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2013

Supplement to PTQ

revamps ptq

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How Can KBR Answer Your Refining Challenges?

Owners of refineries continue to confront

many challenges – rising feedstock prices,

shrinking margins, varying global demands

and a changing regulatory landscape that

includes ever-more stringent specifications

on sulfur and carbon footprints. As refinery

owners debottleneck and enhance existing

facilities, they call on KBR to deliver.

To learn how KBR can address your refining

challenges, go to:

refining.KBR.com/PTQ

refining.kbr.com / PTQ

K13009 © 2013 KBR, Inc. All Rights Reserved.

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©2013. 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 Improving the distillation energy network

SounHoLeeGTC Technology

KwangGilMin GS Caltex Corporation

15 Strategies for improved naphtha processing

MikeArmstrongandMartinBrandt

Jacobs Consultancy

29 Laser surveying a revamp

GaryFarrow

AVEVA

33 Retrofitting a glycol contactor to prevent carryover

AnnePhanikumarandYangQuan

Sulzer Chemtech Switzerland

45 Reactor effluent air cooler safety through design

EricLinandPeterRisse

Chevron Lummus Global

49 Wet scrubbing modifications to reduce emissions

EdwinWeaverandNicholasConfuorto

Belco Technologies Corporation

59 Online cleaning of an integrated distillation unit

MariuszHołowaczandRafałZaprawa Grupa Lotos

MarcelloFerrara ITW

AnoverhaulofitscontrolsystemsledMOLGrouptoimproveenergyefficiencyandprocessstabilityatitsgasdistillationplantinAlgyõ,Hungary,usingadvancedprocesscontroltechnologyfromEmersonProcessManagement. Photo: Emerson Process Management

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Editor Chris Cunningham [email protected]

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Even the best new catalyst only ends up being used if a customer truly believes that it will live up to its claims. As a Senior Principal Scientist, Dave Sherwood is a key part of the CRITERION team for ebullated and fixed bed residue upgrading. He works closely with refiners around the globe to understand the details of their particular feedstock and their operations, and then works with the research and development team to create the very best catalyst solution for the job. He brings together engineers from each of the unique ebullated beds and helps them to compare the data and learn from each other. Thirty years and twenty countries later, he’s still fascinated by the challenge of tailoring the perfect combination of catalysts for each situation, gathering the data to prove it, then presenting the case for a decision. His long track record of satisfied clients continues to grow every year.

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Improving the distillation energy network

Energy costs are the largest percentage of a hydrocarbon plant’s operating expenditures.

This is especially true of the distil-lation process, which requires substantial energy consumption. Concerns over recent high costs and economic pressures continually emphasise the need for efficient distillation design and operation without a loss of performance.

This article illustrates how energy-efficient design can be applied in a distillation unit through optimisation between the distillation column and heat network system. Through a case study, a successful retrofit of an aromatics distillation unit is discussed. Detailed retrofit activities, including complex heat network evaluation, process simula-tion modelling and energy-friendly, high-performance distillation equip-ment implementation, are described.

Strategies for improving thedistillation energy network As continuous distillation requires simultaneous heat input and removal (thus requiring significant energy consumption), complex heat integration becomes more common for modern distillation units to improve unit energy efficiency. Since a distillation column’s degree of separation and enthalpy balance influence each other, it is critical to evaluate and optimise the distilla-tion column and heat exchanger networks together in order to maximise plant economics.

There are numerous strategies to improve the energy efficiency of distillation processes, with the amount of improvement through each strategy varied according to

Energy-efficient design applied to the refit of a distillation unit was achieved through optimisation between the distillation column and heat network system

SOUN HO LEE GTC TechnologyKWANG GIL MIN GS Caltex Corporation

process conditions. The following are common strategies that can be applied to practical energy improvement projects.

Feed temperature Feed temperature is a major factor influencing the overall heat balance of a distillation column system. Increments in the feed enthalpy can help reduce the required energy input from the reboiler at the same degree of separation. Installing a feed preheater is a very common process option to minimise reboiler heat duty. If the feed preheater can be integrated with other valuable process streams (as a heating medium), overall energy efficiency of the distillation system can be improved further. However, increasing the feed temperature does not always improve the over-all energy efficiency of a distillation unit. Excessive feed temperature increments can cause a significant amount of flash of heavy key and non-key components at the distilla-tion column feed zone. In this case, a higher amount of reflux stream is necessary to maintain required overhead distillate purities. This augmented reflux ratio thus requires a higher boil-up ratio. Overall energy efficiency is eventu-ally aggravated.1 Therefore, careful review of the feed temperature and phase is critical to minimise the overall energy consumption of the distillation unit.

Feed location Improper feed location of a distilla-tion column can also increase the reflux/boil-up ratio and energy consumption. An ideal feed

location is a section of the distilla-tion column where the composition of column internal liquid traffic is similar to feed stream composition. In this case, the composition gradi-ent between feed stream and distillation internal fluids is mini-mised. In actual operation of the distillation column, feed composi-tions are often changed from the original design conditions. In cases of significant deviation, discrepancy between column internal liquid composition and feed stream composition can increase, which results in a non-optimum feed loca-tion. Therefore, evaluating feed location is an essential step for successful distillation unit energy improvement.

Inter-condensers and inter-reboilersAdding inter-condensers and/or inter-reboilers can help improve overall energy efficiency. Pumparound, one of the inter- condenser concepts, has been widely applied to numerous petro-leum multi-product fractionators. On the other hand, implementing an intermediate reboiler can reduce the main reboiler duty. As the required temperature of an inter-mediate reboiler is lower than that of the main reboiler, this strategy may allow heat integration with other valuable heat sources that are not as costly or not fully utilised in the plant.

Column operating pressure Relaxation of the column top oper-ating pressure decreases the distillation column’s temperature profile and results in a lower reboiler duty. It has been observed

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4 Revamps 2013 www.eptq.com

The function of the xylene column is to separate the feed mixture to xylene components and heavier C9+ components. This column has two different feed sources. The reformate splitter bottom stream and the toluene column bottom stream (which belongs to the aromatic extraction unit) are introduced as the xylene column bottom feed stream. The bottom feed stream is split equally and charged to two different feed trays. The reformate splitter’s bottom stream is treated at the clay towers to eliminate traced olefin components before charging to the xylene column. Meanwhile, the deheptaniser bottom stream (from the xylene isomerisation unit) is charged as the xylene column top feedstock. This stream is also split and introduced to three different feed trays.

The xylene column overhead vapour stream is split into three parallel streams. Two vapour streams are utilised as the heat source of the extract column reboiler and the raffinate column reboiler, respectively. The condensed xylene column overhead liquid streams are returned to the xylene column receiver. The other vapour stream is supplying heat to the xylene column overhead steam generator, which produces #250 steam. The condensed overhead liquid stream is also returned to the xylene column receiver. The over-head distillate of the xylene column is sent to the paraxylene recovery unit.

In the xylene column reboiler circuit, the xylene column bottom reboiler inlet stream is first trans-ported to the other two distillation column reboilers as heating medi-ums. After providing heat to these reboilers, the xylene column bottom streams are combined and intro-duced to the furnace-type xylene column reboiler.

In the paraxylene recovery unit, the xylene components from the xylene column are separated through the adsorption process. The pre-separated extract stream from the adsorption process is charged to the extract column in order to separate paraxylene from

that numerous commercial distilla-tion columns have been operated with lower operating pressures than their original design values. However, this strategy is not appli-cable to columns operated under an atmospheric pressure range. Column overhead circuit pressure drop and condenser temperature approaches both heavily influence feasibility. In addition, column pressure reduction expands vapour traffic and pushes the limits of existing distillation equipment.

Column pressure drop Reducing column pressure drop can lower reboiler duty at the same degree of separation. The amount of reboiler duty saving relies on operating pressure and enthalpy balance. This strategy is generally feasible when the distillation column is operated under vacuum pressure range. Meanwhile, pres-sure drop improvement does not often provide noticeable energy savings in high-pressure range distillation service.

Column efficiency improvement Column efficiency improvement can reduce the reflux/boil-up ratio at a given degree of separation. This strategy can be delivered by increasing the number of theoreti-cal stages and/or enhancing the efficiency of distillation equipment. The feasibility can be gauged by a dedicated sensitivity analysis. Constructing a column efficiency curve with a reflux ratio is one of the core tools for sensitivity

analysis. A typical curve is shown in Figure 1. This curve visualises column efficiency sensitivity and energy-saving gain. The curve can be categorised by three district zones: steep, moderate and flat sensitivity.2

Column efficiency improvement is usually very feasible when the reflux ratio falls into the steep sensitivity zone and at ratios considerably in excess of the mini-mum reflux ratio. In this scenario, even the small addition of stages, or an increase in distillation equip-ment efficiency, can enhance overall column separation with significant energy reductions.

Improvement gain is diluted in the moderate sensitivity zone. Further detailed feasibility study is necessary through economic analy-sis. The magnitude of energy savings is negligible when reflux ratio variation follows flat motion in the remaining zone.

Case study: unit descriptionThe following is a revamp case study of a xylene mixture separa-tion unit that demonstrates well-thought-out, proven design practices and a selection of the correct, high-efficiency distillation equipment to fulfill the improve-ment in energy efficiency. Figure 2 illustrates the xylene mixture sepa-ration unit’s configurations under discussion. This schematic reveals that the original distillation units have implemented the full heat integration network for energy- efficient operation.

Steep sensitivity zone

Moderate sensitivity zone

Flat sensitivity zone

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Figure 1 Typical column efficiency vs reflux ratio curve

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the desorbent. At the same time, the pre-separated raffinate stream from the adsorption process is charged to the raffinate column to rectify the raffinate components (metaxylene, orthoxylene and ethyl benzene) as a side-cut product.3 A pasteurising section is arranged at the top of the raffinate column to remove moisture from the side-cut product.4

Case study: process evaluation for energy efficiency improvementTo achieve an additional gain in energy efficiency, a dedicated process evaluation was conducted for the unit. Column operating conditions were first compared to the original design conditions. This comparison helps comprehend deviations between the original design and the actual operational environment. It was observed that the actual product purities of the xylene column were higher than the aromatic rundown product requirements. Relaxing the degree of separation of the xylene column can reduce the reflux and boil-up ratio, as well as save fuel consump-tion for the xylene column furnace reboiler. However, the xylene column overhead vapour streams are utilised as the raffinate and extract column reboiler heating mediums, and contribute steam production in the current unit energy network. Lower reflux/boil-up ratios in the xylene column decrease the amount of xylene overhead vapour used as heating medium for the extract/raffinate column and/or steam generation.

Process simulation modelling was utilised as part of the process eval-uation activities to quantify and predict gains in energy efficiency. Equilibrium base simulation soft-ware was utilised for the modelling. Base simulation modelling was first constructed through pertinent unit test run data. Gathered major process stream flow rates were verified via flow meter orifice calculations. Regular stream composition analysis reported bulk compositions for non-key compo-nents such as non-aromatic and C9+ component groups. Preliminary simulation modelling showed that


Table 1

Mixed xylene

Top feed

Bottom feed

Xylene column

Extract column

Raffinate column

Steam generator

C9+

Figure 2 Xylene column heat network configuration

Case parameter Pre-revamp test run Simulation resultsExtract column Column top temperature, °F Base +∆ 0.9°FColumn bottom temperature, °F Base +∆ 1.1°FReboiler return temperature, °F Base +∆ 0°FReflux ratio (to overhead distillate), volume Base +1.1%Reflux temperature, °F Base ∆0Overhead distillate rate, BPD Base ∆0Bottom rate, BPD Base ∆0p-DEB impurity in overhead distillate, wt ppm Base ∆0Xylene impurity in bottom, wt ppm Base ∆0

Raffinate column Column top temperature, °F Base +∆ 6.7°FSide cut draw temperature, °F Base -∆ 4.3°FColumn bottom temperature, °F Base -∆ 0.5°FReboiler return temperature, °F Base -∆ 2.9°FReflux ratio (to side cut), volume Base -0.58%Reflux temperature, °F Base ∆ 0Overhead distillate rate, BPD Base ∆ 0Bottom rate, BPD Base ∆ 0p-DEB impurity in side cut stream, wt ppm Base ∆ 0Xylene impurity in bottom, wt ppm Base ∆ 0

Xylene column Column top temperature, °F Base -∆ 0.9°FColumn bottom temperature, °F Base ∆ 0°FReboiler return temperature, °F Base -∆ 2.25°FReflux ratio (to overhead distillate), BPD Base + 0.04%Reflux temperature, °F Base ∆ 0°FOverhead distillate rate, BPD Base -0.1%Bottom rate, BPD Base -5%C

9+ impurity in overhead distillate, wt ppm Base ∆ 0

Xylene impurity in bottom, wt ppm Base ∆ 0

Pre-revamp test run and base simulation: comparison of results

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streams of both of the columns are preheated by bottom product streams. Adding independent feed preheaters was not feasible due to limited plot and poor economics. The study showed that reducing column pressure drop using the low-pressure drop nature of trays does not deliver reboiler duty savings in both of the columns.

Case study results for column efficiency improvement showed that energy efficiency improvement was feasible. The column efficiency curves were constructed using simulated reflux ratio and theoreti-cal stage values, and base modelling and improved column efficiency points were plotted. Figures 3 and 4 display these curves. An improved column effi-ciency point was predicted through an increased number of trays. Simulated tray efficiency values of

results were reasonably matched to the test run data. As mentioned earlier, the xylene column bottom rate was not matched. Through various sensitivity analyses, the tray efficiencies of the columns were quantified.5

Extensive case studies were performed for the feasibility of energy efficiency improvement in the unit. The case studies focused on energy consumptions in the extract and raffinate columns. Energy improvements in the other two distillation column reboilers in the xylene column reboiler circuit also help reduce fuel consumption in the xylene column furnace reboiler, but magnitude was not significant.

Since both columns are operated under atmospheric pressure, reduc-ing column operating pressure is not applicable. In addition, feed

component assumptions for these component groups varied simula-tion results significantly. To improve accuracy of simulation, detailed component analysis was specially arranged for the test run.

Detailed component analysis was utilised for rigorous simulation modelling. Key component balance closures for the extract and raffinate columns were less than 3%. Reconstructed feed compositions using products were applied for the simulation of the extract and raffi-nate columns. For the xylene column, the given overall mass balance closure was off by 5%. It was found that measured feed rates were more reliable than product rates, and bottom product rate was less reliable through overall unit mass balance investigation. Based on this investigation, simulation modelling for the xylene column focused on matching feed and over-head distillate rates.

A reasonable matching reflux temperature as well as rate is criti-cal to quantify reliable column internal traffic conditions. It has been observed that matching reflux temperature is often overlooked in simulation modelling. Reflux rate is usually metered at a flow meter located on the reflux piping. When the external reflux rate is recycled back to the column, the internal reflux rate will vary, depending on the external reflux temperature. Therefore, a poor matching reflux temperature in the simulation will not predict the actual internal column traffic accurately. This can result in erroneous efficiency assumptions of the existing column or provide a misleading, incorrect result.

Instrumentation for measuring the pressure drop of the extract and raffinate columns was not perti-nent. Matching pressure drop was ignored in simulation modelling; instead, matching column tempera-ture data were focused on. Temperature profiles are more important to predict distillation column energy consumptions. Table 1 summarises the base model simulation results and compares test run data for the three columns. This table depicts that base model

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Revamp design

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Pre-revamp test run

Revamp design

Pre-revamp degree of separation

Figure 3 Extract column sensitivity analysis (revamp design)

Figure 4 Raffinate column sensitivity analysis (revamp design)

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ratio was not matched to the neigh-bouring three-pass tray pass open area ratio. In order to improve flow ratio balancing, the number of passes for the pasteurisation section trays was changed from three to two and a new chimney tray was installed as per pass change.

Case study: post revamp operation reviewThe pre- and post-revamp perfor-mances are summarised and compared in Table 2. As the over-all aromatic unit capacity has been expanded, the column charge rates are also increased. Post-revamp operation verifies that the reflux ratios of the extract and raffinate column are reduced, and these reduced reflux ratios eventually contribute to energy savings in the xylene column reboiler furnace. Measured furnace fuel consump-tion as per the feed rate is substantially improved. Since the paraxylene recovery unit adsor-bent upgrade also contributes to savings in furnace fuel consump-tion, the net energy-saving

tray vapour velocity that can down-grade tray efficiency due to insufficient vapour/liquid contact volume. Various performance- enhancing features of GT-Optim trays were added to improve tray efficiency. Nevertheless, extra indi-vidual tray efficiency improvement was not counted for a conservative approach to revamp design. Applied tray efficiencies for the revamp design were the same as the sieve tray efficiencies obtained through simulation modelling of the pre-re-vamp test run.

Original trays for the pasteurisa-tion section of the raffinate column were designed with a three-pass geometry. A chimney tray was positioned between three-pass pasteurisation section trays and two-pass rectification section trays. It is inherently difficult to achieve a uniform liquid-to-vapour traffic ratio in each section of the three-pass trays. Moreover, the original pasteurisation section trays and the chimney tray did not equip any feature for proper flow ratio balanc-ing. Each chimney pass open area

the base model were maintained for the improved column efficiency case study. Extra individual tray efficiency improvement was not considered. Original trays were arranged with 600 mm (~24”) regu-lar tray spacing. The increased number of trays was predicted through a reduced tray spacing scenario: 450 mm (~18”) regular tray spacing. As a higher tray count can increase the column pressure drop, increased column pressure drop values were applied for case studies of column efficiency improvement. The charts in Figures 3 and 4 show that the reflux ratios of the columns were positioned in the steep sensitivity zone and that enhancing column efficiencies is beneficial to improve energy consumption in both columns. Reduced reboiler duties of the extract and raffinate columns contribute to the xylene column furnace reboiler duty saving in the unit energy network.

Column modificationBased on the case study results, the extract and raffinate columns were modified. The number of trays was increased in both of the columns. At a given column shell height, a higher number of trays requires short tray spacing, causing tray capacity loss. To prevent column capacity reductions, GT-Optim high-performance trays were imple-mented and replaced the original sieve trays in both of the columns. The original xylene column trays remained unchanged. The higher- capacity nature of the GT-Optim tray maintains the desired column capacity with shorter tray spacing. In addition, the efficiency enhance-ment features of these trays can help to maximise column efficiency. Various performance- enhancing features adapted in the trays improve the vapour-liquid contact mechanism and enhance tray efficiency. These include specialised, shaped downcomers, liquid inlet momentum breakers, tray inlet vapour/liquid contact initiation devices and directional valves positioned in the tray periph-ery area. Tray pressure drop was optimised to prevent a “too low”

Case parameter Pre-revamp test run Post-revamp test runExtract column Feed rate, BPD Base +18%Overhead distillate rate, BPD Base +36%Feed temperature, °F Base +∆ 2.3°FReflux temperature, °F Base +∆ 16.9°FColumn top pressure, psi Base +∆ 1.3 psiReflux ratio (to overhead distillate), volume Base -28%p-DEB impurity in overhead distillate, wt ppm Base -∆ 1 ppmXylene impurity in bottom, wt ppm Base +∆ 16 ppm

Raffinate column Feed rate, BPD Base +14%Side cut product rate, BPD Base +28%Feed temperature, °F Base +∆ 5.6°FReflux temperature, °F Base +∆ 0.1°FColumn top pressure, psi Base +∆1.4 psiReflux ratio (to side cut), volume Base -6%p-DEB impurity in side cut, wt ppm Base +∆ 23 ppmXylene impurity in bottom, wt ppm Base +∆ 4 ppm

Xylene column Feed rate,1 BPD Base +30%Unit reboiler fuel consumption,2 EFO BPD/BPD Base -22%Unit reboiler fuel consumption,3 EFO BPD/BPD Base -9%250# steam generation, lb/hr Base +26%

Note 1. Total feed rate. 2. Equivalent furnace fuel oil consumption rate per feed charge rate.3. Simulated fuel consumption saving through the column efficiency improvement.

