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3 LNG accelerating the low-carbon energy transition Rene Gonzalez

5 Dryout design considerations for cryogenic gas plants: Part 1 Scott A Miller, David A Jelf, J A Anguiano and Joe T Lynch Honeywell UOP 11 Optimise sulphur recovery plant emissions during unit upset conditions Jochen Geiger, Michael Gaura and Anantha Kukkuvada AMETEK Process Instruments

17 ‘On the fly’ vs high-performance H₂S selective solvent Ashraf Abufaris BASF Middle East Chemicals LLC Blake Morell BASF Corporation

21 Steam methane reformer tube lifecycle improvement best practices Richard D Roberts and Grant Jacobson Becht 28 Testing and effects of pipeline chemicals on gas processing facilities David Engel and Scott Williams Nexo Solutions 32 Hydrogen recovery from refinery off-gas – Part 1: An overview Zach Foss Divigas

Cover From LNG liquefaction facility projects to advances in hydrogen production technology, bankable investments reflect the gas industry’s role in the transition to net zero emissions by 2050. Photo courtesy of Joshua Sun

©2024. 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 and its supplements the publisher cannot be held responsible for any statements, opinions or views or for any inaccuracies.

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Editor Rene Gonzalez editor@petroleumtechnology.com tel: +1 713 449 5817 Managing Editor Rachel Storry rachel.storry@emap.com Editorial Assistant Lisa Harrison lisa.harrison@emap.com Graphics Peter Harper Business Development Director Paul Mason sales@petroleumtechnology.com tel: +44 7841 699431 Managing Director Richard Watts richard.watts@emap.com Circulation Fran Havard circulation@petroleumtechnology.com ptq PETROLEUM TECHNOLOGY QUARTERLY

LNG accelerating the low-carbon energy transition D emand for liquefied natural gas (LNG) is expected to continue beyond 2040. China has overtaken Japan to become the world’s largest LNG importer, while demand from other Asian countries is also growing, even as domestic gas production in some of these countries is decreasing. Significant infrastructure projects are required to be able to import LNG from the Middle East and other exporting countries. High capital investments needed for regasification terminals do not guarantee secure supply when considering disruptions due to geopolitical conflicts. All this must be factored into the industry’s long-term planning for using LNG to provide grid stability while facilitating a higher share of renewables in the market. In the medium term, potential demand, such as from India, is set to consume new supply coming onto the market by 2025. The question is, who are the exporters that growing markets can depend on, especially when committing to long-term con- tracts? The big winners are the dominant exporters from the Middle East and Russia. This begs the question that as the top global natural gas producer and exporter, the US’s exports of LNG play a vital role in supporting global energy security. Following Russia’s invasion of Ukraine, the US has become the primary supplier of LNG to Europe. However, to what extent will the perceived flexibility and reliability of US LNG in global markets last? With the current US Administration enforcing a moratorium on planned LNG exports to non-Free Trade Agreement countries, compelling postponement or can- cellation of new liquefaction projects, Russia stands to win big by not only influenc - ing Europe and China’s LNG market but elsewhere throughout the globe. Against this backdrop, the US was again the largest supplier of LNG to Europe in 2023, accounting for nearly half of total LNG imports, according to data from Cedigaz. In parallel, Southeast Asia is increasing LNG imports to backfill domestic gas declines. The US EIA Short-Term Energy Outlook (6 February) expected the US benchmark Henry Hub natural gas spot price to average higher in 2024 and 2025 than in 2023 but remain lower than $3.00 per MMBtu. The forecast seems to be based on increases in natural gas prices as demand grows faster than supply in 2024. Developers advanced several LNG projects to the construction phase in 2023, following the signing of sale and purchase agreements (SPA), contracts specifying the terms and conditions of LNG supplies between seller and buyer, underpinning the projects. Developers signed contracts with buyers for almost 22 million metric tons per year of LNG last year, or about 3 Bcf/d of natural gas, according to data from the US Department of Energy (DOE) and company websites. However, the volumes con- tracted in 2023 totalled 52% less than the contracted volumes in SPAs signed in 2022. On 16 January 2024, a record high of 141.5 billion cubic feet (Bcf) of natural gas was consumed in the US Lower 48 states due to the ‘Polar Vortex’, exceeding the previous record set on 23 December 2022, according to estimates from S&P Global Commodity Insights. LNG’s dominance in the transition to net zero emissions and energy security will be enhanced with a range of technologies incorporating advanced analytics, the Internet of Things, and artificial intelligence to enhance efficiency, safety, overall operational performance, and ultimately capacity, including increased ammonia and hydrogen production, such as blue hydrogen. For example, blue hydrogen’s future, as discussed in this issue of Gas 2024 , is expected to become clearer by the end of the year, as the US DOE awards the first billion-dollar subsidies for demonstration projects and the Treasury Department issues guidance for companies that want to claim new tax credits for low-carbon hydrogen production.

