increasing the feed preheat temperature.

The potential for hot utility savings in the SDU is about 0.8 MMkcal/h. The existing heat exchanger network in the slack wax section of the SDU is modified by introducing a new heat exchanger to realise the potential hot utility savings.

Analysis of a GCC for AVU-III shows the potential for MP steam generation is equivalent to 6 MMkcal/h. Steam generation equivalent to 5.1 MMkcal/h is prac-tically achievable by modifying the existing heat exchanger network. However, the 0.9 MMkcal/h balance can be achieved only through extensive modifications, which are not economically viable.

Analysis of the composite curves for the NMPU shows that total cross pinch heat transfer occurring in the existing heat exchanger network is 0.9 MMkcal/h, giving an equivalent energy saving potential. Two addi-tional heat exchangers of 0.4 MMkcal/h and 0.3 MMkcal/h duty are introduced in the heat exchanger network. However, recovery of a 0.2

Figure 6 Base case steam power balance

240

400

360

320

280

200

160

120

80

40

–80 –60 –40 –20 0 20 40 60 80 100

Enthalpy, MMkcal/hr

0

Tem

pera

ture

, ºC

Figure 7 SSSP of refinery


TECHNIP.indd 6 10/3/11 14:00:02

MMkcal/h energy balance requires extensive modifications, which are not economically viable.

Hence, the total energy savings achieved through unit pinch analy-sis amount to 15.9 MMkcal/h out of the potential total savings of 19.6 MMkcal/h. Total energy savings achieved are about 8.8% of the total energy consumed.

Results of furnace efficiency improvement studySince AVU-I/II are old units consuming a major part of the total energy demand, it was decided to perform a furnace efficiency improvement study for all three heater types: the pre-topping column furnaces, crude furnaces and vacuum furnaces (six furnaces in all for two trains of AVUs). The pre-topping column furnace is a single-pass cylindrical furnace with-out a convection section. Hot flue gas from this furnace goes to the convection section of the crude furnace, a two-pass box-type furnace. The vacuum furnace is a four-pass cylindrical furnace. There is a common air preheater (APH), where flue gas from the crude furnace and the vacuum furnace is combined to preheat cold air. Simulation with operating data was performed for these furnaces and their combined efficiency was estab-lished at 88.5%. The objective of this

Figure 8 Improved steam power balance


programme was to improve effi-ciency from 88.5% to 92%. Recommendations from the furnace efficiency improvement study to improve efficiency to 92% include: • Reduction of excess air by install-ing a new oxygen analyser at the stack• Replacement of existing burner oil guns• Replacement of refractory in the shield area of the radiant section of the crude furnace and vacuum furnace • Installation of a differential pres-sure system across the fuel oil line and atomising steam line • Replacement of a duct from the pre-topping column furnace to the crude furnace • Addition of new modules to the existing APH.

Overall energy savings in CO2 emissions through improvements to furnace efficiency amount to 1.2%.

Results of total site integrationThe refinery has five boilers, each with a capacity of 75 t/h of steam production at an operating pressure of 40 kg/cm2, for a total steam production capacity of 375 t/h. The refinery also has three turbogenera-tors rated at 12 MW each. These turbo generators have extraction stages of steam to medium pressure (13 kg/cm2), low pressure (5 kg/cm2) and very low pressure of

0.2 kg/cm2. Different combinations of MP and LP steam can be extracted, depending on process requirements and power demand. Two new gas turbines with 21 MW power generation capacity and 48 t/h steam (40 kg/cm2) generation capacity are proposed to meet the increased steam and power demand of the refinery with the commission-ing of six new process units. The base case steam and power require-ment of the refinery is:• HP steam: 341.6 t/h • MP steam: 116.2 t/h • LP steam: 81 t/h • Power: 48.3 MW

Since the new gas turbines have higher co-generation efficiencies compared to the old combination of boilers and turbogenerators, it was decided to maximise use of the new turbines. Balance steam and power are supplied by the boilers and turbogenerators. A base case steam power balance for the refinery is shown in Figure 6. Standard refinery fuel consumption in boilers and gas turbines for the base case is 18.4 t/h. The objective of total site integration was to reduce the fuel consumption in the boilers and gas turbines.

An SSSP for this refinery was generated using a GCC for each unit. The SSSP for the refinery is shown in Figure 7. Analysis of the SSSP leads to three major recommendations:• Generation of additional MP steam from AVU-III and utilisation of this steam in other units, thus reducing fuel oil consumption in the boilers• Utilisation of excess LP steam in a thermal power station, thus reducing LP steam extraction from the turbogenerators, which, in turn, reduces the HP steam requirement for the turbogenerators and hence fuel oil consumption in the boilers• The HP steam header pressure is increased to 40 kg/cm2 (within the design limit of the existing boiler and turbogenerators) from its exist-ing header pressure of 35 kg/cm2 to obtain higher co-generation, thus reducing the HP steam requirement in the turbogenerators and fuel oil consumption in the boilers.

Implementation of these schemes resulted in a reduction in fuel

TECHNIP.indd 7 10/3/11 14:00:14

consumption in the boilers and gas turbines from 18.4 t/h to 17.8 t/h. An improved steam and power balance for the refinery is shown in Figure 8.

This case study indicates a CO2 emissions reduction of about 8.8% using unit pinch analysis, about 1.2% using a furnace efficiency improvement programme and about 2% using total site integra-tion, for a total reduction in CO2 emissions of about 12%.

Identifying and implementingprojectsEnergy savings reduction schemes developed using the various techniques discussed need to be converted into projects. A preliminary cost benefit analysis is performed at this stage to check the feasibility of each scheme. Those that are economically viable and implementable are selected. Ideally, an inside-out method (start-ing from the implementation of selected schemes developed through column targeting, followed by unit pinch analysis, furnace efficiency improvement, hydrogen network management and total site integra-tion) should be adopted to develop a roadmap for implementation.

However, many of the schemes may be independent and do not require others to be implemented first. Generally, the selected schemes are segregated based on payback period, then priority is assigned to each one depending upon its nature. Some schemes may be implemented in parallel, whereas others may be implemented in sequence only. A roadmap for implementation is developed after analysing the payback period and dependency of the scheme on other schemes.

Potential savings in energy and hence a reduction in CO2 emissions are established for selected schemes. An LP model of the refinery is run again with these modifications to re-establish the refinery’s carbon balance and quantify the reduction in its carbon footprint.

ConclusionsThis integrated programme using process integration techniques can

reduce the carbon footprint of a refinery significantly. The techniques can be applied to existing process units and to grassroots designs. A typical reduction in emissions from a process unit is in the range 8–12%, while a typical reduction through total refinery integration is 3–6% of the total energy consumption. This programme not only helps to achieve CO2 emissions targets, but also to improve profitability by saving energy.

References1 Smith R, State of the art in processintegration,PRES992ndconferenceonprocessintegration,Budapest,May1999.2 Linnhoff B, et al, A user guide on process integration for the efficient use of energy, IchemE,1985.3 Ratan S, van Uffelen R, Curtailing refineryCO

2throughH

2plant,PTQ Gas,2008.

Tanmay Taraphdar isGroupLeader,Refinery&Petrochemicals,ProcessdepartmentatTechnipKTIndia,NewDelhi.HeholdsaMTechinchemengfromIndianInstituteofTechnology.



TECHNIP.indd 8 10/3/11 14:00:34

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A promoter for selective H2S removal:

part II

T he Bayernoil Refinery located in Vohburg, Bavaria, Germany, and the Gas

Treatment Process Technology team of BASF SE, Ludwigshafen, jointly conducted a test with a new promoter system in order to improve the performance and energy consumption of the refin-ery’s amine system. This article provides an overview of the refin-ery’s setup and a summary of the test and the results. A more detailed description of the promoter’s char-acteristics appears in part I of this article (see PTQ Gas 2011).

Bayernoil is among the leading manufacturers of mineral oil prod-ucts in Germany and Europe, with

Refinery trials of an MDEA promoter demonstrate low H2S lean loadings and the

option for enhanced process stability during high-sulphur operation

GERAlD VoRbERG, RAlf NoTz and ToRSTEN KATz BASF SE WiElAND WAcHE and clAuS ScHuNK Bayernoil Raffineriegesellschaft

a high flexibility in processing operations and its range of prod-ucts. Nearly two-thirds of the mineral oil products consumed in Bavaria come from Bayernoil’s process units in Neustadt and Vohburg. The two locations are connected by 11 pipelines and work as one establishment. The majority of the refinery’s supplies arrive via the transalpine pipeline (TAL), which starts at Trieste, Italy, and is routed via Ingolstadt to Karlsruhe for a total run of 759 km. The pipe-line delivers crude supplies from Africa, Russia, Venezuela, Saudi Arabia, Norway and other countries to Bavaria. In the Bayernoil opera-tion plants, the staff process crude

oil into products such as mogas, diesel, jet fuel, LPG, heating fuels and bitumen for commercial and end users.

The plants perform a combination of distillation, conversion and upgrading steps. Figure 1 shows a simplified flow chart of the main processes and products.

Refinery amine systemsAmine systems in refineries may consist of multiple absorber systems connected to one or two regenera-tors. Complexity can increase because, in some cases, two or three amine systems are interconnected, on either the lean or the rich side, or even both sides. Additionally,


Table 1

Oilplatform Distillation

Vacuum distillation

Desulphur-isation

Bitumen

Visbreaker

Bitumen

Solid fuel

LPG

BenzineDieselJet A1Heating oil (extra light)

DieselJet A1Heating oil (extra light)

Cracker Upgrading

Crudeoil

figure 1 Unit setup of Bayernoil

basf.indd 1 11/3/11 10:27:20


from generic DEA to various MDEA-based solvents in 1995. Besides that, several revamps of absorber/desorber internals were carried out.

The amine system has a hold-up of approximately 40 tonnes (88 000 lb) and a maximum reboiler steam feed rate of 5.5 t/h (12 100 lb/h). The stripped acid gas is further processed in two Claus units (25–50 t/day, 55 000–110 000 lb/h) with a recovery rate of >97%. Table 2 gives an overview of individual absorber/regeneration conditions.

Challenge and intentionProcessing bitumenic crudes with high amounts of various sulphur species has a range of specific impacts on individual units, in particular those where the treated gas/liquid specification is crucial; for instance, unit 14 with FCC tail gas absorber E1410 and LPG liquid treater E1411.

During the processing of sour, bitumenic crudes, FCC operation becomes more severe; the FCC reac-tor temperature and charges of offgas containing sour gas increase. In a similar way, LPG treatment is affected by higher sulphur concen-trations, which requires more efficient amine treatment and proper amine/hydrocarbon-phase separation. In the worst case, sulphur and sulphur-containing solvent can carry over to further downstream units such as caustic and Merox treatment, ending up in the debutaniser and the C3 value chain.

Consequently, the processing of sour crudes with high sulphur charges (high-sulphur operation) needs to be controlled and responded to by an increase in the reboiler steam feed rate. In this context, the H2S amine lean loading is one of the guiding values.

Bayernoil and BASF jointly decided to conduct a test run with a new promoter formulation added to the generic MDEA solvent using dosing equipment. The promoter formulation itself was provided as an MDEA-diluted premix. The concentration was adjusted and controlled during the test period. This promoter system is non-

liquid treatment of C3/C4 fractions to obtain LPG and further value products such as ethyl-tert-butyl-ether (ETBE, gasoline additive) has become a substantial part of most refineries’ acid gas removal units (AGRU).

Those AGRUs have been expanded over the course of more than 40 years by adding more absorbers for dedi-cated hydrotreating units. Subsequent debottlenecking measures were partly based on solvent swaps (for instance, from DEA to MDEA) and continuous revamps in downstream Claus units. In today’s world-scale refineries, interconnected AGRUs with more than 15 absorbers are common.

As a further challenge, refineries may not run for long with the same

crude and change the output of dedicated value streams on a weekly or even daily basis. System optimisation is not easy and rigor-ous process simulation of an amine system is difficult and sometimes impossible.

Bayernoil’s amine systemBayernoil’s refinery in Vohburg is a so-called bitumenic refinery, able to process huge amounts of crude fractions with high boiling points in a vacuum distillation unit (see Table 1).

The amine system at Vohburg, called BTV (BeTriebsteil Vohburg, Engl. Plant Site Vohburg), comprises four absorbers in three operational parts (see Figure 2). For upgrading purposes, the solvent was changed

Fraction Boiling range °FButanes and lighter <90Straight run gasoline (LSR)/or light naphtha (LN) 90–190Naphtha or heavy naphtha (HN) 190–380Kerosene 380–520Distillate or atmospheric gas oil (AGO) 520–650Residua 650+Vacuum gas oil (VGO) 650–1000Vacuum residua 1000+

Fractions, boiling ranges

Table 1

Absorb. E602 Absorb. E703 Absorb. E1410 Absorb. E1411 RegeneratorTypical feed gas specification Main component C

1-C

6/H

2 C

1-C

6/H

2 C

1-C

6/Olef. N

2, H

2 C

3-C

6 H

2S/CO

2

H2S inlet, v-% 10–24 6–12 2–4 0.5–1.5 >85

H2S inlet, lb/h 1102–984 330 330–660 220–330 1984–3306

H2S inlet, kg/h 500–900 150 150–300 100–150 900–1500

ConditionsAmine flow, klb/h Up to 37.5 Up to 35.3 Up to 30.9 Up to 11 Up to 116.8Amine flow, t/h Up to 17 Up to 16 Up to 14 Up to 6 Up to 53Temperature, °F 86–104 86–104 86–104 95–104 194–248Temperature, °C 30–40 30–40 30–40 35–40 90–120Pressure, psia 73–87 73–87 58–87 261 21.8Pressure, bara 5–6 5-6 4-6 18 1.5

a) Hydrofiner absorber E602: In the Vohburg refinery, the dehydrogenation unit (hydrofiner type, unit 6) processes up to 120 t/h (264 000 lb/h) of middle distillates (diesel, kerosene). Since the refinery tends to convert more heavy crudes with higher sulphur levels, the sulphur content of the middle distillate pool rose to 0.85 wt% in recent years. Note: Bayernoil’s middle distillate pool is generated 6 km away on the Neustadt refinery site by a new mild hydrocracker with about 250 m3/h feedstock plus more than 240 m3/h feed to the dehydrogenation unit (CHD type). b) Absorber E703: This absorber receives various offgases from the gasoline treater unit (unit 7). Feed streams are delivered from both diesel hydrotreating units and from the platformer-debutaniser. Due of the nature of these products, the H

2S content is lower compared to the other streams.

c) FCC tail gas absorber E1410 and FCC-LPG liquid treater E1411: Processing the offgas streams of FCC unit 14 requires a high degree of flexibility and operational experience. Sharp variations in FCC feed properties and specifications are usual, because about 100 types of crude are processed over the year. For instance, the Conradson carbon number of those crudes varies from 2.5 to 4.0, while the sulphur constant varies between 0.4 and 0.8 wt%. As a consequence, operation of this unit requires daily adjustment.

Individual absorber/regenerator conditions of Bayernoil’s AGE unit

Table 2

basf.indd 2 11/3/11 10:27:29


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Figure 2 Scheme of Bayernoil’s AGRU

LPG

F-1413

M-14-66-E

1

Fuel g

asF

-1102M

-11-66-E1

Fuel g

asF

-1102M

-11-66-E1

Fuel g

asF

-1102M

-11-66-E1

FC

C tailg

asE

-1407

2424242224212417244

Purg

e gas

F-603

24

So

ur gas

F-604

24

Fuel g

asF

-1102M

-11-66-E1

24

6P

C 04

6Q

I 07

6F

C 04

6F

I 07

6T

I 042

6F

C 08

7T

I 17

1215

E-602

1220

E-703

120

E-1410

14P

C 22

14Q

I 40

14T

I 00914

PC

25

7T

I 0217

PC

07

6F

I 036

TI 040

6LC

13

6F

I 216

FI 06

2 1136

FI 08

6T

I 041

7LC14

LC 25

7F

03

7T

O 20

14T

I 023

42

14T

07614

T 077

14F

C 54

41

7F

C 07

43

453433

30

14F

C 12

14F

C 13

14T

I 13

14LC

26

3738

36

35

16

E-1411

14F

I 04

14T

I 021

14T

114

14T

I 113

44

J-1448

Purg

e gas

DH

T-A

2AF

C 02

2AT

I 049

2411

So

ur gas

DH

T-A

2AF

I 022A

TI 050

2412

Pure g

asD

HT-B

2BF

C 02

2BT

I 049

2413

So

ur gas

DH

T-B

2BF

I 022B

TI 050

2414

So

ur gas

E-701

2415

32

J-804AJ-804

F-803C

-810-1/2C

-802-1/2

8LI 03

8T

C 02

8F

I 018

PI 03

M

J-805

J-804AJ-801

31

8T

I 003

2454C

ond

ensatesup

ply

Co

ol w

ater sup

ply

Co

ol w

ater return

C-803-1/2

C-807-1/2

123416

E-801

8T

I 001

8LC

01

8T

C 01

8F

I 02

8P

C 02

So

ur water

E-403

H2 S

flare

H2 S

clean

M

51

8T

C 03

8LC

02

C-804-1/2

F-804

8T

I 06

8F

I 06

8F

I 038

FI 03

8F

I 038

PC

01

8T

I 007

J-803J-803A

C-806

46

52

53

LP steam

supp

ly60

8F

I 038

FC

05

8F

I 038

TI 004

Co

ndensate

return61

LPG

E-1408

8F

I 0314

FC

058

FI 0314

TI 068

23

basf.indd 3 11/3/11 10:27:38

Optimisation phase 2, Mar 2010A further adjustment of the overall amine circulation rate from an aver-age 45 t/h to 42 t/h (99 200–92 600 lb/h) was mainly due to a reduc-tion in amine feed to the FCC tail gas treater E1410. A small drop in H2S lean loading at a slightly reduced reboiler steam feed rate of 3.2–3.5 t/h (7000–7700 lb/h) could be observed. Overall sulphur charge (H2S to Claus) to the refinery was still at a low level (~1 t/h to 2200 lb/h), leading to steam/H2S ratios of 3.2–3.5 tsteam/tH2S.

