My Contributions on Problem Solving of Crude Oil Refining Process

The responses published here are based on my knowledge and experience and don’t have the pretention to be the unique and right argument in all the cases, in all business the diversity of point of view is welcome and this is not different in the downstream industry.

Question 37 - I would like to learn about the usage of Pyrolysis oil produced from Ethylene Cracker in the Delayed Coker Unit. I would like to use this pyrolysis oil which is around 2.5 wt % of total feed as a feed in the Delayed Coker Unit. Do you know any applications like this? If it is, have you encountered any issue to process pyrolysis oil to the coker unit? In addition, what if I use this pyrolysis oil as wash oil stream in the coker fractionator column instead of using as a feed? Do you know any applications or example?

Response - The pyrolysis oil from steam cracking units can be an attractive feed for delayed coking units, especially for those dedicated to producing high quality needle coke due to the high aromaticity of the pyrolysis oil. Despite this advantage, the participation of this stream in the feed can be limited due to the high potential of coking lay down in the fired heaters which will reduce the operational campaign of the processing unit like quoted in the que question.

Despite being a relatively common feed for delayed coking units, unfortunately I don't have experience in operating delayed coking units with this feed stream, but the main side effect can be the acceleration of the coke deposition in the fired heaters which demands sensibility analysis to determine the maximum participation of the pyrolysis oil in the feed in order to balance the economic return between the pyrolysis oil advantages (needle coke production, for example) with the shorter operational life cycle of the processing unit. About the use of pyrolysis oil as wash oil stream in the main fractionator, my point of view is that the pyrolysis oil is much heavy to this service and tends to raise the coking rates in the fractionating and thermal exchanges beds which will reduce the operational life cycle of the delayed coking unit or, at least, reduce the performance of the processing unit.

Question 38 - Is CCR value directly linked with asphaltene content? In hyrocracker feed specification why limit is fixed for both CCR and Asphaltene content? Why are both parameters measured separately?

Response - The CCR analysis is directly related with asphaltenes content which are responsible of the most part of coke laydown tendency from the residue, but another components like resins, aromatics and saturated hydrocarbons can contribute with the stability of the residue and influence over the feed quality of residue hydrotreating or hydrocracking unit as well as with the expected yields of distillates. Normally, it's carried out complimentary analysis to CCR like SARA (Saturate, Aromatics, Resins, and Asphaltenes) and H/C ratio to allow a most adequate analysis of the stability of the feedstock and the coke laydown tendency which will determine fundamental operating parameters like hydrogen partial pressure necessary to ensure an adequate and economic life cycle for the processing unit.

Question 39 - We have a black sludge formation at the interphase of naphtha and water in OVHD accumulator. But all the OVHD parameters like pH, Iron and Chloride are normal. The crude unit has a partial condensation overhead system, and the black sludge is observed in the second boot (Cold reflux boot). There is a CI dosing in the accumulator upstream. What could be the reason for this sludge?

Response - This is a relatively common condition in overhead systems of crude oil distillation units. The black sludge observed in the overhead vessel is probably pickering emulsion stabilized by iron particles which is accumulated in the interface between sour water and naphtha, despite the information that the pH, Iron and Chloride content is controlled in the overhead system it's possible that this system and the atmospheric tower can operate under corrosion situation in the past. When the emulsion is formed in the vessel, this residue cant be removed without the shutdown of the processing unit or through draining the overhead vessel totally which requires a special procedure aiming to minimize the safety risks as well as the damage to the pumps of the overhead drum.

Regarding the corrosion control in the overhead systems, it's important to analyze that the corrosion control parameters is under an adequate range, especially the operating temperature of the overhead system. There are some correlations in the literature which relates the ammonia and chloride concentration in the sour water to determine salt deposition temperature in the top of the tower and this needs to be considered to define the operating temperature of the system.

Question 40 - Can high-grade needle coke (such as P66 or Seadrift) be used for synthetic graphite in EV battery anode materials? If the quality exceeds specifications and needs to be downgraded, is it easy to modify the needle coke manufacturing process? Or is it easy to source lower grade decant oil?

Response - It's possible to produce high qulity needle coke in delayed processing units capable to meet restrict quality requirements, but the production route of needle coke demands some specific operating conditions and feed stream quality to be processed in delayed coking units. The decanted oil from FCC is one of the best alternatives to produce high quality needle coke, to produce a lower-grade needle coke you can blend the decanted oil with vacuum residue with lower content of aromatics or, in refineries relying on solvent deasphalting processing units, add asphaltic residue to the feed stream. It's necessary to carry out some operating tests with different feed stream compositions aiming to determine what is the best composition to reach your desired quality of needle coke.

Question 41 - We have low PH (3 to 4) in the CDU overhead but in same time we have low chloride values ( 3 to 10 ) and already we injected high values of neutralizing amine and corrosion inhibiter. What is the reason that causes this drop in PH value?

Response - It's important analyze the content of chloride salts (MgCl2 and CaCl2) in the processed crude, these salts can suffer hydrolysis and generate hydrogen chloride (HCl) which can cause drastic reduction in the pH. According to the concentration of chloride salts in the crude oil it's possible to minimize this problem injection sodium hydroxide (NaOH) upstream of the desalting vessels aiming to neutralize the hydrochlorides compounds.

Question 42 - What is the basis for maintaining minimum wetting rates in vacuum column (whether based on vacuum charge or design condition?) What will happen if minimum wetting rates are not adhering to?

Response - The response for this question depends on a several parameters like the characteristics of the column internals as well as the mixture which will be separated.

Considering that we are dealing with a vacuum column, there is a great chance that the equipment is operating with packing internals which presents lower pressure drop than the perforated plates. There is a several correlations in the literature capable to give an estimative for the minimum wetting rate of a separation column which relies on the characteristics of the fluids like viscosity, density, and temperature and the characteristics of the packing like applied material, if is stacked or random, geometric form, atc.

An wetting rate below of the minimum will not conduct adequate mass and heat transfer rates, leading to poor performance of the separation column. In services with high temperature with hydrocarbons the low wetting rate can lead to premature coking deposition in the separation section leading a poor fractionating performance, high pressure drop and shorter operating lifecycle.

Question 43 - What is the impact in the product quality if circulating refluxes return temperatures are not maintaining at design temperatures? Is it wise to reduce heat recovery in pre heat network for maintaining design pump around return temp at the expense of pre heat?

Response - The temperature profile of a separation column is a key parameter for an adequate fractionating, for this reason it's expected deleterious effects over the final quality of the products or side streams if the temperature profile is below or above the parameters recommend by design.

Reduce the heat recovery to maintain an adequate temperature profile in the distillation column can be interesting in some cases, bute reveals that you have a problem with your energy balance and recovery of the processing unit. Crude oil distillation units are the major energy consumer is a crude oil refinery and the energy is responsible for higher then 60 % of the operating costs of a crude oil refinery, furthermore the CO2 emissions is raised in an unefficient energy system, based on these data it's not recommended to deoptimize the energy balance of the processing unit even to improve the fractionating quality. In other words, if this is happened it's necessary a energy integration study (maybe through pinch technique) to identify bottlenecks and then propose alternatives to eliminate then.

Question 44 - What are the optimal unit configurations and combinations (e.g., FCC/hydrocracking) for increasing high-margins products, while reducing low-value streams (e.g., HSFO & LSFO)?

Response - The response depends on the characteristics of the processed crude, especially the sulphur content and API grade.

Regulations like IMO 2020 imposed severe restrictions over the refining hardware to process high sulphur crudes and the refiners capable of adding value to heavier and sour crudes reached significant competitive advantage. The synergy between FCC and hydrocracking units gives high flexibility and maximizes the refining margins, especially considering the growing market of petrochemicals but is a capital intensive solution and can be prohibitive for low capital power players.

Refiners processing medium and low sulphur crudes can apply the combination of FCC and solvent deasphalting or delayed coking units and reach significant added value to the processed crude with less capital expense, here it's necessary to consider that the refiner will need to rely with adequate hydroprocessing capacity to treat the intermediate streams and this needs to be considered in the investment analysis.

Another point to be considered is if the refiner needs to meet the market of bottom derivatives like asphalt or fuel oil. For these players it's necessary to consider that a deep conversion refining hardware like reached with the combination of FCC and hydrocracking can led to a lack of bottom barrel streams to produce these derivatives, with consequent opportunity lose (in some cases the refining margin is attractive for bottom derivatives like asphalt) and supply shortage in the market.

Question 45 - In building the petrochemical value chain, how much further can we see the FCC unit being used to increase olefins production with the wide range of feedstocks currently available, including waste plastics-derived pyrolysis oil?