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PTQ (Petroleum Technology Quarterly) (ISSN No: 1632-363X, USPS No: 014-781) is published quarterly plus annual Catalysis edition by EMAP and is distributed in the US by SP/Asendia, 17B South Middlesex Avenue, Monroe NJ 08831. Periodicals postage paid at New Brunswick, NJ. Postmaster: send address changes to PTQ (Petroleum Technology Quarterly), 17B South Middlesex Avenue, Monroe NJ 08831. Back numbers available from the Publisher at $30 per copy inc postage.

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Dryout design considerations for cryogenic gas plants: Part 1

A well-conducted cold plant dryout will help plant commissioning and ensure start-up goes smoothly and does not last any longer than necessary

Scott A Miller, David A Jelf, J A Anguiano and Joe T Lynch Honeywell UOP

A cryogenic plant dryout is a critical step during start-up, but it often does not receive proper atten- tion early in the project that would allow successful execution. For this discussion, ‘cold plant’ refers to a cry- ogenic turboexpander plant for natural gas liquids (NGLs) recovery, which recovers liquids comprised of either ethane and heavier components (often referred to as C₂+ liquids or NGLs) or propane and heavier components (often referred to as C₃+ liquids or liquefied petroleum gas [LPG]). A cold plant dryout can be executed correctly the first time. The owner/operator can be confident that the cold plant is dry prior to cooling down when proper design features are implemented, and guidelines are followed. All gas processors know that water must be removed from the cold plant. However, knowing the best method for remov- ing water; how much water has been removed during the dryout period; and when dryout is complete, can be chal- lenging. Far too often, a dryout is stopped before the cold plant is completely dry. Hydrate problems The solid that forms when water is present in a hydrocar- bon stream is not ‘ice’, but a crystalline structure known as a hydrate. Hydrates can form at conditions where solids would not be expected and will form above the freezing point of water. They are a physical combination of water and other chemical constituents, like those found in natural gas processing, which have an ‘ice-like’ appearance. 1 Hydrates form when enough water is present at the right combination of temperature and pressure and tend to favour systems with low temperature and high pressure. 2 For gas plant owners/operators, this means that when the plant begins to cool down, hydrates will form in process areas where sufficient water is present, restricting or com - pletely obstructing process flow. It is not uncommon to see hydrates obstructing flow through heat exchanger passes or the strainers that pro- tect the heat exchanger from construction dirt and debris. Hydrates can cause enough pressure drop to rip apart strainers, allowing dirt and debris to enter and damage the downstream heat exchanger. In the case of thermosyphon reboilers and side heaters, hydrate formation may restrict the flow through the exchanger and reduce the amount of

heat input to the column enough to prevent achieving the bottoms liquid product purity specification. Hydrates may also form on the cold plant fractiona- tion column trays and packing. The result is a decrease in efficiency, causing low product recovery and potentially off-specification liquid product. Hydrates are also known to plug control valves and plant instrumentation. Water can enter the plant equipment through rain or con- densation from open piping during construction and through water left after hydrostatic testing. The single most impor- tant step that can be taken prior to the start-up of a cold process plant is to drain and blow out as much free water as possible from the piping and equipment. Eventually, all the water must be removed from the cold plant to the parts per million (ppm) level for the cold plant to operate safely and efficiently. Not all of the water can be removed simply by draining low points. The remaining water must be removed by a combination of moving the water to a low point where it can be drained and absorbing the water in a vapour stream so it can be removed from the cold plant equipment and pip- ing. Several options for eliminating this remaining water are presented in the next section. Dryout options The following options are common approaches to drying out a cold plant prior to cooldown. A description of each option is given, discussing its advantages and disadvan- tages, in order from the least cost-effective to the most cost-effective option. Option 1 involves pressure cycling with nitrogen. In this approach, sections of the cold plant are isolated to be pres- surised and depressurised multiple times using nitrogen. This method requires no piping design considerations other than properly locating low-point drains and ensuring the drains are of sufficient number and size to remove water from the system. This dryout approach requires a large quantity of nitro- gen to be available and can be very expensive because of the amount of nitrogen consumed. This dryout option is less likely to be successful if large quantities of free water (puddles of water) are still present in the system. It is more difficult to determine the amount of water remaining in the cold plant after nitrogen purging and whether all water has