Test run, starting 19 March 2010The test run was accompanied by full lean amine analysis (including metals) and operational monitoring.

The initial phase of promoter dosage took around 10 days. In particular, the presence of heat stable salts (HSS by species) was monitored carefully, as this could also affect the H2S lean loading, which might lead to a misinterpre-tation of the results and wrong conclusions.

Within the test review, the follow-ing values were monitored and discussed:

For the entire amine system:• Overall amine circulation rate, t/h• Reboiler steam feed rate, t/h• Sulphur charge to Claus unit, tH2S/h

volatile and is stable at elevated process temperatures. Thus, promoter losses were only expected by removal through the filter system or losses of the amine itself.

Based on pilot tests, a drop in H2S lean loading was expected, particu-larly during high-sulphur operation, as were energy savings or potential capacity increases. As it turned out, a piping connection upstream of the LPG treater E1411 (on the pres-sure side of the LPG pump) was found to be the only feasible dosing point. Prior to the test run, two operational optimisation phases were conducted to exclude opera-tional phenomena or other effects that could not be attributed to the test phase.

Apart from the performance of the entire amine system, the review and test programme also focused on FCC tail gas absorber E1410 and LPG treater E1411 due to their higher sensitivity to changing levels of sulphur charge.

Optimisation of operational conditions prior to the test runEnergy savings (through a reduc-tion in reboiler steam feed), low sulphur emissions and a reduction in solvent losses during high-sulphur operation were major aspects of the envisaged optimisa-tion of the amine system. The following general measures were discussed. General measures for energy savingsThe adaptation of an adequate circu-lation rate is one of the most effective optimisation measures with respect to energy savings. For multi-absorber refinery amine systems, this requires an evaluation of the equilibrium load for each absorber, especially during high-sulphur oper-ation. Of course, hydraulic aspects are a limiting factor and need to be considered too.

General measures for reduction ofamine lossesThe majority of refiners suffer from tremendous solvent losses, which in some cases can exceed the entire hold-up by several times per year. A big portion can be attributed to amine solubility and carryover in LPG liquid treaters. Reasons include


increased solubility in the organic phase with high concentrations of unsaturated species, followed by foaming and poor phase separation in downstream equipment.

From a design perspective, most common countermeasures are sophisticated water-wash systems and separation sections. Further measures are an adjustment of the right amine-to-hydrocarbon ratio followed by a stepwise optimisation of the amine concentration.

Optimisation phase 1 Nov 2009In a first optimisation phase, the overall amine circulation rate was reduced from an average 48 t/h to 45 t/h (105 800–99 200 lb/h), with the majority of the reduction attrib-

uted to the amine flow in the LPG liquid treater E1411. In parallel, the reboiler steam feed rate was reduced from an average 5 t/h to 3.5 t/h (11 000–7700 lb/h, see Figure 3). This reduction, however, was not only linked to the reduced circulation rate but also to a reduced sulphur charge being proc-essed at that time (H2S to Claus ~1 tH2S/h, 2.2 klb/h). Although the H2S lean loading increased slightly during that period, all required H2S specifications were met, but at a reduced sulphur load. MDEA concentration was reduced simulta-neously from an average 40 wt% to 37 wt% for further improvement in LPG treatment.

50

55

45

40

35

30

25 Cir

cu

lati

on

rate

, to

n/h

r

Date

20

7

8

6

5

4

3

2

Reb

oile

r or

steam

feed r

ate

, to

n/h

r

1

Optimisation1 2

Apr 09

May

09

Jun

09

Jul 0

9

Aug 0

9

Sep 0

9

Oct 0

9

Nov 09

Dec 0

9

Jan

10

Feb 1

0

Mar

10

Reboiler steam rateCirculation rate

Figure 3 Optimisation phases: adjustment of circulation and reboiler feed rate

basf.indd 4 11/3/11 10:27:46


• Steam/H2Sratio,tsteam/tH2S• H2S lean loading in amine,molH2S/molamine• HSSconcentration,wt%.

For the FCC tail gas absorber E1410:• H2Schargetoabsorber,kg/hr• H2Soutlet-concentration,vol-ppm• Total acid gas rich loading,molAG/molamine.For the LPG liquid treater E1411:• H2Soutlet-concentration,vol-ppm.

Findings in the test phase for entire amine systemDuringthetestrun,relevantproduc-tion changes such as increasedprocessing of sour crudes weremade. Based on the improvedperformance figures, an adjustment in the reboiler feed rate and anincrease in sulphur charge enabledan identification of potential benefits and savings for the entire refinery.

A substantial drop in H2S leanloading from 0.006 to 0.002 mol/mol was observed. A positiveimpact on the LPG liquid treaterE1411anddownstreamvaluechaingave the opportunity to process more sour crude by increasing the sulphurcharge (H2S toClaus)step-wise from 1 to 2 tH2S/h (2.2–4.4klb/h).Thereboilersteamfeedrateremained at a low level of 2.9–3.2t/h(6400–7000lb/h).SeeFigure4.

Compared to previous operationscenarios, especially those with high sulphur peaks, the ratio betweenreboiler steam feed and processedH2S dropped significantly from an average3.2–3.5to2.5tsteam/tH2S,lead-ing to obvious savings in energy. In addition, the entire operationalstability, with less impact on H2Slean loading, provided a major benefit (see Figure 5).

HSS, analysed by species, showed a consistent picture and obviously did not lead to cross-effects (seeFigure6).

FFC tail gas absorber E1410 Agoodmonitoringandrecordbasis,as well as a high degree of analysis around the FCC tail gas absorberE1410, enabled reliable and consist-ent evaluation of trends during thetest phase. Table 3 shows someabsorber design specifics.

2.4

2.8

2.0

1.6

1.2

0.8

0.4

H2S

to C

laus,

ton/h

r

0.0

0.012

0.014

0.010

0.008

0.006

0.004

0.002

H2S

lean loadin

g,

mol/

mol

0.000

Apr 09

May

09

Jun

09

Jul 0

9

Aug 0

9

Sep 0

9

Oct 0

9

Nov 09

Dec 0

9

Jan

10

Feb 1

0

Mar

10

Apr 10

May

10

Jun

10

Jul 1

0

Aug 1

0

Start promoter testH2S lean loadingH2S to Claus

Date

Figure 4 Processed sulphur (H2S to Claus) and lean loading after test start

6

7

5

4

3

2

1 Ste

am

/H2S

rati

o,

ton/t

on

0

Apr 09

May

09

Jun

09

Jul 0

9

Aug 0

9

Sep 0

9

Oct 0

9

Nov 09

Dec 0

9

Jan

10

Feb 1

0

Mar

10

Apr 10

May

10

Jun

10

Jul 1

0

Aug 1

0

Start promoter test

Date

Figure 5 Steam/H2S ratio

3.0

3.5

2.5

2.0

1.5

1.0

0.5

HS

S,

wt%

0.0

0.012

0.014

0.010

0.008

0.006

0.004

0.002

H2S

lean loadin

g,

mol/

mol

0.000

Apr 09

May

09

Jun

09

Jul 0

9

Aug 0

9

Sep 0

9

Oct 0

9

Nov 09

Dec 0

9

Jan

10

Feb 1

0

Mar

10

Apr 10

May

10

Jun

10

Jul 1

0

Aug 1

0

Start promoter testH2S lean loadingHSS

Date

Figure 6 HSS and H2S lean loading

basf.indd 5 11/3/11 10:27:57

Findings in test phase for FCC tailgas absorber E1410When processing heavier and more sour crudes right after the test start, a higher feed gas flow and, thus, a higher H2S charge have been proc-essed through the FCC tail gas absorber E1410 in a similar way. By taking the decreased H2S lean load-ing and the adjusted solvent flow into account, the following findings were made: • Despite an increased H2S charge to the absorber, the H2S outlet concentration remained low, around 10–30 vppm (see Figure 7)• As a consequence of higher H2S charges, the total acid gas- rich loading increased to levels above 0.25 molH2S/molSolvent (see Figure 8)• A comparison between similar operation windows displays the effect of lower H2S lean loadings; even at high-sulphur operation with a high acid gas rich loading, the equilibrium/driving force at the absorber top is low enough to keep the H2S outlet concentration at a low level (see Figure 9).

FFC LPG treater E1411 LPG treatment and the following C3/C4 value chain are highly sensitive to alternating sulphur charges. As a consequence, the amount of LPG processed in the liquid treater is adjusted continu-ously using the H2S lean loading as leading value. In case of a sulphur breakthrough or even a carryover of loaded amine, sulphur can be analysed along the downstream chain (see Figure 10) and end in products such as ETBE: E1411 → Coalescer → Caustic treater → Merox treater →C3/C4 splitter → ETBE


Internals 20ballasttraysWeirheight 56mm(fixed)Spacing 610mmDiameter 980mmDesignpressure 6baraAminefeedpoint ToptrayMaterial Carbonsteel

E1410 design specifics

Table 3

240

280

320

200

160

120

80

40

H2S

charg

e,

kg/h

r

0

60

70

80

50

40

30

20

10 H2S

outl

et

concentr

ati

on,

vol.−ppm

0

Apr 09

May

09

Jun

09

Jul 0

9

Aug 0

9

Sep 0

9

Oct 0

9

Nov 09

Dec 0

9

Jan

10

Feb 1

0

Mar

10

Apr 10

May

10

Jun

10

Jul 1

0

Aug 1

0

Start promoter testH2S outlet concentrationH2S charge

Date

0.30

0.35

0.40

0.25

0.20

0.15

0.10

0.05

Tota

l acid

gas

rich loadin

g,

mol R

ICH

GA

S/m

ol

SO

LVE

NT

0.00

60

70

80

50

40

30

20

10 H2S

outl

et

concentr

ati

on,

vol.−ppm

0

Apr 09

May

09

Jun

09

Jul 0

9

Aug 0

9

Sep 0

9

Oct 0

9

Nov 09

Dec 0

9

Jan

10

Feb 1

0

Mar

10

Apr 10

May

10

Jun

10

Jul 1

0

Aug 1

0

Start promoter testH2S outlet concentrationTotal acid gas loading

Date

Figure 7 H2SChargetoE1410andrespectiveH

2Soutletconcentrations

Figure 8 TotalacidgasrichloadingandH2Soutletconcentrations

0.30

0.35

0.40

0.25

0.20

0.15

0.10

0.05

Tota

l acid

gas

rich loadin

g,

mol R

ICH

GA

S/m

ol

SO

LVE

NT

0.00

60

70

80

50

40

30

20

10 H2S

outl

et

concentr

ati

on,

vol.−ppm

0

Apr 09

May

09

Jun

09

Jul 0

9

Aug 0

9

Sep 0

9

Apr 10

May

10

Jun

10

Jul 1

0

Aug 1

0

With promoter

Date

H2S outlet concentrationTotal acid gas loading

Figure 9 ComparisonoftotalacidgasrichloadingandH2Soutletconcentrationsduring

high-sulphuroperation

basf.indd 6 11/3/11 10:28:07

Findings in test phase for LPGabsorber E1411Processing increased sulphur charges and subsequent peaks in H2S lean loadings requires control and some reduction in the LPG flow to keep the H2S outlet concentration within specification. Since the start of promoter dosage, high-sulphur charges could be processed without extended control of the LPG flow, thus providing improved process stability. Moreover, sulphur peaks did not result in peaks of H2S outlet concentration in the treated LPG

(see Figure 11) or downstream prod-ucts such as ETBE.

In addition, a lower foaming tendency combined with improved phase separation in downstream equipment was observed. As a consequence, less amine carryover to downstream equipment (caustic and Merox treatment) led to lower amine losses and savings in refill and dumping costs. However, a detailed evaluation of this effect is proceeding by taking the following root causes into account: • Slightly changed process condi-

tions with a variation in the C3-C5 mix (less foaming?)• Ionic character of the promoter system itself (better phase separation?).

Summary of findings for entire test phaseEnergy savings, a reduction in amine losses, higher sulphur throughput and stable operation during crude oil changes are major drivers consid-ered in the optimisation process presented here. By comparing past operation and performance data for amine system operation, FCC tail gas absorber E1410 and LPG liquid treater E1411 with respective data during the six-month test phase, the key findings are:

Optimisation phases prior to testing• Optimisation of circulation rate for each absorber combined with an overall reduction in circulation rate has led to energy savings (reboiler steam feed rate) of more than 25% for low-sulphur operation (~1 tH2S/h, 2.2 klb/h) and 20% for high-sulphur operation (>1.5 tH2S/h, >2.2 klb/h) respectively to keep the treated H2S concentration at low levels• By reducing the amine concentra-tion from 40 wt% to 37 wt%, amine


2.4

2.8

3.2

2.0

1.6

1.2

0.8

0.4

H2S

to C

laus,

ton/h

r

0.0

120

140

160

100

80

60

40

20

H2S

outl

et,

wt–

ppm

0

Apr 09

May

09

Jun

09

Jul 0

9

Aug 0

9

Sep 0

9

Oct 0

9

Nov 09

Dec 0

9

Jan

10

Feb 1

0

Mar

10

Apr 10

May

10

Jun

10

Jul 1

0

Aug 1

0

Start promoter testH2S ClausLPG H2S outlet concentration

Date

Figure 11 LPG H2S outlet concentration and H

2S sulphur charges (H

2S to Claus)

L1406 13.6bar

–1.2%

21.71%

C1437

E1413

Lauge

Caustictreater

KW

C3

C4

12.48 bar

0.00 t/hr

18%NaOH

ES

E1414

Extractor

14.1bar

50.9%

46.0ºC

C1445

LSC PA

F1435

Causticseparator

71.6%

E1416

ES

FG

LSC

Water wash tower

12.5bar

52.3%

z. E1418

DMW

0.16 m3/hr

0.03 Nm3/hr

690.9 kg/hr

J1444/A

38%NaOH

F1467

E1415

F1430

PCV

71.9%

50.6%

45.0%

Oxidation Column

34.5ºC50.0ºC

Disulphideseparator

FA

E1455

LSC

Surge tank to depropaniserF1419

41.6%

37.3%

10.4 bar

10.4 bar

Pressurised gas

J1436

J1436A

C1453

J1447

Dest.A/B

10.5 t/hr

From LPG absorber

To ETBE

Figure 10 C3/C

4 chain downstream liquid treater E1411 and its separator

basf.indd 7 11/3/11 10:28:18

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losses were lowered to some extent. However, amine carryover from the liquid treater to downstream equip-ment could not be avoided.

Six-month test phase with promoter system• By adding a new promoter formulation with characteristics and properties described in part 1 of this article (see PTQ Gas 2011), an apparent decrease in H2S lean load-ing from 0.006 mol/mol to 0.002 mol/mol could be determined, while specific regeneration energy was reduced• Despite processing more sour crudes and achieving lower H2S lean loadings, the ratio between reboiler steam feed and processed H2S dropped from an average 3.2 to 3.5 to 2.5 tsteam/tH2S, even during high-sulphur operation with H2S charges of up to 2 tH2S/h (4.4 klb/h)• There was no indication of a negative effect on equilibrium load and absorber overhead H2S specifi-cation by the promoter system. A detailed view of the FCC tail gas absorber E1410 during high-sulphur operation demonstrated, that, even at high H2S charges (up to 2 tH2S/h, 4.4 klb/h) in combination with a high total acid gas rich loading (>0.25 molH2S/molSolvent), a H2S outlet concentration of 10–30 vppm could be achieved • Similar results were obtained for liquid treatment in the LPG treater E1411; high-sulphur operation did not result in peaks of H2S outlet concentration in the treated LPG nor in downstream products such as ETBE. Reportedly, additional proc-ess reliability and stability has been achieved due to less foaming and improved LPG/amine phase separa-tion. This is still under investigation.

ConclusionsBesides highly selective AGE, H2S selective acid gas removal has become an increasingly important field for world-scale natural gas plant designs. Contrary to AGE techniques, selectivity and thus CO2 slip is limited and requires some adjustability and flexibility.

From a solvent technology perspective, advanced H2S selective acid gas removal can be carried out


through sterically hindered amines or tertiary amines, both in combina-tion with other amines or a promoter system. Additionally, sophisticated design measures are applied to ensure a reliable operation.

Today’s plants require a very high degree of flexibility in turndown rates, changing feed gas specifica-tions and conditions. In addition, very tightly treated gas sulphur specifications in the low ppm range (<5 ppmv) are mandatory.

The high acid gas partial pres-sures in most natural gas applications, and the requirement for adjustable selectivity, make MDEA-based technologies attrac-tive due to competitive solvent prices and generic designs. For opex, capex and specification purposes, acidic promoted systems are in use to overcome the rela-tively high H2S binding energy and to lower residual H2S loading in the amine. However, current acidified state-of-the-art systems also have their limits, as the improved regen-eration will trigger an inverse effect on the entire capacity. Alternatively, the new promoter system described in this article shows different behaviour: a substantial increase in regeneration ability while hardly affecting absorption capacity.

A six-month test with the new promoter system (BASF’s sMDEA+ technology) at the Bayernoil Refinery in Vohburg has been carried out to optimise operation while demonstrating the promoter’s efficiency. A substantial reduction in energy consumption and an improved process stability at high-sulphur operation are some of the findings of the test phase.

