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3 Identifying slug flow key to successful debutaniser revamp Conversion of a stripper tower to debutaniser operation shows how external piping in slug flow regime can negatively impact a tower Daryl W Hanson Valero 9 Yield and energy improvement in atmospheric and vacuum unit Impact of BS III diesel specifications on legacy AVU configuration called for optimising fractionation performance through improved mass transfer efficiency Rahul Srivastava, Manish Rohit, and Harshit Budholiya Indian Oil Corporation Ltd Giuseppe Mosca Sulzer Italy Srl Unmesh Bhagwat, Mukesh Gupta, and Amar Patil Sulzer India Private Limited 15 Cracking the code: FCC turnaround lifecycle best practices – Part 2 The common FCC equipment and recommended engineering practices are reviewed, followed by unit start-up and turnaround report generation Ben Ellebrecht Phillips 66 Wood River
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23 Multi-point temperature mapping of steam methane reformer tubes A retrofit field deployment of ultrasonic waveguide skin temperature sensors in a refinery hydrogen plant Balaji Nambi and Nishanth Raja XYMA Analytics Avoid corrosion failures in refinery process equipment High-velocity thermal spray alloy cladding mitigates corrosion in ammonia scrubbers and amine columns, improving asset integrity and reducing lifecycle costs Shinji Masuzaki Integrated Global Services (IGS) 35 Operational problem solving in heat exchangers and air coolers Selected case studies of strategies developed for resolving heat exchanger operational problems Sadhna Singh and Harshita Gupta Engineers India Limited Cover OMV operates Austria’s only refinery in Schwechat. Today it is one of the largest and most modern inland refineries in Europe and has integrated petrochemical production. Its total annual refinery capacity is 9.6 million tons of crude oil. Courtesy: OMV
©2026 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.
Rethinking Old Problems
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Revamp projects are difficult. Limitations imposed by plot space, congested pipe racks, and outdated equipment, to name a few, present unique challenges. Solutions that rely on excessive margins or comfortable designs lead to overspend. Now more than ever, process designers must find solutions that do more with less. P roven M ethods There is growing awareness that better scope definition earlier in the engineering phase saves time, reduces overall engineering cost, and leads to more successful projects. There is no argument that work completed during Conceptual and Feasibility phases is critical to getting a project on the right path. Engineers at Process Consulting Services, Inc. have developed a proven approach that makes the most of this precious time. At site, PCS engineers coordinate rigorous test runs, much of it through direct field measurements. Data collected is invaluable and often leads to low hanging fruit or hidden gems. Some refinery equipment performs better than design, and for various reasons others perform worse. Good test run data allows seasoned engineers to quickly identify what equipment needs investment and what equipment can be exploited. This way, solutions are developed that direct capital expense in the right areas and overspending is avoided. In one example, pressure drop measurements of a long crude oil transfer pipe showed the line could be reused, saving millions of dollars. Contact us today to learn how PCS’ proven methods can help you do more with less in your next revamp.
Projections for global supply and demand of refined products vary greatly depending on the pace of technological progress and degree of government policy enforcement associated with reducing greenhouse gas emissions. Without major advances in technology, it is hard to imagine a future without conventional fossil fuels over the next decade or two. Based on history, continued rationalization of refining assets is likely. Small, low-complexity refineries will struggle, while large, complex ones will thrive. Capacity creep through gradual improvement of refining units will continue to be a differentiating characteristic for remaining players. Focused revamps will play a critical role. Post-pandemic, inflation and a shortage of skilled construction labor have dramatically increased costs for refinery revamps. It is becoming increasingly difficult for many projects to meet corporate return on investment thresholds.
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Identifying slug flow key to successful debutaniser revamp
Conversion of a stripper tower to debutaniser operation shows how external piping in slug flow regime can negatively impact a tower
Daryl W Hanson Valero
A Valero Gulf Coast refinery recently installed a new alkylation unit that required the removal of butanes from an alkylate (gasoline) product. As part of the revamp, the new debutaniser had to be sited in a very congested area. An out-of-service tower was identified and could be replaced and used in debutaniser service. However, since it was smaller, high-capacity trays were installed. In a previous paper,1 UOP’s proprietary high-capacity Parallel Flow Multiple Downcomer (PFMD) and Enhanced Capacity Multiple Downcomer (ECMD) trays were installed in the tower. The revamped tower diameter was 54in, which is less than the 60in minimum diameter to effectively install a two-pass crossflow tray without special accom- modations (no tray manways, eliminating flow path length minimums). Installation and start-up went as expected, and the tower met the operating goals. However, starting at around 70% of design, tower pressure drop measurements exhibited periodic cycling that had not been seen before. This cycling grew in severity but did not inhibit tower performance above the design rates. Evaluation of this cycling behaviour revealed that a flow regime issue in the feed upcomer pip- ing was causing the problem. Process changes enabled a reduction in cycling severity. Small tower difficulties In small towers, there is naturally a lack of space to add downcomers to handle capacity without sacrificing a lot of efficiency. The tower area simply is not abundant enough to have the luxury of a downcomer addition. There are only a few industry examples of multi-pass trays (two-pass) in towers less than 5ft diameter. The author is aware of one depropaniser where two-pass high-capacity crossflow trays were utilised in a tower that was 42in ID (inner diameter). To handle the capacity, the end user sacrificed efficiency and manway access to install the tray. The efficiency was within the error band for a point efficiency device. When applying high-capacity trays to small columns (<6ft diameter), the practitioner typically has to settle on which variable is more important for their system: capacity or efficiency.