Test run data comparison

Table 2

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pre- and post-revamp cases. Slightly relaxed xylene loss and improved tray efficiency through GT-Optim trays contribute to achieving further reflux ratio savings in post-revamp operating conditions.

It is found that the raffinate column feed structure is changed in the post-revamp operating mode. Sensitivity analysis through simula-tion modelling shows that the post-revamp degree of separation line is substantially shifted by the changed feed compositions, and the pre-revamp degree of separation line is no longer applicable. In Figure 6, the post-revamp degree of the separation curve for the raffi-nate column is significantly shifted by the new feed composition. Although tray efficiencies are substantially improved and higher theoretical stages are achieved, changed feed composition erodes energy savings. Moreover, pre- revamp product purities cannot be maintained in post-revamp operat-ing conditions, therefore purities are a little relaxed.

The post-revamp economic eval-uations are updated and compared to the revamp target evaluations.6

The evaluations are based on 3% inflation, 10% weighted average cost of capital, 15-year deprecia-tion, 1% of the total investment for maintenance, 22% tax bracket and year 2012 average fuel price. Profitability indexes are expressed with regards to payback period, net present value (NPV) and inter-nal rate of return (IRR). These indices are shown in Figures 7-9. The charts show that actual revamp profitability is better than expected.

AcknowledgmentThe paper is updated from an earlier presentation given at the AIChE 2013 Spring meeting’s Distillation Topical Conference/Kister Distillation Symposium 2013, 29 April-2 May 2013, San Antonio, Texas.

GT-OPTIM is a mark of GTC Technology.

References1 Lee S H, et al, Optimize Design for Distillation Feed, Hydrocarbon Processing, June 2011.

operation simulation modelling and compared to the pre-revamp base modelling curves. Degree of sepa-ration lines between pre- and post-revamp operations are shown in Figures 5 and 6. The overall values gained from tray efficiency through simulation modelling are compared in Table 3.

In Figure 5, the extract column degree of separation curves have similar patterns between the

contribution of the column modifi-cations was simulated and is included in Table 2.

As product quality specifications are a little relaxed to maximise the energy saving and column feed structures are changed, it is neces-sary to re-evaluate the performance of the extract and raffinate columns with post-revamp operating condi-tions. The column efficiency curves are constructed using post-revamp

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Pre-revamp test run

Post-revamp test run

Revamp design

Pre-revamp degree of separationPost-revamp degree of separation

Figure 5 Extract column sensitivity analysis (pre- and post-revamp)

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Pre-revamp test run

Post-revamp test run

Revamp design

Pre-revamp degree of separationPost-revamp degree of separation

Figure 6 Raffinate column sensitivity analysis (pre- and post-revamp)

Case section Pre-revamp test run Post-revamp test run Extract column Rectification section, % Base +∆ 3 Stripping section, % Base +∆ 3 Raffinate column Pasteurisation section, % Base +∆ 9 Rectification section, % Base +∆ 7 Stripping section, % Base +∆ 6

Overall tray efficiency comparison

Table 3

gtc rev.indd 6 10/09/2013 10:50


design and troubleshooting for refi ning and aromatic applications. Email: [email protected]

Kwang Gil Min is the Senior Process Engineer for GS Caltex Corporation, Yeosu, Korea, and specialises in process engineering/operation services. He has been assigned various projects in the aromatic complex including reforming, aromatic extraction and xylenes separation units. Email: [email protected]

6 Largeteau D, et al, Challenges and opportunities of 10 ppm sulphur gasoline: part 2, PTQ, Q4 2012.

Soun Ho Lee is the Manager of Refi ning Application for GTC Technology US LLC, Euless, Texas, and specialises in conceptual process design, simulation modelling, energy-saving

2 Hanson D, et al, High capacity distillation revamps, PTQ, Autumn 2001.3 Meyers R, Handbook of Petroleum Refi ning Processes, McGraw-Hill Company, 1986.4 Moczek J S, et al, Control of a distillation column for producing high-purity overheads and bottom streams, I&EC Process Design and Development, 1963.5 Kister H, et al, Sensitivity analysis is key to successful DC5 simulation, Hydrocarbon Processing, October 1998.

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design targetActual

post-revamp

Figure 7 Profi tability index – payback period

Revamp design target

Actual post-revamp

20

15

10

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Figure 8 Profi tability index – net present value

Revamp design target

Actual post-revamp

150

100

50

IRR

, %

0

Figure 9 Profi tability index – internal rate of return

gtc rev.indd 7 10/09/2013 14:46

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Strategies for improved naphtha processing

Changes to the crude oil supply coupled with changes in gaso-line demand and new gasoline

regulations mean that US refiners must re-examine naphtha process-ing throughout the refinery. Many refiners will have access to opportu-nity crudes, including light, tight oils from shale formations in North Dakota, and Texas and Canadian bitumens that contain lighter dilu-ents. This article argues that refiners will need to upgrade their naphtha processing trains to better handle the added light material, comply with gasoline regulations, and adjust to changes in demand for gasoline and distillate.

Refiners in the Midwest have been revamping their facilities over the past 20 years to process Canadian bitumens and heavy crude oils. These refiners largely targeted improved bottoms upgrad-ing and gas oil cracking to accommodate heavier crude oils. In addition, many upgraded their naphtha processing facilities to accommodate the naphtha-based diluents used to transport bitumen via pipeline to their refineries. These naphtha-based diluents used to transport bitumen are becoming lighter.

More recently, the success of gas field fracturing (or “fracking”) is being replicated in the production of “tight oil” from the hydraulic fracturing of shale formations. Crude oils from these formations are lighter than traditional crude oils and have API gravities ranging from 40–70°API.

This article seeks to:• Describe the trends in crude oil production, in new regulations

To take advantage of dilbit from Alberta and light oils from shale formations, refiners must develop a strategy for processing additional naphthas and light components

MIKE ARMSTRONG and MARTIN BRANDTJacobs Consultancy

affecting fuel production, and in product demand that will increase pressure on the naphtha train• Define a strategy to take advan-tage of lower-cost crude oils containing lighter components by identifying naphtha processing bottlenecks in the refinery• Outline the technologies and configuration changes that should be considered as part of a strategy to process opportunity crude oils • Propose a methodology to evalu-ate and rank the various options for improving naphtha processing to take advantage of low-cost crude oils containing light components.

Canadian bitumen and diluent availabilityTo meet pipeline viscosity and density specifications, bitumen is blended with either light naphtha, which historically has been a natu-ral gas condensate, or synthetic crude oil (SCO) from Alberta-based upgraders. If bitumen is blended with diluents, it is called dilbit; if blended with SCO, it is called synbit. Dilbits contain large amounts of resid, heavy gas oil and naphtha with almost no kerosene or diesel range material. Processing the diluent contained in dilbit and processing the light material in tight crude oils may stress the naphtha and light ends processing portions of the refinery compared to processing other crude oils.

Growth in bitumen production is expected to significantly outstrip SCO production from Alberta-based upgraders as well as the availability of natural gas conden-sates used as diluent in shipping bitumen. As a result, bitumen

shippers will need to import increasing amounts of naphtha diluents, which will provide US refiners a means to send diluent recovered from bitumen together with refined naphtha back to the bitumen production sites in Alberta via recently commissioned pipe-lines built for this purpose. We anticipate that this returned diluent will become lighter as refiners seek an outlet for their least valuable light products: typically, butanes, pentanes and hexanes. The result will be that US refiners must be prepared to process dilbit contain-ing lighter diluent.

Tight oilTight oils tend to be high in API gravity. Processing these crude oils rather than heavier crude oils may mean that refineries will need addi-tional naphtha and distillate processing capabilities.

External factors affecting naphtha processingThere are a number of other exter-nal factors that affect naphtha processing capabilities in refineries and drive investment in naphtha processing. These factors are envi-ronmental specifications affecting fuel quality, vehicle efficiency standards and increased use of renewable fuels to supply US trans-portation fuels.

Control of hazardous air pollutantsfrom mobile sources (MSAT2)We have evaluated the impact of MSAT2 regulations for a number of US refiners and helped define their investment strategies for dealing with the required


jacobs.indd 1 10/09/2013 11:08


dynamic modelling tools such as Hysys and Pro-II together with yield estimating capabilities that enable the evaluation of changes in naph-tha processing on catalytic reforming, light naphtha isomerisa-tion, fluid catalytic cracking and hydrotreating. Our approach allows focus on the details in the context of the overall refinery and seeks to understand trade-offs and opportu-nities at the level of each new or existing process unit and at the refinery level.

There is a tendency when dealing with naphtha products to focus too early on the product specifications and then establish individual unit targets based on those specifica-tions. Our experience has led to a more holistic approach in assessing the naphtha train. Particularly when evaluating lighter feedstocks, it is important to take a forward view, starting with the crude unit and then moving down the naph-tha chain. Changes in the crude unit will impact many of the requirements and decisions for downstream processes. We believe the crude unit is where to begin evaluating the impact of processing more naphtha that comes in with the crude oil.

Figure 1 illustrates our strategy for evaluating the impact of a lighter crude slate on the refinery processing capabilities. The first question is: “Is the crude column adequate for the new feed composi-tion?” If no, the preferred new and/or revamp options should be laid out and evaluated, as each option may have a unique impact on the answer to the next major question: “Is naphtha hydrotreating capacity adequate?” Finally, the decisions made about crude unit and straight-run (SR) naphtha treat-ing will likely affect the saturated gas plant and will therefore set the scope for modifications of this unit.

Prior to evaluating expansion of the crude unit or any downstream units, it is important to define a base case complete with economic assumptions (price set) and projected crude slate. This step will require some modelling, preferably both LP and process, to establish how much new opportunity crudes

reduction in gasoline benzene. Our observation is that the options selected for benzene removal — whether extraction, saturation or sales of heart-cut naphtha — offer significant opportunities for process optimisation. For example, small changes in reformer feed composition and severity can create large benefits in terms of pool octane, gasoline yield and benzene saturation hydrogen consumption.

Changes in crude oil composition resulting from changes in diluent composition in dilbit and from increased processing of light, tight crude oil are likely to result in greater naphtha production at the refinery and will require a special-ised review of the full naphtha chain to achieve the following goals:• Ensure benzene removal facilities are still adequate to meet gasoline specifications• Leverage opportunities to improve benzene removal processes • Optimise naphtha reformer performance in line with refinery octane, benzene and hydrogen requirements.

Tier 3 gasoline regulations The EPA’s Tier 3 gasoline regula-tions will require reducing gasoline sulphur levels to an average of 10 ppm. Although it is possible to independently meet the challenges of Tier 3 regulations and take advantage of the opportunities from processing dilbit and lighter crude oils, a combined strategy that addresses Tier 3 and crude slate changes will reveal synergistic solutions that minimise total invest-ment and increase overall naphtha processing and gasoline production flexibility. The design basis for new and/or revamp scopes of work to accommodate the lighter diluent in Canadian dilbit and the light tight oils should consider the following points:• Tier 3 regulations will require additional desulphurisation of FCC gasoline, particularly light FCC naphtha that has typically by-passed FCC naphtha hydrotreat-ers under current regulations• Lower sulphur blending compo-

nents such as light, straight-run naphtha and alkylate will require closer attention and possibly additional treatment to reduce their sulphur levels• The isopentane value will decrease further if RVP specifica-tions are reduced, which will require refiners to remove high- vapour-pressure components from gasoline blends.

Changing demand for gasolineChanges in new light duty vehicle efficiency standards are expected to dampen demand for gasoline. Other federal and state regulations promoting the use of renewable fuels will reduce the demand for crude-based fuels. Market experi-ence with the renewable fuel standards to date has shown greater impact on gasoline than diesel fuel. The result of the renew-able fuel standard has been an increased use of ethanol in gaso-line, which has decreased the demand for refinery gasoline.

Increased refining of light, tight US crudes and Canadian dilbit blends may further drive down the demand for refinery naphtha and gasoline products in favour of kero-sene and diesel. The ability to shift products from naphtha into diesel may become increasingly important to refinery profitability in the future. Likewise, the ability to handle and sell large quantities of C4–C6 material will increase a refin-ery’s competitiveness by allowing the processing of opportunity crude oils containing light components.

Naphtha train optimisationA key step in taking advantage of dilbit and light, tight oils is develop-ing a robust strategy for processing the additional naphthas that accom-pany these crude oils. This means understanding the bottlenecks in the refinery and developing ways to relieve them. In evaluating these types of problems for clients and helping them develop tailored solu-tions, we use tools such as the refinery LP to help identify and value options, as well as tools that allow more detailed examination of crude oil, naphtha fractionation and naphtha processing. We use thermo-

jacobs.indd 2 10/09/2013 11:09

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a refinery can handle. This step may also involve some speculation as to how dilbit characterisation will change in the next five years. Once a base economic case is defined, it is possible to review opportunities around relieving constraints to increase light crude oil/dilbit processing.

Naphtha train optimisation follows a review of primary distil-lation and naphtha hydrotreating. Once the upstream units are defined along with the planned level of chemical (or diluent) naph-tha sales, we can determine if reforming units, benzene/aromatics units or cracked naphtha units require modification. Likewise, we can evaluate whether modifications comply with MSAT2 and Tier 3 gasoline regulations and then progress to the refinery’s gasoline and diesel production targets. Our goal with this exercise is to define a comprehensive set of configuration options in the early stage of project development before delving into detailed process evaluation.

The process configuration review should be part of a larger economic


model that considers crude and product prices, margin improve-ment opportunities and unit utility consumption. Key components of this larger economic evaluation are: • Ability to quickly determine product yields, operating costs and capital costs of refinery process configuration changes• Development of detailed economic models linked to process model outputs• Understanding of market dynam-ics and experience incorporating risks (both in pricing and project execution) in the refinery economic model to assist in good decision-making• Deployment of risk assessment to help understand uncertainties around the impact of changes in crude oil slate, product specifica-tions, product demand and processing changes.

Crude unit modificationsLight crude oils from tight oil formations and light diluent in dilbit may present refiners with unique challenges to the crude unit that will require a focus on the

overhead systems and options to “preflash” the light material. Processing light tight crude oils and light diluents will likely constrain the crude unit in the following ways:• Heat input and recovery: more light material and fewer intermedi-ates will reduce recoverable (above pinch) heat and increase the load on atmospheric heaters• Hydraulics: vaporisation in the crude train and increased heater pressure drop may limit crude unit throughput. Likewise, more light material will increase column load-ing in the atmospheric column• Overhead condensing: increased C5s and lighter material will increase the vapour pressure in overhead condensers, thus requir-ing either more cooling, more off-gas compression or higher column pressure.

Crude unit heat input limitations:prefractionator optionAdding a crude unit prefractionator offers the opportunity to remove C6 and lighter material from the crude. The pressure and temperature of a

CDU overhead, or column or

hydraulicmodifications.

Will cracked stocks be impacted by

adjusted pool yields?

Adjust design to minimise pool and

reforming impacts, and coordinate investments.

Prefractionation required?

Straight-run naphtha splitter modifications?

Is it possible to increase

naphtha sales?

What impacts do upstream modifications have on reformer and benzene unit yields?

Are crude units adequate for

feed?

Is straight-run naphtha

hydrotreating adequate?

Is saturated gas plant adequate?Revamp,

repurpose or new treater.

Yes

Yes

Review integration with CDU/NHT

Yes Yes

No

No

No

No

MSAT requirements Tier-3 specificationsDieselisation

Considerations

Primary distillation and hydrotreating review

Naphtha train optimisation

Figure 1 Strategy for evaluating lighter crude slates

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existing piping and downstream equipment.

Hydraulic constraintsHydraulic constraints associated with the processing of a lighter feedstock typically appear in areas of increased vaporisation:• Flash drum vapour lines• High-temperature exchangers not designed for vaporisation• Charge heaters• Atmospheric column.

As noted above, a prefractionator will remove these limitations by removing the light material upstream of the high-temperature vaporisation points. However, if the hydraulic limit is the only limit, a prefractionator cost will not be justifi ed. The following options will need to be reviewed:• Increased atmospheric column pressure: this change will reduce vaporisation in the heater and vapour traffi c through the column• Feed pump upgrade: higher head pump may remove the hydraulic constraints. Note that this option can be very expensive

• Light straight-run (LSR)/heavy straight-run (HSR) split may require further fractionation.

The design basis for the prefrac-tionator must consider the destination of the overhead product and side product (if there is one). If the overhead product is blended with atmospheric overheads and re-split, little fractionation will be done in the new column. If the material is segregated as a C6- LSR stream for gasoline blending, there may be some requirement to remove low-octane C7 material. Alternatively, if the material is to be sold either as ethylene cracker feed or diluent to Canada, a rough-cut product may be suffi cient, albeit with some economic penalty. Segregating the prefrac and atmos-pheric overhead naphthas reduces the load for the naphtha splitter, but creates the need for a new stabiliser. Routing of prefractionator products will likely present several opportu-nities and challenges that will be unique for every refi nery. Options will need to be evaluated in light of feed composition, operating goals,

prefractionator column can be adjusted to allow heat recovery to the crude unit through overhead exchange.

Adding a prefractionator is a logical consideration for units that are limited by heat input (furnace capacity), although additional heat exchange will be required. A prefractionator can also relieve hydraulic and overhead condensing limits, but is typically more costly than addressing these limitations directly. Recovering product naph-tha at lower temperatures can reduce operating cost, albeit with higher capital cost.

Prefractionator advantages:• Improvement in unit energy effi ciency • Revamp modifi cations outside of the turnaround window• Limited modifi cation to crude and vacuum columns (see Figure 2)• Potential for re-using or modify-ing existing prefl ash vessels.