Consequently, the characteristics of this promoter provide an option not only for grassroots designs but also for revamps to fulfil new H2S regulations or even a capacity increase. An example of a recent design case has been provided to explain the correlation between very low H2S lean loadings and H2S treated gas specification.

Further reading1 Bullin J A, Polasek J, Selective absorptionusing amines, 61st GPA Conference, Tulsa,Oklahoma,1982.

2 HarbisonJL,HandwerkGE,Selectiveremovalof H

2S utilizing generic MDEA, 37th Annual

Laurance Reid Gas Conditioning Conference,Norman,Oklahoma,1987.3 Carey T R, Hermes J E, Rochelle G T, Amodel of acid gas absorption/stripping usingmethyldiethanolamine with added acid, Gas Separation & Purification,Jun1991,vol5.4 KohlA L., Nielsen,Gas Purification, 5th ed,GulfPublishingCorp,1997.5WeilandRH,DingmanJC,Effectofsolventblendformulationonselectivityingastreating,45thAnnual Laurance Reid Gas ConditioningConference,Norman,Oklahoma,1995.6 HuffmasterMA,Strippingrequirementsforselective treating with Sulphinol and aminesystems, 47th Annual Laurance Reid GasConditioningConference,Norman,Oklahoma,1997.7WeilandRH,SivasubramanianMS,DingmanJ C, Effective amine technology: controllingselectivity,increasingslip,andreducingsulphur,53thAnnual Laurance Reid Gas ConditioningConference,Norman,Oklahoma,2003.8A Refinery for Bavaria, official Bayernoilbrochure,July2009.

Gerald VorbergisaSeniorTechnologyManagerinBASF’sGasTreatmentteaminLudwigshafenand Project Leader for Selective Acid GasRemoval.He joinedBASF’sCatalystGroup in1997asaGlobalProductTechnologyManagerand holds a diploma in chemical engineeringfromtheUniversityofAppliedSciences(FHT),Mannheim,Germany.Email: [email protected] NotzisaResearchEngineeratBASFSEinLudwigshafen.Heholds a diploma in processengineering from the University of StuttgartandaPhDinCO

2capturefrompowerplantflue

gas by reactive absorption from the Instituteof Thermodynamics and Thermal ProcessEngineering at the University of Stuttgart.Email: [email protected] KatzisheadoftheGlobalTechnologyTeam at BASF SE and coordinates BASF’snew business development activities in gastreatment.HestudiedmechanicalengineeringattheTechnicalUniversityofAachen,Germany(RWTHAachen)andholdsaPhDinevaporationtechnology.Email: [email protected] Wache is a Process Engineer atBayernoil Refinery in Vohburg, Germany.He holds a diploma in chemistry from theTechnicalUniversity(RWTH)AachenandaPhDin chemical engineering on Fischer-Tropschsynthesis and dehydrogenation of middledistillates from the University of Bayreuth.Email: [email protected] Schunk is Plant Manager at BayernoilRefinery in Vohburg and was formerly LeadProcess Engineer for the development andimplementation of theOATS units. He holdsa diploma in process engineering from theTechnicalUniversityKarlsruhe.Email: [email protected]

basf.indd 8 11/3/11 10:28:28

Background Free or soluble water carrying over from the FCC or Coker Main Fractionator overhead systems can cause foaming problems and flooding in the Absorber-Stripper columns. Column de-signs that incorporate effective water removal can free-up ca-pacity for valuable naphtha or LPG production.

The ProblemEven with bulk water removal in the overhead knockout drums or high pressure receivers, water carryover can go hand-in-hand with capacity creep. Higher charge rates can create superficial velocities that exceed design separation capabilities. Carried-over free water only compounds the problem of soluble water that typically drops out of solution inconveniently in the middle of the column. Without an effective design for removal, water can cycle up in the system – condensing and re-vaporizing mul-tiple times in the column and thus taking up valuable hydraulic capacity. Additionally, when the soluble water begins to drop out of solution, two distinct liquid phases form, and the system is susceptible to foaming. Foaming will cause poor separation of components and increase column pressure drop, hindering capacity.

The SolutionThere are several conventional methods that column designers use to remove water from the Absorber-Stripper including: wa-ter draw recessed pans, water draw collector trays, and exter-nal water draw pumparounds. Water separation outside of the column (whether upstream or as a pumparound) will be covered in a future Sulzer Tower Technical Bulletin.

Water draw sumps or recessed pans are typically added to a center downcomer at the top of the Stripper or in the middle/ bottom of the Absorber column, where the temperatures are cool enough to create a free water phase. In the water draw sump design, the primary separation of water from hydrocarbon must occur within the volume of the recessed pan. The small stream of water is drawn from the column and the hydrocarbon liquid overflows the sump to feed the tray below. The biggest challenge of this design is residence time. Free water droplets formed from soluble water are typically very small and require several minutes of residence time to effectively coalesce and drop out of the hydrocarbon phase. Very rarely can the volume of a sump or pan provide that amount of residence time. Ad-ditionally, when two distinct liquid phases begin to form (water drops out of solution), foaming is likely, which further compli-cates the ability to separate water from hydrocarbon.

Water draw collector trays can offer an order of magnitude more residence time than the recessed pan or sump design; however, a chimney tray with a lot of residence time can take up sig-

Sulzer Chemtech

Tower Technical Bulletin Effective Water Removal Can Create Extra Capacity in Your Absorber-Stripper

nificant vertical space in a column. In a revamp design, where the column capacity has creeped over time, several methods can be used to create additional collector tray residence time. The riser and overflow duct heights can be extended to create a larger liquid volume. The downcomers from the tray above can be extended down into the liquid level on the collector so that the water can be discharged closer to the hydrocarbon-water interface, aiding in coalescence. The draw location can be located to where the liquid’s path of travel across the tray is maximized. Perforated dispersion plates can be installed in the water draw sumps to create a low turbulence zone closest to the draw so that any liquid influx does not disturb separation.

The PayoutCombined with an optimized tray or packing design, effec-tively removing water from the system can allow the refiner to increase FCC or Delayed Coker light product yields or charge rates. Sulzer’s retrofit water draw designs can be tailored to address the refiner’s particular constraints so that the effective removal of water can free up column hydraulic capacity for valu-able hydrocarbon production.

The Sulzer Refinery Applications GroupSulzer Chemtech has over 50 years of operating and designexperience in refinery applications. We understand your process and your economic drivers. Sulzer has the know-how and the technology to provide a scrubber internals design with reliable, high performance.

Sulzer Chemtech, USA, Inc.8505 E. North Belt Drive | Humble, TX 77396 Phone: (281) 604-4100 | Fax: (281) 540-2777 [email protected]

Legal Notice: The information contained in this publication is believed to be accurate and reliable, but is not to be construed as implying any warranty or guarantee of performance. Sulzer Chemtech waives any liability and indemnity for effects resulting from its application.


sulzer.indd 1 10/3/11 11:16:57

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Main fractionator revamp

T he Coryton refinery is a Petroplus-owned oil refinery in Essex, England. It has a

crude throughput capacity of 172 000 b/d and can run up to an additional 70 000 b/d of other feed-stocks. The refinery’s major units include atmospheric and vacuum distillation units, a catalytic reformer, a fluid catalytic cracking unit and several hydrotreating units. Since its start-up in the early 1980s, Foster Wheeler has been involved in seven of its eight major turnarounds.

In January 2009, Foster Wheeler was approached by Petroplus and tasked with replacing a 2m band of shell midway up the 5.2m-diameter main fractionator column during a planned October 2009 turnaround (see Figure 1).

The normal method of replace-ment would involve the removal of the top half of the column, but this was ruled out due to a large quan-tity of pipework on one side of the vessel, crane availability and avail-able plot location. Replacement of the mid-section in situ was, there-fore, the only option.

The novel solution was to replace the section in eight petal pieces. A skid track was devised, which allowed the petals to be landed on one side of the column and then skidded round to the congested side (see Figure 2). Working in parallel on opposite sides of the column, a petal of old shell was removed and replaced with a new petal complete with tray supports and downcomers. The sequence in which the petals were replaced was driven by the design of the stiffen-ers required to reinforce each of the openings left by the removed petal.

Space restrictions on site called for an innovative solution to replace a refinery’s main fractionator mid-section

John Payne and Dan Darby Foster Wheeler


Figure 1 Section of main fractionator to be replaced (shown ringed)

Figure 2 Skid track 3D model

foster wheeler.indd 1 10/3/11 14:12:54


The work was completed safely and on schedule, proving that this unique method is a feasible option for the replacement of column sections.

The challenge Due to a change in duty, high levels of corrosion had occurred on a 2m-high section of the main fractionator column, just above the clad lower section of shell. Non-destructive testing (NDT) during previous turna-rounds had found certain corroded sections to be thinner than required. These sections had been over-laid as a temporary measure. As further corrosion was expected, Petroplus and Foster Wheeler decided that replacement of this section with a new band, complete with cladding, was required.

Finding the right solution There are several methods usually considered when a section of a column has to be replaced. These include replacement of the entire column top with new, or replacement of the section through removal of the top half to grade, insertion of a new band, and then rein-statement of the old top half. The short time frame meant that it was not deemed feasible to purchase an entire new vessel upper-half due to the significant engineering and fabrication time required. Even to just replace the corroded band of column using the conven-tional “lift down top half” approach would require the use of a large crane. Since this work was added to the turnaround scope at such short notice, the crane avail-ability and plot space required for such a lift were not guaranteed.

Added to this, a large amount of pipework is present on the north side of the column, including a 42in main overhead line, three 18in lines, four 8in lines and several small bore lines. Each of these lines would require cutting and bracing and rework should the top half be removed. This could add significantly to the planned turnaround duration. Replacement of the mid-section in situ was, therefore, the most feasible option.

Ideally, the section would be replaced as a prefabri-cated band, complete with internals that could be slotted into place. For this to be possible, the existing band would have to be removed while the entire tower top remained above. The total weight of shell, internals and piping above the section in question was an esti-mated 200 tonnes. Supporting this would require a substantial framework.

One option would be a framework linking the lower half to the top. This was prevented by the metallurgy of the column, the lower half being 1¼ Cr with a 304L cladding. Although welding onto 1¼ Cr is possible, it would require local post-weld heat treatment (PWHT). It was felt that this was an option not worth pursuing since it could have led to further problems and compli-cations during the turnaround. The other option would be to create a framework linking the upper half to grade. A big advantage of this approach would be that the framework could be erected pre-turnaround, with the new section ready to be lifted and slotted into posi-tion. However, this would require a huge amount of


more sour crudes into high-value transportation fuels. Additionally, depending on country emissions requirements, minimum amounts of SOx and NOx may be released via the stack. These hydroprocessing technologies can be combined with gas treating/absorption solutions to help cap these emissions.

In response to the sulphur management issues that are arising within the industry, Shell Global Solutions, Criterion Catalysts & Technologies and other alliances introduced The Sulphur Technology Platform, a comprehensive, customisable and integrated sulphur solution to meet emissions and product requirements and facilitate the processing of heavier, more sour crudes. This sulphur technology platform uses deep flash technology that cuts deeper into the bottom of the barrel, resulting in higher vacuum distillate yields compared to conventional vacuum units. These integrated solutions also include effective revamps, such as hydrocracking units converted to process resid feedstocks, optimised hydroprocessing of gasoline, kerosene and diesel components to low-sulphur products and gas treating technologies to remove refinery emissions.

Ultimately, innovative process and catalyst technology solutions are required for processing heavier, more sour crudes while meeting emission and clean transportation fuels specifications, specifically in relation to sulphur (although benefits in relation to other emissions are often achieved at the same time). The key is managing sulphur levels by employing hydroprocessing, conversion and sulphur recovery technologies. By using integrated technology solutions to optimise sulphur management strategies, we can face today’s refining challenges head on.

he European refining industry is coping with declining domestic demand for fuels,

while the imbalance between product supply and market demand persists. This especially applies with regard to a deficit in diesel supply and an excess in gasoline production. The European gasoline surplus in 2009 exceeded 0.75 Mbdoe (million barrels per day of oil equivalent), while the diesel deficit reached about 0.5 Mbdoe.

The US market presents a recurring gasoline deficit of about 0.7 Mbdoe and, since 2008, the refining industry has been exporting diesel.

What are the main causes of such a situation?Although there is no fundamental shortfall between

the structure of US production compared with the local demand structure when expressed as percentages, the

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steelwork, which, although possi-ble, would be fairly cumbersome and might impede other work taking place during the turnaround. Having evaluated the options for replacement of the band in one single section, Petroplus and Foster Wheeler then considered methods for replacing the band in pieces.

The petal approach Cutting windows into shells is common practice to provide access while replacing cyclones in fluidised catalytic cracking units. The team conceived the idea that this method-ology could be adapted to replace the 2m-high band of fractionator by cutting away the existing shell to leave a window and installing a new petal piece in its place. Each petal piece would be prefabricated with its required tray support rings and downcomers. Significant design work was required to confirm the right number of petals and the work that would be involved.

When a window is cut in any shell, the structure must be appro-priately reinforced to ensure the vessel will not fail (eg, buckle) due to the weakening through removal of the shell plate. This is usually done using stiffening ribs either side of the window and replacing the section modulus of the area removed. It follows that where a larger window is cut, the section modulus removed is larger, and therefore more substantial rein-forcement is required. With large arcs, a circumferential stiffener is also required above and below the window. The vertical stiffeners tie into these circumferential stiffeners and this effectively forms a bridge around the cut.

Since the entire circumference of the vessel was to be replaced, a 360-degree ring would be required as the circumferential stiffening piece (see Figure 3).

Since welding below the lower cut line was not possible, the design of this ring needed special consid-eration. As with all turnarounds, it is preferable to do as much work as possible prior to shutdown. A wedging arrangement was devised for the ring (see Figure 4), and the ring was designed so that it could


Figure 3 360-degree reinforcement ring

Figure 4 Wedging arrangement

Figure 5 Section modulus calculations verified using FEA analysis


use of fewer petal pieces might make the fit-up of new elements to existing elements more difficult should the existing vessel be out of round.

How many petals?Options using six and eight petal pieces were considered. For six petal pieces, each petal would form 60 degrees of the fractionator’s circum-ference, and at 5.2m diameter each petal would be approximately 2.7m wide by 2m high. For eight petal pieces, each petal would form 45 degrees of the fractionator and would be approximately 2m wide by 2m high.

For both six and eight petals, since welding below the lower cut line was not possible, the stiffeners could not be extended below and the usual bridging methodology could not be applied so readily. The section modulus calculations were performed and then verified using finite element analysis (FEA, see Figure 5).

It was determined that both cases were possible, with more substan-tial reinforcements required for the six petal piece option. Although the six petal case would require two vertical welds less than the eight petal case, the new to existing fit-up would be easier using eight petals. The time for the two extra vertical welds could be planned into the turnaround, while any problems due to poor fit-up could not. The use of eight petal pieces also lends itself better to cutting opposing windows in the column, maximis-ing work that can be completed in parallel and therefore possibly reducing the turnaround schedule. It was therefore decided that the band should be replaced in eight petal pieces.

The most obvious way to complete the modification work would be to install the vertical stiff-eners at the location for the petal being replaced, cut out the old shell and crane the replacement petal into position for reinstatement. This process would be repeated for the entire band of shell.

However, a substantial amount of large bore piping is present on the north side of the vessel. This piping


be assembled on-line with the frac-tionator hot and, therefore, of larger diameter than during turnaround.

Although it would have been possible to weld a ring to the shell at the upper cut line (since the shell is carbon steel here), the same wedging arrangement as at the lower cut line was employed. This consistency allowed maximum work to be completed pre-turna-

round, while also removing any possibility of the ring being inad-vertently welded to the shell at the lower cut line, which would bring with it major problems, including the requirement for PWHT of the shell.

While replacement of the band in fewer pieces requires less weld-ing, the bigger windows would require greater reinforcement. The

Figure 6 Skid track running behind pipework

Figure 7 Hillman rollers in channel


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prevented crane access to half of the column and meant that a novel method of installing the petals was required to fit new petals behind the piping. Since circumferential rings are required around the vessel to reinforce the cut-outs, it made sense to make additional use of these rings. A skid track system was designed, which would allow the petals to be landed on the south side and moved around behind the pipe-work for installation (see Figure 6).


Two lugs pre-welded onto each new petal located it behind an angle section resting on the upper ring, holding it in position within the skid track and preventing it from tipping backwards (see Figure 8).

The right sequenceWith a method devised for manoeu-vring the petals around the vessel, development of the sequence began. The sequence in which the petals were replaced was driven by the design of the stiffeners required to reinforce each of the windows left as the old petal was removed (see Figure 9).

Following the installation of the first petal, each subsequent petal would have at least one new petal on one side of it, so stiffeners

Figure 8 Retaining angle section

Figure 10 Petal window reinforcementsPetal A reinforced by two internal stiffeners at existing petals B & H. Petals B & F in parallel; Petal B reinforced by external stiffener on new petal A and internal stiffener on existing petal C; Petal F reinforced by two internal stiffeners at E & G. Petals C & G in parallel; Petal C reinforced by external stiffener on new petal B and internal stiffener on existing petal D; Petal G reinforced by external stiffener on new petal F and internal stiffener on existing petal H. Petals D & H in parallel; Petal D reinforced by external stiffener on new petal C and internal stiffener on existing petal E; Petal H reinforced by external stiffeners on new petals A and G; Petal E reinforced by external stiffeners on new petals D and F

2º

47º

92º

137º

182º

227º

272º

317º

N

AB C

D

E

FG

H

Cut line

E

S

W

Figure 9 Petal sequence diagram

2º

47º

N

92º

137º

182º

227º

272º

317º

AB C

D

E

FG

H

Figure 11 New petal landed in skid track

Two different skidding methodol-ogies were possible: hanging from the upper ring or supporting from the lower ring. The proximity of two large nozzles above the upper cut line meant that hanging would not be possible for a complete ring and so this option was not progressed.