Repurpose of existing tower into debutaniser operation An old methyl tert-butyl ether (MTBE) unit alcohol strip- per tower was identified as a perfect location for the new debutaniser. The tower location would allow close connec- tion to existing piping for the feed and bottoms. The tower was only 54in ID, which was far smaller than the existing tower. The alcohol stripper tower would have to be refur- bished, and its certificates updated as part of the repurpose to debutaniser service. The foundation was checked for a larger tower diameter, but it was deemed too close to its limit. A process evalua- tion was performed, and high-capacity counter-flow trays in the stripping section were the limiting capacity, barely achieving the design loadings. The alcohol stripper (new debutaniser) tower was removed and sent to a local shop for refurbishing. Scope included: removing external platform and piping supports, installing new supports, replacing/removing existing nozzles, and installing new internal tower support rings/attachments. Installing new support rings was required to avoid re-spacing complications in the top of the tower and cor- rosion issues in the bottom. As such, the upper section of the tower had new rings installed at 18in spacing (exist- ing 24in), and the lower section had rings installed at 24in spacing (matching existing).
Vapour ood (Jet) Liquid ood (Downcomer) Observed ood
Constant
Time
Figure 1 Characteristic flood profiles
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Pressure (PSIG) Temperature (˚F)
Butane
7000
DIB #1
6900
6800
Horizontal 300 ft
6700
DIB #2
Vertical 75 ft
Re ux (bpd) Stripping DP (psid)
6600
Note: Data 1 minute averages
6500
1:29
1:37
1:46
1:55
2:03
Horizontal 100 ft
Horizontal 75 ft
Alkylate
Figure 2 Cyclic pressure drop in stripping zone
Figure 3 Debutaniser feed schematic
A combination of the high-capacity PFMD and ECMD trays was installed in the tower to meet the revamp expec- tations. Tray installation was successful and met all require- ments. The revamp success story is documented in the AIChE paper.1 Troubleshooting the revamp After a very successful start-up, the team became very interested in tower performance testing. The author con- siders performance testing to be a critical part of the pro- ject. The performance test not only validates the design but also provides a good source of data for tower benchmark- ing and enables operations to get a good feel for the tower as it approaches limitations. In this case, operations got a ‘three lessons in one’ per- formance test, but the author will only cover the most significant learning from the test. The full results of the per - formance test were presented in the AIChE presentation on which this article is based. The goal of the performance test was to push the tower trays and peripherals to their limit. Throughout the day, the internal recycle rate (tower distillate-to-feed rate) was increased to load up the reactor section. This would allow higher feed rates and increase tower loading (requiring more reflux and reboiler duty for constant separation). After the recycle rate was increased, the tower distillate was placed in flow control while the reboiler steam rate was increased to intentionally load up the tower. Small changes were made to the reboiler duty to artificially load up the tower. At 70% of design rates, the tower entered a distinct pressure-drop cycling mode that intensified as throughput increased to 90% of the design rate. Above 90% of the design rates, the amplitude of the cycling seemed to be constant. Every 540 seconds (nine minutes), the pressure drop would go through a defined cycle. The pressure drop would cycle 110% from high to low off an average. Figure 1 shows the standard responses in dP (ΔP) for a typical downcomer and jet flood situation. Unlike typical jet and downcomer flood, the debutaniser pressure drop would never escalate into a liquid build-up or exhibit the typical characteristics of flood (loss of temperature differ - ential, loss of liquid levels, or massive liquid carryover).