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Residue gas compressor

Air cooler

To sales

To sales , reinjection, or are

Cold plant

Inlet gas

Cold plant

Inlet gas

Dust lters

Dehydrators Dryout f low path

Liquid product

Dehydrators

O-line

Liquid product

Dryout fl ow path

O-line

Figure 1 Once-through dryout

Figure 2 Closed-loop recirculation dryout

been removed from the system compared to the other dry- out approaches. Water content readings must be taken at many more locations to get an accurate assessment of the amount of water remaining in the system. Option 2, the once-through dryout approach flows warm, dehydrated inlet gas through the cold plant, equip- ment, and then to the flare stack, reinjection, or a sales gas pipeline. The dryout path is operated at as low a pressure as possible. The pressure drop through the cold plant is minimised to prevent any Joule-Thomson (J-T) expansion that would cool down the process while drying the plant. The flow rate should be maintained to move any free water to the low-point drains, or to absorb the water in the vapour stream and remove it from the process. A pressure-reduc- ing device (such as a temporary flow orifice or valve) must be included to take the pressure drop upstream of the cold section of the plant. Figure 1 illustrates the main process flow path for this dryout method. Keep in mind that the dryout flow rate may be limited by the flare system’s tolerance for flaring or reinjection system capacity. If the wet gas is sent to a sales gas pipeline, the gas water content should be monitored to ensure it remains below the maximum amount specified. The once-through dryout is an effective approach and has been used on many projects. However, the dryout flow rate can be limited by the flare system or reinjection capacity. For this scenario, the dryout period will most likely be longer in order to remove all the water in the cold plant. Option 3 , the closed-loop re-circulation approach, recir- culates warm, dehydrated gas in a closed loop through the cold plant back to the dehydration system inlet using a residue gas compressor. The dryout loop is operated at as low a pressure as possible without shutting down the res- idue compressor. The pressure drop is minimised through the cold plant to prevent any J-T expansion that would cool down the process while drying the plant. Again, the goal is to minimise pressure drop through the cold plant but maintain a high enough flow rate to ‘sweep’ free water to low-point drains or carry the water away in the gas to be removed by the front-end dehydrators. A recirculation line is required that connects the residue gas line downstream of the residue gas compressors to the inlet gas piping upstream of the dehydrators to make the ‘closed loop’. A pressure-reducing device (such as a temporary orifice

or valve) must be included in the dryout design to take the pressure drop upstream of the cold plant. Figure 2 illustrates the main process flow path for this dryout method. The closed-loop recirculation approach is our recom- mended approach, as it achieves a proper dryout in the shortest time. This dryout option removes both free water and ambient condensation. It does so without excessive amounts of nitrogen, as with nitrogen pressure cycling, or excess flaring, as with a once-through dryout with dehy - drated inlet gas. Nitrogen cycling is only effective at remov- ing ambient condensation within the cold plant and does a poor job of moving free water to the low points where it can be removed from the system. The once-through dryout option must have a location to discharge the wet dryout gas after passing through the cold plant. Monitoring dryout progress using closed-loop recircula- tion is relatively straightforward. Once dryout is complete, transitioning to cooldown can be done quickly with min- imal effort simply by shifting the pressure drop from the dryout pressure-reducing device to the J-T valve. If desired, cooldown can commence using the dryout recirculation gas prior to introducing fresh inlet gas. A well-executed dryout begins during the detailed design phase by determining the dryout features required for a spe- cific cold plant design. A dryout procedure should be prepared prior to dryout, including any cold plant design limitations as well as the specific steps to follow for removing water and monitoring the cold plant water content during dryout. Cold plant design features: Executing dryout Several design features should be considered during the detailed design phase for the cold plant to be included as part of the cold plant design package. Although some of the features listed in this section are required specifically for dry - out, many will be used for alternative purposes as well (such as depressurisation and venting hydrocarbons from process equipment prior to maintenance or repair, isolating equip- ment, and monitoring the dehydrator outlet water content). These features are: • Recirculation piping • Pressure-reducing device and location • Drain valves • Provisions for stagnant process areas • Location to monitor cold plant water content.

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Residue gas compressor

Air cooler

Residue gas compressor

Air cooler

To sales

Location 1

Location 2

Inlet gas

Cold plant

Recirculation piping

Dust lters

Dehydrators

Liquid product

Dryout f low path

O-line

Dryout f low path

O-line

Figure 3 Recirculation piping

Figure 4 Possible locations of pressure-reducing devices for dryout

Recirculation piping A recirculation line connecting the residue gas line down- stream from the residue gas compressor(s) to the inlet gas piping upstream of the dehydrators should be included to complete the closed loop. For designs that include a cold plant bypass, the plant bypass can also be designed to function as the recirculation line during dryout. Figure 3 shows the typical recirculation piping connection points. Pressure-reducing device and location The location of the pressure-reducing device to take the plant pressure drop upstream of the cold plant is dependent on the dehydrator regeneration system design and source of regeneration gas. When selecting the pressure-reducing device location, the cold plant designer must be mindful of the location cho- sen for dehydrator regeneration gas supply and return to ensure the regeneration compressor has enough compres- sion head to return the wet regeneration gas to the plant. For example, in plant designs that use inlet gas for regen- eration, the pressure-reducing device must not be located between the dehydrators and the piping take-off for the regeneration gas supply. The regeneration gas compres- sor will not be able to overcome the pressure difference. Figure 4 shows possible locations to consider. Typically, a temporary flow orifice, manual valve, or auto - mated valve is used as the pressure-reducing device. For example, if a temporary orifice is to be installed at Location 2 in Figure 4, the temporary orifice can be installed down - stream of one of the molecular sieve dust filters inside its isolation valves (see Figure 5 ). Typically, two full-flow dust filters (one filter in service and one spare filter) are included as part of the dehydrator design package, allowing easy orifice installation and removal. The orifice plate should be of suffi - cient thickness to handle the pressure drop, which in many cases will exceed 500 pounds per square inch (psi) or 35 bar. The pressure-reducing device should be sized to take the full plant pressure drop at the maximum dryout flow rate. The maximum dryout flow rate is set by the maximum flow the cold plant J-T valve can pass when fully open with a pressure drop of ~30 psi (2 bar). The dryout gas flow rate will end up being approxi - mately 30% of the cold plant nameplate capacity without