Prefractionator disadvantages:• Potential for a high capital cost with little capacity increase or yield improvement

Heavy naphtha

Stripping steam

Stripping steam

Crude

Make-up water

Cooling waterLight naphtha

Intermediate naphtha

Desalter

Flash drum

Crude distillation column

New prefractionator

Vacuum residual

HVGO

Revamp area

(Flashed vapour line to be blinded or removed)

Cooling water

Vacuum distillation column

Existing crude lines

New equipment

Existing product linesRemoved lines

New lines

Figure 2 Crude unit revamp with prefractionator

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been defined for the upstream units. Outlining configuration options for primary distillation and hydrotreating is dependent upon existing equipment within those units and the overall processing goals. Comparing alternative configurations requires determina-tion of downstream unit yields, net product blend volumes and unit operating costs. To evaluate changes in upstream fractionation, we recommend a full naphtha train simulation, as outlined in Figure 4. An appropriate naphtha chain simulation model should include the following features:• Component characterisation through C8• Naphtha hydrotreating model including product separator and stripper• Isomerisation unit including naphtha stabiliser• Reforming unit model including kinetic yield predictions and sever-ity optimisation• Critical tertiary processing including benzene saturation or aromatics extraction• Refinery gasoline pool blending model including all key blending components to meet RVP, sulphur, octane, volatility and distillation specifications

Naphtha hydrotreating unit (NHT) expansionThe combination of lower-sulphur gasoline specifications resulting from Tier 3 regulations coupled with processing crude oils with

increase the number of mixed phase exchangers in series, because these designs increase the potential for maldistribution and fouling. Likewise, adding more exchangers in parallel often leads to poor distri-bution and subsequent fouling.

Overhead stripperFigure 3 illustrates the addition of an overhead stripper to the crude unit. By sufficiently stripping the heavy naphtha product, a new over-head stripper can allow a refiner to bypass the existing naphtha splitter and stabiliser because all the benzene precursors will be removed. This arrangement improves energy efficiency and HSR C6 content can be tightly controlled. As there is no rectifying section, LSR may see an increase in C7+ material.

Side drawsFor crude units with an extra diesel or jet draw, a refiner may choose to raise the column pressure and pull a heavy naphtha side draw. However, this change will reduce the top temperature in the crude unit and may create salting issues in the top of the column. Thus, in addition to any equipment changes, adding a sidedraw will likely include adding lining to the column, upgrading tray metallurgy and investment to improve desalter performance.

Naphtha train optimisationThe impact on downstream units and product blending can be reviewed once configurations have

if exchanger re-rating is required• Column upgrades: packing and high-capacity trays are typically a lower-cost option to deal with increased column traffic, although extensive column modifications may extend unit shutdowns• Preflash reconfiguration: a preflash vessel acts in a similar method to a preflash column, removing the light material that will create hydraulic limitations at higher temperatures. Existing preflash vessels can be modified or relocated to increase the tempera-ture and allow products to enter the atmospheric distillation column at a more advantageous location, thereby relieving hydraulic constraints.

Crude unit overhead modificationsIf no prefractionator is present, it is likely that the atmospheric column overheads and upper trays are operating at loading constraints. These bottlenecks will need to be relieved to accommodate more light material in the feed to the crude column.

Crude preheat exchangeCrude units not equipped with crude/atmospheric overhead exchange will need to consider this revamp option. Expanding over-head exchange may make sense for units that already have overhead exchange. Lighter crude oils will increase the amount of low-level heat available in overhead exchang-ers. Recovering this heat will be critical to maintaining unit heat performance in the absence of diesel and jet range material. However, a well-integrated crude unit may have limited sinks available for low-level heat recovery.

Two-drum systemSplitting the crude unit overhead into a two-drum system is typically done for corrosion control purposes (for instance, hot reflux, vapour/liquid distribution). For refiners with crude unit overhead limita-tions, a second drum may debottleneck overhead cooling and avoid significant repiping to add new exchanger capacity. However, refiners should avoid revamps that

Light straight run naphtha to stabiliser

Cooling water

Crude

Heavy straight run naphtha to hydrotreater

Atmospheric tower

New naphtha stripper

Figure 3 Crude unit overhead side stripper

jacobs.indd 5 10/09/2013 11:09


designed with an adequate gas plant, crude/naphtha train debot-tlenecks and revamps have typically avoided investment in the saturated gas plant. Often, new hydrotreaters and flare gas recov-ery systems will dump gas into an existing gas plant, further constraining operations, especially during summer months. Losing C3s and C4s to fuel gas, while extremely expensive, is generally tolerated to avoid limiting saturated gas plant capacity and refinery throughput.

Refiners looking to significantly increase C4- material in the crude slate will have to investigate ways to handle the increased load in the saturated gas plant. Figure 5 shows a typical configuration for a new saturated gas plant. Refiners will need to determine whether a new gas plant is required or if existing equipment, potentially across multi-ple plants, can be utilised. Among the key points to investigate are:• Suitable heat source for the stabi-liser (preferably integrated with crude/vacuum unit)• Adequate compression for multi-ple gas sources• Suitable lean oil for the absorber (stabiliser or naphtha splitter bottoms)• Limited products to storage —

light naphtha hydrotreater in the processing scheme.

Heat integrationModifications to the NHT should maximise the temperature of streams that are processed in down-stream units and therefore use waste heat where possible. The largest heat user in the NHT is typically the product stripper/debutaniser.

Design of the naphtha stripperIn many older units, the naphtha stripper has become a depentaniser and dehexaniser. Multiple towers may be required if upstream naph-tha splitting is not adequate.

Naphtha splitterIf light and heavy naphtha hydro-treating is done in a single unit, a single naphtha splitter can separate the stabilised product.

Saturated gas plantIn our experience, refinery satu-rated gas plants are often neglected and under-sized for their current operating requirements. Furthermore, saturated gas plants are often cobbled together using small columns and compressors from multiple obsolete plants. While crude trains are typically

greater light naphtha content will require additional naphtha hydro-treating capacity. The first question is whether the new hydrotreating capacity can be taken from existing units. Coker naphtha hydrotreaters and FCC naphtha hydrotreaters are often the newest units in a refinery and may have spare capacity. However, filling cracked naphtha units may reduce straight-run product value and restrict blending flexibility. Existing naph-tha hydrotreating units can be expanded with additional reactor beds or upgraded catalyst. Availability of fractionation capa-bility is often the biggest limit on existing unit capacity.

When expanding naphtha hydro-treater capacity, a comprehensive design basis should address some key considerations.

Process integration with the saturated gas plantThe NHT can be heavily integrated with the saturated gas plant. An NHT absorber column is commonly used; the NHT stripper can be converted to a de-ethaniser. In addi-tion, stabilisers, liquefied petroleum gas (LPG) splitters and depentanis-ers in the saturated gas plant may all be moved downstream of the

Saturated gas plant

Naphtha splitter

iC4’s

Saturated gas/LPG

OffgasUnstabilised naphtha

Light naphtha

hydrotreater

Coker naphtha

hydrotreater

Naphtha hydrotreater

Depentaniser

C6 isomerate

Reformer Reformate splitter

Benzene saturation

Catalytic naphtha splitter Catalytic

naphtha hydrotreater

Mogas blending

CDU/VDU

LPG sales

C5’s to sales

C5’sIsomerate

Saturated heartcut

Heavy reformate

Light catalytic naphtha

Heavy catalytic naphtha

Alkylate

Figure 4 Naphtha train model key blocks

jacobs.indd 6 10/09/2013 11:09


unit hydraulics, while downstream requirements determine perfor-mance targets. Thus, the adequacy of the existing naphtha splitter is defined by upstream supply and downstream requirements.

strategy chosen for revamping the crude unit train and naphtha hydro-treaters. The goal of the naphtha splitter is to match products to downstream unit requirements. Feed quality and quantity define the

direct to treating or next fractionator.

Naphtha splitterThe revamp strategy for the naph-tha splitter is dependent on the

Absorber stripper

Stabiliser

Cooling water

Flare gas recovery discharge

Overhead naphtha

Crude compressor discharge

Hydrotreater/ reformer offgas

Fuel gas to treating

LPG product to treating/splitter

To naphtha splitter

Diesel pumparound

Heavy gas pumparound

Lean oil (debutaniser bottoms)

Cooling water

Cooling water

Figure 5 Typical saturated gas configuration

increased water in the sulphur plant feed and thus more condensa-tion in the sulphur plant feed NH3 gas knock-out drum (and thus water that has to be recycled back to the sour water stripper).

The design shown in Figure 2 was common in the 1960s. However, it also suffers from the same heat balance drawbacks and needless complications as seen in

www.eptq.com PTQ Q2 2013 101

Figure 1. In other words, a lot of equipment is added to generate refl ux when no fractionation is required between the feed and overhead product. Again, the only purpose of the tower is to strip out the NH3 and the H2S.

Correct stripper designIn 1969, while working for the now vanished Amoco International Oil

(1)

(8)

(15)

Sour water feed

90ºF

Low NH3 water to desalters

Steam

250ºF

E-1

P-1

13 PSIG

Ammonia gas to sulphur plant

190ºF

Medium content NH3 water to hydrotreaters

Figure 4 Two-stage sour water stripper design without feed preheat

Company in the UK, I designed a sour water stripper that eliminated the unnecessary features of the unit shown in Figure 2.

Figure 3 shows the essentials of a correct sour water stripper design. Feed is brought in at ambient conditions (70-100°F, 21-38°C) from the sour water feed tank. To heat the feed from 90°F (32°C) to 250°F (120°C) requires about 16 wt% steam fl ow, or about 1.3-1.4lb of steam per gallon of stripper bottoms, which is close to a typical design stripping steam ratio for sour water strippers. The E-1 feed preheater, refl ux pump (P-2) and the refl ux cooler (E-2) shown in Figure 2 are all eliminated. How, then, does one know that the design shown in Figure 3 will work? Because it was built this way (at the Amoco refi nery in Milford Haven, Wales, UK) in 1970, where it worked just as well as the conventional design shown in Figure 2.

Two-stage sour water stripperFigure 4 shows a sour water strip-per with a side draw-off. The partly stripped sour water is extracted from tray 8 and directed to the hydrotreaters for use as make-up water in the salt (NH4HS) removal step of the reactor effl uent. Completely stripped water from the sour water stripper bottoms is sent to the crude desalter. While

102 PTQ Q1 2013 www.eptq.com

communicate where the module will be installed on the plot plan. Connections between the modules are designed to be similar in config-uration so that construction is relatively straightforward. Ventech estimates that, with modularisation, approximately 70% of a project is already complete even before the modules are shipped from their facility. This greatly decreases field construction time to deliver an operational facility (see Figure 1).

These methods also facilitate easy disassembly and relocation, if necessary, at some point in the future. For example, a remotely located gas processing facility could be easily taken apart and moved to a new natural gas source if an exist-ing supply was depleted in its current location.

Applying modularisation to refin-ery construction has advantages with regard to productivity, prod-uct quality and ensuring the safety of construction personnel. Since the modules are built in a well-lit, climate-controlled environment, work can continue around the clock regardless of weather conditions, for greater productivity and easier quality control. Since module height is restricted, safety is enhanced, as workers build at limited heights within the fabrication facility.

Modularising GTLThe same advantages of modular

construction of refineries are being applied to the construction of distributed GTL plants. The GTL process involves two operations: the conversion of natural gas to a mixture of carbon monoxide (CO) and hydrogen (H2), known as syngas, followed by a Fischer-Tropsch (FT) process to convert the syngas into paraffinic hydrocarbons that can be further refined to produce a wide range of hydrocarbon-based products, includ-ing clean-burning, sulphur-free diesel and jet fuel. Speciality prod-ucts including food-grade waxes, solvents and lubricants can also be produced from the paraffinic hydrocarbons.

Large, commercial-scale GTL plants, including the Sasol Oryx and the Shell Pearl plants (both located in Qatar), have been built at enor-mous capital cost. The Oryx plant, designed for production levels of 34 000 b/d, cost around $1.5 billion to build. The Shell Pearl plant, with an ultimate design capacity of 140 000 b/d of GTL products and 120 b/d of natural gas liquids, cost around $18-19 billion. Conventional GTL plant designs rely on econo-mies of scale to drive positive financial returns and are viable only where there are large supplies of low-priced natural gas.

However, another option being developed — smaller-sized and distributed GTL plants — shows

promise for deriving value from smaller accumulations of unconven-tional gas that would otherwise be left underground, such as shale gas, tight gas, coal bed methane and stranded gas (gas fields located too far from existing pipeline infra-structure). A small, modularised GTL plant has the flexibility to be installed close to the trapped resource and then used to process that resource locally. Associated gas (gas produced along with oil) is another area of opportunity for modularised GTL plants. This gas is typically disposed of either by re-injection, at considerable expense, back into the reservoir or by the wasteful and environmentally damaging practice of flaring, which is subject to increasing regulation. Modularised GTL plants enable this otherwise wasted gas to be converted into additional revenue.

In the larger economic picture, a modular GTL capability can be the key factor that enables the construc-tion of upstream projects that would otherwise be cancelled because of poor results derived from economic models. For exam-ple, some shale gas discoveries are being hampered by high develop-ment costs, which result in marginal economics due to gas prices that are often low. These projects can be enhanced by converting the gas to higher-value clean fuels produced in the GTL process.

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area for most refiners. Any change within the naphtha chain may justify review and re-optimisation of upstream splitter operating targets.

Fluid catalytic crackingnaphtha processingFCC naphtha processing has become critical to meeting product specifications, particularly blended gasoline RVP, octane and sulphur content. Further, Tier 3 gasoline specifications are likely to require expansion of FCC naphtha hydro-treating facilities. Thus any investment to increase straight-run naphtha processing is inexorably tied to FCC naphtha handling. Any issues with C5 disposition, RVP control and benzene slip will also require focus on the FCC naphtha performance.

There is no single answer on how to integrate or coordinate projects between straight-run naphtha and FCC naphtha. Among options to consider are: • Removal of sulphur from light FCC naphtha• Alkylation of amylenes to increase RVP blending flexibility• Processing of light straight-run material in the FCC naphtha hydro-treater (CNH)• Hydrotreating and reforming of intermediate cat naphtha (ICN) heart-cut• Scope for increasing FCC naph-tha desulphurisation• Maximisation of heavy FCC naphtha (HCN) to diesel.

A successful naphtha train opti-misation will require a good understanding of the contribution of FCC naphthas to the refinery gasoline pool.

ConclusionsDiluent used in shipping bitumen as dilbit from Alberta to US and Canadian refineries is expected to become lighter as a result of market changes. Light, tight crude oils from shale formations are consider-ably lighter than most crude oils being processed in US refineries.

Refining these crude oils will apply stress to many refineries’ naphtha and light ends processing capabilities. In addition, external

For reforming units that produce reformate for gasoline blending, a key point of optimisation will be to ensure the right cut points of feeds to the unit. Refineries with more than one reforming unit have more degrees of freedom. As a result, feedstock optimisation to multiple reforming units is more complex but can have a significant benefit.

If the available heavy naphtha exceeds reforming unit capacity, there may be opportunities to remove the C6 and lighter material from the feed to the reforming unit to avoid benzene production as well as remove C10 and heavier material from the reforming unit feed, which may improve catalyst life.

Refiners with aromatics produc-tion facilities, conversely, may have a distinct advantage in dealing with the increased naphtha content of the lighter crudes. Benzene, toluene and xylene (BTX) have always

offered a significant margin upgrade over gasoline. As such, the ability to make BTX from the avail-able naphtha may offer a significant economic opportunity, particularly given the lower energy costs result-ing from the drop in US natural gas prices. Refiners that have the ability to increase aromatics production may be looking for investment in this area, provided a market exists to offload the butanes and pentanes.

Optimisation of the reformate splitter will figure prominently in product blending strategies. Once these strategies are clearly defined (for instance, C5 disposition, RVP or benzene targets), it may be possible to derive significant benefit by adjusting splitter cutpoints and level of fractionation. The balance and optimisation of feedstock to the benzene saturation unit is a new

There are a number of decisions regarding the naphtha splitter that will vary with the design of the crude unit and downstream processing requirements. Among the key considerations are:• Should naphtha be split before or after hydrotreating? Hydrotreating capacity may limit flexibility, though it is likely that the NHT unit will need to be expanded due to an increase in naphtha produc-tion as a result of processing lighter crude oils and dilbit containing lighter diluent• What level of benzene precursor removal is optimal? Benzene satu-ration units and benzene regulations have significantly changed operating strategies regarding naphtha cut-points • Is there an advantage to taking a side cut in the naphtha splitter? Dehexanising and depentanising can be significantly cheaper in a single column, although product quality may suffer• How best to heat integrate the naphtha splitter? Straight-run naphtha splitting is often heat inte-grated with the crude/vacuum column, although some older units have integration with the naphtha hydrotreater and/or reformer. To achieve best-in-class energy perfor-mance, the unit should take advantage of medium-level waste heat sources (eg, 150 psi steam or light diesel pumparound).

Reforming unitThe drop in US gasoline demand, together with greater blending of ethanol, has reduced the need for high-octane reformate to increase the gasoline pool octane. Likewise, low-cost natural gas has reduced the cost of hydrogen from steam meth-ane reforming. As a result, many refiners are reducing the severity of their catalytic reforming units to maximise C5+ yields and minimise benzene production, which provides opportunities to increase catalytic reforming unit capacity with mini-mal investment in product fractionation and unit hydraulics. There may also be an increased bypass of feed around the catalytic reforming unit, provided that pool octane can be maintained.

Optimisation of the reformate splitter will figure prominently in product blending strategies

jacobs.indd 8 10/09/2013 11:10


technology options exist for refi ner-ies to consider and evaluate• A detailed naphtha chain model will simplify evaluation of potential new unit and/or revamp options.

Mike Armstrong has 25 years of operations and process design experience in the refi ning industry, including management of operations activities, technical support, evaluation of project economics, and refi nery confi guration studies. Since joining Jacobs Consultancy in 2005 as a Senior Consultant, he has served as process lead on several revamp and grassroots feasibility studies. He holds a degree in chemical engineering from the University of Illinois and a Master’s in chemical engineering from the Colorado School of Mines.

Martin Brandt is a Director in Jacobs Consultancy’s Chicago offi ce. He has over 23 years of experience, with a focus on providing commercial and technical support to domestic and international clients in refi ning and petrochemical industries. He has served as Project Manager on multiple refi nery revamp studies to increase refi nery capacity, process advantaged crudes, maximise distillate production, and reduce gasoline benzene and sulphur levels. He holds a BS ChE degree from the Illinois Institute of Technology and an MBA from the University of Chicago.

product regulatory requirements and interaction of naphtha process-ing units within the refi nery, along with application of appropriate technologies to achieve objectives.

We offer the following key observations: • Refi neries processing dilbits from Alberta will see increased C4–C6 range material in the dilbit. • Fracking of shale formations will provide increased opportunities for refi ners to process light, tight crude oils. • Regulations and market trends will likely: ■ Increase naphtha desulphurisa-

tion requirements ■ Reduce the need for octane from

the catalytic reforming unit ■ Reduce gasoline demand from

crude oil ■ Shift the gasoline-to-distillate

ratio from refi ning• Changes to refi neries will require investment in primary distillation units, saturated gas plants, and/or naphtha hydrotreating capacity• Multiple confi guration and

factors such as MSAT2, Tier 3, the Renewable Fuel Standard and rising effi ciency standards for light- duty vehicles will put stress on the production of transportation fuels from US refi neries, in particular gasoline.