The observed ‘flood’ was very unusual based on the author’s experience. The author had only seen this type of ‘stitching’ pressure drop when a previous tower was on the verge of dumping. However, it operated this way only for a short (<4 hr) time frame. The debutaniser would operate in this ‘stable’ mode for days, weeks, and even months on end. There was never any evidence of uncontrolled prop- agation resulting in an unstable column that affected downstream column operation or product qualities. After operations got accustomed it, it became a non-issue to maintain. However, discovery work was required to address this unique issue, even though there was no real ‘urgent problem’ to solve. During the cycling, the tower products were sampled at tight intervals. Even though this was not the typical method to flood, the author was interested in the effects of flooding. The samples were taken at tight intervals to ‘catch and identify’ operating deviation. The samples did not reveal any loss in efficiency that would show up as heavy key or light key breakthrough into the respective products. The sample results did not indicate that the tower was losing efficiency or separation performance during cycling compared to normal flatline pressure drop operation. In fact, even the control point temperatures were flatlining during the evaluation period, which suggested normal fractionation. The magnitude of the pressure drop cycling seemed to indicate that a single tray, or at most two, were flooding. This would show up as a stacking of a couple of feet of light hydrocarbon liquid. The pressure drop cycling was about 0.1-0.25 PSID (depending on feed rate). Figure 2 shows a subset of the tower pressure drop data and the corre- sponding flow diagram (see Figure 3 ). By process of elimination, it was quickly ruled out that the lowest tray (highest loaded) was flooding. Typically, three events accompany a stripping section flooding event. First, the reboiler duty and return temperature will fluctu - ate/cycle, and this was not occurring. Second, when the bottom tray floods, the flood frequently propagates up the tower, resulting in an accompanying response of liquid build-up, high pressure drop, and bottoms liquid inventory
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loss. Third, the bottoms liquid level control (flow) becomes erratic. None of these was occurring, and that theory was ruled out as the leading problem. After eliminating the possibility of the lowest tray flood - ing, the author concentrated on the next most likely zone to flood: the feed zone. The feed zone design was reviewed many times and showed no significant issues. Tower feed quality (such as percentage of vaporisation) was dou - ble-checked and did not reveal any problems. Operator rounds Having operators involved during a troubleshooting case is always beneficial and necessary. During one of the trouble - shooting visits, the author noticed a distinct popping sound near the base of the tower. This distinct sound was simi - lar to that produced when wet steam is injected into a hot product side stripper. It was a subtle pop. Then, eight min - utes later, another pop. Then, eight minutes later, another pop. The operator was asked whether he had heard this sound before the revamp, and he did not remember hear - ing it. Searching for the source of the sound led the author to the reboiler. After a while, the author concluded that the sound was not coming from the reboiler. However, what was it? The only consumer of steam was the debutaniser reboiler within 100ft. The sound was not due to a steam issue; it was tracked to the feed line (around the point where the combined feed line turned vertical), which was very close to the reboiler. Figure 3 shows a simple PFD of the debat - aniser feed system. One of the author’s mentors always stressed the need to touch items in the unit. He was always gathering infor - mation to see if they were hot or cold, whether they were pulsing, or whether they were vibrating. In this case, the feed temperature was <250°F and the insulation could be touched without any safety issues. The team could physically feel the line jumping/moving around; the movement aligned with the ‘popping’ sounds they heard. The next step was to find out what was causing the sounds and oscillations in the line. A low-point bleeder was located about 5ft from the origin of the vertical run for the feed pipe, alongside the debutaniser tower. It provided a critical piece of information. A local pressure gauge was installed on the bleeder at grade. During the survey, there was ongoing consultation with the board to make sure that the tower top pressure was constant. During the evaluation, the tower top pres - sure did not move more than 0.50 psi high-to-low. The pressure gauge on the feed showed a rhythmic cycling of 72 psig to 81 psig. After recording a few points, the author recognised that the cycles were within 15 seconds of the peak-to-peak events on the stripping section pressure drop excursions. There was now concrete evidence that the feed circuit pressure was causing a stripping section pressure drop excursion and flooding. Figure 4 s hows the pressure read - ings at the beginning and mid-cycle. Immediately, the team focused on the feed enthalpy and
Feed pressure after slug relieved and tray flooded
Feed pressure when slug forming and tray unflooding
Figure 4 Feed piping pressure at grade during a cycle
vaporisation. Was the feed temperature cycling resulting in problems? In Figure 3, it can be seen that the feed comes from two separate deisobutanisers (DIB#1 and DIB#2). DIB#1 has a feed preheater, while DIB #2 does not. Effluent from the DIB#1 feed preheater did not indicate changes in heat balance that were sufficient to cause the downstream issues. Since DIB#1 was direct from the tower, it was also eliminated since the bottoms temperature (i.e., composi - tion) was very steady. The author knew that since the temperatures of the upstream towers were not changing and the composition of the feed was not changing, the only variable left to set the enthalpy was the feed pressure. What would cause the pressure to change radically from the feed nozzle to grade? The answer was identified after mapping the flow regime. Vertical flow regime maps are not readily understood (upflow versus downflow) and are highly variable due to the effects of static head that accompany vertical flow Analysing feed line flow regimes Flow regimes are readily used when designing highly flashing streams, including reboiler returns. Experience from recent jobs indicates that this evaluation is not fre - quently done on feed streams. Based on the results of this case, it is recommended that all distillation tower feed pip - ing be reviewed for flow regimes. Horizontal flow regimes are well understood, and the flow regime maps are well defined.2 , 3 , 4 Vertical flow regime maps5 , 6 are not readily understood (upflow versus downflow) and are highly variable due to the effects of static head that accompany vertical flow. This is especially true for services that experience large changes in vaporisation, where a discrete element analysis is required. There are also proprietary methods that are used to analyse flow regimes.