introducing too much pressure drop across the fully open J-T valve. Typically, even a single full plant rate residue compressor can be operated on its surge control line at reduced speed to provide a pressure rise of less than the normal operating pressure rise. The pressure-reducing device can then be sized for the expected pressures when running the compressor at reduced speed on its surge line. The surge control valve will then remain in control in parallel with the dryout loop flow for the duration of the dryout procedure. If multiple machines are used, it may be possible to use one machine with its surge control valve closed during dryout. The closed-loop operating pressure on the residue com- pressor suction side must be high enough to clear the low-pressure shutdowns for the compressor. The discharge pressure will be determined by the permissible operating point for the compressor. Drain valves During the model review, identify all low points around major process equipment and add drain valves as required. Hard pipe all low-point drains to a closed-drain system. Design the low-point drains to be easily accessible. The low-point drains will be used periodically to drain any col- lected free water from the system. A one-inch drain is the minimum size recommended. There must not be any inter- nal projection at the drain connection. The turboexpander and booster compressor low points are always inside the machines’ isolation valves and, there- fore, are not available to drain free water during dryout. As a result, the lowest spot around both machines may move

Temporary orifice in place of b lind f lange sp acer

Dust lter

To c old plant

Figure 5 Detail for temporary orifice

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Subcooler

Condenser

Residue gas to compression

Booster

R e f lux accum.

Residue gas to compression

Expander

Gas/gas

Booster

Cold sep.

Inlet gas

Expander

J–T valve

Reux pump

Demethaniser

Gas/gas

Cold sep.

Inlet gas

Reboiler passes

J–T valve

Deethaniser

Stagnant piping and equipment

Liquid product

Reboiler

Figure 6 Ethane recovery plant (main dryout path and stagnant areas)

Stagnant piping and equipment

Liquid product

just outside the isolation points during dryout. In this situa- tion, ensure low-point drains are installed at these locations to drain any free water that may collect during dryout. Provisions for stagnant process areas The following process flow diagrams (PFD) are examples of two typical cold plant designs: one designed for ethane recovery (see Figure 6 ) and one designed for propane recovery (see Figure 7 ). Each PFD identifies major pro - cess equipment and control valves and highlights the main recirculation flow path through the cold plant as well as the stagnant flow areas during dryout. Several piping and equipment loops and the lower sec- tion of the fractionation column below the expander outlet feed are stagnant (no dryout flow). Unless a flow is intro - duced, these areas of the cold plant will remain wet. Design provisions must be included to introduce flow into each stagnant area. A source of dryout gas is located downstream of the dehydrator dust filters (also downstream of the pressure- reducing device used for dryout), but upstream of the cold plant should be selected to provide dryout flow to each stagnant area. As an example, for a 200 million standard cubic feet per day (MMSCFD) or 5.4 x 106 normal cubic metres per hour (Nm³/hr) cold plant, a 4in minimum pipe size is recommended from the dryout gas source to provide adequate flow of dryout gas to multiple stagnant areas at once. As the cold plant nameplate capacity increases, scale up the dryout gas source pipe size accordingly to ensure adequate flow at the recommended 30 psi allowable J-T valve pressure drop during dryout. Part 2: Closed loop dryout method and monitoring UOP Ortloff has witnessed many cold plant dryouts dur - ing commissioning and initial start-up over the years. A well-conducted cold plant dryout will help plant com- missioning and ensure start-up goes smoothly and does not last longer than necessary. There are multiple dryout options which can remove water from the system, as dis- cussed in Part 1 of this two-part article. However, there are