Response - Considering the growing demand by petrochemicals and the operational flexibility of the FCC units is expected than the FCC technologies will be in the core of petrochemical integration efforts from the refiners.

Beyond the traditional actions to maximize the olefins yield in conventional FCC units like higher operating temperature, higher Cat/Oil ratios, and catalyst additives (ZSM-5), is expected growing investments in high severity and petrochemical FCC units. There is some trends that will demand research and development involving FCC units like the renewables co-processing and the use of pyrolysis oil from plastics recycling plants as feed of FCC units which requires a deep analysis regarding the effect on the production profile of the processing unit as well as the catalyst deactivation rate. Regarding the renewables co-processing, the use of ethanol as feed to FCC units aiming to improve the yield of ethylene is one of the most interesting trends which will demand analysis by refiners and technology developers in the next years, in my point of view.

Question 46 - Under what conditions do you see opportunities with the upgradation of distressed refinery products (such as vacuum residue) to higher value outlets? Are these opportunities primarily outside the fuels market?

Response - I believe that the response relies on the market that the refiner is inserted and the crude oil blend which is processed by the refinery. In developing economies still present high demand by transportation fuels like gasoline and diesel and, for these players can be attractive adopt refining routes based on residue upgrading technologies with less capital spending like the combination of delayed coking and hydrotreating to maximize the diesel and gasoline production ensuring high added value to the bottom barrel streams.

Refiners processing low sulphur crudes can still produce low sulfur fuel oil or Bunker (VLSFO) in compliance with IMO 2020 with just dilution of atmospheric residue once this derivative presents high demand due to the environmental regulations in ECA (Environmental Control Areas), but this can be economically attractive just for refiners processing crude blending capable to produce an atmospheric residue with maximum sulphur content close to 0,5 % wt.

Considering the recent forecasts, the developing economies are facing deep changes in the downstream market with growing demand by petrochemicals and hostile scenario for fossil fuels, for these markets closer integration between refining and petrochemical assets is the trend and this is strictly related with the capacity to add value to the bottom barrel streams. In these markets can be attractive the capital investments in deep residue upgrading technologies like hydrocracking and his synergies with FCC units.

Another attractive alternative is the lubricant market once these derivatives present growing demand and high added value, again it's necessary to consider an adequate refining hardware once the Group I lubricants present in contraction market. Refiners which intend to be competitive in the lubricant market need to make capital investments in hydrocracking units capable of producing Group II and III base oils.

Question 47 - Is it possible that refiners could be overlooking some practical solutions to increasing FCC olefins yields, such as in the gas plant/recovery section?

Response - No doubt, there are some relatively easy solutions to improve the yield of petrochemical intermediates in FCC units which can be overlooked by some refiners.

Considering the current market demand, the FCC units can be optimized to produce higher added value derivatives like light olefins, refiners facing gasoline surplus markets can operate the processing unit in maximum olefins operation mode, to minimize the production of cracked naphtha.

In this operation mode the FCC unit operates under high severity translated to high operation temperature (TRX), high catalyst/oil ratio. The catalyst formulation considers higher catalyst activity through addition of ZSM-5 zeolite. There is the possibility of a reduction in the total processing capacity due to the limitations in blowers and cold area capacity.

It’s observed an improvement in the octane number of cracked naphtha despite a lower yield, due to the higher aromatics concentration in the cracked naphtha. In some cases, the refiner can use the cracked naphtha recycle to improve the LPG yield.

In the maximum LPG operation mode, the main restrictions are the cold area processing capacity, metallurgic limits in the hot section of the unit, treating section processing capacity as well as the top systems of the main fractionating column. In markets with falling demand by transportation fuels, this is the most common FCC operation mode.

Through changing the reaction severity, it is possible to maximize the production of petrochemical intermediates, mainly propylene in conventional FCC units.

The use of FCC catalyst additives such as ZSM-5 can increase unit propylene production by up to 9,0%. Despite the higher operating costs, the higher revenues from the higher added value of derivatives should lead to a positive financial result for the refiner, according to current market projections. A relatively common strategy also applied to improve the yield of LPG and propylene in FCC units is the recycling of cracked naphtha leading to an over cracking of the gasoline range molecules.

Nowadays, the falling demand by transportation fuels has made the refiners optimize the FCC units aiming to maximize the propylene yield following the trend of a closer integration with the petrochemical sector. Among the alternatives to maximize the propylene yield in FCC units is the use of ZSM-5 as additive to the FCC catalyst as well as the adjustment of the process variables to most severe conditions including higher temperatures and catalyst circulation rates. Another interesting alternative is to recycle the cracked naphtha to the process unit aiming to improve the LPG and consequently the propylene yield.

The installation of propylene separation units can present a significant capital investment to refiners but considering the last forecasts that reinforces the trend of reduction in the demand by transportation fuels, this investment can be a strategic decision to all players of the downstream industry in the middle term both to ensure higher added value to the processed crude oil and market share. Another possible capital investment aiming to improve the yield of light olefins recovery from FCC units is the use of cryogenic processes in the gas recovery section against the conventional configuration with separation columns, in this case the recovery of ethylene is highly improved.

Question 48 - Do you see growing investor interest in processing plastic waste-derived pyrolysis oil through refinery assets, such as hydrocrackers? Against this backdrop, how prepared are refiners to invest in contaminants removal systems (for pretreatment of the pyrolysis oils)?

Response - We are facing a continuous growth of petrochemicals demand and a great part of these crude oil derivatives have been applied to produce common use plastics. Despite the higher added value and significant economic advantages in comparison with transportation fuels, the main side effect of the growth of plastics consumption is the growth of plastic waste.

Despite the efforts related to the mechanic recycling of plastics, the increasing volumes of plastics waste demand most effective recycling routes to ensure the sustainability of the petrochemical industry through the regeneration of the raw material, in this sense, some technology developers have been dedicated investments and efforts to develop competitive and efficient chemical recycling technologies of plastics.

One of the most applied technologies for plastics recycling is the thermal pyrolysis where the long chain polymeric is converted into smaller hydrocarbon molecules which can be fed to steam cracking units to reach a real circular petrochemical industry. Unfortunately, the thermal process produces chemically unstable feedstock to steam cracking units which raise the coking deposition rates and drastically reduces the operational life cycle of the cracking units. An alternative to the thermal process is the catalytic pyrolysis which is more selective and can produce molecules more stable than the thermal process, but these technologies are still under development.

Another promising chemical recycling route for plastics in the hydrocracking of plastics waste, in this case the chemical principle involves the cracking of carbon-carbon bonds of the polymer under high hydrogen pressure which lead to the production of stable low boiling point hydrocarbons. The hydrocracking route present some advantages in comparison with thermal or catalytic pyrolysis, once the number of aromatics or unsaturated molecules is lower than the achieved in the pyrolysis processes, leading to a most stable feedstock to steam cracking or another downstream processes as well as is more selective, producing gasoline range hydrocarbons which can be easily applied in the highly integrated refining hardware.

The chemical recycling of plastics is a great opportunity to technology developers and scientists, especially related to the development of effective catalysts to promote depolymerization reactions which can ensure the recovery of high added value molecules like BTX. More than that, the chemical recycling of plastics is an urgent necessity to close the sustainability cycle of an essential industry to our society. In my point of view, despite the necessity of better development of the available plastics recycling routes, the capital investment in these technologies is essential to any player which intends to be competitive in the petrochemical market, mainly in the Asian market which is more developed in this sense.

Question 49 - Projected diesel shortages could become a crisis if winter conditions are severe, potentially knocking out already strained power grids. What strategies should refiners rely on to increase distillate-range material?

Response - The response relies deeply of the available refining hardware as well as the processed crude oil.

Generally, it's possible to optimize the crude oil distillation unit to maximize middle distillates as well as gas oils capable to be converted into diesel after post processing in residue upgrading units like hydrocracking or deep hydroprocessing. Another interesting alternative is optimize the blending operations in stockpiling assets aiming to maximize the yield of middle distillates respecting the derivatives specifications, a common operations in some refineries is blending straight run heavy nafta with diesel aiming to improve the produced volume in the diesel pool.

The cracked heavy nafta can also be added to the feed of diesel hydrotreating units to improve the produced volume of diesel, of course, if the processing unit is able for this as well as there is hydrogen availability. Another cracked feed with is sent for diesel pool after adequate hydrotreating is the Light Cycle Oil (LCO) from FCC units, despite the high aromatics concentration of LCO, this stream can help to improve the diesel production through high severity hydrotreatement.