Moving the petals With each petal weighing approxi-mately one tonne, manhandling them around the vessel unassisted would be extremely difficult. The solution was to use two Hillman rollers on the bottom edge of each petal (see Figure 7). The rollers used were small enough to fit inside a channel section that sat on the upper side of the lower ring.


welded to either edge of each new petal would minimise the welding of stiffeners to shell on site. However, to ensure the petal would pass the existing pipework, the stiffeners could not exceed 100mm in depth. This limited their section modulus and therefore the rein-forcement offered, meaning that additional internal stiffeners would be required. The fi rst cut would require two internal stiffeners, while only one internal stiffener was required per cut for each

subsequent window, since each new petal would offer reinforce-ment to the cut made next to it (see Figure 10). The internal tray supports and downcomers prein-stalled on each new petal prevented it from passing an external stiffener, so the sequence began with the installation of the opposite side of the vessel to the landing location — the piece furthest past all the large bore piping. Crane access allowed the petals on the southern side of the vessel to be landed directly in

Figure 12 Old petal removed


Figure 13 New petal skidded round to required location


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position from the crane and, as such, no skid track was required on this side. Figures 11–14 show the landing, skidding and installation of a single petal piece.

In addition to working in parallel on opposite sides of the vessel, the schedule was further reduced through assessing the amount of weld required on each new petal prior to commencing the next cut. By assessing the compressive stress through the weld due to the weight above and any wind loading, it was

Figure 14 Petal installed into window

confirmed that only two runs of weld were required on the two circumferential joins. The welds could then be completed as other work was carried out, minimising the bottleneck effect of welding the petal.

As with any fractionation unit, there were internals to consider. For this section, this included four trays that were supported by two lattice beams. While the trays were fully removed, rather than removing the lattice beams too a

bespoke hanging framework was designed, which utilised the chim-ney tray above, with panels removed to run down slinging supports to the upper beam. The lower lattice beam was hung from the beam above it and the two clamped together using a bolted angle section. With the lower beam still bolted to its foot-stool, this arrangement fully supported the lattice beams while also preventing rotation (see Figures 15 and 16).

A tray below the replaced section provided a working platform with-out the need to scaffold internally.All work was meticulously planned utilising a Solid Edge 3D model. This allowed the extremely tight clearances to be considered in full and prevent any unforeseen issues during the turnaround. An anima-tion of the proposed sequence was produced to clearly demonstrate this unique approach.

Making the plan a reality on site Upon delivery to the site, the petals were labelled and then assembled at grade in a scaffold frame. This allowed a precision survey of the new band of shell to be completed, which showed the exact curvature of each petal at both the upper and lower edges.

Although the skid track had been designed to allow it to be installed pre-turnaround, tempera-tures around the fractionator were deemed too hot to work with the insulation removed. This meant that, when the shutdown commenced, the first task was to install the skid track. Since access to the inside was required to install the internal stiffeners, installation of the skid track took place as the column was being prepared for entry and therefore did not extend the turnaround duration.

Following installation of the inter-nal stiffeners, the first window was cut in the vessel. The sequence was developed to allow the removed petal to be skidded out behind the pipework on the track. However, it was deemed easier and safer to cut the petal into small pieces and remove them piecemeal. This


Figure 15 Lattice beam support (from tray above)


allowed the replacement piece to be landed onto the skid track and moved close to position prior to cutting, again helping to reduce the duration of the turnaround.

As each petal was cut, a precision survey of the window was completed and the results incorpo-rated into the 3D model to assess any potential fit-up problems. The survey could not be completed prior to making the cut, since there was significant deflection of the vessel shell as each section was removed. It is believed that this was due to stresses induced into the shell when the emergency weld overlay was heavily laid during the previous turnaround.

The survey work showed how to get the overall best fit for the petal. For example, making it perfectly flush in one corner may lead to an unacceptable fit at another location. This careful planning helped to ensure the right results first time with no rework required. The sequence previously detailed was closely followed throughout the turnaround (see Figure 17) with full support from the mechanical contractor.

As work progressed sequentially around the column, vertical welds were completed in addition to the circumferential welds. After each weld had been completed and fully inspected, it was overlaid internally up to the start of the cladding aside each weld seam.

The sections of tray supports that straddled the weld seams and the lattice beam footstool on the north side of the tower were then installed and appropriate NDT completed. The lattice beams were bolted down to the footstools, and the bracings and hanging supports removed.

With the work on the shell section complete, final inspection was completed prior to traying out. The skid track was removed as a post- turnaround activity.

ConclusionsThe work was completed safely and on schedule, proving that this unique method is a feasible option for the replacement of column sections, offering substantial cost

Figure 16 Lattice beams brace pieces

Figure 17 Petal G replacement

savings when compared to the traditional heavy lift/revamp method.

Solid Edge is a mark of Siemens.

John Payne is Principal Consultant, Static Equipment, with Foster Wheeler. His experience in vessel engineering includes a variety of


refinery projects and specialisation in new FCC units and turnarounds of existing units. He leads a small team working on a number of FCC projects and providing consultancy services to other Foster Wheeler offices. Dan Darby is a Senior Technology Engineer with Foster Wheeler, providing bespoke engineering solutions to complex problems using 3D modelling and finite element analysis, and is a chartered member of the Institution of Mechanical Engineers.


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Small-scale gas to liquids

A ssociated gas and stranded gas — gas reserves located far from existing pipeline

infrastructure and markets — are potentially abundant sources of energy that are commonly squan-dered. Rather than being transported to refineries for process-ing, stranded gas is often just left in the ground. Associated gas produced along with oil is frequently disposed of by flaring — a wasteful and environmentally unfriendly process that is increas-ing subject to regulation — or by re-injection back into the reservoir at considerable expense.

According to the World Bank, 5.25 trillion cubic feet (tcf, approxi-mately 140 billion m3) of associated gas — the equivalent of 27% of US gas consumption — was flared in 2008. The giant gas flares that light the night sky in Russia, Nigeria, Iran, Iraq, Algeria, Kazakhstan, Libya, Saudi Arabia, Angola and Qatar are a highly visible reminder of this waste. A further 12.5 tcf of gas was re-injected. In addition, there is thought to be as much as 3000–6000 tcf of “stranded” gas —unassociated natural gas already found but without cost-effective access to the world market and, therefore, not yet being produced.1,2 The reason? Cost-effective technolo-gies for capturing these wasted resources are not available.

The available options for captur-ing the value of onshore stranded gas include liquifying or compress-ing the gas (to LNG or CNG), then shipping it in specially designed tankers. Both have serious draw-backs at small to medium scales, particularly in terms of cost. The

Microchannel reactor technology is on trial for the small-scale production of liquids from stranded gas

Andrew Holwell Oxford Catalysts Group

economics dictate that new LNG projects are only economically viable for producing gas volumes greater than 500 mcfd over distances of 4200 km (2500 miles) or more. Although CNG is a good option for transporting smaller volumes with throughputs as low as 100 mcfd, over shorter distances in the range of 1000–2500 km (600–1500 miles) it is too expensive to be used when reserves are more remote.

A third wayFor both stranded and associated gas, gas to liquids (GTL) offers a potentially attractive alternative. Like LNG and CNG, GTL densifies the energy to make it cheaper to transport. In principle, GTL prod-ucts can be transported in the existing petroleum infrastructure. But in order to work efficiently, GTL plants must be designed to work on a very large scale. Conventional GTL technology is only economically viable for large-scale plants producing around 30 0000 b/d of liquid fuel and this requires a very large capital investment.

This has proved to be a consider-able barrier to the progress of the GTL industry. For example, although several larger-scale plants have been developed or announced in recent years, only three have made it off the drawing board:• Sasol’s Oryx plant in Qatar was completed in 2006, but, due to an extended start-up period, did not achieve its nameplate production level of 34 000 b/d until late 2009. Costs rose from an initial estimate of $950 million to $1.5 billion• Chevron’s 34 000 b/d plant at

Escravos in Nigeria will cost an estimated $6 billion and is expected to start up in 2013• Shell’s Pearl GTL plant in Qatar, the world’s largest GTL project, with an ultimate capacity of 140 000 b/d and an estimated price tag of $18–19 billion, is expected to start up in 2011.

But thanks to advances in the development of technology for distributed or small-scale GTL tech-nology, a much more flexible and economical option for capturing associated gas, both on- and offshore — in the form of modular GTL technologies — is on the hori-zon. These systems are designed to operate efficiently and economically when producing just 500 b/d. When combined with petroleum crude, the synthetic crude produced from associated gas can be stored on-board or could be transported to shore along with the produced oil via existing tankers and pipelines, eliminating the need for a separate logistics system to transport the gas to market. Small-scale GTL could also prove useful for capturing shale gas resources now being exploited in the US.

Shrinking the hardware and scalingdown the costThe GTL process involves two operations: steam methane reform-ing (SMR), to convert natural gas into a mixture of carbon monoxide (CO) and hydrogen (H2), known as syngas, followed by Fischer-Tropsch (FT) synthesis to convert the syngas into a liquid fuel (see Figure 1). In SMR, the methane gas is mixed with steam and passed over a cata-lyst to produce a syngas consisting


oxcat.indd 1 10/3/11 14:29:43

of H2 and CO. The reaction is highly endothermic, so requires the input of heat. This can be generated by the combustion of excess H2. The syngas is then converted into vari-ous forms of liquid hydrocarbons via the exothermic (heat-producing) FT process, using a catalyst at elevated temperatures.


For small-scale GTL, the challenge is to find ways to combine and scale down the size and cost of the SMR and FT reaction hardware while still maintaining sufficient capacity. And for offshore installations, whether they are drill ships or floating production storage and offloading units (FPSOs), the equipment also

needs to be able to withstand high-intensity wave motion.

Fixed or slurry bed reactors — the two conventional reactor types currently used in FT plants — only function well and economically at capacities of 30 000/day or higher, and the technology does not scale down efficiently. However, new reactor designs, such as micro- and mini-channel reactors, offer a prac-tical way forward.

Both types of reactor consist of compact, modular fixed-bed designs with process channels that are much smaller and provide a greater surface area than conventional FT reactors. Their small size, lighter weight and lower profile are advan-tages in an offshore environment (see Figure 2).

Mini vs micro Development of small-scale GTL depends on finding ways to inten-sify the SMR and FT processes. This relies on developing ways to enhance heat and mass transfer properties and increase their productivity. Since heat transfer is inversely related to the size of the channels, reducing the channel diameter is an effective way of increasing heat transfer and thus intensifying the process by enabling higher throughput. This is the basic logic behind the approaches being taken by the two main players currently working to develop offshore GTL systems, the UK-based company CompactGTL plc and the US company Velocys, a subsidiary of the UK-based Oxford Catalysts Group. Although both are developing integrated SMR/FT systems and are working on the basis of the same principles, the solutions they have come up with are different.

In essence, both companies are developing modular solutions that combine SMR and FR, and both have found ways to reduce the size of the hardware. In standard SMR and FT processes, the reactions are carried out in 2.5–5cm (1–2in)-diam-eter tubes or channels. In the integrated two-stage system being developed by CompactGTL — which the company says is designed to incorporate modules weighing

Steamreforming

Localnatural gas

Gasrecycle

CO/H2

H2

Products

H2O

Steam

Syntheticcrude

FischerTropsch

Burner

Naturalgas

Air

Figure 1 GTL flow diagram

SMR DC-201A-C

CA-453CA-452

CA-253Pipe rack and layers

Velocys FTR 1.5m ID x 8m T-T

Conventional FTR

3m

3m

9m5m

18mapprox.

60m

4m

2m

Cap: after perpendicular

Figure 2 Profile of an FT microchannel reactor assembly compared to that of a conventional FT plant

oxcat.indd 2 10/3/11 14:29:52

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less than 25 tonnes and producing 200 bbl/day of liquids per module — the SMR and FT reactions are carried out in a series of mini-channels, 1 x 0.5cm (0.39 x 0.20in).

In contrast, the Velocys combined SMR/FT system for offshore GTL takes advantage of microchannel reactor technology to shrink the hardware and intensify the proc-esses even further. Here, reactions take place in microchannels, which have diameters in the millimetre range. For example, the microchan-nel FT reactor system, with a footprint of just 2.4 x 8m (8 x 25ft),


has the capacity to produce around 300 b/d. Several FT microchannel reactors, with footprints of just 0.61 x 0.61cm (24 x 24in) can be combined, or manifolded, in paral-lel to increase production volumes.

The small size of the channels reduces the heat and mass transfer distances, thus accelerating process productivity by 10–1000 times. The enhanced heat transfer properties offered by microchannel reactors make this technology ideally suited to carrying out catalytic reactions that are either highly endothermic (such as SMR) or highly exothermic

(such as FT), where heat must be effi ciently transferred across reactor walls to maintain an optimal and uniform temperature to achieve the highest catalytic activity and the longest catalyst life.

In microchannel SMR reactors, the heat-generating combustion and steam methane reforming processes take place in adjacent channels (see Figure 3). The high heat transfer properties of the microchannels make the process very effi cient. These properties are also used to intensify the FT process.

The basic building blocks of the Velocys microchannel FT reactors consist of reactor blocks containing parallel arrays of microchannels fi lled with FT catalyst interleaved with water-fi lled coolant channels (see Figure 4). Since the reactors are able to dissipate the heat produced by the FT reaction much more quickly than conventional systems, a more active FT catalyst can be used.

The microchannel FT reactors take advantage of a highly active FT catalyst developed by Oxford Catalysts to accelerate FT reactions by a factor of 10–15 compared to conventional reactors. As a result, the microchannel FT reactors exhibit conversion effi ciencies in the range of 70% per pass, a signifi cant improvement over the 50% or less per pass conversion rates achieved in conventional FT plants.

Catalyst keyThe key to the improved perform-ance of Oxford Catalysts’ FT catalyst lies in a patented catalyst preparation method known as organic matrix combustion (OMX).

The OMX method combines the metal salt and an organic component to make a complex that effectively stabilises the metal. On calcination, combustion occurs that fi xes the crystallites at a very small size and in a very narrow range. Since the calcination is quick, the metal crys-tallites do not have time to grow and hence remain at the ideal size for these catalytic reactions.

The OMX method produces crys-tallites of an optimum diameter range that exhibit a terraced surface. These are both features that enhance catalyst activity. OMX also produces

Microchannel process technology module

Boiling Heat Transfer

10 times higher heat flux than conventional reactors

High Heat Flux

0.01-0.20”

0.01-0.20”

Microchannel process technology module

Boiling Heat Transfer

Figure 3 Schematic of the SMR microchannel reactor

Cross-flow design

Partial boiling water coolantProcess length ≈ 0.6 mProcess microchannels = 40Coolant length ≈ 0.3 mCoolant microchannels = 425Nominal capacity ≈ 7 litres/day

Figure 4 Microchannel reactor schematic

oxcat.indd 3 10/3/11 14:30:02

fewer very small crystallites that could sinter at an early stage of operation. This results in greater catalyst stabil-ity. Less stable crystallites tend to deactivate quickly, reducing the activity of the catalysts.

Technology on trialBoth the CompactGTL and Velocys technologies have reached the trial stage. According to CompactGTL, the company entered into a joint development testing agree-ment in 2006 with the Brazilian state oil company, Petrobras, to deliver a 20 b/d pilot plant to be tested onshore at the Petrobras Aracaju site in Brazil. The cost of the pilot plant construction and testing project is being funded by Petrobras. The trial was due to begin during the second half of 2010. Industry reports currently suggest that although the CompactGTL skid is now in situ at the Petrobras site in Aracaju, it is not yet operat-ing, although trials are expected to start soon. However, a fully integrated pilot plant at the CompactGTL site at Wilton in Teeside, UK, has been operating continuously and successfully since mid-2008 and the company expects its first commercial plant to begin operation in 2012.

Meanwhile, in March 2010, Velocys entered into a joint demonstration and testing agreement with offshore facil-ity developer Modec, global engineering firm Toyo Engineering and Petrobras, to build and operate a 5–10 b/d microchannel GTL demonstration plant at the Petrobras facility in Fortaleza, Brazil.

Assembly is complete and the plant is due to be deliv-ered in Q1 2011. It will be operated for nine months, starting in Q3 2011. Following a successful demonstra-tion, it is expected that the first commercial deployment will be on an FPSO to mitigate flaring of associated gas resulting from the development of offshore oil fields. Under the terms of this agreement, the total cost, esti-mated at several tens of millions of dollars, will be covered by Toyo Engineering and Modec, while Petrobras will be responsible for the installation and operating costs of the demonstration plant. This demonstration plant, which is designed to accelerate SMR 200-fold and FT reactions by a factor of 10–15, is expected to be up and running during 2011.

ConclusionsThe significant investments in large-scale GTL plants such as Pearl and Oryx demonstrate belief in the potential for GTL to establish itself as major technology to capture the value of large stranded gas deposits. By greatly reducing the size and cost of chemical processing hardware, micro- and mini-channel technology has the potential to extend the use of GTL to capture value from small deposits too, as well as to eliminate flaring or re-injection of associated gas. The trials being undertaken by CompactGTL and Velocys suggest that it may well be possible to reap the advantages of small-scale GTL sooner rather than later.