Figure 7 Propane recovery plant (main dryout path and stagnant areas)

drawbacks to some of them. The closed-loop recirculation dryout effectively removes water from the system while easily monitoring dryout progress This allows for a quick transition to plant cooldown after dryout is completed. Some of the design challenges, the method for implement- ing a closed-loop recirculation dryout, and monitoring pro- gress will be discussed in further detail in PTQ Q2 2024. References 1 Engineering Data Book , 12th ed., Gas Processors Suppliers Association, Tulsa, OK, 2004, Section 20, pp.9. 2 Carrol J, An Introduction to Gas Hydrates , 2008, [Online] Retrieved May 13, 2010, www.telusplanet.net/public/jcarroll/HYDR.HTM. 3 Engineering Data Book , 12th ed., Gas Processors Suppliers Association, Tulsa, OK, 2004, Section 20, pp.19. 4 Engineering Data Book , 12th ed., Gas Processors Suppliers Association, Tulsa, OK, 2004, Section 20, pp.23. Scott A Miller is a Principal Process Engineer with Honeywell UOP, supporting Ortloff Cryogenic Gas Processing Technologies. He has more than 23 years of experience in the petrochemical and gas pro - cessing industries, of which 15 are in the design and operation of NGL/ LPG recovery plants. Email: scott.a.miller@honeywell.com David A Jelf is a Principal Instrument and Controls Engineer with Honeywell UOP, supporting Ortloff Cryogenic Gas Processing Technologies. He has more than 40 years of experience in the design and operation of NGL/LPG recovery plants. J A Anguiano is a former Engineering Manager with Honeywell UOP, previously supporting Ortloff Cryogenic Gas Processing Technologies. He has more than 25 years of experience in the design and operation of NGL/LPG recovery plants. Joe T Lynch is a retired Engineering and Licensing Manager from Honeywell UOP, previously supporting Ortloff Cryogenic Gas Processing Technologies. He has more than 45 years of experience in the design and operation of NGL/LPG recovery plants.

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Optimise sulphur recovery plant emissions during unit upset conditions

Knowledge, understanding, and awareness training are essential to maintain process instrumentation for optimal information gathering

Jochen Geiger, Michael Gaura and Anantha Kukkuvada AMETEK Process Instruments

W orking conditions for refineries are changing, and such changes might be challenging. In several geographical areas, the historical supplier of crude oil has changed, resulting in new crude oil compositions. The method of supply may also have changed from pipe - lines, with stable composition, towards ocean vessel sup - ply, with different compositions. Regardless, final product quality from refineries needs to be stable, and quantities need to be ever increasing. At the same time, environmental impacts are being more strictly regulated, increasing monitoring and reporting require - ments. Emissions always need to be reduced. In summary, more flexibility from refiners is required. All of this puts more load and attention on sulphur recov - ery units (SRUs), which are expected to operate continu - ously and in an optimised manner. As discussed in several papers, such SRUs require some special attention with regard to safety and operational monitoring. On its own, elemental sulphur is not an issue. It is one of the most common elements (by mass) on earth and is essential to life. Sulphur is present in living organisms, including humans, and can be used in fertiliser production. It

even has medicinal applications. However, sulphur can also be harmful to both humans and the environment. Hydrogen sulphide (H₂S) concentrations of just a few parts per million (ppm) in ambient air can lead to illness and death at sus - tained levels greater than 100 ppm. Sulphur dioxide (SO₂) is also toxic to humans at very low concentrations but has received significant attention based on its negative impacts on the environment. Specifically, it can destroy vegetation and wildlife and contribute to the production of acid rain. The basic chemistry of sulphur recovery has been known for more than 90 years. In today’s world, the capacity of modern SRUs can range from tens to thousands of tons of sulphur production per day. The most significant improve - ments have been made to the overall recovery efficiency. Although it was sufficient to operate at 80-85% in the 1970s and between 95% and 99% in the 1990s, the cur - rent requirement is to operate at recovery efficiency rates of 99.9+%. Considering that we rely on the same chemistry as 90 years ago, except for tail gas clean-up units, such efficiency improvements only became possible using reliable process instrumentation (see Figure 1 ).

Stack gas emissions monitoring SO No (mass ow) x

Claus sulphur recovery unit

AT7

Tail gas treating unit

Absorber outlet H HS (COS)

Quench column o-gas H HS

Acid gas feed forward Total hydrocarbons HS NH CO

AT5

AT6

Claus tail gas HS SO (COS CS)