Question 50 - How can the refining industry supply the aviation industry’s growing demand for sustainable aviation fuel (SAF)? What are the most efficient pathways?

Response - This is one of the hot points of the downstream industry nowadays. The biofuels and renewables co-processing have a fundamental role in the energy transition and decarbonisation of refining industry and we are seeing attractive processing routes capable to reduce the carbon intensity of the fossil fuels like the co-processing of renewable raw materials in hydrotreating units to produce less carbon intensive diesel and jet fuel, for example. In the petrochemical industry, the ethanol to olefins route is another promising route which already presents commercial production plants.

The use of total renewable feedstock can be attractive under specific scenarios, but it's always important to consider the source of the renewable raw materials in order to avoid the competition with food industry as well as the pressure over the agribusiness especially in regions with restrictions of available arable lands. These restrictions can be also related to the biofuels production through esterification which are normally blended with the fossil fuels before commercialization.

Another interesting processing route is the Gas to Liquid (GTL) hydrocarbons production route applying biomass as feedstock, again it's necessary to consider the availability of the renewable raw material and the politician and social impact of this alternative. In the technical point of view, this alternative can produce high quality and low contaminants liquid hydrocarbons.

Lastly, but not less important, any effort to energy transition of the downstream industry needs to consider the hydrogen source applied to the process. We are seeing an increasing pressure in the last decades to reduce the environmental footprint of the fossil fuels and great part of the obtained results was achieved through the hydrorefining, leading to a growing dependence of hydrogen which, until this moment, is industrially produced through natural gas reforming process that produce great amount of carbon dioxide (CO2) emissions. The processing of renewable raw material requires even more hydrogen to achieve the goal of high quality and less contaminant fuels production, in other words, the sustainability cycle only will be closed if the hydrogen applied to the process is renewable or there is efficient carbon capture technologies which are still incipient in the market.

In summary, there are available processing routes and technologies capable of supplying the market of renewable fuels, but it's necessary to consider all impacts of the production chain as well as if the sustainability cycle is really closed.

Question 51 - What are water partial pressure & chloride partial pressure in the fixed bed catalyst of Naphtha Reforming Unit ? And how can be controlled ?

Response - The management of water/chloride relation is a key parameter for catalytic reforming units aiming to ensure an adequate balance between the acidic and metal functions of the catalyst. Normally, fresh catalytic reforming catalysts presents close to 1,0 %  wt of chloride, to maintain this chloride concentration it's necessary to control the water concentration aiming to allow an effective interaction between the alumina (catalyst support) and the chloride, reaching then a good performance of acidic sites of the catalyst which is responsible by the cracking reactions.

According to the literature, several factors impact the chlorides concentration in catalytic reforming catalysts. The reactor temperature and surface area of the support can directly affect the chloride concentration in the catalyst and are the most relevant factors. Still according to the literature, fixed bed catalytic reforming reactors should operate keeping the water to chloride molar ratio between 15 to 25 in the recycle gas aiming the keep the activity of the catalyst, to control this parameter it's necessary to install sample facilities or online monitoring systems in adequate points aiming to keep this parameter according to the licensor specifications. It's possible to find in the specialized literature chlorides equilibrium curves capable of helping the refiners to control the water to chloride ratio in the catalyst under the specifications defined by the licensors.

Question 52 - What is the best MOC (Material of Construction) for the NMP recovery and Dehydration portions of a Solvent Extraction System? We are finding that acid in the feed oil concentrates in the recirculating NMP and that this is degrading the 304 SS process vessels. Is 316 SS a good choice or will we need to go to more exotic alloys? The recirculating NMP can hit a pH of 4.0 and sometimes down to 3.7

Response - Unfortunately, the concentration of acidic compounds in NMP (N-Methyl Pyrrolidone) is a relatively common issue in lube oil dearomatization units.

Before considering changing the material of construction of the process equipment, please verify the possibility to raise the frequency of purge and make up the NMP aiming to reduce the concentration of acidic compounds in the solvent. Additionally, some references describe the use of sacrificial metals like magnesium and zinc installed in the critical of solvent recovery and dehydration sections of dearomatization units as an effective way to deal with corrosion issues in these processing units, I believe that this can be most economically attractive face to change the MOC (Material of Construction) of the processing unit.

The use of stainless steel like AISI 316 can be interesting, but it's expensive and can led to other issues once these materials are very sensitive to stress corrosion due chlorides and it's very difficult to ensure the absence of these compounds in a processing unit (a simple cooling water leakage can contaminate the system with chlorides). I believe that it can be interesting to make an economic analysis comparing the cost of replacing the construction material of the processing unit face to raise the frequency of solvent change or make up flow rate in order to reduce the acid concentration in the recirculating solvent (NMP).

Question 53 - Are there any DCU units that processes more than 50%wt of SDA in fresh feed? If so, do you have any problems with increased foaming or fouling?

Response - Unfortunately, I don't know a Delayed Coking unit that processes this perceptual of SDA residue, but please consider these factors about the foam formation in delayed coking reactors:

1 - Feedstock's characteristics: The paraffinic feeds tends to present high foam level in the reactor than aromatic feeds once the paraffinic compounds cracking faster than aromatics compounds, creating a flow of gas through the liquids in the reactor. Another parameters of the feedstock's which can cause foam production is the presence of high sodium and solids concentration in the feed;

2 - Sudden depressurization of the reactors: This disturb can cause an excessive velocity in the reactor, favouring the foam formation;

3 - Inadequate heating of the feed: Some refiners can try to reduce the temperature to minimize the coking issues in the fired heaters and reduces excessively the feedstock temperature leading to the increasing of the foam in the reactors. It's necessary to make a balance between the coking of the fired heaters and foaming formation in the reactors;

4 - Excessive velocity in the reactors: The high velocity in the delayed coking reactors can be caused by an excessive flow rate of the feed as well as reduced pressure of the reactors;

As described above, acting in the temperature and pressure of the reactors it's possible to minimize the foaming formation. Higher temperature and pressure tends to reduce the foaming production in delayed coking reactors, but it's necessary to consider the another aspects once the increasing of pressure and temperature affects directly the quality of the produced coke.

Question 54 - What's the philosophy of desalting system in Crude Distillation Unit, with respect to High Voltage & Demulsifier?

Response - The desalting of crude oil is one of the most important processes in a refinery to ensure the reliability and the operational availability of the refining hardware. During the crude oil extraction processes the petroleum drag sediments and water beyond inorganic salts (carbonates, calcium, chlorides, etc.) which are responsible for fouling heat exchangers leading to efficiency reduction, raise in energy consumption and reduce the operation campaign of the process units.

The presence of the dissolved salts in the crude oil is still responsible for catalysts deactivation in conversion process units (FCC and Hydrotreating), furthermore, these compounds can accumulate in the top of atmospheric crude distillation columns leading to corrosion and loss in separation efficiency. The desalting process involves the mixture of crude oil with water aiming at the dissolution of the salts considering the higher solubility of these compounds in the aqueous phase.

The crude oil is pumped from the storage tanks through the heating battery where it is heated and mixed with dilution water, the mixture is made by a mixing valve that promotes an intense mixture through pressure drop. A major part of water is under the free form and is removed by decantation due to the difference of density between the aqueous and oil phases, however, part of the water is emulsified in the oil phase and are required actions to break the emulsion and allow the decantation of this water and the dissolved salts.

The emulsion breaking is carried out with the application of high-intensity electric field (close to 3,0 kV/cm) that provokes the polarization of water droplets, his agglutination and consequently his decantation. Desalting heavy crude oils is a greater challenge to refiners once the lower difference of density between the aqueous and oil phases makes the separation hard, beyond the higher content of compounds which stabilize the emulsions in heavier crudes (asphaltenes), in these cases the refiners operate under higher desalting temperatures and are used demulsifiers to facilitate the emulsion breaking.

Demulsifiers are normally a combination of surfactants with hydrophilic and hydrophobic bands in the same molecule which normally have their formulation protected by patents and his dosage needs to be accompanied by a specialist (chemical vendor). Regarding the electrical field, higher electrical intensity tends to improve the desalting efficiency considering the other variables fixed once improve the mixing effect and intensity of water droplets, collision with consequent coalescence and decantation, but it's necessary to consider that there is an optimal point for achieve this effect, once mixing in excess can promote collisions but without adequate conditions of coalescence.