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securely inside the slug flow ‘zone.’ At grade, where the pressure is higher and vapor- isation lower, the operating point would move to the left and stay within slug flow and could possibly enter into the very narrow bubble region before going single phase (no vaporisation). Strategies to get out of the slug flow region are shown in Figure 6 : reduce diameter to increase mass velocity (and pressure drop), reduce vapor- isation and transition to single phase, or increase vaporisa- tion to push into annular/mist flow. • Reduce diameter : For new installations, diameter reduc- tion has been used by the author to prevent slug flow and even in revamp scenarios to eliminate pulsing/vibrating up-comer sections of pipe. In this case, piping changes were not possible since it would incur significant eco- nomic loss to an operating unit • Reduce vaporisation: Calculations indicated that moving into the shallow region of 0 to 0.2 on the 1/ Xtt axis was not possible with the hot feed from the DIB#2 tower. A short field trial of this option was attempted and
1000
Reduce diameter
Design basis
100
Mist
Increase vaporisation
Reduce vaporisation
10
Annular
Bubble
Slug
1
0.1
1
10
100
1/Xtt
Figure 5 Design point shown on the Fair vertical flow map
1000
Design basis Alt operation 1 Alt operation 2
100
Mist
10
Annular
Bubble
Slug
1
0.1
1
10
100
1/Xtt
Figure 6 Fair map with alternate operating points (circle is lower feed temperature, square is higher feed temperature)
The Fair map4 (i.e., John R. Fair flow regime map) for vertical upflow has been used in industry extensively for predicting the flow regime of gas-liquid mixtures in vertical pipes. Several other maps have been developed since its inception, but largely, they are based on the Fair map. Many refining companies have their own flow regime maps that they base their designs on. It is important for the engineer to understand that there are different characteristics for flow in horizontal, vertical up-flow, and vertical down-flow. The flow regime maps are not interchangeable. For the practitioner, it is critical to understand that the lines on the flow regime charts are not well-defined but rather are transitional. The lines on the charts should be considered broad and thick, not narrow, distinct lines. Figure 5 shows the Fair vertical flow map with the design point for the debutaniser feed line illustrated. The point shown is for the design point at the entry into the tower (i.e., pressure = tower pressure + feed distribu- tor pressure drop). It is evident that the design point is
was not successful • Increase vaporisation : Calculations supported the idea to increase the feed preheat. The DIB#1 feed preheater was bypassed, resulting in a higher temperature feed, and the feed temperature was increased by 15°F. The resultant Fair map is shown in Figure 6 . This option was successful. In a new design, all of the strategies above could be used to stay out of slug flow. Reducing diameter and increasing vaporisation would move the operating point diagonally up to the right. To further test the trays and provide some operator security, ways to eliminate the slug flow were reviewed. Increasing the feed temperature is shown as the red box (Alt Operation 1), while reducing the feed tem- perature is shown as the orange circle (Alt Operation 2). Increasing the feed temperature by increments of 5°F was trialled to decrease slugging. After the feed temperature change, the stitching pressure drop profile immediately disappeared. Figure 7 illustrates the resolution of the pressure drop with increasing feed temperature.