References1 http://tinyurl.com/Flare-Gas-Stats2 http://tinyurl.com/FlareGasRegsWorldBank)

Andrew HolwellisBusinessDevelopmentManageratOxfordCatalystsLtd.Email: [email protected]


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Simulation of a visbreaking unit

T he visbreaking unit utilises vacuum residue as a feed and converts it into fuel oil. In

this study, the visbreaking unit of Tehran refinery was simulated and then a parametric sensitivity analy-sis was carried out. KBC’s Petro-Sim simulator was used in this study. Initially, the simulator was vali-dated using actual plant test runs and, after tuning, the simulations provided errors of less than 3%. Using the validated simulator, the sensitivity of the yield of fuel oil, gasoline and fuel oil viscosity to variations in furnace temperature (reaction temperature) was investi-gated. The validated simulator can be used to optimise the unit’s oper-ating conditions, to obtain the required product specifications or to study possible changes in the feed conditions, such as the use of diluents.

Visbreaking is a non-catalytic thermal process that converts atmos-pheric or vacuum residues via thermal cracking to gas, naphtha, distillates and visbroken residue. Atmospheric and vacuum residues are typically charged to a visbreaker to reduce fuel oil viscosity and increase the distillate yield in the refinery. The process will typically achieve conversion to gas, gasoline and distillates of 10–50%, depending on the severity and feedstock char-acteristics. Visbreaking reduces the quantity of cutter stock required to meet the fuel oil specifications and, depending upon the sulphur specifi-cations, can decrease fuel oil production by 20%. Additionally, this process can be attractive when it comes to producing feedstock for catalytic cracking plants.1 The

Simulation of a commercial visbreaking unit supports optimisation of the unit’s performance

S Reza Seif Mohaddecy, SepehR Sadighi, oMid ghabuli and Mahdi RaShidzadehResearch Institute of Petroleum Industry

process severity is controlled by the interchangeable operational varia-bles (being essentially a first-order reaction) such as temperature and residence time.2

There are two types of commer-cial visbreaking units: the coil or furnace type3 and the soaker

process. The coil visbreaker is oper-ated at high temperatures (885–930°F, 473–500°C) and low residence times (one to three minutes), while in a soaker unit, by adding an adiabatic drum after the coil furnace, the product is held for a longer time so that the coil is kept


Furnace

Feed480ºC

Quench

Fractionator

Gas oil stripper

Overheaddrum

Gas

Gasoline

Gas oil

Visbroken residue

CW

figure 1 Coil visbreaker

Furnace

Soakerdrum

Feed450ºC

Quench

Fractionator

Gas oil stripper

Overheaddrum

Gas

Gasoline

Gas oil

Visbroken residue

CW

430ºC

figure 2 Soaker visbreaker

ripi.indd 1 10/3/11 14:34:38

at a relatively lower temperature (800–830°F, 427–443°C). Therefore, the heater duty and, in turn, the fuel consumption is only 70% of that for the coil visbreaking proc-ess.4 Worldwide, about 200 visbreaking units are in operation, and Europe alone accounts for about 55% of the total visbreaking capacity.4 Process flows of coil and soaker units are shown in Figures 1 and 2.

The product yields and properties are similar, but the soaker opera-tion, with its lower furnace outlet temperatures, has the advantages of


lower energy consumption and longer run times before having to shut down to remove coke from the furnace tubes. Run times of 3–6 months are common for furnace visbreakers, and 6–18 months is usual for soaker visbreakers. This apparent advantage for soaker visbreakers is at least partially balanced by the greater difficulty encountered in cleaning the soaking drum.5

To effectively design and perfect the control of any process, a simu-lation of the process is needed to predict product yields and qualities

against variables such as space velocity and temperature. The aim of this research was to develop a simple yield predictor model, according to a process simulation, to predict the products with the highest added value — gas, LPG, gasoline, diesel and visbroken fuel oil — in a commercial soaker unit. The main advantage of this work is the investigation of the influence of operating conditions on the yield of products such as LPG and gasoline. The soaker visbreaking unit of the Tehran refinery has been simulated, and the effects of operating varia-bles on the yield and quality of products have been studied.

Process descriptionThe vacuum residuum, which is stored in two tanks at 93°C, is charged to the unit. It picks up heat from the partly cooled product in

the cold charge heat exchanger and accumulates in the charge surge drum. The charge from the surge drum splits and goes through two parallel coils of the heater. The flow through each coil is on flow control. In the hip section of each coil is a steam injection point. The visbreak-ing furnace is constructed in two sections, which are fired independently.

After the coil furnace, the two hot streams converge in a transfer line, then the mixed product is fed into the soaker drum. A quench stream of cooled product is added on flow control, and the combined stream enters the flash section of the flash fractionators. In the flash section, operating at 80 psig pressure, much of the gas, gasoline and distillate

Variable ValueNumber of tubes 128Number of convection tubes 76Number of radiation tubes 52Tube length, m 18.745Outside diameter, m 0.114

Specifications of the coil of the visbreaking unit

Table 1

Variable ValueOutside diameter, m 2.405Length, m 16.5

Specifications of the soaker of the visbreaking unit

Table 2

Variable ValueFeed rate, kg/hr 13 2500Feed density, kg/m3 1006Feet temperature, °C 93Feed pressure, bar 11.89Distillation analysis (ASTM D1160)IBP, °C 2035 vol%, °C 40910 vol%, °C 45720 vol%, °C 50330 vol%, °C 54350 vol%, °C 585Nitrogen content, wt% 0.4Sulphur content, wt% 3.19Asphaltic content, wt% 5.1Kinematic viscosity (100°C), cSt 430Nickel content, ppm 53Vanadium content, ppm 135

Specifications of the feed

Table 3

Furnace Soaker

Str

ipp

er

Sta

bili

ser

Frac

tio

nato

r

Feed

Steam

Lightgas

Light gas

LPG

Tar

Gasoline

Figure 3 Block flow diagram of visbreaking process

Worldwide, about 200 visbreaking units are in operation, and Europe alone accounts for about 55% of the total visbreaking capacity

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formed during the cracking process flashes off. To split some of the light gas content in the fuel oil and gasoline products, stripper and stabiliser columns are used. A simplified process flow diagram of this configuration is shown in Figure 3.

The specifications of the coil and the soaker drum at the Tehran refinery are shown in Tables 1 and 2. The output product from the soaker drum is quenched by the cooled product to prevent more


cracking reactions after the soaker and so inhibit coke formation. The combined stream is transferred to the fractionation tower and side strippers to separate the visbreak-ing products.

Process simulation and validationPetro-Sim can simulate catalytic and non-catalytic processes on an industrial scale.6 It can simulate a visbreaking unit with or without a soaker drum and, in this study, it was used for the simulation and sensitivity analysis of the Tehran refinery’s visbreaking unit.

The soaker-visbreaker unit was simulated as a case study (see Figure 4). This unit was designed to visbreak 20 000 b/d of a mixture of

Table 6

Variable ValueInlet temperature, °C 345.8Outlet temperature, °C 440.5Inlet pressure, bar 7Outlet pressure, bar 31Number of tubes 128Number of tubes (convection zone) 76Number of tubes (radiation zone) 52

Specifications of the furnace

Table 4

Variable ValueRate, kg/hr 150Temperature, °C 316Pressure, bar 44.82

Specifications of the injected steam

Table 5

Variable ValueFlow rate, barrel/day 901Density 0.001CompositionMethane, vol% 36.9Ethane, vol% 24.38Propane, vol% 20.56Isobutene, vol% 4.94n-butane, vol% 5.03Isopentane, vol% 0.77n-pentane, vol% 0.52Hydrogen sulphide, vol% 6.91

Specifications of gas production

Variable ValueFlow rate, barrel/day 1222Density 0.744Sulphur, wt% 3.4Distillation analysis (ASTM D86)IBP, °C 485 vol%, °C 6710 vol%, °C 7630 vol%, °C 11050 vol%, °C 14170 vol%, °C 16390 vol%, °C 18495 vol%, °C 190FBP, °C 201

Specifications of gasoline production

Table 7

Variable ValueFlow rate, barrel/day 18180Density 0.9995Distillation analysis (ASTM D1160)IBP, °C 4525 vol%, °C 50210 vol%, °C 52820 vol%, °C 55930 vol%, °C 584Sulphur content, wt% 3.4Asphaltic content, wt% 8.3Kinematic viscosity (100°C), cSt 80Nickel content, wt% 0.004Vanadium content, wt% 0.0153

Specifications of fuel oil production

Table 8

VBFeed

VBsteam 1

FractionatorV-302

Furnace 301A

VBsteam 2

Fueloil

Water To Visbreaker heaterFurnace

301B

E301 E302

E306

Steam

C1 C2

Off gas

LPG

Gasoline

StripperV-303

StabiliserV-306

R

R

Figure 4 Simulation of visbreaking unit at Tehran refinery

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as a cutter blend with the fuel oil. A comparison of operating data

from the Tehran refinery and from the simulation runs was made to evaluate the simulation of the visbreaking unit (see Tables 9 and 10). These results confirmed the ability of a simulation to predict the desired outputs.

Influence of furnace outlettemperature on product flowsThe effect of increasing the furnace outlet temperature on the flow rates of products at a constant inlet feed rate (132 500 kg/hr) and operating conditions was investigated. According to the results of this exercise (see Figures 5 and 6),

increasing the furnace outlet temperature leads to a decrease in the rate of production of fuel oil and an increase in the rate of gaso-line production.

The effect of increasing the furnace outlet temperature on the viscosity of fuel oil was also inves-tigated and the results are shown in Figure 7.

ConclusionOperating data from the Tehran refinery’s visbreaking unit was gathered to calibrate a simulation of the unit in Petro-Sim. Following confirmation of the results of the simulation, the effect of increasing


vacuum residuum and slop vacuum gas oil, which are both taken from the vacuum tower. The composition of the fresh feed can vary slightly with time from start of run to end of run.

To prepare a simulation of the visbreaking unit, data were gath-ered during a test run of the Tehran unit. The data are shown in Tables 3–8.

As Figure 4 shows, off-gases including C1 and C2, as well as LPG, gasoline and tar are the output streams from the visbreak-ing plant. It is possible to take the gas oil product from the stripper tower, but it is usually blocked so that the gas oil can be mixed

Variable Simulation Actual Rate, barrel/day 887.8 901Hydrogen sulphide, vol% 6.57 6.91

Comparison of gas product between actual data and simulation results

Table 9

Variable Simulation Actual Rate, barrel/day 1230 1222Hydrogen sulphide, vol% 3.322 3.4

Comparison of gasoline product between actual data and simulation results

Table 10

Variable Simulation Actual Rate, barrel/day 18 190 18 180Hydrogen sulphide, vol% 3.1 3.4Kinetic viscosity (100°C), cSt 80.23 79

Comparison of fuel oil product between actual data and simulation results

Table 11


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the furnace outlet temperature on the rate of production of fuel oil and gasoline, and on fuel oil viscos-ity, was investigated. A sensitivity analysis for these values showed that increasing the furnace tempera-ture leads to an increase in the gasoline production rate and a decrease in the fuel oil’s production rate and viscosity. These results and other constraints, such as product quality and furnace operating temperature, can be used to opti-mise the unit.

Since the simulation showed high accuracy when compared with real operating data, the results of an optimisation based on variations in operating conditions and feed have proved to be practical and acceptable.

References 1 Benito A M, Martinez M T, Fernandez I,MirandaJL,Visbreakingofanasphalteniccoalresidue,Fuel,74,1995.2 Kataria K L, Kulkarni R P, PanditA B, JoshiJ B, KumarM, Kinetic studies of low severityvisbreaking,Ind. Eng. Chem. Res.,43,2004.3 Wiehe IA, Process Chemistry of Petroleum Macromolecules,CRCPress,2008.4 JoshiJB,PanditAB,KatariaKL,KulkarniRP,SawarkarAN,Petroleumresidueupgradingviavisbreaking:a review, Ind. Eng. Chem. Res.,47,2008.5 Upgrading Process of Heavy Oil, JCCPTechnicalTrainingCourse,Jun2005.6 Petro-Sim User Guide, KBC AdvancedTechnologies,KBCProfimatic.

S Reza Seif Mohaddecy isaSeniorResearcherin the Catalytic Reaction EngineeringDepartment at the Catalyst Research Centre,ResearchInstituteofPetroleumIndustry(RIPI),Tehran,Iran.Email: Seifsr @ ripi.irSepehr Sadighi worksintheFacultyofChemicalandNaturalResourcesEngineering,UniversityofTechnology,JohorBahru,Malaysia.Omid Ghabuli is a Senior Researcher in theCatalyst Synthesis Department, CatalystResearchCentre,RIPI.Mahdi Rashidzadeh is Head of the CatalystResearchCenter,RIPI.


dp

b,

etarw

olF

Temp, ºF

Figure 5 Sensitivityofproducedfueloilvsthefurnaceoutlettemperature

1200

1300

1100

814 816 818 820 822 824 826

Flo

w r

ate

, b

pd

Temp, ºF

1000

Figure 6 Sensitivityofproducedgasolinevsfurnaceoutlettemperature

79.85

80.05

79.95

79.75

79.65

79.80

80.00

79.90

79.70

79.60

814 816 818 820 822 824 826

Vis

cosi

ty,

cS

t

Temp, ºF

79.55

Figure 7 Sensitivityoffueloilviscosityvsfurnaceoutlettemperature

The simulation showed high accuracy when compared with real operating data

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Modelling for ULSD optimisation

The Chevron Pembroke oil refinery is a complex and large (220 000 b/d) processing

site. This case study examines the improvements achieved by a project with a high return on investment, which resulted in better operation of the process units involved in middle distillate production and higher ultra low-sulphur diesel (ULSD) output. This article describes how as much as a 10% increase in middle distillate production can be achieved essentially without invest-ment in process units or equipment, mainly through the upgrading of cracked feeds, higher average distil-late cut points, optimisation of the process unit and diesel rundown blending. These significant improve-ments, which are estimated at $10 million per year (minimum), have been realised through a team effort involving various departments of the Pembroke refinery, in particular the following groups of people:• Operations organisation, includ-ing white oils, black oils and cracking• Planning and scheduling teams• Process engineering group• Control and information system department, where the process control team resides • Apex Optimisation, supplier of medium-term closed loop optimisa-tion technology.

The Pembroke refinery blends middle distillates directly from the process unit to hydrotreaters. The day-to-day operation of the two downstream hydrotreating units (HTUs) is challenging as through-put has to be maximised subject to a variety of process constraints and the availability of the various feed

On-line coordination and optimisation of refinery process units led to a 10% increase in middle distillate production

Klas Dahlgren Apex Optimisation/Dynaproc an rigDen Chevron henriK TernDrup Apex Optimisation

components, which include kero-sene, several straight-run gas oil streams and FCC product streams such as HHCN and light cycle gas oil (LCGO). The decision-making process for these blends involves several refinery areas and console operators in different control rooms across the site.

Hence, as part of the improve-ment programme, a new large-scale, multi-unit coordination tool (GDOT) was implemented. The GDOT software supplied by Apex Optimisation is used within

Chevron Pembroke for medium-term optimisation problems. This system, which is basically an on-line refinery linear programming (LP) model, runs in closed loop and has been in service since late 2006 at Pembroke with essentially 100% utilisation even during significant changes in crude slates.

This article describes the issues, challenges and constraints that the Pembroke refinery faces when ULSD becomes the most valuable product most of the time.

Like many other ULSD-producing refineries, the Pembroke site blends middle distillates directly from the process unit rundown lines prior to hydrotreating. The main advantages of this approach, compared to a conventional batch blending system, are lower tank storage and manpower requirements, and the swing cuts of the upstream process unit can be optimised in real-time to operate the hydrotreaters at multiple ULSD quality constraints. However, the rundown blending approach also results in a more challenging day-to-day operation of the downstream HTUs, especially if the throughput is to be maximised subject to a variety of process constraints, taking into account the availability of the various feed components.

The Pembroke diesel system has two HTUs, HTU1 and HTU2, which are fed by a rundown blending header. The configuration of the Pembroke refinery diesel system is shown in Figure 1.

The operations department with the console operators is organised into three areas: black oils (crude and vacuum distillation), white oils (hydrotreaters and naphtha process-ing) and the cracking area. Traditionally, the scheduling department advises the area opera-tors, through a daily schedule, on how to set diesel blending compo-nent flow rates and middle distillate component cut points from the crude and vacuum distillation units. Now, using the medium-term opti-misation tool, the coordination of the units is done automatically on-line and, instead of fixing diesel component flow rates, the scheduler


Day-to-day operation of the two downstream hTus is challenging as throughput has to be maximised subject to a variety of process constraints

dynaproc.indd 1 10/3/11 14:41:33

specifies the product specifications and the key swing cut points to be optimised by the system.

Maximising ULSD production Feed quality management is one of the keys to maximising the perform-ance of a HTU, subject to constraints. A highly constrained HTU can be very sensitive to incre-mental changes in the individual component flows of the feedstock. Therefore, the challenge is not just to push the rate through the unit to the maximum, but to establish the optimum blend that enables throughput to be maximised subject to product quality constraints. A different feedstock composition will significantly change the hydrot-reater operation, which will have an impact on the maximum possi-ble feed rate dictated by unit


constraints. The following list high-lights some of the difficulties with feedstock components and constraints observed at the Pembroke refinery’s HTUs• Maximisation of cracked feed (eg, LCGO) results in higher reactor temperatures and high hydrogen consumption• Maximisation of kerosene and light cracked feeds results in constraints on the operation of the product stripper columns. In the past, this has caused operational problems, including a positive doctor test requiring costly diesel reprocessing, which led operations to put conservative limits on the throughput of the unit• Maximisation of heavy feedstock, such as the back end swing cuts from the crude and vacuum distillation units, requires high reac-

tor temperatures to meet the sulphur specification, which aggra-vates HTU heater constraints.