To take advantage of dilbit from Alberta and light, tight crude oils from shale formations, while also managing regulatory and market changes, refi ners must develop a robust strategy for processing the additional naphthas and light components that accompany these crude oils. Doing so means under-standing the bottlenecks and processing limits in the refi nery and developing ways to relieve them.

We have outlined an approach that focuses on processing the addi-tional naphtha in the context of the overall refi nery, while at the same time evaluating trade-offs and opportunities at the level of each existing and new process unit. Our process requires an understanding of market trends, feed properties,

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Sulzer Chemtech

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

Background Initial liquid distribution directly affects the efficiency of structured and random packing beds. As such, some form of flow test should be conducted to validate distributor performance prior to column startup. Steady state water flow testing is the most common method of testing distributors. It can be conducted in the shop or in the field to verify the design, manufacture and installation.

Three Common Types of Liquid Distributor Flow Tests• Liquid Range Test with Level Measurement: Water is fed

to the distributor at the maximum and minimum design operating rates. At each rate, level measurements are taken at various locations within the distributor, typically at the end of each trough. Relative standard deviation of the level measurements must fit within the manufacturer specifications (normally 5-10% of the calculated flow rates). The flow patterns within the distributor should appear uniform, without excessive localized momentum. Sulzer conducts this mandatory flow test for all distributors.

• Area Measurement: This test is more common for large diameter and spray header distributors. Identical pans are placed at several locations under the distributor to collect flows from multiple distribution points over a constant timed collection period. The liquid flow rate is then calculated for each pan (area). The standard deviation of the flows to each pan is used to determine the macro-scale uniformity of liquid distribution across the distributor. These tests are typically conducted at minimum, design, and maximum flow rates.

• Point Measurement: This test is similar to the area measurement test except that flows are measured from individual distribution points at predetermined locations for a micro-scale evaluation.

The standard deviation calculated from these samples must be within design standards. Typically, a minimum of 10% of the total number of individual distribution points are measured.

ConclusionsFlow tests can quickly reveal maldistribution that could adversely impact column performance. By identifying issues early, simple adjustments can be directly made to the equipment while on the test stand, prior to column startup, avoiding costly shutdowns.

Final Note – InstallationNo matter how well a distributor is designed and verified during manufacture, poor installation can substantially compromise column performance. All distributors must be properly oriented and leveled within the tolerances shown on the assembly drawings provided by the equipment manufacturer. Always verify distributor installation in the field prior to startup.

Distributor on Test Stand for Level Measurement

Tower Technical Bulletin Safeguard Packed Column Efficiency with Proper Liquid Distribution

The Sulzer Refinery Applications GroupSulzer Chemtech has over 150 years of in-house operating and design experience in process applications. We understand your process and your economic drivers. Sulzer has the know-how and the technology to design internals with reliable, high performance.

Distributor Ready for Point Measurement

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Laser surveying a revamp

Plant revamp projects are big business, but they are also particularly challenging. The

new installation must arrive on time and fit right first time, or the contractor runs the risk of contrac-tual penalties. Achieving this demands not only efficient business processes, but also detailed and reliable knowledge of the true condition of the existing facility.

This article describes current capabilities and best practice in the use of 3D laser surveying for de-risking, planning and executing revamp projects.

In an ideal world, every operat-ing plant would have a complete set of detailed and up-to-date engi-neering drawings or, better, a 3D CAD model that accurately describes it. In practice, not even the most capable of owner-opera-tors would claim to have achieved this, although a few do get close to it. Design documentation rarely reflects the true as-built state of a plant, even at handover. Several years later, the divergence can be significant, as modifications and repairs accumulate without always being properly reflected in drawing updates. When tendering for a revamp project, therefore, any wise engineering contractor (EPC) must assume that no existing design information is trustworthy. Accurate surveying is essential.

Like the CAD tools used in plant design, surveying technologies have come a long way. Tape meas-ures and theodolites gave way to photogrammetry, which, properly used, enabled a 3D model of an area to be compiled. However, recent rapid advances have made

Efficient project workflow in a revamp can be achieved using 3D design technology that integrates with laser scan and engineering data sources

GARY FARROWAVEVA

3D laser scanning the tool of choice for surveying an operating plant. Today, there are a number of high-quality scanner systems and service providers to draw on. Equally important, there is also very sophisticated software for exploiting the rich data that laser scanners generate.

Equipment compactness and usability have advanced as rapidly as the now ubiquitous digital camera. Scanners are quick and easy to set up and use, and can capture a hemispherical 3D representation of their surroundings to an accuracy of a few millimetres over distances of tens of metres. However, just as a camera can take either a blurry snapshot or a studio portrait, according to the skill of the user, so the apparent simplicity of laser

scanning can raise expectations of its capability that the unwary user, or their client, may be disappointed with. It is a tool and, as with many tools, there are right and wrong ways to use it.

Our team at LFM Software regu-larly encounters users of laser scanning who have experienced problems or sub-optimal results. Almost invariably, the causes lie in not appreciating a few basic princi-ples. A business paper summarises these.1 Our top tips for a successful laser-surveyed revamp project follow.

Plan the projectJust like professional photography, success lies in preparation. When planning a laser survey, especially when only a tight time window is


Figure 1 AVEVA Everything3D: Bubble View and 3D model working together to enhance the value of captured laser data

rev aveva.indd 1 10/09/2013 11:20


different times using different scan-ners, there is a risk that these datasets may not all be compatible with the design software. At AVEVA, we have created applica-tions and interfaces that can handle all the commonly used scan data formats, so that multiple individual surveys can be integrated into a single model. Users should either confirm that their own design solu-tion can achieve this or, if not, restrict surveying to a known compatible scanner.

Much has been achieved in reverse engineering 3D point clouds into CAD model objects. Software can recognise simple geometric solids such as cylinders and create the corresponding 3D model. However, if the data repre-sent a pipe, which has attributes and connectivity on a P&ID, there is relatively little value in model-ling it as a solid cylinder. Current state-of-the-art software can infer that the data really do represent a pipe, can offer the designer a short-list of its possible specifications from a design catalogue, and then reverse engineer an intelligent pipe object in the design system, accu-rately co-located with the original scan representation (see Figure 2). This can save considerable costly effort and brings reverse engineer-ing of large plants into the realm of the practicable and affordable.

Create accurate deliverablesThe point of laser surveying for revamp projects is to enable new fabrication to fit accurately in place and an installation sequence to be planned to achieve this with mini-mum disruption to the plant. In turn, this requires that the laser data can be used in a 3D design system in such a way that a designer can create a new design that mates accurately with in-situ installation and does not clash with anything. Not all design solutions are created equal, either in their ability to work with scan data, or in their handling of clashes, or in the quality of their deliverables for fabrication and installation. Clash handling can vary widely; a single pipe model might clash with perhaps thousands of individual

available, it is essential to coordi-nate the survey with other, unrelated plant activities to ensure that the work can be performed in the time available. For example, one cannot be expected to survey an area if a contractor’s crane is parked there for an essential main-tenance task. Where areas to be scanned are normally off-limits, their being made accessible may be time-limited, so one should plan the survey sequence to accommo-date this. A common pitfall is to attempt a more extensive scan than time or circumstances permit and be forced to take shortcuts.

Conversely, just as an owner- operator tries to maximise the work performed during a shutdown, so it may be worth taking the opportu-nity to scan more than just the area of immediate interest, to anticipate future requirements. Forward-looking owner-operators are increasingly commissioning whole-plant surveys for this reason.

Plan the scanWhile currently available software can often work wonders with sub-optimal scan data, there are limits. High-quality survey data require a well-planned scan sequence. An experienced user will plan the number and positions of individual scans so as to achieve adequate overlaps between succes-sive “point clouds”. This will ensure that the point clouds can be

stitched together accurately on a common coordinate system and that areas shadowed in one view are scanned in at least one other (see Figure 1). While there are some guiding principles to optimum scan distribution, each situation is unique and a pre-survey site inspection can be time well spent to achieve a quality result.

Laser scanning is safe and non- invasive, but one should neverthe-less plan to avoid or restrict personnel movements in the area. Structural vibration can degrade

scan accuracy, while it is not uncommon to find a scan including a detailed image of half a person standing in front of some important object, which may require re-scanning.

Use the right toolsThere is little difficulty in selecting scanner hardware; all the leading suppliers’ systems are excellent. However, if scans are performed at

Figure 2 AVEVA Laser Modeller displaying intensity window and 3D window while modelling structural elements. The user is being prompted for catalogue-based structural components

While currently available software can often work wonders with sub-optimal scan data, there are limits

rev aveva.indd 2 10/09/2013 11:20

scan points representing one as-built pipe. Evidently, that is only one clash, but some systems will register it as thousands of clashes, making it diffi cult to identify the real clash and to resolve it effectively.

Ensuring that an accurately designed pipe spool actually fi ts correctly on site requires accurate fabrication. Leading 3D design solutions such as our PDMS or AVEVA Everything3D can not only generate fully detailed fabrication drawings automatically, they can also perform manufacturability checks at the design stage to help maximise fabrication quality. It is also now possible to scan a completed fabrication and compare it against the design model, to quickly verify its accuracy and resolve any errors early.

Owner-operators considering placing revamp projects should review a contractor’s capabilities in these areas. The most capable typi-cally achieve less than 1% design-related rework costs, even on complex projects. Our vision of plant design for lean construction goes further than this; our goal is to use laser scanning, among other tools, to completely eliminate rework in construction. This vision is discussed further in a business paper.2

Plan the installationScheduling the revamp installation involves similar considerations to planning the original survey. EPCs and owner-operators must collabo-rate closely to achieve a well-executed installation. Here, the power of 3D design can add considerable value.

Reverse engineering objects that are to be removed enables the crea-tion of accurate demolition drawings and the determination of weight and centre of gravity, which informs the correct use of handling equipment. Model animation enables planners to evaluate proposed task sequences and to check that, for example, items can be moved safely in the available space constraints. Design review is arguably even more important for revamps than for new-build


projects; revamps take place on plants that may contain hazardous chemicals, temperatures or pressures. Evidently, these consid-erations imply the need for well-specifi ed, specialist plant design software.

Business processThe introduction to this article referred to the need for effi cient business processes. That is hardly a great insight, but, in the engineer-ing industry, business processes are inextricably linked with the engi-neering and design technologies that generate project information. Best practice is therefore to select solutions that can share informa-tion effi ciently and reliably. An effi cient revamp project workfl ow can thus be achieved using 3D design technology that integrates both with laser scan data sources and with engineering data sources, so that engineering and design information can be kept synchro-nised as the project progresses. From this, new design can automat-ically generate accurate materials requirements that feed into an enterprise resource management (ERM) system.

Such engineering, design and information management technolo-gies now exist and are in use on a wide variety of new-build projects. Their ability to support effi cient business processes becomes even more important on revamp projects with their need for on-time, right-fi rst-time, low-risk installation.

References1 Lighting the Way, www.aveva.com/publications2 Plant Design for Lean Construction: Innovation for a new era in plant design, www.aveva.com/publications

Gary Farrow is Vice President 3D Data Capture with AVEVA in Cambridge, UK. He works with customers in the use of 3D data capture technology to increase productivity and to advance the performance of AVEVA’s LFM software. A mechanical engineer, he has been involved in 3D laser scanning from its inception in the late 1990s, initially undertaking projects delivering data and 3D models, including a huge scanning project for Fluor/TCO in Kazakhstan.

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Retrofi tting a glycol contactor to prevent carryover

The gas production rate on an offshore platform was constrained due to a high

level of glycol carryover from glycol contactors. The existing column internals were investigated and it was found that the area of the mesh pad in the top of the column was around half of the column area. In addition, the gas outlet location and associated piping arrangement appeared to create vapour maldistribution. A high-capacity Shell Swirltube separator and Sulzer structured packing internals were proposed in conjunction with a new gas outlet arrangement and the existing separator support system. CFD studies were carried out to evaluate the possible vapour distribution, based on which a revamp proposal was fi ne-tuned. New internals were manufactured and installed on site in 2012. Recently, the design was validated in high-rate trial runs, with signifi cantly lower triethylene glycol (TEG) carryover.

All together, there are three trains on the platform. The facility dehy-drates the gas and condensate separately, and then recombines the dried products before transporta-tion via trunkline(s) to the gas plant. Subsequently, the gas plant processes these fl uids to produce LNG, condensate and LPGs.

In 2006, a customer in the Asia Pacifi c region decided to proceed with the design of a new bridge-linked facility to its plant. The compression installed on the new plant would allow 20 000 t/d of gas throughput per train. Specifi cation of the glycol content of the export gas plant must be less

A revamp of glycol contactor column internals, following CFD studies, resolved problems with gas production caused by glycol carryover

ANNE PHANIKUMAR and YANG QUAN Sulzer Chemtech Switzerland

than 4 m3/d to prevent onshore upsets.

During a trial in 2007, dehydration capacity as well as TEG losses were evaluated at various fl ow rates. The gas rate was increased to around 18 000 t/d per train, at which point it was found that the glycol losses reached the maximum limit of 4m3/d (total of three trains). Dew points were found to increase and were worse at the higher rates. Condensate dehydration was estab-lished to be adequate. Overall, dehydration performance was found to be acceptable for the fi rst and second trunklines, where the satura-tion specifi cation had basically not been affected and remained at around 25% and 19%, respectively, which was less than the integrity limit of 75% and the operating limit of 65%. Therefore, it was concluded that the dehydration of both gas and

condensate was suffi cient at 18 000 t/d. This gas throughput became constrained by glycol carryover losses. It was decided by the customer to revamp the third train in early 2012 and leave the other trains to a later date. This would allow confi dence and experience to be gained with minimum exposure (a single train rather than all three trains). Sulzer was selected to revamp the internals of the contac-tor to minimise the TEG losses at higher gas rates in the customer’s complex.

In the late 1980s, to increase throughput, the original bubble cap trays in the glycol contactor were upgraded to structured packing (GEMPAK 3A, which is comparable to Sulzer Mellapak 350Y). In the column top, a mesh pad was installed to prevent glycol losses. It covered about half of the column area (see Figure 1). A lean glycol distributor was installed above the structured packing, and in the bottom there was a chimney tray to collect the rich TEG before being directed to the glycol regeneration system (see Figure 2). This was the column internal confi guration to be revamped. To avoid having to stress release the column after revamp, welding to the column wall was not permitted. The existing separator support system had to be reused for the various revamp options.

Gas-liquid separator in the column’stop section The location and arrangement of the gas outlet was relatively unique, as the gas outlet nozzle is located later-ally below the mesh pad. An elbow-shaped internal pipe passes

www.eptq.com Ravamps 2013 33

N2

Figure 1 Sketch of the existing mesh pad and the gas outlet arrangement

sulzer.indd 1 10/09/2013 12:36

34 Ravamps 2013 www.eptq.com

the trial in 2007 could immediately be attributed to the small mesh pad area, which was only half of the packing cross-sectional area. The arrangement of the existing mesh pad was perfectly symmetrical, but the space between the column head and mesh pad was very narrow, causing a poor infl ow to the gas outlet piping above the mesh pad and an uneven vapour distribution through the mesh pad. This maldis-tribution also contributed to the observed excessive entrainment of TEG.

Any revamp proposal was limited by two key factors:• The gas outlet could not be relocated• Hot work to the column wall was prohibited.

Therefore, the existing shroud had to be reused, and it was neces-sary to install high-capacity demisting equipment to compen-sate for the limited area available for a separator.

There are several types of gas-liquid separators on the market, and they can usually be classifi ed into three categories: mesh pad, vane pack and axial cyclone types. The selection of different types depends on required capacity, pressure drop and effi ciency (indicated by cut-off size). Table 1 shows a simple summary of the selection criteria. It has to be noted that values in the table are indicative only. They are subject to individual separator prod-ucts and operation systems.

With increasing operating pres-sures, vane packs, even if they are equipped with a coalescer in the front, suffer higher effi ciency loss than axial cyclones. The operating pressure of this TEG contactor is 108 bar. Axial cyclones are the best option for having up to three times higher capacity than mesh pads per unit area. They can compensate for the limitation caused by the smaller area available for separation. A Shell Swirltube separator was used in the investigation of various options.

Direct welding to the column wall was forbidden, so it was proposed that the existing shroud be extended downwards by weld-ing additional parts onto it. To

through the centre of the mesh pad. This pad was supported by a shroud welded to the column head. Usually, the gas- loading capacity of

structured packing in TEG contac-tors is higher than that of a wire mesh pad. Obviously, the excessive glycol entrainment observed during

New Sulzer HC separator

New Sulzer mesh pad

New Sulzer liquid distributor

New Sulzer structured packing

Existing packing support

Existing chimney tray

Existing feed inlet

Bed 1

TL

N2 M2

N1 M1

Figure 2 Proposed Sulzer KnitMesh coalescer and Shell Swirltube separator

∆P

∆H

∆P

∆H

Liquid backup to counterbalance ∆P

Liquid cannot drain if ∆P is too high

Figure 3 Principle of liquid backup in a drain pipe

Mesh pad Vane pack Centrifugal deviceCapacity (per unit area) 1 1-3 3Pressure drop Base Low HighCut-off size (glycol) Down to 3-5 mm Down to 20-30 mm Down to 10 mmOperating pressure limit - Up to 50 bar in particular - if hydrocarbon liquids are present

Performance of various types of gas-liquid separators

Table 1

New Shell Swirltube separator

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Even the most seasoned refinery managers may see four to five FCCU turnarounds throughout their career. AltairStrickland’s managers and craftsmen have performed, on average, four to five FCCU turnarounds per year since 1976. This experience is the kind of advantage you need to help you manage and execute a successful project.

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ous sections to make sure a human, with tools, can access and work effectively and safely in tightly confined spaces.

Our crews, from craft to top management, have worked on so many turnarounds and revamps that they can identify and quickly correct a problem, and work around a snag or lagging schedule.

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maximise the number of Swirltubes and to take into account the dimen-sions of the standardised cyclone modules, the outlet piping had to be shifted from the centre to the side of the shroud. Internal piping for the dry gas outlet nozzle needed to be cut and welded onto the extended shroud. A schematic of the proposed new arrangement is shown in Figure 2. This arrange-ment posed some uncertainty with regard to vapour distribution across the new Swirltube deck due to sudden changes in the vapour fl ow direction of 90 degrees after passing through the Swirltube separator and the asymmetry of the new arrangement in general. The typical pressure drop of a Swirltube separator is much higher than for a mesh pad (30 mbar versus 0.5 bar). This helps to improve vapour distribution but, because the concern for maldistribution could not be eliminated, CFD studies were conducted to understand and quantify the risk.

There is another aspect to consider in the design of the axial cyclones. The liquid drops sepa-rated by the Swirltubes are collected in liquid collection cham-bers, from where the liquid is drawn off through a pipe system. The downpipe is subject to liquid backup caused by the pressure drop through the Swirltube deck. If the drainpipe’s vertical height is not suffi cient to accommodate the liquid backup, the liquid may fl ow back onto the deck and cause fail-ure of the separator.