Temperature

AT1

Reduction gas generator

AT3

Cobalt molybdenum reduction reactor

Air

Converters & condensers

Thermal oxidiser

Reactor furnace & Waste heat exchanger

AT4

Quench column

Absorber & stripper columns

Reduction reactor o-gas H SO

Sulphur pit sweep gas HS SO

AT2

Sulphur pit

Figure 1 Claus SRU

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be caused by rapid changes in hydrocarbons in the feed gas. We can observe such changes in both streams, acid and SWS gas. Hydrocarbons need a significantly higher amount of oxygen (O₂) than H₂S to react in the Claus fur - nace, as shown in Table 1 . Depending on SRU design, the sudden occurrence (or sudden disappearance) of hydrocarbons can cause opera- tional issues. The potential impact can even affect bypass recovery processing steps with potentially drastic conse- quences. In the case where only the tail gas clean-up unit gets bypassed, the overall sulphur dioxide (SO₂) emissions could potentially increase by a factor of 100. Normal SO₂ concentrations after final incineration, in combination with an amine-based tail gas treatment unit (TGTU), are 200- 400 ppmv SO₂, which will jump to 10,000/20,000 ppmv in cases of TGTU bypass. Even when not approaching such dramatic process con- ditions, it can be said that a closed feed gas control loop can reduce emissions by 10-15% of the overall emission rate. Flaring events can be reduced by 65-95%. 1 Considering that sources of hydrocarbon upsets are caused upstream from the SRU, another new challenge will be caused by fre- quent changes in crude composition, with most refineries not designed for such changes in crude composition. A reliable understanding of the feed gas composition (H₂S, NH₃, THC, CO₂) will make process control more effec - tive. Fast responses to compositional changes will allow a fast reaction on the acid gas-to-air ratio, resulting in an overall emissions reduction. The AT2 sulphur pit is a safety measurement where H₂S must be maintained below the explosive limit (lower explo- sive limit [LEL] 3.25 vol%). SO₂ should be measured to pro - vide an early warning of sulphur fire. A smouldering sulphur fire is not uncommon; exposed iron pipe reacts with sulphur from pyrites, a slow but exo- thermic reaction that can result in an ignition source. For example, with H₂S as the fuel source and pyrophoric fire as the ignition source, waiting only for oxygen will make the fire triangulate. At the AT3 tail gas analyser, also called the air demand analyser, the entire process control of the modified Claus process is guided by this measurement. The leading com- ponents to be measured here are H₂S and SO₂. Their components can be of interest but do not require process control. Taking a step back, what does the process control

Ratio of O₂ needed per mole HC compared to moles of H2S

Compound moles O₂ per mole HC

Ratio of O₂ needed per mole HC compared to mole of H₂S

Methane

4 7

Ethane

3.5

Propane Butane Pentane Hexane

10 13 16 19

6.5

9.5

Table 1

The challenges are: • Understanding of the environmental impact of each sin- gle instrument. • Knowing the potential improvement of using the best instrument combination. • Keeping the instruments in operational conditions. • Mitigating upset conditions by understanding ‘unex- pected’ instrument behaviours. The following discussion will investigate each sample point and review from a control aspect and the potential environmental impact. When considering AT1 feed gas to the Claus reaction furnace, also known as the feed gas analyser (refer to AT1 in Figure 1), the bespoke feed gas (in refineries) can consist of two different source streams. The first is the acid gas stream, and the second is the sour water stripper (SWS) gas. While acid gas should be more stable, the SWS is known to be a potential source of upsets. Against this backdrop, we should not fall into the belief that no conditions can cause upset conditions for SRUs originating from acid gas. Refinery operating conditions Refineries need to operate their SRUs with both acid and SWS gas. This might not be the case for every single unit (train), but as a general operating condition, it can be taken as given. One of the problems of SWS gas comes from ammonia (NH₃) as part of the SWS stream composi - tion, requiring operation of the Claus reaction furnace at a higher temperature (<1,200°C) in order react and destroy the ammonia before entering the heat exchanger. The key aspect of any feed gas measurement and control is to mitigate upset conditions. Such upset conditions can

Tail gas

Claus beds & condensers

Burner

Reaction furnace

Acid gas

To tail gas clean-up unit or incinerator

AT- 301

FT- 201

Main air

Trim air

FV-101

FV-102

Sulphur

FT- 101

FT- 102

Combustion air

Figure 2 Feedback plus ratio control

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clean-up units are therefore required. The most installed these days are amine-based TGTUs. The purpose of the TGTU is to convert all remaining sulphur components carried over from the modified Claus unit into H₂S. This happens in the catalytic section of the TGTU process. The reduction reactor utilises a cobalt-mo - lybdenum catalyst (also mentioned as CoMo bed). The cobalt-molybdenum catalysed reactions are shown in Figure 3 . The first step of the reaction does require hydrogen (H₂). H2 can be present as a byproduct of the modified Claus reaction but can also be generated by an inline reduction burner or provided from an external source. To ensure that the ‘sulphur compounds to H₂S’ reaction is complete, an excess of H2 is required after the reduction reactor, regard - less of the H2 source. In the second process step, the H₂S needs to be sepa - rated and returned to the inlet of the modified Claus reac - tion furnace. This step is based on an amine absorber/ regenerator system. The overall sulphur recovery efficiency (modified Claus reaction and TGTU) is required and predicted to be at 99.9+%. This is confirmed by measuring the total mass emission of sulphur dioxide (SO₂) at the exhaust of the final thermal reactor or ‘stack’. Two points of interest for close process control include the quench column and absorber outlet. Considering the AT5 quench tower outlet measurement (H₂ and H₂S), since the TGTU process was first intro - duced as the Shell Claus Off-Gas Treating (SCOT) process, H2 measurement has been expected at this point. It was included in the original system design. As previously men - tioned, the purpose of measuring the H₂ here is to ensure that excess H₂ is coming out of the CoMo reactor. H₂S is also measured at this point so that operators and the automated control system understand the amount of H₂S that will be entering the absorber. Sample gas measurements at this point of the process are easier to handle than AT4 measurements at the quench tower inlet because of the lower temperature of the process gas. Any particulates will also have been removed in the quench tower. The primary measurement at the AT6 absorber outlet H₂ and H₂S and carbonyl sulphide (COS)/CS₂ point was defined as a single H₂S measurement to ensure the perfor - mance of the amine absorber. By gaining knowledge about the application and availability of multicomponent instru - ments, additional measurements became interesting. Knowing the importance of excess H2 in the TGTU and recognising that an additional measured component does not add significant cost to an analyser, a redundant H₂ measurement should be added at this point. By adding the H₂ measurement, redundancy can be achieved without sig - nificant extra costs. The same can be said about adding a COS and/or CS₂ measurement; both can be used to determine the condition of the CoMo bed catalyst. If the COS and CS₂ values are increasing, the CoMo catalyst needs to be replaced, or other operational variables such as flow rate or temperature need