There are essentially 18 variables available to control the production rates and the qualities of the three middle distillate products of the refinery: kerosene, diesel and gas oil. During the winter period, it is typically best to run at minimum flash points on all three products and at maximum cloud points on gas oil and diesel, while also meet-ing production rate targets on one or two of the three streams. In the summer period, the 95% point or density typically replace the cloud point as the back end constraint on the diesel.

The sulphur content is controlled within the diesel hydrotreater, but other diesel qualities such as density, cloud point, flash point

HTU 1

ULSD

HTU 2

ULSD

FCCU

Gasoil import

Gasoil export

HCCN

LCGO

VBD

VBU

VGO

VD

HGO (HD)

LGO (LD)

Kero

Naphtha

CDU

Vacuum unit

Figure 1 Pembroke refinery diesel system

dynaproc.indd 2 10/3/11 14:41:44

and distillation must be controlled upstream of the hydrotreater; that is, by the side stream cuts and the feed blends.

Given that both HTUs are heavily constrained, there is a strong incen-tive to utilise all available hydrotreater capacity and to avoid reprocessing as a result of off-spec production. Hence, the optimum operating strategy for the diesel system can typically be summarised as follows:• Always keep diesel production on grade with minimum giveaway• Fill the hydrotreater capacities with available feedstock subject to constraints• Maximise cracked feed over straight-run middle distillate• Maximise heavy feed components over lighter components.

Operator training and coordination issuesOne of the important challenges is to train the CDU, cracking and VDU operators so that they are more aware of the operating objec-


tives and constraints of downstream units when making moves on their units. In large and complex refi ner-ies, operators traditionally control to the targets that have been speci-fi ed for their particular unit and are not necessarily aware of operational constraints on downstream units and any opportunities to minimise giveaway on the product rundown lines.

The training required for console operators is mainly related to an understanding of the concept of on-

line coordination of multiple process units, intermediate product fl ow rates and quality targets. The GDOT tool can signifi cantly improve this decision-making proc-ess, making sure that the upstream units are optimised in a coordinated manner and keeping all units within acceptable operating ranges. This enables the operators to work better together and achieve an improved overall performance of the refi nery.

Chevron has selected the GDOT

Figure 2 Diesel production from HTU1 and HTU2, barrels per stream day


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it was decided to let the four console operators use the interface of the 12 MPC controllers (based on Aspen DMCplus) instead. Also, GDOT has been configured to track the status and the limits of the MPC variables. By using the existing interface, operators did not need to be retrained, resulting in a smoother transition and faster operator acceptance. However, the user interfaces of the individual MPC

controllers will obviously not give the complete picture of a current optimised solution for the entire diesel production system. Hence, customised database displays have been made available to operations and production planning staff.

Robustness and maintenance What maintenance effort is required for an advanced dynamic coordina-tion system like GDOT? Our experience has been that once commissioned, it is important to allow some time for fine-tuning of the system for different operational scenarios, some of which may not have been considered in the origi-nal design. Also, during the transition phase, where a team effort is required to use the system to gradually move the operation towards the global optimum, addi-tional training, discussions and possibly further adjustments may be required. After that, however, the installation requires almost no maintenance.

Project executionProcess data and models available from the existing MPC systems were sufficient to develop most of the models required for GDOT. This meant that the project was able to proceed with minimum impact on the refinery’s operation. The preparatory site work included the installation and configuration of a server on the process control network.

The building of the GDOT model is typically done by the vendor (Apex Optimisation) in their offices, following a two-week kick off meet-ing and information-gathering visit, including interviews with refinery departments such as operations, planning and process engineering.

The second visit is typically dedi-cated to the software installation and the loading of the model and optimiser, which is then typically put on-line in an advisory (open loop) mode. The following visits thereafter are dedicated to closed commissioning and fine-tuning.

During the first year of system operation in closed loop, follow-up visits and remote monitoring are performed to make sure the


tool for coordination of production areas such as the diesel system. It is important to note that the system does not require any specialised staff and is maintained by the same process control engineers who are responsible for the multi- variable predictive control (MPC) applications.

GDOT provides an on-line console interface for engineers and operators. For this project, however,

42

42.5

41.5

41

40.5

2004 2005 2006 2007 2008 2009 201040

Figure 3 Kero flash point,°C

0.845

0.850

0.840

0.835

Jan

05

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05

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Figure 4 Final tank density

360

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Figure 5 Tank 95% distillation, °C

Spec change to EU ULSD only Spec

Spec

dynaproc.indd 4 10/3/11 14:42:04

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optimiser is performing at its best under all common constraint scenarios and for the typical operat-ing strategies, such as different modes of operation.

Post-audit resultsThe key objective of this project was to optimise the middle distil-late cut-points, the uplift of cracked feed and the middle distillate blending to achieve a higher yield of ULSD and gas oil sale. It should be recognised that the entire Pembroke organisation has contrib-uted to the significant improvement in the operation of the refinery. GDOT is merely a tool that the organisation is using to consistently implement a more profitable oper-ating strategy. Figures 2 and 3, which show data from 2004 to 2010, should give an understanding of the improvements achieved.


Analysing quality data from the final diesel shipment indicates how well the system and organisation perform over a longer period. It is important when analysing long-term data to consider significant product specification changes. One of those changes was in 2006 when UK ULSD was changed to EU ULSD 0.845 density specification. Figures 4–7 show the final quality improvements for ULSD diesel sold to customers.

The change in specification from UK ULSD to EU ULSD involved an improvement in sulphur of 1 ppm. This might not sound much, but at this level it is equivalent to a change of 4°C on average reactor bed temperature and significantly increases catalyst lifespan.

ConclusionsThe diesel production improvement project has been a success, with

overall benefits valued at $10 million, including a large increase in diesel production. The hard work of many people from various areas of the Pembroke refinery has contributed significantly to this success.

The main benefit of the GDOT system is that it allows operational instructions and strategies to be consistently implemented, minute by minute, day and night, driving the units towards more profitable operation and improving the competitive position of the refinery. The system deals with daily opera-tional issues. The GDOT modelling approach, using dynamic non-linear models, is capable of adapting to all expected and unexpected operating scenarios and has proven to be very robust. Uptime statistics are excellent.

The ULSD improvements project did not require any major unit upgrades or revamps and took nine months from start to finish. The payback for this improvement project was achieved in a few weeks. Another conclusion from this project is that multi-unit coor-dination systems should be considered as one of the next logical steps for a refinery to improve operation further with an existing configuration.

Klas Dahlgren is a Principal Control and Advanced Dynamic Optimisation consultant with Apex Optimisation, Aberdeen, UK. He has 25 years’ experience of day-to-day process industry operations, implementing advanced control and optimisation system solutions. Email: [email protected]

An Rigden is Process Control Team Leader at Chevron Pembroke refinery, UK. She joined the control group seven years ago, having previously worked mainly in business planning and refinery optimisation, and has more than 10 years’ experience with linear programming models. She has master’s degrees in chemical engineering and engineering management and is a chartered engineer and member of the Institute of Chemical Engineers. Email: [email protected]

Henrik Terndrup is a Principal Consultant with Apex Optimisation, with more than 25 years’ experience in the operation and optimisation of industrial plants, including applications of advanced dynamic optimisation. Email: [email protected]

–4

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Figure 6 Final tank cloud point

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Figure 7 Final tank sulphur, ppm

Winterspec

Spec7 ppm

dynaproc.indd 5 11/3/11 14:49:03

NPRA 2011 Conference Schedule

Join colleagues and customers at NPRA’s industry-leading conferences.

Meet the Right Peoplefor Bus iness Success

Register today at npra.org/meetings or call us at (202) 457-0480 for more information.

Security Conference and Exhibition

March 1–2InterContinental HotelHouston, TX

National Safety Conference and Exhibition

May 10–11Omni Fort WorthFort Worth, TX

Labor Relations/Human Resources Conference

May 25–26Convention CenterDenver, CO

Reliability & Maintenance Conference and Exhibition

May 24–27Convention CenterDenver, CO

Q&A and Technology Forum

October 9–12JW Marriott JW Marriott San Antonio Hill Country Resort & SpaSan Antonio, TX

Environmental Conference

October 24–25New Orleans MarriottNew Orleans, LA

International Lubricants and Waxes Meeting

November 10–11Hilton Post OakHouston, TX

Annual MeetingMarch 20–22 « Marriot Rivercenter « San Antonio, TX

International Petrochemical ConferenceMarch 27–29 « Grand Hyatt « San Antonio, TX

Public Policy Conference

May 5–6Washington Court HotelWashington, DC

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Catalyst development is key to increasing plant efficiency andmaintaining profit margins. The Global Catalyst TechnologyForum will promote the exchange of ideas and informationbetween catalyst companies, process licensors and operatingcompanies to help develop the right partnerships for long termsuccess.

The Bottom of the Barrel Technology Conferences havegained a reputation as the most important residue upgradingevents in the industry. This year the event will include a strongfocus on coking, including delayed coker technology, cokerrevamp projects, profit-making uses for delayed coke andstorage/transportation challenges of coke.

The Green Refining and Petrochemicals Forum will provide anexcellent opportunity for Oil companies from all sectors to beupdated on and discuss the environmental issues surroundingour industry and will focus on the latest technologies, issues,trends, regulations and strategies for reducing greenhouse gasemissions.

Visit our website for further detailswww.europetro.com

Tel: +44 (0)20 7357 8394 · Email: [email protected]

organised by

supported by

MEETING REFINING & PETROCHEMICAL CHALLENGESTHROUGH NEW CATALYST & TECHNOLOGY INNOVATIONS

THE LEADING GLOBAL FORUM FOR RESIDUE UPGRADING

MEETING THE INDUSTRY’S ENVIRONMENTAL CHALLENGES

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A liquefaction and shipping site for LNG has automated its allocation of hydrocarbons to multiple customers. By replacing a chiefly manual approach to hydrocarbon allocation Atlantic LNG, based in Trinidad & Tobago, has increased the speed and flexibility of a complex change management task. Atlantic has installed an application from EnergySys to automate its hydrocarbon allocation processes for four LNG trains more efficiently and flexibly. The new approach is also said to provide a detailed audit trail in support of a process that involves multimillion-dollar decisions daily.

One of the largest LNG liquefaction plants in the world, Atlantic purchases gas from suppliers and sells freight on board to customers from three of four lique-faction trains at its Point Fortin site, and supplies LNG on contract to buyers that are subsidiaries of BP, BG, Repsol YPF and GDF Suez.

The liquefaction trains are owned separately, but a joint use and operating agreement among owners allows them to be run in combination, sharing common assets and costs. The company’s total production capacity is 15 million t/y, employing an improved version of the ConocoPhillips Optimized Cascade Process. Plate-fin heat exchangers support the cooling process, with propane, ethylene and methane used as refrigerants. GE Frame 5D gas turbine drives are

employed and parallel compressor units are used for two-train reliability within the single-train design. Atlantic LNG is the largest supplier of LNG to the US.

The hydrocarbon allocation team at Atlantic had previously used a manual approach, which took several hours to complete and relied on data entry to move information between systems. EnergySys applied an automated system based on its hydrocarbon alloca-tion system. All metering data is automatically extracted and loaded into the application, where it is processed based on custom business rules and reports that are generated automatically.

Atlantic wanted spreadsheets for encoding its busi-ness rules and for reporting, and was reluctant to adopt a hard-coded system that could not easily be main-tained by its commercial team. Using a spreadsheet

Reduce gasoline cutpoint

Venturi steam traps solve condensate pressure problem

engine and reporting tools that are integral to the basic application, Atlantic can review and modify the busi-ness rules that are used to produce reports without specialist programming knowledge while still following a detailed change control process.

The new system automatically creates statements that allocate hydrocarbons to each gas supply contract and sends them by email to the finance department to attach pricing information. The new approach enables segregation of duties between the commercial and finance teams.

Changes in local legislation, in line with an EU direc-tive on cogeneration, have led Galp Energia to upgrade one of Portugal’s largest cogeneration sites, serving its Matosinhos refinery. The upgraded plant will also supply Portugal’s power grid. In line with the move, GE has supplied two Frame 6B gas turbines for the Matosinhos cogeneration plant near Porto. Replacing older oil-fired technology at the site, the gas turbines will increase the plant’s efficiency and reduce its envi-ronmental impact in line with the Portuguese Government’s regulation to promote efficiency and reduce CO2 emissions.

The Frame 6B gas turbines are expected to provide all of the steam requirements for the nearby refinery. The Matosinhos cogeneration turnkey project is being developed for Galp Energia by the consortium Ensulmeci-Efacec Cogeraçao do Porto, ACE, formed by two of the major Portuguese engineering, procurement and construction companies.

The modernisation of Galp Energia’s refineries supports growing interest in Portugal to implement cogeneration as a more efficient, cleaner way to produce electricity and meet process steam needs. According to the refiner, the use of 6B technology at the Matosinhos plant will enable it to avoid emissions of over 400 000 t/y of CO2.

The new gas turbines are the third and fourth GE 6B gas turbines selected by Galp Energia to modernise power and steam production for its refineries in Portugal. The first two 6B machines were installed at a cogeneration plant that supports the Sines petrochemi-cal industrial park. Both the Sines and Matosinhos cogeneration plants have power capacities of 80 MWe.

Scope of supply for the Matosinhos project also includes gas turbine auxiliary equipment, technical advisory and training services, during construction and commissioning. A contractual service agreement also has been signed, providing for ongoing maintenance of the GE gas turbines.


Technology in Action

Gas turbines raise cogeneration efficiency

LNG site automates hydrocarbon allocation

The new approach enables segregation of duties between the commercial and finance teams

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All change for BP’s US refiningBP is to restructure its refining and marketing business in the US and divest two of its local refineries. It intends to seek buyers for the Texas City refinery and the Carson refin-ery near Los Angeles, California, together with its associated inte-grated marketing business in southern California, Arizona and Nevada. Subject to regulatory and other approvals, BP plans to complete the sales by the end of 2012, halving its US refining capacity.

The company says it plans to focus future downstream invest-ment in the US on further improving and upgrading its other, more refining and marketing networks in the country, based around the Whiting, Indiana, and Cherry Point, Washington, refiner-ies and its 50% interest in the Toledo, Ohio, refinery. According to the company, these refineries have greater flexibility to refine a range of crude oils, including heavy grades, and on average are more capable of diesel production.

BP says it intends to sell both the Texas City and Carson refineries with its marketing network as going concerns and expects signifi-cant interest in them. The sales will be subject to regulatory and other approvals.

The Carson refinery is the heart of an integrated fuels chain, cover-ing southern California, Arizona and Nevada. The refinery has 265 000 b/d refining capacity and supplies some 25% of the gasoline demand in Los Angeles. It became part of BP through the 2000 acquisi-tion of ARCO, and employs 1200 staff and 500 contractors.

The assets associated with the Carson refinery include BP’s inter-ests in a cogeneration plant on the refinery site, crude and product terminals, as well as marketing interests. As part of this sale, BP expects to divest the ARCO brand (although retaining brand rights for

P

northern California, Oregon and Washington) and to retain owner-ship of and license the ampm brand.

The Texas City refinery became part of BP with the 1998 merger with Amoco. It is a complex refin-ery with 475 000 b/d refining capacity, the third biggest refinery in the US, with gasoline manufac-turing capability equivalent to approximately 3% of US produc-tion. The refinery employs 2200 staff, and numbers of contractors can vary between 2000 and 4000.

During the last few years, over $1 billion has been invested in moder-nising and improving the plant. However, Texas City lacks integra-tion into its marketing operations, says BP. The assets to be sold off associated with the sale of Texas City also include the cogeneration plant. BP intends to keep the Texas City Chemicals complex adjoining the refinery.

BP is currently in the process of carrying out a number of invest-ments in its other US refineries, including a programme to transform its 405 000 b/d capacity Whiting refinery, increasing its capability to process heavy Canadian crude; a clean diesel upgrading project at its 240 000 b/d Cherry Point refinery; and the addition of a continuous catalytic reformer to its 160 000 b/d capacity Toledo refinery, a joint venture with Husky Energy.

Benzene reduction start-upValero has successfully started up advanced reformate splitters at three of its US refineries. The three Mobile Source Air Toxic (MSAT) II Benzene Concentration Units —located at Valero’s Port Arthur and Sunray, Texas, and Memphis, Tennessee, refineries — incorporate KBR’s advanced reformate splitter Dividing Wall Column (DWC) tower designs. A fourth unit is located at Valero’s St Charles refin-ery in Norco, Louisiana, and will be commissioned later this year.

The units are designed to

Industry News

concentrate and remove benzene from gasoline streams to meet US regulations limiting the benzene content of gasoline. The concept was developed by Valero, while the DWC towers were designed and optimised by KBR for each project and have enabled Valero to meet their regulatory requirements.

KBR has been awarded a contract by Saudi Aramco to provide front-end engineering and design (FEED) and project management services (PMS) for its grassroots Jazan refin-ery, a 400 000 b/d facility to be located in Jazan, Saudi Arabia.

KBR will provide FEED and PMS services to develop the process design, layout, integration and opti-misation of the facility, develop equipment and material specifica-tions, prepare EPC bid packages and develop an estimate for the construction of the refinery. KBR will also assist Saudi Aramco in overseeing, managing and directing the work-related activities for all phases of the Jazan Refinery and Marine Terminal Project.