The customer’s previous experi-ence showed that there was signifi cant condensate carryover from the upstream production separator to the glycol contactor. It was therefore expected that condensate could be entrained together with TEG into the packing section and further upwards to the demisting device itself. Due to the huge difference in density between TEG and condensate, and the limited height between the separa-tor and TEG distributor, the downpipes for the separator should extend to the rich glycol sump to cope with the condensate backup. Equation 1 describes how the liquid

www.eptq.com Ravamps 2013 37

A

B

C

0.00.20.40.60.81.01.21.41.61.82.0

Velocity, m/s

Figure 4a CFD model confi guration showing the high-capacity separator (green), the mistmat (pink), the liquid distributor (cyan) and the gas exit nozzle (side draw) arrangement

Figure 4b CFD vapour fl ow trajectories represented by the velocity magnitude, where blue represents the low velocity and red represents the velocity peak of the scale

Figure 4c CFD vapour fl ow vector plot represented by the velocity magnitude, where blue represents the low velocity and red represents the velocity peak of the scale

sulzer.indd 3 10/09/2013 11:29


top conditions of the TEG contactor and asymmetric confi guration of the proposed arrangement caused concerns over possible maldistribu-tion of the vapour and uneven loading of the individual swirltubes of the deck. CFD studies were carried out for four operating load cases. These cases represented minimum and maximum loadings corresponding to pre- and post-revamping conditions to verify the risk by looking into the velocity profi les across internals and column cross-section at various elevations of the tower.

The CFD model confi guration is illustrated in Figure 4a. The vapour fl ow through the column is simulated numerically using a general-purpose commercial fl uid dynamics code.

To model the turbulence behav-iour of the fl ow, the standard k-ε model is used. The pressure drop over the mistmat is described by means of the resistance law:

∆p = k H ρ v2/2, (2)

Where k is dependent on the oper-ating condition and the type of mistmat.

Porous media has replaced the mesh pad in the CFD model confi g-uration. The liquid distributor and high-capacity separator are modelled as accurately as possible to represent the actual pressure drop through these devices. Separator drain downpipes and the liquid distributor feed pipe are not included in this study, as they are not needed to simulate the vapour fl ow inside the vessel. The projected horizontal cross-sectional area perpendicular to the vapour fl ow direction of the liquid distrib-utor feed pipe (T-shaped pipe) and separator downpipes with respect to the column cross-sectional area is less than 2%.

CFD analysis results of the load case corresponding to 20 000 t/d are used to validate the models against the available plant test data for this load case. Figures 4b and 4c demon-strate the CFD vapour fl ow trajectories and vector plots repre-sented by velocity magnitude,

mat above the deck serves to catch these droplets. Unfortunately, the available space above the separator would not be suffi cient to fi t a secondary KnitMesh for the glycol contactors, as this involves extend-ing the primary gas outlet, and the resultant space between the outlet and the vessel head would become too small. The fi nal proposal consisted of a coalescer mistmat plus a deck. It is worth noting that the overall glycol entrainment could be decreased further if using a secondary KnitMesh in this case.

A new distributor was proposed to eliminate a concern of entrain-ment caused by the distributor itself. The revamp also included the installation of a better support grid suitable for the expected higher loads.

CFD verifi cations CFD tools can provide more detailed insights using the actual tower confi guration and process loading conditions.1 The cramped

backup is calculated, where H is the backup height of fl uid in the drainpipe, ρ is the density of the fl uid, and g is the gravitational acceleration constant. Figure 3 demonstrates the principles of this consideration:

(1)

A typical arrangement of the Swirltube separator consists of a Sulzer KnitMesh mistmat placed in front of and at the back of the Swirltube deck. They both serve different purposes. The one below the deck is the coalescer mistmat and serves to coalesce the small droplets coming from below. To help collection and drainage of the collected droplets by the swirltubes, some purge gas guides the liquids into the liquid collection chambers. A small amount of liquid droplets might be carried with this purge gas. The secondary KnitMesh mist-

0.00.20.40.60.81.01.21.41.61.82.0

Velocity, m/s

0.60.81.01.21.41.61.82.0

Velocity, m/s

0.00.20.40.6

A

C

B

D

Figures 5a-d CFD vapour fl ow velocity vector plots: a) above the liquid distributor; b) below the mistmat; c) across the high-capacity separator ;and d) at the downcomer entry after the separator

gPH

fluidfluid ×

Δ=ρ

sulzer.indd 4 10/09/2013 11:30

Sure we can do the standard tower open/clean/inspect/close work but it’s those tough and challenging jobs that have helped us earn our stripes.

We recently completed a revamp on one of the largest vacuum towers in the western hemisphere, much to our customer’s satisfaction. We’ve mastered a resection method that is an excellent and cost saving alternative when footprints are tight.

Not long ago we tackled a 114-vessel project at a natural gas processing plant. Our team blinded, opened, cleaned, inspected and repaired all of the vessels in just 14,642 man-hours; less time than had been scheduled and well under budget.

When performing any number of services, we don’t overlook anything from external pipe fl anges, complicated vessel internals, feed/draw arrangements, section replacements, nozzle and strip lining installation/repair. We strive for “Zero Injury.” As for quality, well, that’s why our customers invite us back again and again and give us their annual maintenance/service contracts.

We thrive on accepting challenges then exceeding expectations but we are just as agile with a single tower project as we are with plant-wide turnarounds. We are quick on our feet and can mobilize swiftly for emergencies.

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Change of packingThe capacity possible through a contactor is determined by the maximum gas load factor. If this gas load factor is exceeded, signifi -cant entrainment will occur before the column becomes inoperable. Sulzer MellapakPlus 452Y was chosen as a replacement for the existing packing. This packing allowed a much higher capacity than the originally installed pack-ing, up to a throughput of 25 000 t/d. The higher capacity of this packing is attributed to the use of smooth bends within the packing rather than 90-degree directional changes at the packing layer inter-face (see Figure 7). By avoiding directional changes of 90 degrees by vapour, the liquid hold-up and subsequent vapour pressure drop over the packing are reduced, thereby allowing higher vapour throughput for the same pressure drop. This effectively increases the maximum gas rate. By replacing the packing and operating well within its maximum capacity limit, the liquid entrained to the top of the column should be minimised, reducing the liquid load on the Shell Swirltube separator. The combination of the new packing and Swirltubes was seen as essen-tial to minimise glycol losses.

Since the packing bed height was constrained by the existing column dimensions, checks were performed to confi rm that the packing height would be suffi cient for the given dehydration duty. These checks confi rmed that the packing height required was just within the height available. Further improvements in dehydration performance are possi-ble by focusing on the regeneration system, primarily TEG purity and, to a lesser extent, circulation rate. This is not the subject of this article. It may be noted that the dehydra-tion performance of Sulzer MellapakPlus 452Y packing is intrinsically better than the one of the existing Gempak 3A for the same TEG purity and vapour load.

Bottom arrangementAt the bottom section, there is an inlet sparger, immediately above which a chimney tray is installed.

• Pressure drop values across the mistmat and the high-capacity separator are within the design specifi cation• The cramped top conditions of the TEG contactor do not detrimen-tally affect high-capacity separator performance• The vapour vertical velocity vari-ation across the high-capacity separator was found to be very low. Based on Sulzer’s experience with comparable high-capacity separators, the measured values show a very good vapour distribu-tion quality. Due to the even vapour distribution across the device, the required liquid separa-tion can be achieved.

where blue represents the low velocity and red represents the velocity peak of the scale. It can be clearly observed from the vector plot that there are small eddies, turbulent fl ows, observed in the exit nozzle duct and the top of the downcomer due to the sudden change in the vapour fl ow direction by 90 degrees and the cramped top conditions. They do not have a seri-ous impact on the separation effi ciency of the column, as they are not propagated towards the high-ca-pacity separator. The vapour distribution across the tower cross section at various elevations has also been investigated.

Figures 5a-d show the CFD vapour fl ow velocity vector plots above the liquid distributor, below the mistmat, across the high-capacity separator and at the down-comer entry after the separator. As vapour distribution across the high-capacity separator has a more signifi cant impact on the gas-liquid separation effi ciency of the tower, statistical analysis was used to determine the vapour distribution quality. This quality is calculated by comparing the vertical velocity values at points on a regular grid cut plane across the high-capacity separator.

Figure 6 shows the CFD vapour fl ow vertical velocity plot across the high-capacity separator used for analysis.

Based on this investigation, the following conclusions may be drawn on the model confi guration:

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Y-component of velocity, m/s

Figure 6 CFD vapour fl ow vertical velocity plot across the high-capacity separator

Sulzer Mellapak Sulzer MellapakPlus

Figure 7 Progressive change of the corrugation angle at the both ends of the element

sulzer.indd 5 10/09/2013 11:29

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ConclusionsA TEG contactor’s capacity was constrained by a conventional mesh pad in the top of the column. A high-capacity Swirltube separator was designed to adapt to the exist-ing supporting system rather than risk welding to the vessel wall. In addition, the packing was replaced with Sulzer MellapakPlus M452Y and a new lean glycol distributor was installed. A subsequent test run demonstrated that these inter-nal modifications allowed an increase in gas throughput from 18 000 t/d to greater than 20 000 t/d, while reducing TEG carryover from 1.3 m3/day to around 0.4 m3/day. An increase (worsening) in gas dewpoint was observed due to not-yet-optimised TEG regenera-tion, but the overall performance of the unit remained within trunkline saturation limits.

GEMPAK is a registered trademark of Koch-Glitsch.Swirltube is a registered trademark of Shell.

References 1 Debangsu R, Arora A, Phanikumar A, CFD study of VDU feed inlet and wash bed, PTQ, Revamps 2009. 2 Kister H Z, Distillation Operation, McGraw-Hill, Inc, p79.3 Suess P, Analysis of gas entries of packed columns for two phase flow, International Conference and Exhibition on Distillation and Absorption, Birmingham, UK, 7-9 Sept 1992. 4 Kister H Z, Distillation Design, p291.

Anne Phanikumar is a Subject Matter Specialist, Computational Fluid Dynamics and Finite Element Technology, with the Mass Transfer Technology business unit of Sulzer Chemtech. He represented Asia Pacific in various Sulzer global task force teams for operational improvements. He holds a Master of Science degree from Nanyang Technological University, Singapore. Email: [email protected] Quan is a Sulzer Technical Expert, Trays and Separators. He represented Asia Pacific and China in various Sulzer’s expert teams, and now is responsible for global product and application management. He holds a PhD in environmental engineering from National University of Singapore. Email: [email protected]

where liquid is aerated by rising vapour. On the other hand, for a normal chimney with partial draw-off, the downcomer can be sized with a velocity larger than 0.5 m/s. Whether the existing downpipes have the capacity to handle rich TEG depends on the severity of aeration. The liquid height built up on this chimney is estimated at 15 mm, lower than the 70 mm riser height, which means no interaction between rising vapour and descending liquid is expected on this tray. The liquid-handling capacity of the downpipes should be sufficient.

Results of trial runs The maximum gas throughput achieved was 20 500 t/d through the modified glycol contactor (T300). The key observations from the test run were: • Glycol carryover with Train 3 was acceptable, approximately

0.4 m3/day on average, which is well below the 1.3 m3/day per train limit• Gas dewpoint was -2.6°C, limited by TEG regeneration capacity, not linked to the packing• Overall trunkline saturation, gas and condensate, was at 28%, well below the integrity limit of 75% and the operating limit of 65%.

This was achieved with the following key glycol regeneration system parameters:• Lean glycol purity = 99% w/w• Lean glycol circulation rate = 4.4 l/s• Glycol reboiler bottom tempera-ture = 190°C.

In addition, the success of the test runs proves that the chimney tray and liquid-handling capacity of the downpipes are sufficient, which provides a good reference for industry.

There is no possibility of revamp-ing the bottom arrangement due to narrow space and the hot work required. Evaluation of these inter-nals was necessary, as they could be the bottleneck at an increased feed rate of 20 000 t/d.

The general description of this type of chimney tray or vapour distributor can be found in the literature.2 However, it is rarely seen in industry nowadays due to the high pressure drop caused.

The chimney tray collects rich TEG and then discharges it to the vessel sump via downpipes (see Figure 2). Vapour released from the gas inlet sparger rises up to the packing bed, around 500 mm above the inlet, via risers that are small in diameter and low in height. The chimney tray basi-cally serves as a vapour distributor. The total open area of the risers is very low, only 8% of the column area.

The momentum calculated for the gas inlet nozzle is 21 600 Pa0.5 at 20 000 t/d, which is very high. The back pressure of the inlet could be higher than the pressure drop in the entire packed section, and the suction effect immediately above the inlet could lead to severe non-uniform pressure and velocity profiles, and eventually reverse the flow in the packing beds.3 The chimney tray can bring in additional pressure drop to the bottom section of the vessel, smooth the pressure profile inside the column, and hence eliminate problems linked to the high inlet momentum.

It was found that if all rich TEG drains down via downpipes, assuming a 4.6 l/s rich glycol lean TEG flow rate, the liquid velocity is about 0.18 m/s. For a normal chordal downcomer of a tray, the downcomer area should not have a liquid velocity higher than 0.13 m/s, based on vapour and liquid physical properties in a TEG contactor, to avoid downcomer flooding. In our case, this maxi-mum velocity must be derated, as degassing in the tiny downpipes becomes much more difficult.4 It is worth noting that the above guide-line is applicable to a normal tray,

The maximum gas throughput achieved was 20 500 t/d through the modified glycol contactor

sulzer.indd 6 10/09/2013 11:29

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Reactor effluent air cooler safety through design

In hydroprocessing units, the reac-tor effluent cooler (REAC) is one of the most vital pieces of equip-

ment, and any hindrance to its smooth operation immediately impacts the whole high-pressure loop. Older REAC designs used carbon steel, but these required low concentrations of ammonium bisul-phide and/or polysulphide sulphide injection together with frequent and thorough inspections. As feeds in most heavy oil hydroprocessing service have become more laden with sulphur and nitrogen, the concentrations of ammonium bisul-phide with economic levels of water injection have risen to a point where carbon steel tubes have routinely been substituted by alloy tubing.

Duplex 2205 and Alloy 825 are used, with the former being very popular because it is relatively less expensive. Initially, there were several problems associated with Duplex 2205, which were a result of poor fabrication techniques. These included a rapid cooling rate asso-ciated with thick header boxes, which could result in high ferrite and thus poor corrosion resistance; welding of thick tubes to tube sheets with joint leaks; and lack of control of welding, resulting in high hardness and thus susceptibil-ity to sulphide stress cracking. Many of these initial problems seem to have been overcome. However, problems persist around the REAC, primarily because of a lack of attention to detail and not adhering to licensor specifications. In this article, we will illustrate real REAC problems from recent projects and consider the remedies that were recommended.

Quality-controlled replacement of carbon steel with Duplex 2205 for revamps can increase the service life and reliability of the REAC in the high-pressure loop

ERIC LIN and PETER RISSEChevron Lummus Global

BackgroundIn the hydroprocessing industry, the REAC is one of the most important components in the high-pressure recycle gas loop. It typically provides the final cooling solution before separating the vapour (recycle gas) from the oil effluent and the sour water. The outlet temperature directly impacts recycle gas molecular weight as larger hydrocarbon molecules drop out of the vapour phase. The same mechanism also affects the hydro-gen partial pressure, which directly impacts reactor catalyst life. The importance of this piece of high-pressure equipment cannot be overstated. However, operating under high pressures and low temperatures can bring a host of issues into the equation, not the least of which includes ammonium bisulphide (NH4HS) and ammo-nium chloride (NH4Cl) precipitation (leading to pressure drop build-up), corrosion and/or erosion-corrosion. Some refiners have also experi-enced cracking in REAC welds, leading to fires, and loss of both time and money.

Tube metallurgy has a huge impact on a REAC’s life expec-tancy. Materials currently used in REAC systems include carbon steel, Type 400 series stainless steels, Type 300 series stainless steels, duplex stainless steel Alloys 3RE60 and 2205, Alloy 800, Alloy 825, Alloy 625 and Alloy C-276. Early refiners used carbon steel and this was found to be effective with NH4HS concentrations up to 3 wt%. As sour opportunity crudes became available, the resulting NH4HS crept up into the double digits. In

cases where the resulting NH4HS was between 3-8 wt%, polysulphide injection was a very cost-effective solution for the continued use of carbon steel (especially for revamps). However, over time, the persistent plugging/operation problems, frequency of inspection and the foul smell of the polysul-phide liquid greatly diminished its use. Increasingly sour feeds demanded more corrosion-resistant materials such as nickel-based Alloys 625, 800 and 825. Capable of handling up to 15 wt% NH4HS, Alloy 825’s corrosion resistance is normally matched by an equally hefty price tag. As the cost of mate-rials continued to rise, a more economical alternative with compa-rable protection against corrosion was sought.

Duplex stainless steels are often successfully used in these systems because they offer advantages from both the ferritic and austenitic stainless steel families. They are often cost effective due to their higher strength and reduced alloy element content compared to other higher alloys (up to one-third the cost of Alloy 825). However, since these materials consist of dual-phase microstructure, heat-treating, fabrication and welding techniques need to be carefully reviewed and monitored to assure that the balanced microstructure is not compromised. In the past, Duplex 3RE60 was used, but it had inferior corrosion resistance and toughness at the welds (it is no longer availa-ble). The most commonly used grade today is Duplex 2205.

Early implementations of duplex REACs failed to show significant


rev clg.indd 1 10/09/2013 11:34


welding guidelines for Duplex 2205 above and beyond those specified in API TR 938C• Using single-point water injection instead of multi-point water injection.

Symmetrical piping for REAC inlet tubes is critical for reactor effluent. It is impossible to imple-ment any type of flow control for two-phase flow, so the use of symmetrical piping will ideally provide even distribution of gas and liquid phases as it enters the air cooler. The danger of uneven phase distribution is that tubes favouring vapour will have more NH4HS deposits, and tubes favour-ing liquid will have more of the injection water. These deposits will eventually plug the tubes and increase the rate of corrosion. Consequently, unplugged tubes will see higher than acceptable velocities, increasing the rate of erosion/corrosion. It is possible for piping to be symmetrical and still be unbalanced. Balanced headers require additional splits, but will prevent flow that follows the “path of least resistance”.

During basic engineering, CLG specifies tube velocity limitations based on the tube metallurgy selected along with NH4HS concentration:• For carbon steel tubes and piping with less than 3 wt% NH4HS, 10-20 ft/s is allowed, with 15 ft/s preferred• For 2205 duplex stainless steel tubes and piping with 3-12 wt% NH4HS, 10-30 ft/s is allowed, with 25 ft/s preferred• For Alloy 825 tubes and piping with up to 15 wt% NH4HS, 10-40 ft/s is allowed, with 35 ft/s preferred.

In all cases, tube velocities falling below 10 ft/s become quite hazard-ous, as phase separation and corrosion can result.