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Analysis Operating data analysis yielded interesting results. Public domain flow regime articles are prevalent in the literature for off-shore facilities. In fact, off-shore operators will often install slug catchers in the piping systems to avoid equipment damage, as it is impossible for them to always control the flow regime. Figure 8 illustrates the physical phe- nomena that occur on the tray immediately below the feed tray when it has to resolve the incremental flow caused by slugging. A time study on the feed line was used to determine the operating issues with the piping system. It was obvious that the slug size and pressure drop in the bottom section were inversely related. When one was high, the other was low. There was a distinct cycle that occurred between the tray being relieved of liquid load and the next cycle of slug forming. As the next slug was ready to break loose, the tray was at normal pressure drop and could accept the slug without going into a full flooding event. There are many methods in industry to estimate the slug length. The author does not have the expertise to determine the validity of each method. It was realised that many of the methods are based on a large difference in phase densities (>50 lb/cf), whereas this system has a relatively close liquid and vapour density difference of 32 lb/cf. Quick estimates range from six to 12 pipe diameters for the height of the slug. Since the tower is 54in diameter and the downcomers are close to 10in width, one can readily visualise that it would only take a liquid slug length of 12 pipe diameters to overwhelm the liquid-handling capacity of the top tray downcomer. As the tray is already operating close to 80% capacity, this slug would easily throw the tray to the liquid hold-up flooding point. It is believed that this simple anal - ysis was sufficient to substantiate tower operation and the results of the increased feed temperature trial to resolve the slug flow conditions. Conclusions Revamping a tower to a high-capacity tray is not a triv - ial issue. When the tower diameter is small and limiting, resolving every bottleneck is important. Public domain case studies are abundant with reboiler problems, feed zone problems, and transition problems that result in tray capac- ity loss. This is the first article that has identified an external pip - ing system flow regime that has a negative impact on tray stability and capacity. Resolving the feed piping slug flow regime eliminated the premature flooding instability that this tower was exhibiting and allowed for higher tower capacity. 7:58 7:49 7:40 6500 6600 6700 6900 6800 7000 Re ux (bpd) Stripping DP (psid) Note: Data 1 minute averages
+15 deg F feed temperature
8:06
8:15
8:24
8:32
8:41
8:49
Figure 7 Pressure drop resolution by increasing feed temperature
Tray ood resolves
Tray oods
Slug loosing
Slug forming
Time
Point of tray ood in ection Point of critical slug release
Relative size of slug Tower dP
References 1 Hanson, D., B. Miller, and B. Nowak, Results of an Alkylation Debutanizer Revamp with PFMD™ trays, Presented at the 2019 AIChE Spring Distillation Conference, April 1-4, 2019, New Orleans. 2 Kern, R., How to Size Process Piping for Two-Phase Flow, Hydrocarbon Processing , October 1969. 3 Daniels, L., Dealing with Two Phase Flows, Chemical Engineering, June 1995. 4 Fair, J. What You Need to Design Thermosiphon Reboilers, Petroleum Refiner , February 1960. 5 Anderson, R., Designing for Two Phase Flow – Part 1, Chemical Engineering, December 6, 1965. 6 Street, J., Unsteady State Gas-Liquid Slug Flow Through Vertical Pipe, AIChE Journal, Vol. 11, No 4, July 1965. Daryl Hanson is a technology advisor for Valero Energy in San Antonio. He is in the Strategic Technology and Development group where he oversees large projects, advanced optimisation, and troubleshooting in the area of fractionation and acquisition review. He manages the frac - tionation group and the heat transfer group, including fired heaters. Hanson has more than 30 years of experience and has published many technical papers. Before Valero, he worked for Glitsch, Koch-Glitsch, and Process Consulting Services. Hanson has a BS in chemical engi- neering from Texas A&M University. Figure 8 Operating regimes for slug flow interaction with tower internals
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Yield and energy improvement in atmospheric and vacuum unit Impact of BS III diesel specifications on legacy AVU configuration called for optimising fractionation performance through improved mass transfer efficiency
Rahul Srivastava, Manish Rohit, and Harshit Budholiya Indian Oil Corporation Ltd Giuseppe Mosca Sulzer Italy Srl Unmesh Bhagwat, Mukesh Gupta, and Amar Patil Sulzer India Private Limited
I ndia’s transition from Bharat Stage (BS) II to BS III fuel specifications marked a significant tightening of die - sel-quality norms, driven by national emission-reduction objectives. Under BS III, the sulphur limit in high-speed die- sel (HSD) was reduced from 500 ppm to 350 ppm, with phased implementation that began in 13 cities in April 2005 and extended nationwide by April 2010. Alongside sulphur reduction, a less visible but equally impactful change was introduced in the diesel distillation specification, wherein the T95 recovery temperature (ASTM D86) was lowered from 370°C to 360°C to limit heavy-end hydrocarbons and associated particulate emissions. While these regulatory changes were environmentally progressive, they created a structural mismatch with the design intent and hydraulic capability of several legacy crude distillation units (CDUs) in India, including the atmos- pheric and vacuum unit (AVU) at the Indian Oil Corporation Ltd Mathura refinery. Legacy design constraints The Mathura refinery AVU was commissioned in 1982 with a design capacity of 6 MMTPA and was later expanded to 8 MMTPA. The CDU fractionation scheme was designed under pre-BS-III conditions, when diesel T95 limits were more relaxed, enabling recovery of heavier gasoil compo- nents directly into the straight-run diesel pool. Although a pre-fractionator column was added post- Preheat-II to manage increased vapour traffic following throughput augmentation, the CDU main column internals, particularly in the light gasoil and heavy gasoil (LGO-HGO) section, were last revamped in 2004, prior to BS III imple- mentation. The packing hydraulic design in this section was not intended for sharp LGO-HGO separation under stricter diesel end-point control. Regulatory impact on fractionation strategy With BS III implementation, the diesel T95 constraint of 360°C had to be strictly maintained at the CDU draw itself, especially given Mathura refinery’s single CDU configura - tion and absence of crude segregation or parallel process- ing flexibility. This eliminated the earlier operational latitude of routing the HGO cut, drawn just above reduced crude oil (RCO), into the diesel pool.