The Jazan Refinery and Marine Terminal Project is intended to provide a foundation industry for the Jazan Economic City develop-ment and to provide additional refined products to meet growth in domestic demand within Saudi Arabia. Work on the project began in February.

Mass transfer, separation deal renewedShell Global Solutions International and Sulzer Chemtech have extended their strategic alliance agreement, originally established in 2000. Under the previous agreement, Sulzer Chemtech became the worldwide licensee for Shell Global Solutions’ high-capacity tray and phase separa-tion technology. The agreement proved mutually beneficial, say both parties, who have agreed to continue and extend their collaboration to include support for, and formalise joint development of, new mass


ind news copy 7.indd 1 11/3/11 11:49:33

transfer and separation equipment. The two companies say that the

extended alliance can result in improved utilisation of resources, covering process requirement and utilisation to manufacturing and marketing.

Shell Global Solutions recently signed three licence agreements with the state-owned North Refineries Company of Iraq in Kirkuk, north-ern Iraq. The company will provide a process licence and basic engineer-ing package for a kerosene hydrotreater, a diesel hydrotreater and a vacuum gas oil hydrocracker unit as part of the deal. Each agree-ment includes the grant of a licence to Shell proprietary technology and the provision of engineering services. Agreements for the supply of cata-lysts and reactor internals are expected to be signed in the future as part of the deal.

Aromatics performance studySolomon Associates has launched the Worldwide Aromatics Perfor-mance Analysis (aromatics study) intended to provide manufacturers with a view of their performance across all of their aromatics units, whether located inside or outside their refineries. The study aims to enable manufacturers to identify performance gaps as a basis for improving manufacturing opera-tions efficiency.

The aromatics study includes peer group performance analysis from a standpoint of region, size, complex-ity and global status for more than 100 aromatics complexes around the world. Individual plant rank-ings are provided within each peer group based on key performance indicators covering principal aspects of the manufacturing process.

The study will analyse manufac-turing data, with an emphasis on operations and maintenance. Focus areas include energy consumption, process efficiency, maintenance costs, operating and mechanical reliability, and staffing efficiency.

Catalyst regeneration for IndiaEurecat is to build a new catalyst processing plant in India, to be located in the Jhagadiya Industrial

Estate, Bharuch District, Gujarat State. The plant will be operated by Eurecat India Catalyst Services Private Ltd, a subsidiary of France-based Eurecat, which is a joint venture of IFP Investissements and Albemarle Corporation.

The plant is expected to begin commissioning during Q2 2012. It will be dedicated to the ex-situ regeneration of spent catalysts and associated services, such as hydro-carbons stripping or react catalyst rejuvenation facilities, as well as providing field services tools for industrial reactor turnarounds.

Eurecat says the plant will enable it to offer catalyst management for local customers in India, aiming at a maximum reuse of their catalysts in a fast-growing market, with large installed oil refining capacities, petrochemical and gas treatment plants.

Eurecat has plants in France, Italy, the US and Saudi Arabia, and licensed operations in Japan and Russia.

Mass transfer award for RuwaisGTC Technology has been awarded a contract to design and supply mass transfer equipment for the Ruwais Refinery Expansion Project, a grassroots expansion to increase refining capacity. The refinery is owned by The Abu Dhabi Oil Refining Company (Takreer) and the refining facility is located in Ruwais, United Arab Emirates. The expansion project includes 21 process units, off-sites and utilities.

GTC Technology Korea, a subsidiary of GTC Technology International, was selected to provide mass transfer equipment for the crude distillation unit, satu-rated gas plant and residue catalytic cracking unit, which is the largest unit of its kind in the world. The contract includes a giant-sized pre-flash column, crude column and residue fluidised catalytic cracking (RFCC) main fractionator.

20-year deal for LNG supply to ChinaAustralia Pacific LNG (APLNG) and China’s Sinopec have signed an agreement for the supply of up to 4.3 million t/y of LNG for 20 years.



Compressor suction drums:

I think I’ve got liquid carryover.What can I do about it?

Read more on this topic atwww.amistco.com

It happens in petrochemical

else that the gas approaching

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The agreement specifies terms under which Sinopec will take a 15% equity stake in APLNG, with ConocoPhillips and Origin Energy each reducing its interest to 42.5%. In addition to the signing of the heads of agreement, the APLNG project received Federal Environ-mental approval. ConocoPhillips says it expects the full go-ahead for the project by mid-2011, with the first LNG cargo to be delivered in 2015.

The Australia Pacific LNG project includes the development of substantial coal seam gas resources in the Surat and Bowen basins over a 30-year period, a 450 km trans-mission pipeline and a multi-train LNG facility on Curtis Island, near Gladstone. Initial plans for the LNG facility are focused on developing two LNG trains, each with a name-plate capacity of approximately 4.5 million t/y of LNG.

Origin and ConocoPhillips have agreed that final investment decision payments on the first two trains for the project will be deferred until ConocoPhillips achieves an agreed cash internal rate of return on proj-ect investment.

Energy demand to grow a third in two decadesExpanding prosperity for a growing world population will drive an increase in energy demand of about 35% by 2030, compared to 2005, even with significant efficiency gains, and natural gas will emerge as the second-largest energy source behind oil, according to ExxonMobil Corporation in its latest edition of Outlook for Energy: A View to 2030. Supplies of shale gas and other unconventional energy sources will be vital in meeting this demand. The growing use of natural gas and other less carbon-intensive energy supplies, combined with greater energy efficiency, will help mitigate environmental impacts of increased energy demand. According to the Outlook, global energy-related CO2 emissions growth will be lower than the projected average rate of growth in energy demand.

The Outlook for Energy is devel-oped annually to help guide ExxonMobil’s investment decisions.

The company shares the findings publicly to increase understanding of the world’s energy needs and challenges. The Outlook is the result of an analysis of 100 countries, 15 demand sectors and 20 fuel types, and is underpinned by economic and population projections, as well as expectations of significant energy efficiency improvements and tech-nology advancements.

Rising electricity demand — and the choice of fuels used to generate that electricity — represent a key area that will have an effect on the global energy landscape over the next two decades. According to the Outlook, global electricity demand will rise by more than 80% through 2030, from 2005 levels. In the non-OECD alone, demand will soar by more than 150% as economic and social development improve and more people gain access to electricity.

According to the Outlook, efforts to ensure reliable, affordable energy, while also limiting greenhouse gas emissions, will lead to policies in many countries that put a cost on CO2 emissions. As a result, abun-dant supplies of natural gas will become increasingly competitive as an economic source of electric power, as its use results in up to 60% fewer CO2 emissions than coal in generating electricity. Demand for natural gas for power genera-tion is expected to rise by about 85% from 2005 to 2030, when natu-ral gas will provide more than a quarter of the world’s electricity needs. Natural gas demand is rising in every region of the world, but growth is strongest in non-OECD countries, particularly China, where demand in 2030 will be approxi-mately six times what it was in 2005.

Sasol picks Linde for supportSasol Technology of South Africa has appointed the Engineering Division of The Linde Group as preferred licensor’s engineering contractors for a major portion of Sasol’s Fixed Bed Dry Bottom (FBDB) gasification technology for an initial term of 10 years.

Linde’s mandate covers the down-stream aspects of the technology

(raw gas cooling, raw gas shift, byproduct processing and overall integration of the gas island). Hatch has been appointed for the remain-ing upstream aspects of the technology (coal delivery, gasifica-tion proprietary equipment and ash handling).

Linde’s experience in hydrogen and synthesis gas production and gas processing technologies will be used to provide engineering consul-tation and related engineering services to Sasol. Linde will assist Sasol’s existing coal-based gas production operations with the development of optimisation and engineering studies, and will be responsible for process engineering work for new gasification plants. Furthermore, Linde will assist Sasol with the on-going development of its gasification technology.

Since the foundation of Sasol, Linde’s Engineering Division has cooperated with the company, supporting its R&D and providing engineering, procurement and construction services. To date, Linde has been awarded over 70 Sasol contracts, ranging from conceptual engineering work to complete supply of plants on a lump sum turnkey basis.

Sohar expansion awardCB&I will provide the FEED and PMS for the Sohar refinery expan-sion project in Oman. The contract, valued at over $40 million, was awarded by Oman Refineries and Petrochemical Company.

The project will increase the capacity of the existing Sohar refin-ery from 116 000 to 187 000 barrels per stream day by installing various clean fuels units, as well as increas-ing capacity and debottlenecking existing units in the refinery. The Sohar refinery, which was commis-sioned in 2006, was built to process the feedstock of long residue that is produced at the Mina Al-Fahal refinery in Oman and blended with crude oil.

In conjunction with the upgrade, several new units will be added and integrated into the existing refinery complex. They include a 71 500 b/d crude distillation unit, a 96 800 b/d vacuum distillation unit,



a 66 400 b/d once-through hydro-cracker unit, a 42 400 b/d solvent deasphalting unit, a sulphur recov-ery unit, sour water stripper units, an amine regeneration unit and an isomerisation unit. Utility genera-tion and off-site storage facilities will be expanded.

The upgrade is intended to contribute to improvements in Sohar refi nery’s performance, output and quality of feedstock for downstream petrochemical projects Aromatics Oman and Oman Polypropylene. In addition to enhancing the feed qual-ity of the residual fl uid catalytic cracker, the upgrade will enable Sohar Refi nery to supply feedstock naphtha to Aromatics Oman. The modernised plant will also be able to meet feedstock commitments in the form of polymer-grade propyl-ene to Oman Polypropylene’s complex located next door. Additionally, bitumen will be added for the fi rst time to Sohar Refi nery’s range of products and byproducts. When the expanded and upgraded refi nery is eventually brought into operation, its blended refi ned prod-ucts will comply with Euro VI emissions standards.

Expansion plans to boost fuel supplyIranian and Chinese companies are reported to be close to signing a $7 billion deal to cooperate in the development of two oil refi neries in the southwestern Iranian city of Abadan and the central city of Isfahan. They are together expected to produce 30 million litres of gaso-line and diesel per day, according to local news sources. The project to expand the Abadan oil refi nery will start as soon as the contract is signed, but studies for the Isfahan oil refi nery project are not yet completed.

Some $4.1 billion is needed for expanding the Isfahan oil refi nery and over $2.8 billion is necessary for the Abadan oil refi nery. Over $850 million has been invested in the two undertakings. When all of the projects are completed, a total of 40 million litres of gasoline and 18 million litres of diesel will be added to the country’s production capacity.

Currently, operations are under way for the optimisation and expansion of gasoline production in 10 Iranian refi neries with an invest-ment of over $13 billion. Sinopec is currently working on a project for the expansion of the Imam Khomeini oil refi nery. The fi rst phase of the project recently went on stream.

Compressors for Jubail petrochemsInternational Polymers Company of

Jubail, Saudi Arabia (Sipchem) has selected Burckhardt Compression to deliver a hyper-compressor for its world-scale ethylene vinyl acetate (EVA) plant. The tubular high-pressure, low-density polyethylene process technology is licensed by ExxonMobil. Hyper-compressors are high-pressure reciprocating compressors for LDPE plants with discharge pressures up to 3500 bar (50 000 psi).

IPC is a Joint Venture between Sipchem and Hanwha Corporation

www.ptqenquiry.com for further informationwww.eptq.com PTQ Q2 2011 123


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of South Korea. The 200 000 t/y plant will be built at Sipchem’s site in Jubail as part of a third-phase project. The project is targeted to be commissioned in 2013.

Outlook for wax is slackChanges in lubricant refining tech-nology have led to a decline in the supply of petroleum wax, which has traditionally dominated the global wax industry. As petroleum waxes fall behind, new formula-tions have emerged to compete for the vacated market space. Global wax supply in 2010 is estimated at 9.6 billion lb according to a report* from consultant Kline & Company.

The fate of petroleum wax supplies is largely outside the control of the wax industry. Wax supplies are tied to regional base-stock quality requirements, and this is independent of wax supply and demand. Today, petroleum waxes represent 85% of the global wax supply, dropping below 90% for the first time in decades.

A combination of tight wax supplies and rising crude oil costs has caused a steady rise in wax prices during the past five years. This has created market openings for higher-cost synthetic, animal and natural vegetable waxes, and provided a platform for the research, development and market-ing of hydrogenated vegetable waxes.

Fighting to take the place of petroleum waxes are synthetic and vegetable waxes. Synthetics, mainly represented by Fischer-Tropsch and polyethylene waxes, have the advantage of closely resembling the physical properties of petroleum waxes and currently represent 11% of the global wax supply. On the other hand, vegetable waxes can fulfill the need for softer waxes, which are in deficit due to the reduction in slack wax supplies. Their use has already begun in applications including board sizing, candles and fire logs. Vegetable and animal waxes account for 4% of the global wax supply.

Wax supply has eroded signifi-cantly in North America and Western Europe due to a rapid Group I supply rationalisation. In

the past, the decline in supplies from North America and Western Europe were compensated by growth in supplies from China. However, this is no longer the case, as the Chinese lubricant industry starts demanding more Group II and III basestocks, leading to a flat or declining Group I supply.

Wax consumption is expected to grow at an average annual growth rate of more than 2% from 2010 to 2020. Different regions and different product applications will enjoy varying growth rates. Ultimately, insufficient growth in supply or even a decline in supply will constrain growth in demand.*Global Wax Industry 2010: Market Analysis and Opportunities

Safety management planThe UK’s Energy Institute (EI) has developed a common High Level Framework for Process Safety Management across all energy industry sectors. It aims to provide a simple and systematic approach, defining the key issues organisa-tions need to get right to assure the integrity of their operations, incor-porating technical, maintenance and operational, as well as human and organisational factors.

The impact safety incidents in the energy industry have had upon the share prices of the involved compa-nies has caused some institutional investors to question the security of their investments in the high-hazard industries, says the EI.

The publication aims to assist senior executives and managers to understand how well they are iden-tifying and managing the significant risks within their organisations, which, if not appropriately managed, could result in a major incident.

Mexico’s algal ambitionOriginOil is to take part in a pilot-scale algae-to-fuels project to be funded by the Mexican govern-ment. The project will demonstrate industrial algae production, paving the way for investment by the Mexican government in large-scale jet fuels production.

The aim of Mexico’s Manhattan Project is to produce 1% of the

nation’s jet fuel from algae in less than five years. By the end of this decade, the project aims to produce nearly 20 times that amount, which would place Mexico among the leading biofuels-producing nations.

The project’s operator is Genesis Ventures of Ensenada, Baja, California, which has received an Economy Ministry grant through The National Council for Science and Technology for its first site. Genesis will develop the site as a model for numerous additional projects to be located next to large sources of CO2, such as thermal power stations. The operator will rely heavily on OriginOil’s exper-tise in feeding and sanitising algae cultures, and its core harvesting and extraction technology.

Call for support in SA refining The South African government should come to the rescue of the refining industry if it wants to keep one of its main sources of income, according to Philip Lloyd, a profes-sor with the Cape Peninsula Institute of Technology. The indus-try is putting R50 billion a year into the state’s coffers and is struggling. The government has to do some-thing if it wants to keep its golden goose, he says.

South Africa’s refineries are old and have been subject to continu-ous modernisation and upgrades, to such an extent that nothing more can be done to improve them, Lloyd says. The problem is that the demand for gasoline and diesel, and kerosene to a lesser extent, continues to grow, but refineries are reluctant to invest money in new infrastructure. The main reason is that refining margins are very low. This situation will remain for the coming years. As a result, and because constraints in the electricity supply have had an impact on certain projects, refining production is dwindling.

Current plans in South Africa to build a 400 000 b/d refinery are absurd, Lloyd says. The country only needs about 100 000 b/d. According to the government, the rest will be exported to other coun-tries in the region. This does not make sense, Lloyd says. The region



is already over-supplied and is making its own refined products. Angola, for instance, is planning to increase the capacity of its Luanda refinery from 40 000 to 100 000 b/d. In addition, it wants to build a new 200 000 b/d refinery in Lobito.

Other countries such as Mozambique and Kenya also have refinery projects in the pipeline. The Kenyan KPR plant will see an increase in its refining capacity from 72 000 to 100 000 b/d. In Tanzania, a 200 000 b/d project is to replace an old refinery.

Sino-Kuwaiti megarefinery wins go-aheadChina’s top economic planning body, the National Development and Reform Commission (NDRC), has granted final approval for Kuwait to build a long-awaited refinery and petrochemical complex on Donghai Island of Zhanjiang in Guangdong Province.

The joint venture between Kuwait Petroleum Company (KPC) and Sinopec entails a 15 million t/y (300 000 b/d) refinery, a 1 million t/y ethylene plant and related utili-ties, as well as support facilities, such as a crude jetty, product oil and chemical jetties, a bulk jetty and oil product pipelines to an initial station.

The Sino-Kuwaiti refinery project will be China’s largest ever joint venture of its kind. The $9 billion project is expected to boost the city’s revenues by $1.8 billion. The city is also planning to build a 30 km2 petrochemical park along with the refinery. The project involves a downstream marketing network, including retail petrol stations in and around the province, where the economy is growing at more than 12% annually.

Kuwait will supply all of the crude feedstock for the integrated plant. KPC and China’s down-stream major Sinopec will each hold a 50% stake in the joint venture, with KPC planning to give 20% of its share to its potential partners.

Kuwait Petroleum International (KPI), KPC’s international refining and market unit, has been repre-senting Kuwait in lengthy negotiations with the Chinese side

of the deal. Kuwait will become the second Arab oil producer to build a refinery in China, following Saudi Arabia, which put a refining and petrochemical joint venture into operation last year in southeast China’s Fujian Province.