For the industrial use of Duplex 2205, API TR 938C gives a great starting point for fabrication and welding guidelines. CLG has crafted a standard specification that addresses certain areas in more depth, and creates greater account-ability for both the fabricator and the welder. For example, CLG

reliability improvements, and a few units failed by hydrogen embrittle-ment cracking or sulphide stress cracking (SSC). Advancements in steel manufacturing have mini-mised microstructural deterioration during fabrication. Weld proce-dures and practices have been developed to assure balanced ferrite and austenite content, thereby improving reliability. API TR 938C provides guidance on materials and fabrication practices to achieve good corrosion resist-ance in duplex stainless steels.

Case studiesMany of the reported incidents involving Duplex 2205 REAC fail-ure have a common thread: pressures greater than 1000 psig and a NH4HS concentration of 6 wt% or greater. The following two recent case studies highlight these conditions.

A licensee in Asia (“Company A”) had a carbon steel REAC running for a number of years. An expansion of the unit required a REAC metallurgy upgrade to Duplex 2205. The design pressure of the REAC was ~2400 psig and the NH4HS concentration was expected to be 5 wt% during the modelling phase. A new Duplex 2205 REAC was installed and ran without issue for about two years, then fire erupted due to REAC fail-ure, and nearby equipment and piping were damaged. Cracks were observed on the weld joints between the top plate and tube sheet, as well as the bottom plate and tube sheet of the floating header. Fin tubes were deformed and the walkway was nearly unus-able. An investigation was commissioned to find the root cause of the REAC failure. The investigation discovered that during the REAC fabrication, Charpy impact testing was conducted at 0ºC instead of -40ºC (as specified by CLG). This over-sight led to inadequate toughness and low ductility for the welds as well as less than favourable micro-structural phase balance. Hardness values in the heat-affected zones (HAZ) and at the weld of failed specimens were higher than those

recommended (in the range of 313-359 HV10 vs 310 HV10 maxi-mum). As a result, CLG concluded that REAC failure in this case was due to sulphide stress cracking (SSC).

Another licensee based in the US (“Company B”) had a two-stage hydrocracker and hydrotreater ready for startup. The unit required a Duplex 2205 REAC. The design pressure of the REAC was ~2400 psig and the NH4HS concentration was expected to be 8 wt% based on modelling. During initial startup, a decision was made to skip the high-pressure tightness testing. A fire erupted from cracks formed in the REAC outlet temperature indi-cator thermowells in two separate trains. These cracks were found to be a result of improper manufactur-ing practices (of the failed duplex

material) and would have been detected during a high-pressure tightness test. By not following standard operation practices, the unit startup was delayed.

REAC safety through designIssues that plague the industry’s REACs can be mitigated or elimi-nated altogether by considering “Safety Through Design”. In preparing a basic engineering pack-age for licensees, CLG highlights certain areas and design considera-tions that are considered strong recommendations. These considera-tions include:• Using balanced, symmetrical piping• Limiting tube velocities based on metallurgy, NH4HS concentration and H2S partial pressure• Specifying fabrication and

Weld procedures and practices have been developed to assure balanced ferrite and austenite content, thereby improving reliability

rev clg.indd 2 10/09/2013 11:34

specifies a maximum hardness rating of 310 HV10 versus API’s requirement of 320 HV10. There is also tighter control of weld consumable and filler metals. Qualifications for Weld Procedure Specifications (WPS) and Procedure Qualification Record (PQR) are also much more stringent. As the two case studies illustrate, Duplex 2205 REAC failure can be attributed to not respecting the guidelines specific to duplex material.

One final design consideration is whether to use single- or multi-point water injection before the REAC. With the numerous valves involved in a multi-point water injection system, water manage-ment is difficult. Even with symmetrical piping, a multi-point water injection system may have some tubes favouring more water than others. This scenario can be dangerous, as some tube tempera-tures fall below the NH4HS precipitation temperature. Single-point water injection ensures that


the global water requirement is met from the start, well before the efflu-ent splits into their separate headers (the injection point is typi-cally 10 pipe diameters upstream of the split). Another benefit of single-point injection is that the recycle gas compressor spill back line can be positioned downstream of the injection point so that it is washed with injection water. This setup minimises NH4Cl and NH4HS formation in the dead-leg piping.

SummaryThe combination of high pressures and corrosive environments can spell disaster for REACs in the refining industry. Fortunately, measures such as using balanced symmetrical inlet piping, limiting tube velocities based on metallurgy and NH4HS concentration, working with suppliers and welders who have good quality control with Duplex 2205, and the use of single-point water injection can increase the service life and reliability of the

most important air cooler in the high-pressure loop.

Eric W Lin is a Principal Process Engineer with Chevron Lummus Global in Bloomfield, New Jersey. His responsibilities include leading basic engineering designs and proposals for both grassroots and revamp projects. He has design experience with crude/vacuum distillation, as well as high-pressure hydrocracking, hydrotreating and dewaxing units. With more than 15 years of experience in the refining industry, he holds a BS degree in chemical engineering from Columbia University.

Peter J Risse is a Senior Staff Materials Engineer with Chevron Lummus Global in Richmond, California. He and his group are responsible for materials selection for new capital projects, review of reactor fabrication, and aid with unit troubleshooting, fixed equipment asset reliability and remaining life predictions in the hydroprocessing units. With more than 30 years of experience in the refining industry, he holds a BS degree in chemical engineering from the University of Rhode Island and a MS degree in materials science from the University of California at Berkeley.


and DHP units. He holds a degree in chemical engineering from Middle East Technical University, Turkey, and is a certified Energy Supervisor for industrial plants. Email: [email protected] Asim Ay is CCR/NHT/ISOM Units Process Superintendent with Tüpras Kirikkale refinery. He holds a degree in chemical engineering from Middle East Technical University, Turkey. Email: [email protected] Kahraman is CCR/NHT/ISOM Units Process Chief Engineer with Tüpras Kirikkale refinery. His six years of refinery experience includes the process side of hydrocracker and hydrogen production plants, sulphur recovery, NHT, ISOM, CCR and DHP units. He holds a degree in chemical engineering from Middle East Technical University, Turkey. Email: [email protected] Selmen is Axens’ Technology Manager for Naphtha Hydrotreatment and Reforming Technologies. He has worked mainly with bottom-of the-barrel technologies, specialising in heavy crude oil upgrading. He has also been involved in the process design of aromatics complexes and NHT, as well as reforming units startup and troubleshooting. He holds an engineering degree from the ENSGTI engineering school and a DEA in refinery process modelling from IFP School. Email: [email protected]

year without a major interruption, indicating that the vibration prob-lem was correctly identified. This case shows that a good licensor and refinery relationship is essential for solving problems that require both technological and operational experience.Osman Kubilay Karan is the Hydroprocessing Units Process Superintendent with Tüpras Kirikkale refinery. His 25 years of refinery experience includes the operational and process sides of crude, vacuum units, hydrocracker, hydrogen production plants, CCR

measurements are taken at the outlet stream of this adsorber and the results are 0 ppm HCl, whereas inlet concentration averages 30 ppm HCl.

ConclusionCurrently, the H2-rich gas compres-sors are running smoothly without any problem. The compressors were recently opened by the mechanical maintenance group and no green oil formation was found, although they ran for almost one

Reduction chamber

Reaction section

Dedicated chlorine trap

Separator drum

H2 to reduction

H2 from reduction

To booster compressor

Recycle gas

Catalyst streamProcess stream

Figure 3 PFD after adsorber drum


refi ner and particularly the ability to anticipate market needs in differ-ent regions as constraints evolve.

PolyFuel and PolyNaphtha are trademarks of Axens.

Marielle Gagnière is Technology Manager for hydroprocessing and olefi ns technologies downstream FCC, especially oligomerisation and etherifi cation technologies, in Axens’ Marketing, Technology and Technical Assistance Department. She is an engineering graduate from the Ecole Nationale Supérieure de Chimie de Paris, and holds a post-graduate engineering degree from the IFP School.Annick Pucci is Deputy Product Line Manager in the fi eld of light ends hydrotreatment and a specialist in refi ning olefi ns processing, particularly for FCC effl uent upgrading. She holds a bachelor’s degree in chemical engineering from Ecole Nationale Supérieure des Industries Chimiques de Nancy, France.Emilie Rousseau is a Strategic Marketing Engineer in Axens’ Marketing Department. She holds a chemical engineering degree from the Ecole Nationale Supérieure des Ingénieurs en Arts Chimiques et Technologiques de Toulouse, a master’s in chemical engineering from Imperial College in London and a master’s in energy economics and corporate management from IFP School.

It is in Europe where the difference between refi nery yield structure and market demand is critical, especially since conven-tional refi ning tools do not have the fl exibility to reduce excess gasoline production and to increase the amount of middle distillates. Moreover, with European refi neries facing increasing diffi culty in fi nd-ing export markets for their excess gasoline and given the tensions in middle distillate supply, PolyFuel should fulfi l a primordial role in adjusting the gasoline-distillate production to better fi t market demand.

In other regions, new tendencies such as shale oil and shale gas are revolutionising the US market, providing additional light products and consequently infl uencing market balance and prices. Today in the US, as a result of the impact of shale gas on the cost of LPG, PolyFuel is already profi table for a mixed feed of LPG and C5/C6 cut.

The fl exibility of the new process offers many advantages to the

The new technology is profi table taking into account today’s US market prices even if the middle distillate price is lower than the gasoline price. Indeed, as a result of shale gas production, LPG prices are low. Adding LPG (C3 and/or C4) cut in a PolyFuel unit lowers feedstock costs and contributes to increased profi tability, while maximising middle distillates production in the refi nery.

To reach 15% IRR for PolyFuel with prices based in 2012 in the US Gulf Coast, the middle distillate price can be $96/t lower than gaso-line. If the middle distillate price were equal to the gasoline price ($1129/t) and the LPG price kept at $636/t, the IRR would reach 28%.

ConclusionWith the world market for middle distillates growing and a reduced demand for gasoline in certain regions, the new process for olefi nic gasoline oligomerisation allows the refi nery scheme to be adapted to a maximum distillate mode.

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Wet scrubbing modifi cations to reduce emissions

Many FCC units have wet scrubbing systems installed on them to reduce air emis-

sions. A large percentage of these were designed to address only particulate and SOx emissions. Some of them may have been designed to achieve older emission mandates that are higher than the ones presently being mandated by regulatory authorities. A modifi ca-tion of these systems may be required to reduce particulate and/or SOx emissions and possibly to also include a reduction in NOx emissions. This article addresses how these reductions can be achieved, considering the optimisa-tion of economics and system outage time to fi t within turna-round schedules.

First, a reduction in SOx emis-sions is discussed. Methods to lower SOx levels from a current wet scrubbing system are reviewed along with the amount of emissions reductions that can be achieved. Next, particulate emissions reduc-tions are discussed. The modifi cations required to reduce particulate emissions in an existing wet scrubbing system are detailed. Finally, the addition of NOx control to an existing wet scrubbing system is discussed. Several approaches are presented, along with the merits and potential issues associ-ated with the different approaches.

A basic wet scrubbing system For many years, refi ners have chosen to use the EDV wet scrub-bing system to control both particulate and SO2 emissions. With this system, particulate and sulphur emissions are removed

Modifi cations to existing wet scrubbing systems for FCC units can readily be made to achieve additional reductions in particulates, SO

2 or NOx

EDWIN WEAVER and NICHOLAS CONFUORTOBelco Technologies Corporation

simultaneously and effi ciently. This technology is well proven in providing fl exibility to handle added capacity that may result from FCC unit expansions or to increase reduction effi ciency as regulatory pressures increase and in providing uninterrupted opera-tion/performance exceeding that of FCC units. Each refi ner’s specifi c reasons for choosing wet scrubbing differ, but these have generally been related to environmental compliance as well as relative costs, reliability and fl exibility of wet scrubbing compared to other emis-sion control options.

The EDV wet scrubbing system is the state-of-the-art approach for controlling particulate and SO2 emissions from oil refi nery FCC

units, boilers and heaters. One arrangement of this system is shown in Figure 1.

The system treats hot fl ue gas-containing particulates (such as FCC catalyst fi nes) and SO2, and discharges cleaned gas to the atmosphere through an integral stack. At the scrubber inlet, FCC fl ue gas is quenched and saturated by multiple water sprays in the spray tower’s horizontal quench section. Normally, the fl ue gas enters the wet scrubber after pass-ing through a heat recovery device, such as boiler tubes or a fl ue gas cooler. However, the system can also be designed to accept the full fl ue gas temperature directly from the high-temperature fl ue gas source without any heat reduction


Absorber vessel re-circulation pumps

Quench

Filtering module re-circulation pumps

Filtering modules

Absorber vessel

Droplet separators

Figure 1 EDV wet scrubbing system

belco rev.indd 1 10/09/2013 11:42


Makeup water is added to the system to replace the water lost to evaporation in the quench zone and also the water purged from the system. Captured pollutants, including suspended catalyst fi nes and dissolved sulphites/sulphates (NaHSO3, Na2SO3 and Na2SO4 in cases when sodium-based reagents are used) resulting from the reduc-tion in particulate and SOx, are purged from the spray tower recy-cle loop to maintain the proper balance. The treatment of this purge stream is addressed later.

In order to remove very fi ne particulate, fl ue gas leaving the spray tower is distributed to a bank of parallel fi ltering modules. Within each module, the fl ue gas fi rst accelerates (compresses) and then decelerates (expands). This action causes the water that is present in the fl ue gas as moisture to condense onto the fi ne particulate and the acid mist (mostly H2SO4 from condensation of SO3 in the saturated fl ue gas), increasing both their size and mass by converting them to relatively large droplets. These droplets are then removed from the fl ue gas by proprietary F nozzles located at the exit of the fi ltering modules. Additional condensation occurs on the walls of the fi ltering modules, keeping them continuously clean. Some agglom-eration also takes place in the fi ltering modules, further enhanc-ing particulate removal.

As mentioned above, a F nozzle is located at the exit of the fi ltering module. This nozzle sprays counter-current to the gas fl ow and provides the mechanism for the collection of the fi ne particulate and mist, which has been enlarged by condensation and agglomerated. This device has the advantage of being able to remove fi ne particulate and acid mist with an extremely low-pressure drop and no internal components that can wear or cause unscheduled shutdowns. It is also relatively insensitive to fl uctuations in gas fl ow. It is illustrated in Figure 3.

Prior to being discharged to the atmosphere through a stack, the fl ue gas enters the system’s droplet

in cases where a CO boiler on an FCC unit application requires to be bypassed. The fl ue gas from the FCC unit can be diverted directly to the EDV wet scrubbing system without any concerns and without having to make any adjustments to the operation. This not only results in a more reliable and simpler operation, but also allows the plant to continuously reduce emissions even during bypass and upset conditions.

The EDV wet scrubbing system utilises proprietary nozzles to produce high-density water curtains through which the gas must pass. Each nozzle sprays water droplets that move in a cross-fl ow pattern relative to the fl ue gas. These cover the entire gas stream and uniformly fl ush the vessel’s surfaces clean. The spray nozzles are non-clogging and are designed to handle highly concentrated slurries.

SO2 absorption and particulate removal begins at the quench section and continues as the fl ue gas rises up through the main spray tower, where the gas is again contacted with high-density water curtains produced by additional spray nozzles. The spray tower

itself is an open tower with multi-ple levels of Belco spray nozzles. Since it is an open tower, there is nothing to clog or plug in the event of a process upset. In fact, this design has handled numerous process upsets and reversals with-out any concern.

The scrubbing liquid is controlled to a neutral pH with reagent addi-tion to drive SO2 absorption. Caustic soda (NaOH) is typically used as the alkaline reagent. However, other alkalis such as soda ash and magnesium hydroxide have also been utilised with excel-lent results in terms of performance and reliability. In non-refi nery applications and in extremely few refi nery applications, lime has also been used. However, for FCC unit applications and other refi nery applications where three-to-seven-year continuous operation is required, the use of lime as a reagent is strongly discouraged.

In the EDV wet scrubbing system, multiple levels of spray nozzles provide suffi cient stages of gas/liquid contact to reduce both partic-ulate and SO2. An illustration of the spray tower and the spray nozzles is shown in Figure 2.

Belco spray nozzlesSpray tower

Flue gas inlet

Liquid for re-circulation

Liquid to spray nozzles

Figure 2 Spray tower and Belco spray nozzles

belco rev.indd 2 10/09/2013 11:42

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separators. These separate and collect entrained water droplets, allowing the fl ue gas to exit the stack free of such water droplets. For boiler and heater applications, chevron-type droplet separators may be used as an alternate because of the relatively low partic-ulate loading and upsets that are normally associated with FCC unit operations.

For FCC applications, the EDV system normally uses large tubes with fi xed spin vanes as droplet separators (called Cyclolabs). The gas entering each separator passes through a fi xed spin vane, where centrifugal acceleration causes free water droplets to impinge on the separator’s walls. Collected water droplets fl ush the walls uniformly clean and drain to the bottom. Collected water is recycled for fl ue gas cleaning in the fi ltering modules or spray tower. This device is also illustrated in Figure 3.

Optimisation of SO2 emissions

FCC sulphur emissions are in the form of SO2 and SO3 (also referred to as SOx). These vary signifi cantly, depending on the feed sulphur content and the FCC unit’s design. In the FCC reactor, 70% to 95% of the incoming feed sulphur is trans-ferred to the acid gas and product side in the form of H2S. The remaining 5% to 30% of the incom-ing feed sulphur is attached to the coke and is oxidised to SOx, which is emitted with the regenerator fl ue gas. The sulphur distribution is dependent on the sulphur species contained in the feed and, in particular, the amount of thio-phenic sulphur. SO2 can range from 150 to 3000 parts per million on a dry volume basis (ppmvd), whereas SO3 typically varies from 2% to 10% of the SO2 content. Of course, for


other combustion sources such as boilers and heaters, 100% of the sulphur content in the oil being used is converted to SOx in the fl ue gas.

Although most recent wet scrub-bing systems are designed to achieve an SO2 outlet of 25 ppm or less, there are older units designed for higher outlet emission levels. Also, as regulatory authorities trend towards mandating lower emission levels, there may be a need to increase the SO2 emissions reduction capability from an exist-ing wet scrubbing system. Generally speaking, there are a few approaches that can be considered:

Adjustment of pHMost caustic-based wet scrubbing systems operate at a slightly acidic

pH, usually around 6.8. Raising the pH will lower SO2 emissions and is a simple adjustment. However, there is a limitation, in that pH levels above approximately 7.4 can result in the formation of carbonates and cause build-ups and/or plugging of the liquid piping lines, so caution must be exercised in taking this approach.

Adding additional liquid togas contactSince the scrubber design being discussed has a staged approach for liquid to gas contact, multiple levels of spray nozzles are used to achieve the desired SO2 perfor-mance. It is possible to add additional level(s) of spray nozzles to increase the liquid to gas contact, which will lower the SO2 outlet emissions. Although the addition of spray nozzles is relatively simple, adding them will require an increase in the amount of liquid being circulated in the wet scrub-ber, so the pump capacity and confi guration will have to be evalu-ated, along with the changes required to the piping confi guration.