It's important to consider the whole desalting process and all operating variables and not only the demulsifier and electrical field. The desalting temperature is a key parameter of the process once impact the viscosity of the crude and consequently the sedimentation velocity, it's important to realize a study including all operating variables like content of dilution water, pressure drop in the mixture valve, electrical field and desalting temperature. It's important to consider the compatibility of the crude oils processed, which can lead to asphaltenes precipitation in some cases, especially in blends of high paraffinic crudes with heavier crudes.

Question 55 - What are the other methods to reduce feed foaming in DCU reactor, apart from the use of anti-foaming agents, increasing pressure and temperature?

Response - The foam formation in the delayed coking reactors can be caused by a series of factors like:

1 - Feedstocks characteristics: The paraffinic feeds tends to present high foam level in the reactor than aromatic feeds once the paraffinic compounds cracking faster than aromatics compounds, creating a flow of gas through the liquids in the reactor. Another parameters of the feedstocks which can cause foam production is the presence of high sodium and solids concentration in the feed;

2 - Sudden depressurization of the reactors: This disturb can cause an excessive velocity in the reactor, favouring the foam formation;

3 - Inadequate heating of the feed: Some refiners can try to reduce the temperature to minimize the coking issues in the fired heaters and reduces excessively the feedstock temperature leading to the increasing of the foam in the reactors. It's necessary to make a balance between the coking of the fired heaters and foaming formation in the reactors;

4 - Excessive velocity in the reactors: The high velocity in the delayed coking reactors can be caused by an excessive flowrate of the feed as well as reduced pressure of the reactors;

As described above, acting in the temperature and pressure of the reactors it's possible to minimize the foaming formation. Higher temperature and pressure tend to reduce the foaming production in delayed coking reactors, but it's necessary to consider another aspects once the increasing of pressure and temperature affects directly the quality of the produced coke.

Question 56 - 1) Increasing of the water content (H2O) of the Heavy Naphtha in the Storage Tanks … and the moisture in the Recycle Gas. What are the causes?

2) If the required concentration of chlorine (Cl) in the feed is (2 ppm). Provided that: Unit capacity (Reforming Unit) = 40 m³/hr (Density of Heavy Naphtha= 0.742 gm/cc ; Density of PDC (Propylene Dichloride C3H5Cl2) added = 1.15 gm/cc)

So, what's the amount of PDC to be added so that we obtain the above concentration (2 ppm)?

Response - I'm understanding that the heavy naphtha is fed to a catalytic reforming unit and was observed higher water content in the feedstock tank of the unit recently. About this issue it's recommended to check the separation quality in the top vessels of the crude distillation unit, especially the straight run naphtha split column in order to ensure that water is not being dragged to the heavy naphtha and accumulating in the storage tank. The water excess in the naphtha feed can be raising the moisture in the recycle gas.

Regarding the second question, based on informed data it's necessary to feed close to 95 ml/h of PDC (Propylene Dichloride) to ensure the desired chloride concentration in the feedstock of the catalytic reforming unit.

Question 57 - We have a problem in the specifications of the feed of the Reforming Unit (Heavy Naphtha: Sulfur content = 2.6 must be <1; Dr. test = Slightly H2S must be negative)

Knowing that the operating conditions of the Hydrodesulfurization unit are normal.

Any solutions welcome.

Response - Considering your information that the Hydrodesulfurisation is OK, we need to check the characteristics of the feed. Some relevant questions need to be answered like if your catalytic reforming is operating with only straight run naphtha, in some refineries it's common to apply hydrotreated coker naphtha as feed for catalytic reforming units.

I will consider that your processing unit is operating wit only straight run naphtha, in this case, it's necessary to check the sulphur content of the straight run naphtha to check that this is respecting the maximum hydrodesulfurisation capacity of the hydroprocessing unit. Maybe the change of processed crude to heavier or sourer crudes can be affected the sulphur content of the straight run naphtha.

Another point to check is the fractionating quality int he atmospheric column of the crude distillation unit, in some cases the salt formation in the tower internals can led to the poor fractionating quality in some sections and kerosene can be dragged to the naphtha, raising the sulphur content of this stream. My suggestion is to carry out a characterisation of the naphtha fed to hydrodesulfurisation with analysis of total sulphur and boiling range to identify possible contamination with heavy hydrocarbons.

Question 58 - We are having problems with the volatile material in the carbon of the DCU. To lower the MV we have increased the temperature of the furnace, and increased the steam flow during the vaporization of the bed; any other recommendations for this?

Response - The unit is using a motorized ball valve as switch valve? In this case, it's common to use many steam purge lines to avoid coke deposition over the seal faces of the valve and then it's necessary to evaluate and monitor the steam purges procedure to avoid an excessive decrease in the feed stream temperature to the reactors which can lead a high volatile matter in the produced green coke.

Question 59 - Currently we are facing problem of higher CO (Carbon Monoxide) and CO2 contents in the Net Gas (Net Gas is H2, produced in CCR Platforming Unit). This CO is affecting UOP's Penex Unit Catalyst. Please help finding out the source of CO & CO2 and guide how to reduce both in the Net Gas (our main focus is on reduction of Carbon Monoxide as it is a permanent poison for Penex Catalyst).

Response - It's important to understand better the refining scheme adopted in your refining asset, from the question it seems that the Net gas from CCR Platforming is directly fed to the Isomerization Unit (PENEX Process). If the net gas is not fed previously to a PSA unit or another hydrogen purification unit, it's expected a high concentration of CO and CO2 which are poison to isomerization catalyst. In this case, it’s necessary to study the economic viability to install a PSA or MEA treating unit aiming to reduce the CO and CO2 concentration in the Net Gas sent to the Isomerization unit.

Question 60 - How can we control the undesirable reactions in catalytic naphtha reforming (SRR Unit) (fixed bed reactors) in order to enhance the Octane Number and Hydrogen Purity?

Response - The most common side reactions in Catalytic Reforming Units is the hydrocracking which are favored by lower temperature (the hydrocracking reactions are exothermic), higher hydrogen pressure and lower space velocities. In this sense, a good way to minimize the risk of hydrocracking occurrence is to operate under higher temperature, lower hydrogen pressure and high space velocities. In other words, it's important to avoid the operation of the catalytic reforming with low feed flow rate and under low severity.

Question 61 - Crude Tank (Floating Double Deck Pontoon) is preparing to take crude oil (API 29) for the first time. Is it safe to put crude directly into the tank without Oxygen freeing? If water is first taken into the tank up to the floating level and then charge the crude, is it safer?

Response - It's possible to ensure a safety operation controlling the speed in the feed nozzle of the tank, there is some references which presents the maximum flow rate to ensure the filling the tank with controlled speed in order to minimize the vapour formation in the tank, of course, this flow rate depends on the diameter of the feed nozzle of the tank.

Question 62 - We faced problem in the particulate contamination specification of jet fuel product in the two tanks , and from our survey the specification from the unit of hydrobon unit is ok ,on the other hand in the refinery we refine sweet crude with API 42 and we got this problem for first time, in your view can you advice me?

Response -It's important consider in the analysis the information of the last cleaning of the internals of the storage tanks. It's important keep a routine of cleaning the storage tanks each five years in order to avoid the contamination of derivatives with water, sludge and corrosion products, this is specially true for jet fuel tanks.

Question 63 - Our desalter is facing a rag layer issue when we process cabinda crude. The brine turns black. It seems like our current emulsion breaker can not solve this problem. Is there any ideas or recommendations?

Response - According to the datasheet of the Cabinda crude oil, this is a light  and sweet crude oil which probably contains high amounts of paraffinic hydrocarbons. To realize an adequate analysis it's important to know if the refinery is processing only the Cabinda crude or under blending with heavier crudes, in this case we can saw chemical instability between the crudes leading to asphaltenes precipitation which stabilize emulsions reducing the separation efficiency in  the desalter vessels and provoke the change in the brine colour. In this case, it's possible to solve the problem by applying a crude stabiliser agent which is dosed independently of the emulsion breaker agent.

Another approach is analyze the incompatibility between the Cabinda crude with the another crude oils processed by the refinery and take anticipatory actions like reduce the processed flow rate in the crude oil distillation unit to ensure a higher residence time in the dessalters when processing a crude blending with high incompatibility potential, or avoid to process crude oils chemical incompatible with the Cabinda crude.

Question 64 - What are types of Corrosion Inhibitor (Filmer) and neutralizing amine used in crude distillation units. What is the philosophy of their work and the concentration of injection?

Response - According to the characteristics of the processed crude, the performance of desalters needs to be optimized, especially considering the concentration of magnesium and calcium chloride salts which tends to suffer hydrolysis and generate HCl which will concentrates in the overhead system of the atmospheric column, this is a special concern to refiners processing heavy crudes, slop blended with crude oil, and opportunity crudes.