Final approval from the NDRC clears the way for KPC to achieve its crude oil export target of 500 000 b/d to China. Kuwait’s exports to China in 2010 stood at 198 000 b/d, up 39% from the previous year. The joint venture also reflects Kuwait’s strategy to expand refining and marketing outlets in high-growth strategic markets, such as China, India and Vietnam, according to KPC.

The alliance gained the NDRC’s preliminary approval for the project in May 2010, when the commission officially designated Donghai Island as the site for the complex, followed by the approval of the Ministry of Environmental Protection in September 2010.

Based on an initial deal in 2005, KPC and Sinopec originally planned to locate the facility in the Nansha district of provincial capital Guangzhou, but due to concern over environment impact on a densely populated area the Guangdong provincial government proposed the relocation of the plant. The two sides agreed to move the project to a less crowded area within the same province and then, in August 2009, the joint venture opted for Zhanjiang City, about 400 km southwest of Guangzhou, as the new site.

Uzbeks to upgrade refineriesUzbekistan plans to spend $305 million on upgrades to two major oil refineries to boost its product output. State holding company Uzbekneftegaz intends to raise production with upgrades to the country’s oil refineries within its national Programme on Priorities of Industrial Development for 2011–2015.

The Ferghana refinery in eastern Uzbekistan, which produces high-quality diesel and fuel oils, will receive the largest slice of invest-ment, with three projects amounting to $205 million. Uzbekneftegaz will

spend $99 million on modernising the facility’s distillation processes. On completion, the plant is expected to produce 550 000 t/y of oil products.

A second project, at Uzbekistan’s largest petroleum refinery, costing $6 million, will boost output with the introduction of new processing technology. Planners will also spend $100 million on upgrading and reconstructing hardware at the Ferghana refinery, which was commissioned in 1959 during the Soviet era. This third project is expected to increase production of light oil products by 95%, according to Uzbek authorities. The plant in Ferghana region currently produces over 60 types of oil products with a total output of 6 million t/y.

Uzbekneftegaz will also moder-nise and reconstruct hardware at its Bukhara refinery in central Uzbekistan with a $100 million project that will see output of light oil products increase by up to 95%. The Bukhara refinery became oper-ational in 1997 and currently produces 10 types of oil products.

Oiling the wheels of automationExxonMobil has installed an enter-prise control system (ECS) at its lubricants plant in Beaumont, Texas, to help manage its major processes. The aim is to apply real-time process automation across the site to improve the flexibility of its production processes and work-flows, and to deal with issues of ageing within the existing control systems. The new installation, based on Invensys Operations Management’s InFusionT software, also integrates existing systems for enterprise resource planning, batch process, and final packaging and shipping systems.

The installation is part of a sweeping modernisation plan to achieve what ExxonMobil calls “the lubricant plant of the future” at the Lubricants & Specialties site. Operating as a system of systems, the InFusion ECS integrates with legacy applications and third-party solutions to improve the plant’s operations and performance in real-time.



Fluidised bed dryingUhde Services of Haltern, Germany, has signed a framework agreement with RWE Power for the use of fluidised-bed drying technology with internal waste heat utilisation. Uhde Services will join Linde KCA Dresden (Linde Group) as the second supplier of this technology developed by RWE Power.

The aim is to use the technology at lignite-fired power stations, as well as for coal gasification and liquefaction projects. Electricity and also synthesis gas can be produced more energy efficiently, with fewer CO2 emissions, using the fluidised-bed drying [email protected]

Valve controller interfacesEmerson’s Fisher Fieldvue digital valve controllers with Hart and Foundation Fieldbus communica-tions can be integrated into any process control system that supports FDT/DTM technology, says the developer. Using ValveLink DTM software, users can communicate with any Fieldvue digital valve controller to perform start-up, commissioning and diagnostic activities.

ValveLink software has been certi-fied by the FDT Group for compliance with FDT standards. It has also been tested and certified for use with FDT-compliant host systems, including Honeywell, Invensys and Yokogawa.

Fieldvue digital valve controllers use AMS ValveLink Snap-on to integrate into Hart and Foundation Fieldbus host systems, such as Emerson’s DeltaV and Ovation digital automation systems. The addition of DTM support provides another avenue for integration for Fieldvue instrument users.

The role of Fieldvue digital valve controllers is to maintain control valve performance, diagnose the assembly and enable predictive maintenance. [email protected]

Supply chain software Aspen Technology has released new local language versions of manufac-turing and supply chain products in the 7.2 release of aspenONE V7. Local language versions are now available for Aspen PIMS, Aspen Petroleum Scheduler, Aspen Plant Scheduler, Aspen IP.21 Process Browser and other products in the aspenONE Manufacturing & Supply Chain suite.

Local language software makes it possible for users to train employ-ees and standardise their operations. Local language versions of aspe-nONE products are available in Chinese, French, German, Italian, Japanese, Korean, Portuguese, Russian and Spanish, in addition to [email protected]

Hart supports EDDLThe Hart Communication Found-ation has announced support for the latest enhancements to Electronic Device Description Language (EDDL) incorporated in the second edition of the International Electrotechnical Commission’s IEC 61804-3 stan-dard. EDDL is the key industry standard for integrating real-time diagnostic and asset management information from intelligent field devices for optimum data and device interoperability with auto-mation systems.

Recent improvements to the EDDL standard enhance the inte-gration of intelligent devices with automation systems, specifically:• Support for offline configuration with default parameter values suggested by the device manufac-turer to simplify and speed device commissioning• Support for Unicode character sets, enabling parameter labels, diagnostics and device manufac-turer expert help text to be displayed in many different languages including Chinese, Japanese and Russian


New Products

• Ability to display all device diag-nostics and setup information with rich user-friendly graphics for easier and faster completion of commis-sioning and maintenance tasks• New capabilities to support the display of illustrations based in user-preferred language, such as images with explaining text to convey know-how from the manu-facturer for help in the interpretation of advanced diagnostics, guide setup and for troubleshooting.

With these enhancements, the EDDL standard remains backward compatible and stable for the lifecy-cle needs of industrial automation. Installed devices and systems remain compatible, EDDL files load and perform on systems in the same way, and updates continue without the problems of executable software.

EDDL is a text-based language standard to describe the character-istics of intelligent field devices for integration with systems. The Hart Communication Protocol was the first to implement EDDL, enabling suppliers to define their products in a single format readable by host applications for handheld commu-nicators, control systems, PCs and other process interface devices. [email protected]

Simulator updateInvensys Operations Management has released the latest version of its SimSci-Esscor PRO/II simulation software, a steady-state process simulator used to design, analyse and improve chemical processes.

Version 9 of PRO/II software includes conditional formatting to identify different process stream properties, customisable tabs that enable sorting of unit operations by type, dockable windows and improved toolbars. It also lowers the cost of ownership, says Invensys, by supporting Microsoft App-V for application virtualisation, allowing better use of existing computer hardware resources. Simulation

new prods copy 4.indd 1 11/3/11 11:47:50

Fluidised bed dryingUhde Services of Haltern, Germany, has signed a framework agreement with RWE Power for the use of fluidised-bed drying technology with internal waste heat utilisation. Uhde Services will join Linde KCA Dresden (Linde Group) as the second supplier of this technology developed by RWE Power.

The aim is to use the technology at lignite-fired power stations, as well as for coal gasification and liquefaction projects. Electricity and also synthesis gas can be produced more energy efficiently, with fewer CO2 emissions, using the fluidised-bed drying [email protected]

Valve controller interfacesEmerson’s Fisher Fieldvue digital valve controllers with Hart and Foundation Fieldbus communica-tions can be integrated into any process control system that supports FDT/DTM technology, says the developer. Using ValveLink DTM software, users can communicate with any Fieldvue digital valve controller to perform start-up, commissioning and diagnostic activities.

ValveLink software has been certi-fied by the FDT Group for compliance with FDT standards. It has also been tested and certified for use with FDT-compliant host systems, including Honeywell, Invensys and Yokogawa.

Fieldvue digital valve controllers use AMS ValveLink Snap-on to integrate into Hart and Foundation Fieldbus host systems, such as Emerson’s DeltaV and Ovation digital automation systems. The addition of DTM support provides another avenue for integration for Fieldvue instrument users.

The role of Fieldvue digital valve controllers is to maintain control valve performance, diagnose the assembly and enable predictive maintenance. [email protected]

Supply chain software Aspen Technology has released new local language versions of manufac-turing and supply chain products in the 7.2 release of aspenONE V7. Local language versions are now available for Aspen PIMS, Aspen Petroleum Scheduler, Aspen Plant Scheduler, Aspen IP.21 Process Browser and other products in the aspenONE Manufacturing & Supply Chain suite.

Local language software makes it possible for users to train employ-ees and standardise their operations. Local language versions of aspe-nONE products are available in Chinese, French, German, Italian, Japanese, Korean, Portuguese, Russian and Spanish, in addition to [email protected]

Hart supports EDDLThe Hart Communication Found-ation has announced support for the latest enhancements to Electronic Device Description Language (EDDL) incorporated in the second edition of the International Electrotechnical Commission’s IEC 61804-3 stan-dard. EDDL is the key industry standard for integrating real-time diagnostic and asset management information from intelligent field devices for optimum data and device interoperability with auto-mation systems.

Recent improvements to the EDDL standard enhance the inte-gration of intelligent devices with automation systems, specifically:• Support for offline configuration with default parameter values suggested by the device manufac-turer to simplify and speed device commissioning• Support for Unicode character sets, enabling parameter labels, diagnostics and device manufac-turer expert help text to be displayed in many different languages including Chinese, Japanese and Russian


New Products

• Ability to display all device diag-nostics and setup information with rich user-friendly graphics for easier and faster completion of commis-sioning and maintenance tasks• New capabilities to support the display of illustrations based in user-preferred language, such as images with explaining text to convey know-how from the manu-facturer for help in the interpretation of advanced diagnostics, guide setup and for troubleshooting.

With these enhancements, the EDDL standard remains backward compatible and stable for the lifecy-cle needs of industrial automation. Installed devices and systems remain compatible, EDDL files load and perform on systems in the same way, and updates continue without the problems of executable software.

EDDL is a text-based language standard to describe the character-istics of intelligent field devices for integration with systems. The Hart Communication Protocol was the first to implement EDDL, enabling suppliers to define their products in a single format readable by host applications for handheld commu-nicators, control systems, PCs and other process interface devices. [email protected]

Simulator updateInvensys Operations Management has released the latest version of its SimSci-Esscor PRO/II simulation software, a steady-state process simulator used to design, analyse and improve chemical processes.

Version 9 of PRO/II software includes conditional formatting to identify different process stream properties, customisable tabs that enable sorting of unit operations by type, dockable windows and improved toolbars. It also lowers the cost of ownership, says Invensys, by supporting Microsoft App-V for application virtualisation, allowing better use of existing computer hardware resources. Simulation


a single 19in sub-rack unit, to measure dew points as low as -100°C. Measurements of moisture in process gases and liquids can be combined in a single analyser system as the Promet I.S channels for moisture in gas can be combined together with a sister product for moisture in liquid measurement —the Liquidew I.S.

Each channel functions indepen-dently of the others so that maintenance can be carried out on one channel while the others continue to operate as normal..

Unit conversions from dew point to a range of alternative moisture content units allow the user to select a preferred unit of measure-ment. The front panel interface enables the user to scroll through the setup menus to configure the analyser. Four user-adjustable alarm points and two analogue 4-20 mA outputs are provided, as well as a digital RS485 RTU for connection to external [email protected]

Overload indicatorWIKA has developed an overload indicator for two pressure gauge series, so that unexpected overpres-sures in industrial processes can be documented beyond any doubt. According to WIKA, the advantage of the new design over a pressure gauge with a drag pointer lies in the guaranteed security of the indi-cation. The status of drag pointers can alter due to external influences such as vibration, and thus the explicit confirmation of overpres-sure cannot be ensured. In addition, the new overload indicator is protected against any manipulation after the fact.

The core of the new instrument option is a spring pin. This is fitted at the dial’s scale value, which the user has defined as the highest pressure. The indicator, an alumin-ium pin, is carried by the instrument pointer; if the designated maximum pressure is exceeded, it is locked permanently in the red zone.

WIKA supplies the overload indi-cators for model 23x.50 and 212.20 Bourdon tube pressure gauges in nominal sizes of 100 and [email protected]

capabilities have also been enhanced through an update of the thermody-namic correlations, including a link to the National Institute of Standards and Technology’s Reference Fluid Thermodynamic and Transport Properties (REFPROP) application. The database provides critical prop-erty values needed to evaluate fluids and optimise related equipment and processes. [email protected]

Multi-gas detectorIndustrial Scientific has introduced the Ventis MX4 multi-gas detector. It is a lightweight, configurable instrument that is available with or without an integral pump and is compatible with iNet, Industrial Scientific’s Gas Detection as a Service solution.

The Ventis detects one to four gases including oxygen, combusti-ble gases (LEL or CH4) and any two of the following gases: CO, H2S, NO2 and SO2. It is designed for confined space monitoring and/or continuous personal monitoring in potentially hazardous environ-ments. In confined space applications, the Ventis can be used to draw samples from up to 30m with the integral pump. The gas detector alerts users in dangerous conditions through an audible alarm, ultra-bright LED visual alarms and a vibrating alarm.

Among the battery options is an extended range lithium-ion battery, which enables up to 20 hours of personal monitoring when used with the no-pump version. [email protected]

Process moisture analyserMichell Instruments has launched a new Promet I.S. process moisture analyser that can be retrofitted in almost all existing AlOx or ceramic transmitter-based trace moisture or dew point measuring analyser systems. The sensors, which form the basis of the Promet I.S analyser, may be installed within the hazard-ous area directly into an existing sampling system.

The Promet I.S. multi-channel format enables the control of up to four pressure compensated mois-ture measurement channels within


Pipeline heatingChromalox has developed the Skin-Effect Heating System for pipelines up to 15 miles in length. Since it uses a single power supply, the system provides a cost-effective alternative for freeze protection, temperature maintenance and heat-up, particularly in remote areas where installation and maintenance are more complex and expensive, says Chromalox.

Typical applications include tank farms and storage terminals, as well as long-distance piping between processing facilities in oil and gas, refining, chemical, and similar industries. The system includes a small steel tube, containing a skin-effect electric cable, which is banded onto the pipe. The cable and tube transfer conductive heating directly to the wall of the process pipe. Supply connections are made in special boxes. This method elimi-nates the need for additional heating equipment, such as heat exchangers, circulation heaters or conventional heat trace, says Chromalox.

In addition to reducing installa-tion and maintenance costs, the system reduces the number of heat-ing circuits required, resulting in lower control panel costs and simpler heating operation, accord-ing to the developer.

Heat input can be adjusted up to 300°F (150°C) to handle a range of thermal viscosities, including water, crude or refined oils, steam and a variety of chemicals. The Skin-Effect Heating System can be used in pipes above ground, buried or submerged. With no moving parts, the system requires little maintenance. The system can be used with potentially corrosive processes and can be certi-fied for hazardous areas [email protected]

Compressor range extendedBurckhardt Compression has extended its Process Gas Compressor API 618 product line for large plants in the hydrocarbon processing industry. It now covers compressor duties requiring a rod load up to 1500 kN/335 000 lb and power up to 31 000 kW/42 100 hp. [email protected]


ABB Global Consulting 39

Aerzener 73

Albemarle Catalysts Company 14

Amistco Separation Products 92, 121

Axens OBC

BASF Catalyst 19

Borsig Group 101

Bryan Research & Engineering 64

Burckhardt Compression 77, 83

Cansolv Technologies 50

CB&I 7

Curtiss-Wright Flow Control Corporation 115

Delta Valve 68

Det Norske Veritas 105

DuPont Belco Clean Air Technologies 102

Emerson Process Management 8

Enersul 56

Euro Petroleum Consultants 118

ExxonMobil Research and Engineering Company 113

Flexim 93

Flexitallic 21

Foster Wheeler 25

Grabner Instruments 20

Haldor Topsøe IFC

Hoerbiger Kompressortechnik Holding 108

ITW 43, 110

Johnson Screens 31

KBC Advanced Technologies 2

KBR 4

KIDExtractor 24

Koch-Glitsch 46

Lewis Pumps 74

Linde 67

Lurgi 23

Marsulex Refinery Services 96

Merichem Company 61, 63

Michell Instruments 123

MPR Services 17, 70

Newtons 10

NPRA Reliability & Maintenance Conference 117

OHL Gutermuth Industrial Valves 12

Paharpur Cooling Towers 84

Prosim 45

Rentech Boiler Systems 32

Sabin Metal Corporation 107

Saint-Gobain NorPro 36

Samson 27

Spectro Analytical Instruments 55

Sulzer Chemtech 86

Swagelok Company 40

Thermo Scientific 53

TPS-Technitube 16

Tricat Catalyst Service 99

Uhde 11, 13

UOP 28

Yokogawa Europe 88

World Refining Association IBC

Alphabetical list of advertisers

For more information on these advertisers, go to www.ptqenquiry.com


ad index copy 9.indd 1 12/3/11 20:45:34

Executive Speaker Panel Includes: Tom Crotty, Group Director, INEOS

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Graeme Burnett, Senior Vice President, Asia and Middle East, Total Petrochemicals

Fred du Plessis, Executive Advisor, Corporate Strategy & Planning, SABIC

Maxim Savchenko, Head of Strategic Planning & Business Development, SIBUR

Dr. Bernd Blumenberg, President, BASF-YPC Company

Saleh Ali Zahrani, Manager, Global Technology market, Licensing, SABIC

Vladimir Jacobson, Commercial Manager C4’s, Dow Europe

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