SO + 3H

HS + 2HO

S + H

HO + CO

H + CO

COS + HO CS + 2HO

CO + HS CO + 2HS

Figure 3 Cobalt-molybdenum catalysed reactions

of a modified Claus unit look like? The process consists of two chemical reaction steps: 3H₂S + 3/2O₂ → SO₂ + 2H₂S + H₂O SO₂ + 2 H₂S ↔ 3Sx + 2H₂O First is a thermal reaction (Claus furnace) followed by a catalytical reaction. The key parameter is the air (oxygen) flow to the thermal reactor. This setup will result in a spe - cific H₂S-to-SO₂ ratio. For a modified Claus unit, the ratio is described as 2:1. Depending on the tail gas treater down - stream of the Claus unit, this ratio can be different. As shown in Figure 2 , the tail gas analyser controls the trim air valve, which typically represents 10% of the air - flow to the reaction furnace. The main air valve is set by the acid gas-to-air ratio calculated on the best available infor - mation about the gas entering the reactor. This brings us back to the previous arguments: improved process control by getting a reliable feed gas composition analysis. The two measurements (feed and tail gas) in combination will result in better process control. Significant sulphur recovery rate improvements in the modified Claus plant are possible by using modern process instruments and combining them into a full process control scheme Modified Claus unit control It is essential to understand the potential impact on the process control scheme for each of the installed measure - ments. ‘Nice to have’ does not make the difference; relia - bility and safety are the key. Significant sulphur recovery rate improvements in the modified Claus plant are possible by using modern process instruments and combining them into a full process control scheme. What is important to consider with any addition of pro - cess instrumentation, besides safety, is that all instrumen - tation must provide data quickly and reliably. Downtime is not acceptable. Today’s environmental regulations do not allow sulphur recovery rates <99%. Sulphur emissions will still be too high. A standalone modified Claus unit will not achieve the required sulphur conversion rates. Downstream tail gas

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to be adjusted, or the modified Claus unit has some opera - tional concern that needs to be addressed. Looking into the combination of the two previously described sample points, a redundant measurement of H₂ will ensure optimal performance of the TGTU. It should reduce the cost of replacing contaminated amine and min - imise downtime. If H₂S is measured at the quench tower outlet and the absorber outlet, it will be possible to measure and control absorber efficiency online on a 24/7 basis. Process control of the amine recycling/regeneration is also possible based on the H₂S ‘IN’ and ‘OUT’ measurements. The final quality control parameter is the SO₂ mass emis - sion to be measured at the thermal oxidiser outlet. Only mass emission can tell the true rate of sulphur recovery. Knowing the H₂S entering the process and the amount of SO₂ leaving the oxidiser will provide an accurate value. Summary Modern process instrumentation has the potential to help reduce the emissions from an SRU. The initial investment is not negligible, but the application challenges are high with regard to measurement reliability and safety aspects (the SRU has the most toxic gas mix - tures of the entire refinery). Knowledge, understanding, and awareness training are essential to maintain the instru - ments; without the instrument signals, the emission targets are not achievable. To achieve such goals, technical disciplines must work together, starting at the processing unit’s design phase.

Misunderstandings about piping can result in unsafe installations. Choosing the ‘wrong’ instrument may result in not achiev - ing the best possible process performance. Untrained local staff may result in instruments not working. Finally, the average lifetime of an instrument installed in an SRU is 15-20 years. Choosing a supplier under these circum - stances can be challenging. References 1 Molenaar, van Son – Worley, Keep your sulphur recovery unit online and efficient, Sulphur Magazine , No. 408. 2 Design, safety and operational aspects of SRU analysers, CRU Sulphur Magazine , No. 400. Jochen Geiger is Director Sales and Service EMEA and India at AMETEK Process Instruments, based in Cologne, Germany, and specialising in industrial applications and instrumentation (process analysers) in the chemical industry. He has provided training courses on chemical plants, refineries, and oil and gas production sites around the globe. He holds a Masters in electronics and is Process Analyser Specialist. Email: Jochen.geiger@ametek.com Michael Gaura is Senior Product Manager, USA at Ametek Process Instruments USA/Canada. He holds a BS in biology and chemistry from Purdue University and a Masters from Oklahoma State University. Anantha Kukkuvada is Regional Sales Manager India, Sri Lanka, Bangladesh and Africa, at Ametek Process Instruments India. He has 20 years of experience in sales and marketing in the field of analysers and field instruments. He holds an MBA in business administration and management from Indian Institute of Management Bangalore.