Lowering the scrubber operating temperatureThe reduction in SO2 in a wet scrubbing system is dependent on several factors. One of these factors is the scrubber operating tempera-ture. Generally, the operating temperature of the wet scrubber will be the saturation temperature of the fl ue gas. The saturation temperature primarily depends upon the inlet temperature to the wet scrubber. Lowering the fl ue gas inlet temperature to the wet scrub-ber will lower the saturation temperature, causing some improvement in SO2 removal. There

Filtering module

Cyclolab droplet

separator

Figure 3 Filtering module for fi ne particulate control and droplet separators

Increase pH Increase L/G ratio Lower scrubber inlet temperature Sub-cool systemRelative cost Minimal-only additional caustic Moderate Varies HighSO

2 reduction Varies Signifi cant Moderate Signifi cant

Advantages No capital investment Relatively simple Energy recovery Reduction in water consumptionConcerns pH too high Pump and piping modifi cations Upstream corrosion if Design of cooling device temperature too low

Approaches to SO2 reduction

Table 1

belco rev.indd 3 11/09/2013 12:25


models and other scrubbers). The addition of fi ltering modules to the scrubbing system will remove much of the fi ne particulate that is less than 3 microns in size. This fi ne particulate can be a signifi cant amount. This modifi cation involves installing the fi ltering modules downstream of the absorber section, and also installing the necessary pumps and piping asso-ciated with the fi ltering modules.

Conversion of fi ltering modulesfrom EDV 5000 to EDV 6000 The performance of the fi ltering modules (the device that removes the fi ne particulate) can be enhanced by adding a liquid spray at the entrance to the fi ltering module. This modifi es the fi ltering module from a 5000 design to a 6000 design. In addition to the spray nozzles, this change requires pump and piping modifi cations due to doubling the number of spray nozzles in the fi ltering modules. A 5000 fi ltering module design and a 6000 fi ltering module design are illustrated in Figure 4.

Increase fi ltering module spraypressureAnother modifi cation that can be made to reduce particulate emis-sions is to increase the fl ow rate and pressure of the fi ltering module sprays. This is a fairly simple modi-fi cation, but, of the options that have been described, will have the least impact in terms of particulate emissions reduction. Each applica-tion will need to be carefully evaluated by Belco.

A summary of the possible approaches to reducing particulate emissions is shown in Table 2.

Control of NOx emissions with awet scrubbing systemNOx emissions have also become an increasing focus for concern. These emissions contribute signifi -cantly to a variety of environmental problems, including ozone forma-tion (photochemical smog), acid rain and elevated fi ne particulate levels. There are many federal, state and local regulatory drivers addressing these issues for many different stationary sources. As

heaters that use heavy oils or resides, also contribute to particu-late emissions, but to a much lower degree. The particulates from these sources are products of combus-tion; they contain heavy metals and are typically very fi ne in particle size.

Older wet scrubbing systems may have been designed for particulate emission rates as high as 1.0 lb particulate per 1000 lb of coke burned. Most regulatory mandates are now for an emission rate of no more than 0.5 lb particulate per 1000 lb of coke burned and, in some cases, even less. Therefore, there may be a need to reduce particulate emission levels. This can be accom-plished in several possible ways:

Addition of fi ltering modules to awet scrubbing systemSome existing units do not have fi ltering modules (EDV 1000

may be a side benefi t to this approach in that the additional heat recovery upstream of the scrubber should have some positive benefi ts in the overall energy recovery of the FCC system.

Additional lowering of the scrub-ber’s operating temperature can be achieved by sub-cooling of the fl ue gas. This is achieved by adding a cooling device on the wet scrubber liquid recirculation loop, which lowers the temperature of the circu-lating liquid and thus the operating temperature of the wet scrubber, improving performance. An addi-tional benefi t of this approach is that sub-cooling recovers some of the liquid that is used to quench the fl ue gas, reducing system over-all water consumption. However, the cooling device required will operate in a harsh environment with catalyst fi nes that circulate in the wet scrubbing liquid, so careful consideration of the design of this device is necessary to ensure a reli-able system operation.

A summary of the possible approaches to reducing SO2 emis-sions is shown in Table 1.

Optimisation of particulate emissionsFor a FCC unit, particulate (cata-lyst) emissions vary, depending on the number of stages of internal and external cyclones. Although cyclones are effective in collecting the greater constituent of catalyst recirculated in the FCC regenerator, the attrition of catalyst causes a signifi cant amount of fi ner catalyst to escape the cyclone system with relative ease. Typically, emissions will range from 200 to 650 milligrams per normal cubic metre of gas (mg/Nm3). Other combus-tion sources, such as boilers and

Model 5000 Model 6000

Figure 4 Filtering module designs

Add fi ltering Convert Model Increase liquid modules 5000 to 6000 spray pressureRelative cost High Moderate LowParticulate reduction Signifi cant Moderate SomeAdvantages Most signifi cant particulate Relatively simple Extremely simple emissions reductionConcerns Most scrubbers already Pump and piping Amount of particulate have fi ltering modules modifi cations that will be reduced

Approaches to particulate reduction

Table 2

belco rev.indd 4 10/09/2013 11:42


such, the market for NOx control systems has been growing steadily over the past several years.

Since the FCC unit is one of the largest single sources of emissions, in terms of tons emitted per year, it is a primary area of focus for NOx reduction. Uncontrolled NOx emissions from the regenerator can vary greatly and depend on many variables such as the feed to the FCC unit, regenerator design (partial or full burn) and the design of the secondary combus-tion device (CO boiler), if applicable. Uncontrolled NOx emissions could range from 50 ppm to 400 ppm, although most facilities see uncontrolled levels in the range 75 ppm to 150 ppm.

LoTOx technology for NOx removalThe LoTOx process is a selective, low-temperature oxidation technol-ogy that uses ozone to oxidise NOx to water-soluble nitric pentoxide (N2O5), which, inside the wet scrub-ber, forms nitric acid that is subsequently scrubbed by the

N2O5 conversion to HNO3 and scrubbing

NO, NO2 conversion to N2O5

conversion to and scrubbing

conversion

Removed in scrubber purge

Ozone injection

Figure 5 Simplifi ed LoTOx schematic

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medium. Sulphite or sulphurous acid is an ozone scavenger that results in a negligible ozone concentration in the treated flue gas stream leaving the scrubber. An illustration of the LoTOx process in a wet scrubbing system is shown in Figure 5.

Chemistry of the LoTOx process The LoTOx process uses ozone to oxidise NO and NO2 to N2O5, which is highly soluble. In contact with the wet scrubbing liquid, N2O5 is easily and quickly absorbed, converted to HNO3 and then neutralised to NaNO3. Many reac-tions occur, but for the sake of brevity the process can be summa-rised in the following reactions:

NO + O3 → NO

2 + O

2

2NO2 + O

3 → N

2O

5 + O

2

N2O

5 + H

2O → 2 HNO

3

HNO3 + NaOH → NaNO

3 + H

2O

N2O5 is an extremely soluble gas and reacts with water instantane-ously. As a result, it is easily removed by the system even before the SO2. N2O5 is estimated to be at least 100 times more soluble than SO2.

Application of the LoTOx process to existing wet scrubbersThe primary challenge in applying the LoTOx process to an existing wet scrubbing system is finding the optimal configuration to get the residence time required for the ozone to react with NOx to form water-soluble N2O5. Several approaches have been successfully utilised on existing wet scrubbing systems. The approaches that have been utilised are described below:

Modify the existing vessel to addresidence timeIn this approach, empty space is added to the vessel between the quench area and the area where the spray nozzles are located. This provides the proper temperature zone for the LoTOx reactions to occur. It also provides for an area that is free from excess water drop-

nitrogen into the aqueous phase in the scrubber is rapid and irreversi-ble, allowing nearly complete removal of NOx.

The rapid reaction rate of ozone with NOx makes ozone highly selective for the treatment of NOx in the presence of other compounds

such as CO and SOx, resulting in high ozone utilisation efficiency for NOx removal with no reaction with CO and SOx at the design retention time. SO2 absorption in the aqueous medium within the scrubber forms sulphite or sulphurous acid, and any unreacted or excess ozone in the oxidised flue gas readily absorbs into the scrubber’s aqueous

scrubber nozzles and neutralised by the scrubber’s alkali reagent.

The process operates optimally below 300°F (150°C); therefore, it does not require heat input to maintain operational efficiency (no flue gas reheating) and can enable maximum heat recovery from hot flue gases. Ozone is produced in response to the amount of NOx present in the flue gas generated by the combustion process and the final NOx emissions required. The low operating temperature allows stable and consistent control regardless of variation in flow, load or NOx content. There are no adverse effects of acid gases or particulates on the performance of the LoTOx process. Ozone, once mixed with flue gas, rapidly reacts with insoluble NO and NO2 mole-cules to form higher oxides such as N2O5. These higher oxides of nitro-gen are highly soluble and rapidly react with moisture present in the flue gas stream to form dilute oxy acids such as nitric acid. The conversion of higher oxides of

Clarifier

Air blower

Oxidation towers

Effluent cooler

Effluent filters

Settling bins Sump

Scrubber purge

Effluent discharge

Figure 6 Belco purge treatment unit

Modify existing vessel Add vessel upstream Add vessel downstreamRelative cost Least Significant SignificantNOx reduction Excellent Excellent ExcellentAdvantages Least cost, no Existing liquid sprays No potential impact on existing additional plot space used to absorb NOx/N

2O

5 wet scrubber operation

Concerns Foundations and vessel Space needed between Need the proper amount design may not be FCC unit and existing of sulphites in new adequate wet scrubber vessel

Approaches to NOx reduction

Table 3

The LoTOx process uses ozone to oxidise NO and NO

2 to

N2O

5 which is highly

soluble

belco rev.indd 6 11/09/2013 12:25


SummaryModifications of existing wet scrubbing systems can readily be made to achieve additional reductions in particulates, SO2 or NOx. This is best accomplished through a detailed evaluation of the existing system once the new requirements for particulate, SO2 or NOx have been clearly defined. This approach will lead to the most logical and cost-effective modification to achieve the desired emissions levels.

LoTOx is a trademark of the Linde Group.

Edwin Weaver is Technology Director at Belco Technologies Corporation. He has over 30 years’ experience in the field of air pollution control, and has been the author of numerous technical papers. The topics include the control of air emissions in the oil refining process, the design and operation of various types of air pollution control systems for a variety of processes, and developments in air pollution control technologies. He also holds a patent pertaining to air pollution control systems. He holds a BS in civil engineering from Lehigh University, is a member of the NPRA environmental committee and a registered professional engineer in the State of New Jersey.

Nicholas Confuorto is Vice President of Technology, Sales and Marketing, at Belco Technologies Corporation in Parsippany, New Jersey, USA, responsible for the company’s global licensing, business development, sales and marketing activities. He is an engineering graduate of Columbia University, and has worked for more than 30 years in the design and implementation of emission control systems. He has written numerous technical articles in this field.

lets, which inhibit the NOx reactions. It utilises the existing liquid sprays to absorb the N2O5 and convert it to nitric acid. Although seemingly the simplest approach, it is usually limited by the fact that 20-50ft (6-15m) of scrubber height is added, which may require extensive foundation modifications and/or reinforcement of the scrubber vessel walls.

Add the NOx reaction zone upstreamof the existing wet scrubberIn this approach, a vessel is added upstream of the existing wet scrubber. The flue gas is quenched as it enters the vessel, ozone is injected and the reactions occur to convert NOx to water-soluble N2O5 before the flue gas reaches the liquid sprays in the existing scrub-ber vessel, where the NOx is removed.

Add NOx reaction zone downstreamof the existing wet scrubberIn this configuration, the NOx removal is accomplished in a new vessel placed after the existing wet scrubber. A primary advantage to this approach is that the flue gas properties upstream of the existing scrubber are unchanged. With some scrubber designs, it is impor-tant to keep the existing scrubber inlet conditions to avoid any impact on the existing performance for SO2 and particulate removal. This approach achieves that specific requirement. The scrubber liquid from the existing scrubber is used to provide sulphites in the NOx reaction vessel to remove excess ozone from the flue gas before it exits the system. A secondary advan-tage of this configuration is that some additional SO2 removal is achieved in the new NOx reaction vessel.

A summary of the possible approaches to reducing NOx emissions is shown in Table 3.

Treatment of scrubber purgeCaptured pollutants, including suspended catalyst fines and dissolved sulphites/sulphates resulting from SOx and NOx removal, are purged from the spray tower recycle loop. The liquid that is purged from the scrub-ber is typically processed in a purge treatment unit (PTU). Here, a clarifier removes suspended solids and generates a concentrated slurry that is dewatered in settling bins. Water, now free of suspended solids, over-flows from the clarifier and is oxidised in a series of vessels using air and agitation. Oxidation converts sulphites in the purge to sulphates, reducing any chemi-cal oxygen demand (COD) prior to discharge. In some systems, a secondary filter is used to further reduce total suspended solids, and a cooler may be utilised to control the discharge temperature of the purge stream. A typical arrangement of a PTU is shown in Figure 6.

If modifications are made to reduce the particulate emissions in the wet scrubbing system, the clarifier and effluent filters, if applicable, will also need to be examined to determine if modifications to this part of the purge treatment system will be required. If SO2 emissions are reduced, the oxidation portion of the PTU will need to be examined to determine its adequacy for controlling COD to the desired level.

“adequate organisation of documentation, both in terms of management and data traceability”.

China’s Jiutai Energy (Zhungeer) has licensed UOP’s methanol-to-olefins technology to convert methanol from coal into key petrochemicals. The Hydro MTO process converts methanol from gasified coal or natu-ral gas to produce ethylene and propylene. The technology enables producers in China, the world’s largest miner of coal, to tap local resources, rather than more expensive petroleum, to produce petrochemicals.

Jiutai will produce 600 000 t/y of ethylene and propylene at its facility in Ordos City, Inner Mongolia Province, China. In addition to technology licensing, UOP will provide basic engineering, catalysts, adsor-bents, speciality equipment, technical services and training for the project, which is expected to start up in 2014.

The methanol-to-olefins (MTO) process was jointly developed by UOP and INEOS and converts methanol from crude oil and non-crude oil sources such as coal or natural gas to ethylene and propylene. The process is based on UOP’s catalysts and is said to provide high yields with minimal byproducts. MTO can also vary the relative quantities of propylene and ethylene it produces, so producers can adjust plant designs to market demand.

UOP announced a similar project with China’s Wison (Nanjing) Clean Energy Company, which licensed the first commercial-scale installation of the MTO process that combines the Hydro MTO process and the Total Petrochemicals/UOP Olefin Cracking Process. That project is expected to produce 295 000 t/y of ethylene and propylene.

JPNOR Engenharia, an engineering, procurement and construction company based in Brazil, used Aveva Plant software for the 3D engineering and design of Complexo Petroquímico do Estado do Rio de Janeiro (Comperj), a major refinery and petrochemical construction project located in the city of Itaboraí, outside Rio de Janeiro.

The applications used were PDMS, Review and Global. JPNOR says that it selected the Aveva software because of the flexibility and capacity of its object-based data structures. PDMS and Review were adopted as tools used at the individual project sites, while Global helped engineers to perform integration tasks across multiple sites in execution, visualisation and inspection.

The integration of PDMS models from over 20 designers and EPCs enabled JPNOR to provide daily progress updates to customers including Petrobras.

Construction of the Comperj refinery and petrochem-ical complex is estimated to cost $8.4 billion and marks

developers, the process delivers a fuel with improved performance over more conventional biodiesel and petroleum-based diesel, including a cetane value of 80 compared with a cetane range of 40-60 in retail diesel.

Eni says that the decision to convert the Venice site to the production of renewable biofuels is a response to the European Union’s Renewable Energy Directive, which mandates that, by 2020, 20% of its 27 member states’ energy must come from renewable sources, and greenhouse gases must be reduced by 20%.

The Venice refinery has a recent history of establish-ing its green credentials. As part of an energy conservation programme called Stella Polare, it was the first in Europe to obtain ISO 16001 Certification for its energy management system.

The refinery was selected for a pilot programme to define methods to be extended to all of Eni’s industrial plants.

The primary objective of the project is to identify actions that will lead to a reduction in direct and indi-rect consumption as well as energy-saving measures through the adoption of innovative techniques. The certifying body for ISO 16001, DNV, highlighted elements of the Venice site’s activities such as manage-ment initiatives and activities aimed at reducing energy consumption with a resulting fall in CO2 emis-sions. DNV also noted that the Venice refinery distinguished itself for its “excellent strategic evalua-tion in identifying roles and responsibilities for the management of improvement objectives” and for the

Designing for a return to petrochemicals

Petrochems from coal

6 PTQ Q2 2013 www.eptq.com

president truly meant what he said and his words were more than political rhetoric, we could see an end to the anti-fossil fuel environment that has persisted throughout his first term in office. I am hopeful that in 2013 this administration will work to advance a true “all-of-the-above” energy strategy that recognises the importance of all of our domestic energy resources and fuel and petrochemical manufacturers in rebuilding our nation’s economy.

The US, in combination with our ally and friend Canada, has an abundance of natural resources capable of bringing the nation to economic prosperity and North American energy independence within the next 10 years. This can be done by expanding shale development on federal lands, stopping the attacks on hydraulic fracturing that threaten affordable feedstocks necessary for all manufacturing, and increasing offshore drilling permits.

Further, by approving the Keystone XL Pipeline, the president could send a strong signal to the nation that he is serious about creating thousands of domestic jobs and improving our economy. More importantly, the president would put the world on notice that the US is charting a future to include energy independence and security for the country, with a goal of ultimately becoming a global energy provider.

Sadly, without a significant practical and political course correction, growth is not possible for fuel and petrochemical manufacturers in the current regulatory environment. In January 2011, President Obama ordered a review of federal regulations to be eliminated because they hinder economic growth and job creation. Two years later we are still waiting. A commitment by the administration is needed to fix what continues to be a regulatory nightmare for the refining and petrochemical industry and an unnecessary, costly burden for the American public. For example, federal regulations require fuel and petrochemical manufacturers to spend billions of dollars to reduce greenhouse gas (GHG) emissions, even though the EPA has acknowledged the reductions would bring little or no environmental benefit. These same regulations increase energy costs, result in job loss, and harm the US economic and national security.

AFPM will continue to support sensible and beneficial environmental regulation, but we believe that America’s national interest would best be served by comprehensive and objective cost-benefit analyses of regulations to determine which make sense and which do more harm than good.

Today, it is clear that the vast majority of Americans want to develop our own natural resources and promote manufacturing jobs. Our nation is blessed with an abundance of energy resources that could revitalise job growth and our economy, enhance our national security, and ensure a strong fuel and petrochemical manufacturing industry. The decisions made by the Obama administration in 2013 will determine the future of the industry I represent and, more importantly, the entire nation.