Consequently, the HGO stream became a swing cut, dynamically routed to either the HSD or VGO pool depend- ing on available specification headroom and product mar - gin. However, the HGO stream contained more than 50% diesel-range material, rendering it unsuitable for consistent routing to the VGO pool. This was due to significant overlap between LGO and HGO, driven by: • Existing packing and internals limitations. • Low liquid-to-vapour (L/V) ratios caused by high vapour traffic. Limitation imposed on secondary units Slippage of HSD material into the VGO pool resulted in avoidable loading of high-severity secondary conversion units, namely the fluid catalytic cracking (FCC) unit and once-through hydrocracker unit (OHCU), both of which are energy-intensive units with inherently lower liquid product efficiency for diesel-range feeds. When diesel-range material is routed to the hydrocracker, only about 65-70% is converted into lighter products such as HSD and kerosene, while the balance remains as uncon - verted oil (UCO). This UCO is subsequently routed to the FCC unit for further processing. In the FCC, processing of this UCO leads to intrinsic yield losses, with approximately ~5 wt% of feed converted to coke and ~4-5 wt% converted to fuel gas. These losses reduce overall liquid yield and increase energy consump- tion and catalyst stress. Consequently, routing diesel-range material through the hydrocracker to FCC pathway results in a compounded conversion penalty, higher hydrogen and fuel usage, and increased coke generation, representing a clear economic loss when compared to direct treatment routes. In contrast, processing the same diesel-range material in hydrotreating units such as diesel hydrodesulphurisation (DHDS) or diesel hydrotreating (DHDT) offers approxi- mately 99% (~99%) liquid yield conversion to on-spec- ification HSD or naphtha, with significantly lower energy intensity and minimal yield loss. Fast-track remedy To overcome these bespoke issues, the refinery imple - mented a fast-track remedy: routing the HGO from the
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Atmospheric distillation unit
Vacuum distillation unit
O -gas
Overhead
LVGO
LN /HN fractionation
LVGO pumparound
Light naphtha
HN /ATF fractionation
Diesel pool
Steam
LVGO / MVGO fractionation
Heavy naphtha
MVGO
ATF pumparound
Steam
ATF / HK fractionation
LDO / MVGO fractionation
Aviation turbine fuel
VGO pool
HK pumparound
HVGO pumparound
HVGO
Steam
Heavy kero
LGO pumparound
Wash section
Diesel pool
SLOP wax
LGO / HGO fractionation
Steam
Light gas o il
Heavy gasoil reprocessed at vacuum column
Crude
Steam
Flash zone
Stripping section
Stripping section
Steam
Steam
Atmospheric residue
Vacuum residue
Figure 1 AVU flow diagram – before
CDU to the downstream vacuum column via the heavy vac- uum gasoil (HVGO) pumparound (PA) return stream (see Figure 1 ). It allows recovery of its light portion with the vacuum diesel cut, taken from the top of the vacuum column. The heavy portion is recovered with the HVGO, taken from the middle section of the vacuum column. This practice, while minimising the negative impacts on the downstream unit, introduced unintended systematic penalties at the vacuum tower: • Increased vapour loading to the three beds above the HVGO PA of the vacuum tower. • The light vacuum gasoil (LVGO) PA got overloaded and was close to incipient flooding. • Increased gasoil entrainment to the vacuum distillation unit (VDU) overhead hot-well, placing additional load on the vacuum system. • The desired degree of vacuum could not be achieved anymore. • Poor fractionation between vacuum diesel and medium vacuum gasoil (MVGO). • The pressure drops between the top of the tower and the flash zone increased consistently up to 35 mmHg. • The pressure at the flash-zone increased consistently as well. • The HVGO yield was drastically reduced: it could not be compensated by increasing the coil outlet temperature (COT) of the heater.
• Proportionally to the reduced HVGO recovery, the yield of the vacuum residue (VR) increased. • The front end of the VR became lighter than required. • To meet the required viscosity and penetration of the VR and make it suitable for the downstream bitumen plant, the furnace COT had to be increased up to 410°C, with conse- quent increased operational severity and required energy. These effects collectively constrained crude throughput, increased energy intensity, and imposed additional stress on downstream conversion units. The AVU performance was jointly studied to identify unit bottlenecks and the required modifications to achieve desired distillates yields and qualities while minimising the needed energy. Objectives The refinery established clear operational and commercial objectives to improve performance: u Reduce SRGO slippage into the VGO pool and maximise distillate recovery by: • Improving fractionation between LGO and HGO; before revamp, the overlap was higher than 60°C. • Rerouting HGO directly to the VGO pool, bypassing the vacuum column, by making the front-end of the HGO distil- lation curve heavier, so it is suitable for the VGO pool. v Stabilise VR quality across diverse crude blends (i.e., VG-30 and VG-40 grade bitumen production in high-sul - phur crude). w Reduce VR yield to improve crude-to-distillate conversion.