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Midstream gets the full treatment. For H2S, oil and gas, and mercaptan treatment, we go the distance. Every midstream project has its own challenges. Merichem meets those unique challenges with speed and exibility. We oer a range of both standardized and customized solutions to meet your needs while surpassing your expectations.

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‘On the fly’ vs high-performance H2 S selective solvent

A focus on highly selective designs in low-pressure tail gas treating units using BASF’s proprietary technology in comparison to generic MDEA solutions

Ashraf Abufaris BASF Middle East Chemicals L LC Blake Morell BASF Corporation

S elective removal of hydrogen sulphide (H 2 S) has become an important topic over the last two decades. Selective designs are tailored either on maximum or controlled H 2 S selectivity, depending on the application. The reaction equilibrium in sulphur recovery units (SRU or Claus section) prevents complete conversion of the sulphur species in feed gas to elemental sulphur. Typically, an SRU with two to three Claus reactors can only achieve 93-98% sulphur recovery efficiency. However, higher recoveries of 99.8% and above are achievable if the remaining sulphur compounds in the SRU tail gas are hydrogenated to H 2 S, which is removed in a selective amine unit (tail gas treating unit, TGTU). Selection of the proper amine technology for the TGTU is essential to make these projects economically and environ- mentally viable. Use of a highly H 2 S selective solvent, such as BASF’s proprietary OASE yellow, can provide benefits by optimising the capital investment or reducing the oper- ating cost. Various parameters during the design phase influence the H 2 S selectivity (and consequently carbon dioxide [CO 2] slip) in TGTUs, such as absorber height, amine circulation rate, and absorber internals in the mass transfer zone. However, one of the most effective levers is the amine tem- perature itself. H 2 S selectivity of generic solvents rapidly deteriorates once amine temperature exceeds 45°C. A key benefit of the OASE yellow selective solvent is a maintained H 2 S selectivity even in high ambient temperature environ- ments and subsequent high lean amine temperatures of up to 50°C. This avoids installing/operating costly chillers for solvent cooling and makes the design reliable, robust, and flexible for various operational scenarios. Against this backdrop, key parameters for these selective designs are discussed, followed by operational start-up data from OASE yellow solvent swaps. Design options to influence H2S selectivity A number of factors influence H2 S removal in the presence of CO 2 . Adjusting these parameters plays a critical role in unit optimisation throughout the design, commissioning, start-up, and operation phases:  Amine type Historically, methyldiethanolamine (MDEA) has been

widely used in H 2 S selective applications in the industry. However, recent stricter sulphur dioxide (SO 2) emission targets that meet the World Bank standard of 150 mg/ Nm 3 often require additional chemistry to further boost the MDEA performance of other amines to achieve tight treated gas H 2S specifications. Besides the performance- related characteristics, properties such as volatility, stability, acid gas loading capacity, and of course commercial aspects are important selection criteria.  Lean amine temperature Selective treatment with amine-based solvents generally takes advantage of the rapid reaction of H 2 S compared to the kinetically hindered reaction of CO 2 : CO 2 first must react with water to form carbonic acid before the solvent can absorb the CO 2 . Thus, tertiary amines such as MDEA are often used for selective applications as they cannot form carbamates (the only fast reaction with CO 2 ). The following reactions of tertiary amines take place in aqueous solutions: Reaction of water and amine (fast) R1R2R3N + H 2 O  R1R2R3NH + + OHˉ 2 H 2 O  H 3 O + + OHˉ

H 2 S reaction (fast) H 2 S + H 2 O  HSˉ + H3 O +

CO 2 reactions (overall reaction: slow): CO 2 + 2H 2 O  HCO 3ˉ + H3 O + (slow) HCO 3ˉ + OHˉ  H 2 O + CO 32 - (fast)

In this reaction system, CO 2 co-absorption and, thus, H 2 S selectivity are heavily influenced by reaction conditions. This means higher pressure and temperature, as well as a higher CO 2 /H 2 S ratio in the feed gas, favour CO 2 co-absorption and lower H 2 S selectivity, especially at lean amine temperatures above 50°C, which are typical for the Middle East region. The CO 2 reaction accelerates and strongly competes with the H 2 S reaction. As a result, a cooling system/chiller is often part of the design in these climates to achieve H 2 S selectiv- ity with an MDEA/acidified MDEA solution.  Mass transfer Besides lean amine temperature and feed gas pressure

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