Some refiners adopt an online injection of corrosion inhibitor in the overhead systems to control the salts precipitation. Normally, the corrosion inhibitor applied is a kind of filmic amine like alkylene polyamines (ethylene diamine - EDA, for example) which produces a protective film inside of the critical areas of the atmospheric column and neutralize the acid affect of decomposed salts from the crude oil.

Regarding the concentration of the corrosion inhibitor this vary case by case due to the characteristics of the processed crude, crude oils with high contaminants content and poor desalting performance like heavy and slop crudes tends to present higher concentration of NH3, HCl and amines in the desalted crude oil which raises the probability of corrosion and demands higher flow rates of corrosion inhibitor, it's recommended to look for specialized advice from a chemical treating company which will optimize the system to control de corrosion process under minimum flow rate of corrosion inhibitor as well as indicate what is the best corrosion inhibitor for your processed crude oil blend.

Question 65 - How can I know the water mole fraction in the overhead stream in the CDU? I need it to know the optimum temperature of the top refluxes.

Response - I believe that you can carry out a mass balance using the stripping steam flow rate and eventually another water injection to the atmospheric column to determine the water concentration in the overhead system indirectly, if you have a safe and adequate sample point you can make a chemical analysis to determine it. But it's important considering that the corrosion in the overhead system of the atmospheric tower is strictly related with the performance of the desalting system.

According to the characteristics of the processed crude, the performance of desalters needs to be optimized, especially considering the concentration of magnesium and calcium chloride salts which tends to suffer hydrolysis and generate HCl which will concentrates in the overhead system of the atmospheric column, this is a special concern to refiners processing heavy crudes, slop blendend with crude oil, and opportunity crudes. The adequate reflux temperature is fundamental to control the corrosion and fouling in the top plates of the atmospheric column which is caused by low reflux temperature and presence of NH3, HCl and amines which are absorbed by the water which vaporizes along the downward of the reflux leading to the salt fouling in the top plates which causes severe pressure drop leading to poor fractionating performance and can limit the the operating cycle of the processing unit.

Some refiners adopt the amine injection in the overhead systems to control the salts precipitation, especially H2S scavengers but the side effect here is the trend of deposition of corrosive salts under low temperature.

The literature highlights that the design of overhead systems needs to consider the probability of the corrosion and fouling in the top section of the atmospheric column, if the probability is high a overhead system with two separating vessels needs to be considered once to avoid low reflux temperature which can cause water condensation inside the column. A very good reference about this topic is the article published by Mr. Tony Barletta and Mr. Steve White in the Q3 2007 issue of PTQ Magazine.

Question 66 - For a desalter system with low performance, we made an RCA that revealed multiple issues to the desalter hardware, including that the equipment is undersized and electro-coalescence not happening, while grid is still ON. Could anybody advise equipment retrofitters to shift the existing desalter almost electrodynamic type to low velocity type ?

Response - Well, we have two types of electrostatic treaters which are known as low velocity type and high velocity type. The low velocity type is normally applied in platforms or in upstream assets aiming to promote the water separation from the produced crude oil, in this system the emulsion is fed in the top of the separation vessel along the length of the vessel and the emulsion is fed in the aqueous phase which due to the density gap presents an upward flow in direction to the electrodes which are normally positioned above of the centre line of the separating vessel. In these dessalters the crude oil is dispersed in the aqueous phase and suffers a kind of wash which remove salts and solid particles present in the oleous phase along of this upward flow in direction of the electrodes there is a small coalescence effect due to the low intensity electric field between the lower electrode and the water-oil interface. When the emulsion reaches the high intensity electric field occurs the remaining coalescence. These dessalters can suffer with poor performance along their lifecycle due to the raise in the water concentration in the crude oil along the time due to the depletion of the crude oil well and according to the strategy applied to improve the flowrate production of the well, and this can be one of the sources of the issue mentioned in the question.

In this case, the revamp of the desalting system for a high velocity desalter can be thought as a solution. In these dessalters, the emulsion is fed directly between the electrodes.

The limitation here is precisely the amount of water in the emulsion, for water concentration above 10 % the protection system can trip the desalting system due to the high electric current level and this limitation can be prohibitive to upstream assets due to the raise of water concentration in the crude oil along the time as explained above. For this reason, it's important to consider the service that you are planned to your desalting system, if the system is operating in an upstream asset, the revamp for a high velocity desalter can be a bad idea considering the trend of elevation in the water content in the crude oil along the time.

My suggestion would be to analyze the installation of a third electrode grid as well as consider the installation of a mud wash system in order to remove the sediments in the separating vessels bottom which can significantly reduce the residence time of the emulsion in the dessalter and lead to a poor performance of the whole system.

Question 67 - Can we add a valve Before Flash zone in CDU, to use it when we have a maintenance on the heater without the need to empty the tower?

Response - I believe that it's possible to add a valve in this region considering some criteria's.

The valve will need to be kept totally opened in order to avoid cavitation and consequently short operating lifecycle, especially if the crude oil distillation unit does not have a pre-flash tower. Another key point is that it's important to carry out a balance between the benefits and costs, taking into account that the valve to be installed can be significantly big according to the capacity of the processing unit as well as will add more leak points in a critical section of the processing unit which can lead a severe risks scenarios like fire and process safety accidents.

Question 68 - Regional shifts in higher refinery capacity seem to correspond with the need for more intensive water treatment programmes involving wastewater recycle processes while protecting heat exchangers and linked assets from fouling and corrosion. At what level of investment have you seen refinery operators commit to plant water quality while reducing its consumption?

Response - This is a key question to the sustainability of the crude oil refining industry as well as to the whole downstream business. We are seeing great efforts in the last decades to improve the sustainability of the oil & gas industry and no doubt, the adequate water management in crude oil refineries is a key part of the strategy to achieve this goal.

I believe that the investments in wastewater recycling in crude oil refineries tends to increase in the next years in compliance with stricter regulations aiming to minimize the water consumption which needs to be preserved for noblest purposes like human consumption. The water consumption can be a survive question for some refiners that can inclusive lead a refinery to be closed, an interesting article about this topic was published in the Q4 2008 edition of PTQ Magazine by KBC Company, the article details some actions and management program to optimize the water consumption in crude oil refineries.

Question 69 - I am currently working on fixed bed platforming unit. The unit is Semi-regenerative catalytic reforming. The unit was commissioned on 1989.  The unit is designed for Arabian light crude heavy naphtha. We have processed, Iranian light naphtha, Murban naphtha. Currently we are processing Arabian super light naphtha. I see that, the reactor delta T's are high enough (more than the one when process Arabian light), but gases production increased, hydrogen purity is between 70 to 43% and reformate yield is OK but RON decreases. The water chloride management is in range. With this, recycle gas compressor discharge temperature and pressure increases as gas production increases. The discharge temperature are so high (200-230F). Common discharge header pressure also increases. As recycle gas flow increases, reactor effluent trim cooler which is before HP separator temperature increases. The temperature must be in range of 100-105F. The current temperatures range between 120-140F. The gases production decreases and hydrogen purity also decrease. While sulphur and nitrogen are in limit and moisture and HCL is also in limit

What could the reasons of above query. What action should we take to resolve these problems.

Response - This is a relatively common situation faced by refiners which operate with semi-regenerative catalytic reforming processing units. To describe the phenomena that is occurring in this processing unit it's important to remember some concepts of the naphtha catalytic reforming process. In my response I'm considering that both naphtha are free of contaminants which can reduce the activity of the catalytic beds, especially considering which is a semi-regenerative processing unit.

One of the most relevant reactions which is carried out in the catalytic reforming of naphtha is the paraffin dehydrocyclization which involves the conversion of paraffin's in aromatics which contributes significantly to the octane index of the reformed naphtha. Unfortunately, these reactions are extremely slow and it is necessary to offer adequate residence time to ensure that the paraffin dehydrocyclization reactions occur.

Considering the scenario presented in your question, I understand that the change of naphtha from AL (Arabian Light Crude) for ASL (Arabian Super Light Crude) is raising the paraffin's content in the feed of the catalytic reforming unit and a highly paraffinic feed is very hard to processing, especially in a semi-regenerative unit.

Long chain paraffins (as tends to be the case of ASL crude naphtha) tends to suffer hydrocracking which involves the reaction of the paraffins with hydrogen to produce methane, ethane, and propane. These side reactions can be responsible for the reduction in the octane index and hydrogen purity, which is mentioned in your question, especially considering that paraffin hydrocracking is a quick reaction in comparison with paraffin dehydrocyclization.