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Steam
Heavy kero
LGO pumparound
Hold down grid
LGO / HGO fractionation MellapakPlus
Steam
Hold down bar
Light gasoil
Steam
Tie rod
Flash zone
Heavy gasoil
Stripping section
Support grid
Steam
Atmospheric residue
Figure 2 HGO section modification
Support bar
x Improve the VGO tail-end distillation curve for a superior quality of feed to the downstream units. y Reduce flash zone pressure by lowering the pressure drop in the vacuum column. z Integrity issues in the vacuum column: • The HVGO PA packing and the wash bed were found dis - lodged during every turnaround. • Rapid vapour-traffic shifts during block-mode crude changes (low-sulphur ↔ high-sulphur) caused mechanical disturbances. • Loss of bed integrity degraded mass-transfer efficiency, worsening VGO-VR overlap.
Implemented solutions (2024 turnaround) IOCL has provided test run data to build the process sim - ulation model of the existing running units. A deep pro - cess and hydraulic study was performed jointly by Sulzer and IOCL to identify the bottlenecks in both columns and develop remedies to achieve the goals. A focused column Figure 3 Post-revamp HVGO wash section sandwiched bed design with tie rods
Atmospheric distillation unit
Vacuum distillation unit
O -gas
Overhead
LVGO to diesel pool
LVGO PA Mellapak
LN /HN fractionation
Light naphtha
HN /ATF fractionation
Steam
LVGO / MVGO fractionation MellapakPlus
Heavy naphtha
MVGO
ATF pumparound
MVGO / HVGO fractionation bed removed
Steam
ATF / HK fractionation
Aviation turbine fuel
VGO pool
HK pumparound
HVGO PA packing reused
HVGO
Steam
Heavy kero
LGO pump a round
Wash section MellapakPlus Mellagrid AF
Diesel pool
SLOP wax
LGO / HGO fractionation MellapakPlus
Steam
Light gasoil
Heavy gasoil routed directly to VGO pool
Flash zone
Steam
Crude
Stripping section
Stripping section
Steam
Steam
Atmospheric residue
Vacuum residue
Figure 4 AVU flow diagram – post-revamp
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Vacuum furnace COT (˚C)
VGO combined pool 5% (˚C)
420
250 275 300 325 375 350 400
368 372
408 407 408 407 408 408 408 408 407 409
359
407
350 353
350
343 338 340
334 339 332
401 400
325
400
380
Figure 5 VGO 5% (ASTM D1160)
Figure 6 Lower fuel consumption and CO₂ emissions
internals revamp was executed to restore mass-transfer efficiency and reduce energy losses. CDU main fractionator column (11C1) The existing conventional structured packing and liquid dis - tributor were replaced with proprietary High Performance Mellapak Plus packing, along with the liquid distributor in the LGO-HGO section, as shown in Figure 2 . This modi - fication substantially increases fractionation efficiency to reroute heavy gasoil directly to the VGO pool, bypassing the vacuum column. VDU column (12C1) The LVGO pumparound section was upgraded by replacing the existing conventional random packing with Mellapak structured packing and spray distributor, reducing the excessive pressure drop in the LVGO PA and minimising gasoil entrainment to the overhead hot-well. The MVGO/HVGO fractionation bed was removed to reduce pressure drop. This bed became useless now that the diesel is recovered at the CDU main fractionator. The wash-bed packing was replaced with Mellapak Plus and Mellagrid Anti Fouling (MG40 AF) in a sandwich configu - ration. Tie-rods were installed across the wash bed to pre - vent dislodgement during vapour-traffic surges, as shown in Figure 3 . The revamped PFD of the unit (see Figure 4 ) shows all the upgraded sections of the main fractionator and vacuum column, as well as the rerouting of the HCO and MVGO to the VGO pool. Results and benefits The revamp delivered significant gains in separation effi - ciency, product quality, and energy savings: u Distillate recovery: • Incremental increase in diesel yield Y-o-Y by approximately 2 wt%, providing direct refinery margin improvement. v Product separation: • LGO-HGO overlap reduced from 60°C to 30°C (ASTM D86), delivering sharper CDU separation. • HGO recovery at 360°C (ASTM D86) reduced from 45 vol% to ~15 vol%, making HGO heavier and routed to the VGO pool on a consistent basis. This directly indicates a reduction in diesel dropping into the VGO pool. • HVGO 5% (ASTM D1160) and VGO 5% (ASTM D1160)