Dmitry BalandinChief Financial OfficerGazprom Neftekhim Salavat

We are living through difficult economic times. Despite this, the market

for oil and gas production in Russia remains stable. The main source of this stability is high oil prices, which allow OJSC “Gazprom Neftekhim Salavat” and other major market participants to achieve positive margins. The domestic market still tends to be influenced by the activities of government and the regulatory authorities, in particular the measures government takes to reduce domestic fuel prices. During the second half of 2011, the government implemented a new export duty system (so-called “60-66”), which benefits the upstream industry, but is less advantageous for companies within the downstream sector. Thanks to our broad market appeal and high-quality offerings, Gazprom Neftekhim Salavat is finding the export and domestic markets to be profitable.

There are challenges that Russian process industry companies operating in the oil and gas market need to overcome in the next 12 months. One key challenge is the high volatility of oil prices and crack spreads for oil


nitrogen removal. The extent of FCC yield improve-ments were often a function of desired operational cycle life and available hydrogen for the pretreat units. Hydroprocessing catalyst systems were developed util-ising cobalt molybdenum (CoMo) and nickel molybdenum (NiMo), depending on these objectives and constraints.

In today’s clean fuel operations, much investment has been made in ULSD and FCC naphtha HDS, with few refiners now achieving environmental compliance via previously designed pretreat units. Additional results of the global drive towards clean fuels are continued advances in catalyst technology that have provided significant gains in both HDS and HDN performance. These technology gains are being utilised to drive new FCC pretreat designs to very high levels of performance and have provided refiners with the option of revisiting how best to maximise the value of existing FCC pretreat units. This has resulted in many units shifting catalyst system designs in order to provide higher levels of nitrogen removal and aromatic saturation by using more high-activity NiMo catalysts, maximising FCC conversion capability.

If distillate maximisation is desired, many FCC pretreat units can be revamped to effectively operate in a MHC mode. This more severe operation is performed with higher reactor temperatures and often by modify-ing the catalyst system to include a more active

conversion catalyst such as an amorphous silica-alumina (ASA) or zeolite.

Depending on the conversion and distillate selectiv-ity required, all alumina, alumina/ASA or alumina/ zeolite stacked systems can be considered. Higher conversions can be achieved by using alumina/ASA stacks and even higher by using alumina/zeolite stacks compared to a total alumina system. In specifying a MHC catalyst system, the balance of hydrotreating versus cracking catalyst and the potential addition of reactor volume is largely influenced by feed qualities and the desired level of conversion. As many of the feeds processed are high in contaminant metals, sulphur and nitrogen, the pretreat section is required to remove these contaminates to ensure a sufficient cycle life can be maintained while both meeting any product targets and minimising nitrogen slip into the cracking section of the reactor. Feed quality and the reactor and catalyst system specified determine the ultimate sulphur and nitrogen removal capability for a given cycle life. HDS functionality can remain an important criteria for some MHC units depending on existing product specifications, which are dependent on site refinery constraints and capabilities. However, HDN capability is often more important, as it influ-ences cracking catalyst selection and performance due to the remaining nitrogen heteroatoms, which reduce cracking reactions. As mentioned, zeolite-containing products can provide the highest levels of conversion, but they tend to be the most sensitive to nitrogen slip, reducing their long-term effectiveness in such cases. Amorphous silica alumina (ASA) cracking catalysts provide increased levels of nitrogen tolerance with a lower level of conversion capability and, for units with limited HDN capability, conventional pretreat catalyst can be operated in a MHC mode but with a reduced conversion capability.

When using desalter water for coke cutting, you should at least consider some major problems related to water composition: entrained salts; entrained caus-tic; and entrained sludge and sediments. Entrained inorganics might have an impact on the metallurgies of the cutting tool, coke drum and lines, together with potential salts precipitation in the cutting tool’s nozzles. Some mitigation of such phenomena might be found in acidification and/or anti-scalant injection, but all of that needs to be carefully evaluated.

Also consider that caustic, oil carry-under and chem-icals might act as emulsifiers and can potentially stabilise the frothing of coke particles, preventing/limiting their precipitation in the water recovery system and creating coke particles carry-over in the cutting tool, which, in turns, creates plugging and possible erosion/corrosion.

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Online cleaning of an integrated distillation unit

Grupa Lotos is a vertically inte-grated oil company based in Gdansk, Poland, and is a

producer of high-quality fuels, motor and industrial lubricants, bitumen and waxes. Gdansk refin-ery is the main asset of Grupa Lotos. Located on the Baltic Sea coast, it is a complex oil refinery with an annual processing capacity of 10.5 million tonnes (approximately 210 000 b/d) and is the second larg-est refinery in Poland. Refinery conversion is achieved by two hydrocracking units, one of them being fed by deasphalted vacuum residue. Other major process units include catalytic hydrotreaters, two catalytic reformers, two hydrogen generation units and a solvent-based lubricants plant.

Grupa Lotos has completed the implementation of a refinery upgrading and modernisation project known as the “10+ Programme”, within which new process units have been built, such as diesel hydrodesulphurisation, mild hydrocracking, ROSE solvent deasphalting, steam reforming, and an atmospheric and vacuum distil-lation unit. As a result of the completion of this programme, the overall crude distillation capacity in the refinery has reached 10.5 million t/y (210 000 b/d).

The company focuses on improv-ing its operational excellence and carefully evaluates new opportuni-ties to reduce overall operational costs. In this connection, it wanted to validate ITW Online Cleaning technology, in order to take advan-tage of its benefits both during the run of the units or a scheduled turnaround.

Online cleaning technology was validated during a scheduled turnaround at the second largest refinery in Poland

Mariusz HOŁOwacz and raFaŁ zaPrawa Grupa LotosMarcellO Ferrara ITW

The opportunity to validate was a scheduled turnaround of the CDU/VDU. The validation could there-fore include visual inspection of each piece of equipment. The CDU/VDU is a new, integrated crude distillation unit/vacuum distillation unit. This is one of the main units built in the 10+ Programme, with a crude oil capac-ity of 4.5 million t/y. The unit started up in March 2010 and is designed for processing heavy crude with a high content of naph-thenic acids (opportunity crudes).

The CDU/VDU mainly consists of four sections: desalination and stabilisation of crude oil, atmos-pheric distillation, vacuum distillation and naphtha stabiliser.

Crude oil goes into the desalina-tion section, then goes to the pre-flash and crude tower, where it is separated into lighter and heav-ier fractions. The lighter fraction is routed to the gasoline stabilisation section, wherein LPG and fuel gas are separated. Heavy and medium fractions are separated into kero-sene and diesel fractions: LGO, MGO and HGO. Atmospheric resi-due is heated in another furnace and goes to the vacuum distillation column, wherein VGO and LVGO are separated. Vacuum gas oils are processed in a hydrocracking plant designed for diesel processing. Heavy residues, such as vacuum residue and the slop wax fraction, are directed to the SDA/ROSE and bitumen unit.

ITW is a service company with an applied research background, which has developed and patented, among other things, a technology for online cleaning production

units. This technology transforms sludge and coke-like deposits into a fully reusable product. It has been successfully used for online clean-ing of refinery and petrochemical production units. In the case of polymeric fouling, a novel polymer modification chemistry is used to perform online cleaning. Different refinery and petrochemical units have been cleaned online in as little as 24 hours on an oil-to-oil basis.

advantages of online cleaningAn array of technologies is currently available for solving problems related to equipment cleaning, including ITW’s patented technol-ogy for cleaning heat exchangers and process equipment in a closed loop. Equipment can be cleaned online with this technology without having to extract the heat exchanger bundles. This is particularly benefi-cial in severe applications such as visbreakers and columns with struc-tured packing internals. Crude distillation units (CDUs) and vacuum units have also been successfully cleaned online, as well as many other refinery units. An entire unit can be cleaned in 24 hours. The advantages of online cleaning over hydroblasting include:• Reduced downtime• Closed loop operation, without the need for extracting the bundles• Simultaneously cleaning of multi-ple units• No waste generation• No emissions.

The ITW cleaning method involves the injection of an oil-based chemical that transforms sludge into a fully reusable prod-uct, removing any sludge and


itw.indd 1 10/09/2013 17:02


sational costs, huge production loss costs must also be added, followed by equipment downtime. Even more relevant are the costs associ-ated with the decrease in performance during the run. Moreover, all these operations can suffer from different problems, such as:• Time for isolating/inserting the equipment• Waste generation• Hydrocarbon emissions• Bundle damage•Difficultiesinbundleextraction•Difficultiesinunbolting• Hazard to workers performing mechanical operations• Hazard from scaffolding• Hazard from spills of carcino-genic products• Hazard to people working nearby moving equipment/vehicles• Hazard from bundle transporta-tion/cleaning apparatus/cranes.

Due to continuous improvements in process safety and environmen-tal performances, these problems are now only acceptable when there really is no other choice.

While hydroblasting/mechanical cleaning were the only reliable technologies available in the market, Gdansk refinery searchedfor an improvement over the conventional way of doing the job.

ITW proposed an alternative solu-tion to hydroblasting by providing an online cleaning technology.

Job executionOnline cleaning started after the unit was initially inventoried with a carrier and the patented MEG F550F chemistry was injected into the loop. The closed circulation loop was arranged as agreed between Lotos and ITW. The desalters were not included in the loop. The carrier had the character-istics shown in Tables 1 and 2, and Figure 1.

The patented cleaning method essentially consists of the following steps:• Inventory the unit with the carrier• Inject ITW chemical MEG F550F• Circulate for the desired amount of time (with a 24-hour oil-to-oil basis in mind)

deposits from metal surfaces. The washing solution can be reutilised/reprocessed because the stabilising properties of the chemical avoid the reaggregation of deposits, eliminat-ing any potential precipitation during storage and/or fouling of process equipment. Moreover, the patented chemicals used in ITW’s method are compatible with hydro-carbon processing and combustion because they:• Do not contain any metals• Do not contain any compounds based on P, B, S, As, Bi, Si or Pb• Do not contain any halogen• Do not contain any compounds that might be harmful to plant metallurgy• Do not contain any carcinogenic compounds• Do not contain any compounds that in operating dosages might interfere with biological waste treatment processes.

After online cleaning, the heat exchangers canbe immediatelyputback on stream without extractingthe bundle. However, if visual inspection is to be performed, a cleanbundlecanbeextracted.Hardcoke will be dry if it is present, as the technology eliminates all coke binders, and so is friable and easily removed. Any coke left will be in very negligible quantities compared to an alternative cleaning solution.

Current operationsThe standard method used by Gdansk Refinery was to drain theunit and steam out the exchangers(and any other equipment) so that the exchanger heads could beremoved. At this point, the exchanger bundles were extracted,transported to the washing area and hydroblasted. After hydro-blasting, the bundles were transported back to the unit and remounted. For the vacuum and atmospheric towers, coke was removed by mechanical cleaning.

Although a “normal” operation for a refinery, the hydroblasting/mechanicalcleaningofanexchangerinvolves at least 20 different opera-tions, with related hazards.

Technology for cleaning heat exchangers includes the followingprocedures:

• Flushing• Isolating equipment from the process and blinding• Removing hydrocarbons• Steaming out for gas removal• Scaffolding• Unbolting• Removing heads and distributor•Extractingthebundle• Transporting the bundle from plant to washing area• Hydroblasting• Transporting the bundle from washing area to plant• Inserting the bundle• Inserting new gaskets• Installing covers and distributor• Bolting• Removing blinding• Removing scaffolding• Air removal and purging• Performing pressure tests• Inserting the apparatus in the process.

All of these operations have their related costs. The bidding paid costs are to be added to the custom-er’s organisational costs in terms of:• Awarding (bid organisation, eval-uation and award)• Planning• Co-ordination•Controlduringexecution.

To these operational and organi-

Density @15°C, g/cm3 0.8334 Viscosity @ 15°C, mm2/s 6.204 Viscosity @ 40°C, mm2/s 2.228 CCR, wt% 0.014 Sediments, wt% 0.015 Ash, wt% 0 Ni, ppm <0.05 V, ppm <0.05 Cl, ppm <5 Acid number, mg KOH/g 0.256

Analytical characteristics of the carrier

Table 1

IBP, °C 177 5%, °C 203 10%, °C 211 30%, °C 225 50%, °C 234 70%, °C 244 90%, °C 260 95%, °C 267 FBP, °C 279

Carrier distillation

Table 2

itw.indd 2 10/09/2013 17:05

• De-inventory the unit by pump-ing out the washing fluid.

Samples of the washing fluid were pulled, and the data were utilised to monitor the progress of the job and approximate the quan-tity of heavy sludge/solids that were dissolved and stabilised into the washing fluids. A sample of the carrier oil was taken prior to inject-ing the ITW chemical. Then, subsequent samples were taken prior to de-inventory of the unit. An example of the monitoring results is shown in Table 3.

ITW uses a proprietary software to correlate the analytical results for monitoring operations and calcu-late the actual amount of deposits that have been removed from the system. It is to be noted, that the original sludge/coke-like material has been fully converted into a stable and reprocessable fluid. Washing fluids have been routed into a crude tank.

ResultsBased on a monitoring programme, it was calculated that a minimum of 20 000 kg of deposits were removed from the CDU/VDU Unit 120. In order to validate the achieved results, visual inspection of exchangers and columns was performed on the cleaned equip-ment. All of the equipment was very clean and free of any organic deposits.

Figures 2-7 show some of the results in the equipment inspected. All of the pictures were taken on opening the equipment, without any mechanical cleaning being performed. In some exchangers, a very minor and negligible inorganic deposit was left, which would have no impact on performance when starting up the unit just after


270

290

250

230

210

190

170

Boili

ng p

oin

t, º

C

150IBP 5 10 30 50 70 90 95 FBP

Distillation, %

IBP = Initial boiling pointFBP = Final boiling point

Figure 1 Carrier distillation characteristics

Parameter Carrier Washing fluidDensity @ 15°C, g/cm3 0.8334 0.8437Viscosity @ 15°C, mm2/s 6.204 7.734Viscosity @ 40°C, mm2/s 2.228 2.967CCR, wt% 0.014 1.09Sediments, wt% 0.015 0.06

Typical washing fluid analysis after online cleaning vs carrier

Table 3

Figure 2 Vacuum tower bed 1

Figure 3 Vacuum tower bottom

itw.indd 3 10/09/2013 17:05


cleaning. As a result, the refinery has decided not to schedule mechanical cleaning during future turnaround activities or for recover-ing unit performance.

Ideally, the best way to quantify the improvement is to immediately start the unit following cleaning and to examine heat exchanger performance (furnace inlet temper-ature improvement, duty improvement, pressure drop improvement and so on). This would be useful in understanding the future possibilities of eliminat-ing mechanical cleaning and dramatically speeding up the shut-down process (especially when the unit is processing opportunity crudes that are known in the indus-try to cause fouling problems).

Most of the equipment was not hydroblasted/mechanically cleaned after online cleaning. Some bundles were scheduled to be hydroblasted anyway, but in general there was no need to hydroblast them and, if any minor deposits had been left, these would have been removed by the normal flow, as they were negligible in quantity (1-2 kilograms) and inconsistent as material.

It should be noted that washing fluids were routed into a crude tank and fully reprocessed upon the unit’s start-up. No problems have been reported at all in the CDU/VDU, or in any of the down-stream units and/or in the finished products specification. This further confirmed that the washing fluids are fully reusable and reprocessable.

The results open up a new oppor-tunity for Lotos, as online cleaning can be applied during a plant run, to recover the unit’s performance and, in preparation of a turna-round, to reduce downtime and the turnaround’s scope of work.

Proactive online cleaningGrupa Lotos’s validation also aimed to evaluate the different options for applying online clean-ing, including a proactive option. Proactive application of online cleaning is a departure from current operating procedures.

Refinery/petrochemical unit runtime is currently dictated by

Figure 4 Vacuum tower feed inlet and bottom trays

Figure 5 Vacuum residue – crude exchanger (upstream atmospheric tower furnace)

Figure 6 Crude exchanger (upstream pre-flash tower)

itw.indd 4 10/09/2013 17:05


processes might increase in the future, and the fouling of units might become a more serious concern, potentially impacting a large number of production days. By regular application of online cleaning, the unit can always be operated under clean conditions. In addition, the vast majority of mechanical work, which takes up a significant number of days during unit shutdowns, can be replaced with an online cleaning process that requires no opening of equipment or man entry, and can be carried out in as little as 24 hours oil-to-oil.

AcknowledgementSpecial thanks to all the Lotos team for the support and valuable assistance provided during the entire time that ITW Online Cleaning has been applied on site.

Mariusz Hołowacz has been Crude Oil Distillation Complex Manager with Grupa Lotos in Gdansk, Poland, since 2008 and, since 2010, Manager of the new distillation unit 120.Rafał Zaprawa was VDU and propane deasphalting unit Shift Team Manager with Grupa Lotos in Gdansk, Poland, from 2001-2008, then Deputy Manager of the crude oil distillation complex. Marcello Ferrara is the Chairman of ITW. With 27 years’ experience in the petroleum business, including oil exploration and production, refining, petrochemicals, transportation, and energy production, he holds a PhD in industrial chemistry and international patents for new processes and additive compositions for environmental control and for improving petroleum/petrochemical processes. Email: [email protected]

market considerations and by the downtime required for a mechani-cal cleaning turnaround. With a downtime of 15-20 days, it is much more economical to run the units under non-optimised conditions rather than lose production. This, however, results in energy losses, giveaway and capacity reduction, which negatively impact unit economics.

The introduction of online clean-ing, by cleaning the unit in 24 hours on an oil-to-oil basis, allows for the recovery of losses and the operation of units under improved and more reliable conditions. Furthermore, turnarounds can be avoided or rescheduled with reduced downtime.

In the case of turnaround improvement, an additional reduction in downtime can be achieved by applying ITW’s improved degassing/decontamina-tion technology.

ConclusionThe results of ITW Online Cleaning have opened up new possibilities for Lotos, whereby online cleaning can be applied during a plant run, to recover a unit’s performance and improve the level of operational excellence, and in preparation for a turnaround, to reduce downtime and the turnaround’s scope of work.

These options are increasingly important, as the amount of oppor-tunity crudes the company

Figure 7 HGO — crude exchanger

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Flexim 63 www.flexim.com

Foster Wheeler 19 www.fwc.com

Gas Asia Summit 55 www.gasasiasummit.com

Grabner Instruments 57 www.grabner-instruments.com

Hoerbiger Kompressortechnik Holding 27 www.hoerbiger.com

Hydroprocessing Associates 58 www.hpa-usa.com

ITW Technologies 25 & 41 www.itwtechnologies.com

KBR IFC www.kbr.com/Technologies

Koch-Glitsch 17 www.koch-glitsch.com

Linde 6 www.linde-worldwide.com

Linde engineering North America 14 www.lindeus-engineering.com

Onis International 53 www.onis.fr

Operational excellence 2013 IBC www.opex.biz

Optimised Gas Treating 47 www.ogtrt.com

Process Consulting Services 8 &10 www.revamps.com

Sabin Metal Corporation 32 www.sabinmetal.com

Sulzer Chemtech 28 www.sulzerchemtech.com

Tiger Towers 39www.tigertowerservices.com UOP 22 www.uop.com

Zeeco 48www.zeeco.com

Alphabetical list of advertisers

For more information on these advertisers, go to www.ptqenquiry.com


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