To minimize this problem, it's possible to consider use a blend of naphthas in order to control the paraffins content the catalytic reforming feed, a very good factor to control the quality of the feed is the N + 2 A (Naftenics (%Vol) and Aromatics (% Vol)) parameter which should be controlled in the range required by the processing unit licensor. Another key factor is the initial boiling point of the naphtha feed, the IBP upper to 160 F is recommend for semi-regenerative catalytic reforming units once avoid the paraffin's hydrocracking which normally is favoured in high-pressure naphtha reforming which is a characteristic of semi-regenerative processing units.

My suggestion is to carry out a complete characterization of the naphtha feed from ASL crude to determine the paraffin content, IBP and N + 2 A parameters which should help to understand what is occurring in your processing unit. As mentioned above, a naphtha blending with a heavier naphtha can help to solve this situation.

Question 70 - We are experiencing a pungent foul smell in the polypropylene product pellets. The source of the same has remained untraceable having attempted various changes in the process, as required. Can someone please share a similar experience and the possible troubleshooting options?

Response - The literature presents some case studies related to this issue in plastics production, especially when recycled material is applied in the production process. In this case, the smell is attributed to VOCs (Volatile Organic Compounds) which are generated by the degradation of high molecular substances that were absorbed to the plastics during the use and these molecules that potentially cause odours are not removed during the wash step. For recycled plastics an alternative to minimize this issue is to include a degassing and filtration steps in order to avoid the odour production.

Generally, the plastics production even without recycled materials can face the odour issue in the granulation step which is normally associated with three factors:

1 - Decomposition of small molecules

2 - Use of miscellaneous materials

3 - Degradation of the material due to the multiple processing

The literature quotes that it's possible to deal with odour issues through the use of gas adsorbent and antibacterial agents in the plastics aiming to reduce the as well as exhaust system to eliminate residual odours from bacterial agents. Another critical factor to control the odour in the granulated plastic is the quality of the applied resin, poor quality resins have higher concentration of residual monomer which will lead to odours in the final polymer.

The other strategies to control the odour in the plastics is the use of adsorbent like zeolite material, use of an antibacterial agent, a desorption agent like activated carbon or simply add a fragrance to reduce the impact of the unpleasant odors.

Question 71 - What is the expected volumetric efficiency in the diesel product treating only SRGO? (It is understood that it is less than 103.4% due to the decrease in the content of aromatics and olefins)

Response - Volume swells in hydrotreaters are strictly related with the process severity applied once the volume gain is determined by aromatics and olefins saturation. In other words, higher hydrogen partial pressure and LHSV (Liquid Hourly Space Velocity) tends to raise the volume gain in typical diesel hydrotreaters, obviously the catalyst is also responsible for the volume swell.

Considering this fact, it's important to understand that volume swell is directly related with hydrogen consumption and it's important to quote that aromatics saturation is a reversible reaction under higher temperature which is a characteristic of hydrotreaters processing highly sour feeds, in this case the volume swell tends to be lower.

In summary, it's very difficult to precise the volume swell for a hydrotreater considering that this parameter depends on feed quality (more aromatics can lead to higher volume swell), catalyst, LHSV, hydrogen partial pressure, etc.

Question 72 – According to the Inspection Guidelines for Corrosion Control in Hydroprocessing Reactor Effluent Air Cooler (REAC), we need to ensure that at least 25% of the wash water is liquid. My question is how do we calculate it practically?

Response - This a fundamental issue to ensure adequate management of hydroprocessing assets, according to the literature and the API RP 932-B between 20 to 25 % of the wash water injected to the process need to remain in the liquid phase to ensure a real capacity to remove the NH4HS (ammonium bisulfide) and NH4Cl (ammonium chloride).

The wash water flow rate is calculated based on the feed flow rate of the processing unit, the literature quotes a minimum flow rate of 5,0% of the feed stream, but this depends on the design of the hydroprocessing unit. Most severe hydrotreating units processing heavier and high contaminants content of sulfur, nitrogen, and chloride tend to demand higher flow rates of wash water. In this sense, hydroprocessing units processing cracked feeds and residue will demand more wash water.

It's important to consider that a good parameter to estimate if the wash water flow rate is adequate is the concentration of NH4HS in the sour water which can be measured in the separator vessel, again according to the specialized literature, the concentration should be around 6,0 to 8,0 % (maximum). Another way to verify the quantity of wash water injected is to measure the free water flow rate downstream of the injection point.

Further the discussed above it's important to consider a verification of another important topics related to the wash water injection system as described below:

- The presence of oxygen in the wash water can cause corrosion and the oxygen concentration in the wash water should be below than 50 ppbw;

- It's important to ensure symmetry in the piping arrangement of the air coolers in order to ensure adequate wash water distribution and non-flowing sections which accelerate corrosion;

- The velocity in the tubes needs to controlled aiming to avoid the corrosion-erosion phenomena, according to the literature the velocity in tubes should be controlled in the range of 3,0 to 6,0 m/s;

- At last, taking into account the chloride concentration in the feed. Chloride can lead to corrosion due to HCl formation in aqueous phase and accelerate the NH4Cl corrosion and fouling.

Question 73 - We have a problem in one of the main towers (capacity = 150000 bbl/day) in the company I'm working for. There is an inclination in the tower which may affect the efficiency of separation. So, what's the maximum allowable inclination so that no effect in the separation efficiency may occur?

Response - It's important to consider that distillation columns with industrial applications with high capacity like described in your question normally have an inclination grade. The fractionating efficiency will be affected according to the diameter of the distillation column, once higher diameters will lead to a most severe disequilibrium in the liquid holdup in the fractionating stages which will produce an accumulation of liquid in the inclined section of the fractionating stage with higher pressure drop, preferential flowing of the vapor phase and then a reduction in the fractionating efficiency. This phenomenon is called vapor-liquid channeling by specialized literature and a good reference about this topic is the book "Working Guide to Process Equipment" by Norman P. Lieberman and Elizabeth T. Lieberman (Fourth edition, 2014). Based on this reference, it's possible to conclude that is not rare to identify a 1 ft out of level in a distillation column with 14 ft diameter ( I believe that 3 % of inclination (based on the column diameter) would be acceptable) and this out of level can be compensated through raising the pressure drop of the liquid flowing through the orifice holes which demand the shutdown of the distillation processing unit. It's important to quote that any inclination will affect the fractionating performance due to the mechanism mentioned above, it's possible to find a tolerated point of performance reduction, but the inclination should be solved as soon as possible to return the processing unit to their optimized point. Despite this, I believe that the most important concern should be the stability of the structure considering all efforts and wind load to ensure the process safety requirements.

Question 74 - We are an Indian refinery and recently commissioned our full conversion VGO hydrocracker unit. With in 2 months after start-up we observed higher COT's in one of the heater passes. Our heater is 4 pass heaters. What could be reason for this?

We carried out flushing with high gas/liquid flow rate with jerks as well. still dp across the pass is very high. Finally, we wanted to carry out pigging as issue still persists.

What could be reasons for this and how to avoid such scenarios in future. Request to share your ideas and similar experiences.

Response - This phenomenon can be related to the start-up procedure of the hydrocracking unit. It's important to check if the licensor recommendations and procedures were totally accomplished specially related to the heating rate of the feed and the feed quality, especially related to the hydrogen quantity and presence of heavier compounds that could accelerate the coking process of the heater pass including the presence of coking agent like sodium. Another reason can be mechanical damage in the heating coil which can produce preferential flow in some passes and accelerate the coking process in the low velocity pass.

Another action can be checked if there is no flame impingement over the furnace tubes which can be caused by low air to the flame or burner tip fouling.

If these points were checked and everything is considered normal, considering the experience of the refining industry related with similar scenarios (Irving St. John Refinery in 1998), it's recommended to shut down the hydrocracking unit to investigate the causes and avoid a potential process safety accident due to the tube rupture during furnace operation.

Question 75 - I am currently working on a decommissioned polyethylene plant. Can anyone suggest a good reference, and or, databases or applications for estimating cost of turnaround, pre-comm, comm and re-start?

Response - This is a very important topic which unfortunately it's difficult to find detailed references, but there are very interesting and high level publications in the specialized literature. It's possible to find a short list of references about this topic:

1:  Book "Turnaround Management" by Tom Lenahan published in 1999.

2: Book "Estimating Construction Costs" by Peurifoy, R.L. and Oberlender, G.D. published in 1989.

3: Book "Process Plants: Shutdown and Turnaround Management" by Trinath Sahoo published in 2013.

4: Article "Cost Estimating for Turnarounds" by Gordon Lawrence published in the Q1 2012 of PTQ Magazine.