increased by 20°C, indicating reduced loss of gasoil in VGO, for superior feed quality to the downstream units. • VGO 5% (ASTM D1160, as shown in Figure 5 ). • HVGO 95% (ASTM D1160) and VGO 95% (ASTM D1160) increased, respectively, by 12°C and 10°C, indicat - ing a deeper cut point at the vacuum tower and additional feed for the FCC unit/OHCU. w VR quality and yield: • VR yield reduced Y-o-Y by approximately 2 wt%, increas - ing crude conversion. • Achieved stable VR viscosity and penetration across crude blends, ensuring reliable bitumen production. x Crude throughput flexibility: • Lower VR yield allowed for the debottlenecking of the downstream processing units, enabling the refinery to have a higher and more flexible crude processing slate. y Energy savings: • Vacuum column pressure drop reduced from 35 mmHg to 25 mmHg. • VDU furnace COT reduced by approximately 8°C (from 408°C to 400°C), lowering fuel consumption and CO₂ emis - sions (see Figure 6 ). z Reliability enhancement: • New wash-bed configuration with tie-rod reinforcement eliminated recurring packing dislodgement. • Improved mechanical stability ensured sustained mass-transfer efficiency under varying feed conditions. Conclusion The IOCL Mathura refinery and Sulzer performed a sys - tematic analysis that combined plant-match diagnostics, hydraulic review, mass-transfer evaluation, and strategic hardware upgrades, resulting in noteworthy yield, quality, and energy improvements in the AVU. The 2024 revamp not only enhanced distillate recovery and VR stability but also optimised the operating window, reduced energy demand, and alleviated load on down - stream units. This case demonstrates how targeted inter - nal revamps can unlock significant performance gains in legacy crude-distillation units, especially under diverse and dynamic crude-processing environments. Job execution Sulzer Tower Field Services executed all site activities on a turnkey basis within the timeframe of the planned
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turnaround; the job was safely completed in 22 days against the planned span of 25 days.
calculations, sizing of column internals, commissioning, revamps, and troubleshooting of distillation columns for enhancing plant per- formance and reliability. Bhagwat holds a BE in chemical engineering from University of Pune. Amar Patil is Deputy Manager, Application & Technology at Sulzer India Pvt. Ltd. He specialises in digital twins, operator training simu - lators, and refinery process simulations, supporting troubleshooting, optimisation, and revamp of refining units and contributes to simula - tion-driven solutions for enhancing plant performance and reliability. Patil holds an MTech in chemical engineering from ICT, Mumbai. Mukesh Gupta is Deputy General Manager, Sales, at Sulzer India Pvt. Ltd. He is a sales professional with more than 30 years of extensive experience in technical sales across the process industry, including oil and gas, chemicals, and fertilisers. He has expertise in managing key accounts with hands-on experience in process industry applications and industrial equipment and solutions.
Literature cited 1 Mosca, G, Ramani, P, Dharmaraj, G, Retrofitting a vacuum tower and the heater’s transfer line to maximise throughput and operational reli - ability , IRPC Europe Conference June 2018. 2 Mosca, G, Arora, A, Squeezing the bottom of the barrel while improving operational reliability , BBTC ME&A Technology Conference October 2016, Manama, Bahrain. Rahul Srivastava is General Manager (Technical Services) at Indian Oil Corporation Ltd. (IOCL), Mathura refinery. In this role, he oversees technical evaluations, process optimisations, and complete technical services department to ensure maximum operational efficiency and reliability. Srivastava holds a B.Tech in chemical engineering from Jadavpur University Kolkata.
Manish Rohit is Chief Manager (Technical Services) at IOCL, Mathura refinery. He is responsible for managing day-to-day technical service operations, troubleshooting process issues, driv- ing energy conservation initiatives, and implementing refinery modernisation projects. Rohit holds a B.Tech in chemical engineering from The Maharaja Sayajirao University of Baroda. Harshit Budholiya is Process Manager at IOCL, Mathura refinery. His responsi - bilities include monitoring process units, troubleshooting process issues, driving energy conservation initiatives, and ensuring strict compliance with safety and environmental standards. Budholiya holds a BE (Hons) in chemical engineer - ing from Birla Institute of Technology and Science – Pilani (BITS-Pilani). Giuseppe Mosca is Global Technical Expert and Head of Engineering Solutions at Sulzer Chemtech. He is involved with process simulations, revamping proposals, troubleshooting, commissioning, and start-up assistance of fractionation equipment. He is the author and co-author of more than 50 papers published at international confer- ences and in magazines. Mosca received a MS degree in chemical engineering from the University La Sapienza in Rome, Italy, and is a registered professional engineer of the district of L’Aquila, Italy, and a senior member of AIChE. Unmesh Bhagwat is Senior Manager, Application & Technology at Sulzer India Pvt. Ltd. He contributes as the Refinery Regional Application Manager for India and Middle East. His core expertise includes process simulation, hydraulic
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