Question 76 - One of our Unit has four furnaces viz. F-1, F-2, F-3 & F-4. There are two fire boxes: one for F-1/F-2 and one for F-3/F-4. Both fire boxes have two parallel convention banks and a common stack. The furnaces have the issue of low steam generation from convection banks, high stack temperature, BFW flow mal distribution in convection banks and hence a lower efficiency than design. Sketch of the furnaces with current & design flows are provided below for reference. What could be the possible reasons and remedial solutions.

Response - The question does not specify what kind of processing unit provides the flue gases (Catalytic Cracking, Distillation, etc.), but I believe that it's possible to consider some verifications generally.

It's important to consider if the furnaces have soot blowers, if the response is yes, how is the availability of these devices? The deposition of soot or ash in the furnace wall and specially in the convection tubes (which are generally finned) can significantly reduce the thermal exchange in the furnaces, leading to a poor thermal efficiency and the main side effect is a high stack temperature like mentioned in the question. This is especially true in furnaces using FCC flue gases due to the carryover of catalyst fines in the flue gas which will be deposited in the furnaces wall and convection tubes.

Another key point is the distribution of BFW (Boiler Feed Water), which seems to suffer with a poor distribution. This is a key factor to the performance of the furnace and the flow and symmetry criteria needs to be accomplished in order to reach the expected design. A good strategy to identify poor flow performance of gases and BFW is a CFD (Computational Fluid Dynamics) study using a simulated model of the furnace.

A general review of the design with verifications of fin distribution on the convection tubes, their area, and the diameter of the convection tubes facing the eventual operating changes of the process conditions of the furnace over the years (capacity raising) should help to understand the poor performance of your furnace system.

Question 77 - What is the best way to reduce HCGO end point, apart of increasing the flow on the sprays?

Response - Based on your question it seems that your refinery directs the HCGO (Heavy Coker Gas Oil) to a hydrocracking unit once normally these streams are normally only injected under low flow rates to FCC units due to their poor crackability once the stream already suffered thermal cracking. I believe that it's possible to achieve this goal (reduction of the end point of the Heavy Coker Gas Oil) through raising the flow rate of the inferior pumparound of the fractionator, this will lead to a reduction in the bottom temperature of the tower which can limit the feed stream of the delayed coking unit if the fired heaters do not have sufficient thermal charge to compensate this reduction.

Obviously, if the main fractionator have side stripper column to HCGO stream, it's possible to cut or reduce the flow rate of stripping steam to help to reduce the end point of the stream, but normally these system are not installed or are not operated once the main fractionating tower is able to reach high flash point for this stream being unnecessary to operate the stripper column, furthermore, this will improve the energetic efficiency of the delayed coking unit. Another key parameter to reduce the end point of HCGO is to raise the recycle ratio of the delayed coking unit, but again this can limit the capacity of the delayed coking unit. In summary it's not easy to achieve this goal without improving the internal recycling, an alternative is to blend the HCGO stream with lighter and cleaner streams like straight run gas oil before feeding the consumption unit (hydrocracker or FCC).

A very good article about this topic was published by Mr. Scott W. Golden in the 2003 issue of the Revamps Magazine (a supplement of PTQ Magazine).

Question 78 - I have three questions: 1) What reactions are source of high hydrogen purity at naphtha platforming (catalytic reforming) unit? Either naphthene's or paraffins reactions?

2) The NHT-Platformer unit was down for couple of days. The startup was executed after 1 week. After startup it was observed that, system pressure was not meeting the set point of 350 psi. But in previous history, wherever startup was performed, at turndown capacity the system pressure of 350 psi was meet. This is it’s not even reaching to system pressure of 350 psi. While hydrogen purity is also decreased up to 60%. HCL is recycle gas is almost ranges between 0.5 to 1 ppm and H2S is also 1 ppm. During startup at 880F, the HCL was found in traces while H2S was found 2ppm. Platformer catalyst is UOP-R56.

3) Furthermore, about NHT, it's not removing sulfur properly even though we have done skimming recently. Due to low hydrogen purity of platformer, the ratio is limited to 340 to 350. While ratio must be 380. The reaction temperature is 630 F (which is 5 F higher then EOR for catalyst). Still, it's not removing properly, the stripper is operating and minimum pressure and maximum bottom temperature.  NHT catalyst is UOP-HYT-1119 Can you tell me about this to improve sulfur removal currently?

Response - Regarding your questions, I have the following answers and suggestions:

1 - The main reactions of the naphtha catalytic reforming process is the naphthene dehydrogenation to aromatics, alkyl cyclopentanes dehydroisomerization to aromatics, alkanes dehydrocyclization to aromatics, n-alkanes isomerization to branched alkanes, and cracking reactions to lower carbon products. The first three reactions are responsible to produce hydrogen, the paraffin dehydrocyclization which involves the conversion of paraffin's in aromatics which contributes significantly to the octane index of the reformed naphtha. Unfortunately, these reactions are extremely slow and it is necessary to offer adequate residence time to ensure that the paraffin dehydrocyclization reactions occur.

Long chain paraffins tend to suffer hydrocracking which involves the reaction of the paraffins with hydrogen to produce methane, ethane, and propane. These side reactions can be responsible for the reduction in the octane index and hydrogen purity, which is mentioned in your question, especially considering that paraffin hydrocracking is a quick reaction in comparison with paraffin dehydrocyclization.

2 - It's important to analyze how was the shutdown procedure of the processing unit, despite to processing a high quality naphtha as feed stream an inadequate shutdown procedure can lead to a coke deposition over the catalyst, considering the informed catalyst type (UOP-R56) your processing unit is a semi-regenerative reforming unit which demands even more attention with the risk of catalyst bed coking during the shutdown and stop period, it's important to check the operating manual of the processing unit to verify the licensors orientation regarding the attention points before to carry out a planned shutdown of the processing unit. For semi-regenerative processing units, it's fundamental an adequate management of water/chloride ratio aiming to ensure an adequate balance between the acidic and metal functions of the catalyst. Normally, fresh catalytic reforming catalysts presents close to 1,0 % wt of chloride, to maintain this chloride concentration it's necessary to control the water concentration aiming to allow an effective interaction between the alumina (catalyst support) and the chloride, reaching then a good performance of acidic sites of the catalyst which is responsible by the cracking reactions.

According to the literature, several factors impact the chlorides concentration in catalytic reforming catalysts. The reactor temperature and surface area of the support can directly affect the chloride concentration in the catalyst and are the most relevant factors. Still according to the literature, fixed bed catalytic reforming reactors should operate keeping the water to chloride molar ratio between 15 to 25 in the recycle gas aiming the keep the activity of the catalyst, to control this parameter it's necessary to install sample facilities or online monitoring systems in adequate points aiming to keep this parameter according to the licensor specifications. It's possible to find in the specialized literature chlorides equilibrium curves capable of helping the refiners to control the water to chloride ratio in the catalyst under the specifications defined by the licensors. The presence of only traces of HCl in the recycle gas can indicate that the chloride concentration over the catalyst is low, reducing the acidic function leading to an unbalance between the acidic and metal functions of the catalyst. A suggestion is to carry out a regeneration procedure of the catalyst beds in order to recovery the catalytic activity as recommended to the licensor, another point to check is the paraffinicity of the feed, high paraffins concentration can lead to the reduction of hydrogen purity and RON of the reformate if the severity of the processing unit was not adjusted, as previously mentioned, the paraffin dehydrocyclization are extremely slow and it is necessary to offer adequate residence time to ensure that the paraffin dehydrocyclization reactions occur. Long chain paraffins tend to suffer hydrocracking which involves the reaction of the paraffins with hydrogen to produce methane, ethane, and propane.

3 - Considering the given information regarding the catalyst (UOP-HYT-1119), this is a Ni-Mo catalyst with high activity and high hydrogen consumption. The poor performance related to the hydrodesulphurization can be related to the low hydrogen/feed ratio caused by the low hydrogen purity, leading to an insufficient hydrogen partial pressure to the hydrotreating reactions. Another point which needs to be checked is the quality of the feed, if the final boiling point of the feed was increased, it will be observed difficulties to remove nitrogen and sulfur.

Question 79 - In our CDU we have stabilizer and splitter columns, stabilizer for separating LPG from Naphtha, after the annual maintenance, we have a problem in the boot of the overhead drum of stabilizer column we have a Black water and high iron number so what's the problem that makes this black water?

Response - The black sludge and water observed in the overhead vessel is probably pickering emulsion stabilized by iron particles which is accumulated in the interface between sour water and naphtha. It's important to check the pH, Iron and Chloride content in the overhead system, a lack of control of these parameters can lead to a severe corrosion process in this system as well as in the atmospheric tower. When the emulsion is formed in the vessel, this residue can’t be removed without the shutdown of the processing unit or through draining the overhead vessel totally which requires a special procedure aiming to minimize the safety risks as well as the damage to the pumps of the overhead drum. Regarding the corrosion control in the overhead systems it's important to analyze that the corrosion control parameters is under an adequate range, especially the operating temperature of the overhead system. There are some correlations in the literature which relates the ammonia and chloride concentration in the sour water to determine salt deposition temperature in the top of the tower and this needs to be considered to define the operating temperature of the system.

Question 80 - How to prevent gumming or carbon formation in pre-reformer catalyst aside from maintaining an inlet temperature? Does hydrogen recycle in the pre-reformer beneficial in preventing gum formation?

Response - This is a common concern for operators of hydrogen generation units, especially for those applying naphtha as feed.

Well, the available routes to control the carbon deposition of the pre reforming catalyst are listed below: 1: Operating respecting the temperature ranges (considering the question, I understand that this is being respected); 2: Increase the steam to carbon ratio; 3: Improve the basicity of the catalyst through addition of promoters like potassium; 4: Improve the partial pressure of hydrogen in the reactor; 5: Maximize the gas flow rate in the reactor aiming to improve the heat flux and avoid hot points which tends to accelerate the carbon laydown over the catalyst.

Considering the action 4, recycle hydrogen to the reactor can reduce the carbon laydown over the catalyst but it's necessary to study adequately this strategy aiming to avoid the contaminants concentration (mainly CO) in the reactor which will lead the side and deleterious effects like the Boudouard and Beggs reactions presented below:

2CO = C + CO2 (Boudouard Reaction)

CO + H2 = C + H2O (Beggs Reaction)

Question 81 - The Hydrogen Production Unit (SMR) is designed for both natural gas and naphtha. The sulfur in naphtha feed is in limit. My queries are: 1) What will be effect of PONA of naphtha on hydrogen production and methane production?

2) What will happen if naphtha feed contains more amount of naphthenes (more then permissible)?

3) What is possiblity of conversion of naphthenes into methane?

Response - I believe the question is considering that the operating conditions of the SMR unit is considered constant and only the feed characteristics is changed (from natural gas to naphtha), in this case we have the following scenarios:

1: With high content of heavier hydrocarbons (aromatics and naphthenes) the production of methane tends to be high and the hydrogen falls in the syngas produced;

2: Keeping the same operating conditions (temperature, pressure and steam/carbon ratio, the methane concentration will be high and the hydrogen production will fall;

3: To improve the conversion for heavier feeds it's necessary to improve the steam/carbon ratio as well as maximise the temperature and reduce the operating pressure as low as possible. Despite these alternatives, maximising temperature and reducing operating pressure can be difficult to carry out in practice, the best way to improve the conversion in this case is to maximise the steam/carbon ratio. Here, it's fundamental to taking into account that operating with heavier feeds (naphthenic naphtha) will raise significantly the carbon deposition issues over the catalyst once the side reactions like hydrocarbons pyrolysis tends to take place and you can face severe troubles with operating lifecycle of SMR unit. It' important to carry out a sensibility study to determine which is economically attractive to install a naphtha splitter upstream the SMR unit and operate the unit with lighter feed.

Question 82 - We having caustic regeneration facility in our LPG treating unit. During caustic regneration, Disulphide oil (DSO) will be removed by absorbing with help of naphtha in CFC. We are experiencing higher levels DSO in the regenerated caustic.

My questions are :

1. Does LPG quality will be impacted if regenerating caustic is higher levels of DSO. (Design says NIL)?

2. What the possible ways to reduce the DSO content in regenerated caustic?

3. Is there any correlation on DSO based on LPG inlet mercaptans?

Response - Based on the information in your question it seems that the LPG sweetening system is overloaded possibly due to an increase of overall sulfur content in the feed stream. In this case, it's important to verify the performance of the amine treating system upstream of the LPG sweetening unit once higher H2S content in the LPG will overload the caustic treating and this will lead to a higher production of disulphide oil (DSO). This higher production of DSO can impact the LPG quality if the caustic regeneration system is operating under poor performance or inadequate make up caustic flow rate is carried out.Another key question is regarding the solvent applied in the CFC (Continuous Film Contactor) section, the naphtha applied as solvent to wash the DSO needs to present low sulfur content (the ideal is apply hydrotreated naphtha) once the caustic can react with the mercaptans present in the solvent producing mercaptides and raise the sulfur content in the LPG.Among the operating variables which can help to reduce the DSO amount in the regenerated caustic are the temperature and the air flow rate to the oxidizer.If was observed a significant change in the processed crude oil slate for a high sulfur content crude oil and this will be permanent, should be considered the installation of a caustic pre-wash vessel in the LPG sweetening unit in order to reduce the H2S content in the feed and then improve the performance of the caustic treating.

Question 83 - A CO2 removal unit with MDEA solvent has experienced severe foaming issue in the regenerator which has led to solvent loss to the regenerator vent. When foaming happens in the regenerator, observed high/ fluctuating DP across the packed bed and wash water tray (bubble cap trays) on the regenerator section. On the absorber side, the DP maintain stable, and no solvent carry over to downstream vessel during foaming incident. However, the foaming in regenerator has led to inefficient solvent regeneration and caused high CO2 breakthrough at the absorber overhead (treated gas).

The process gas to the absorber is mainly coming from the syngas produced from upstream Steam Methane Reformer unit with Natural Gas feed. The Process Gas to absorber mainly composed of CH4, CO, CO2 (7-8 mol%) and H2.

Analysis on the cause of the foaming incident in the MDEA regenerator is suspected due to solvent contamination with particulate matters as contamination due to long chain hydrocarbon is not possible in this process. MDEA solvent has been analyzed with the TSS has been observed between 3 - 10 mg/mL. The solvent appearance remains clear without any coloration which would indicate solvent degradation. Lab analysis has also shown low HSS and no signs of solvent degradation.

Hence, the way forward to avoid the foaming in the regenerator is to replace the filter element of the side stream filter from 10 micron nominal to 5 microns absolute. This is to ensure small particulate matters are sufficiently filtered during normal operation with 5 micron absolute filter.

However, would like to check if following parameters can also cause amine foaming inside the regenerator: Can over stripping from the regenerator reboiler caused turbulence and foaming especially at the rich amine inlet to the regenerator feed gallery? Rich amine at inlet of regenerator is located below of the wash water trays (bubble cap trays). Some amine would expected to be entrained with CO2 to the bubble cap trays. Should the antifoam be injected at the regenerator reflux line to break down the foam which could build up at the bubble cap tray section?

Appreciate your feedback/ thoughts on this.

Response - Regarding the first question, I understand that over stripping is related to superheating of the steam fed to the reboiler of the regenerating tower. In this case, the main effect or risk is the thermal degradation of the amine, some references quote that the liquid temperature in the reboiler should be kept below 125 oC with a low pressure steam temperature close to 150 oC. I understand that a possible overheating can cause turbulence which favors the foam formation, but in this case the thermal degradation of the amine solution is a more significant risk. Furthermore, the stripping columns have a lower tendency of foam formation due to the operating conditions, the lower pressure and higher temperature which reduces the surface tension of the amine solution minimizing the filming tendency of the surfactants compounds. Another key point to be verified is regarding the presence of oxygen in the amine solution which can lead to the formation of carboxylic acids raising the rate of chemical degradation of the amine solution, the antifoam and amine tanks can be a possible point of oxygen to the amine system.   About the second question, it seems that it can be a good strategy but it's necessary to consider if this will be an additional antifoam injection or if the point of antifoam injection will change for this point. In this case, it's possible to reach an over injection of antifoam which can cause more foam. Another key point which is not mentioned in the question is related to the operating temperature of the absorbing tower, the literature recommends that the temperature of the amine solution should be 5 to 10 oC above the inlet temperature of the gas. A very good article about this topic was published in the September 2007 Issue of PTQ Magazine by Mr. Stephen A. von Phul and Mr. Arthur L. Cummings.

Question 84 - In a crude processing unit, we have two stages of separation (gas, oil & water) and two stages of desalter. I need to know if we should heat the crude between two separators or before the first desalter? and why?

Response – Well, this is a very interesting question, an increase in the process temperature has two antagonic effects which need to be balanced aiming to maximize the separation and desalting efficiency. The crude oil heating will reduce the density and viscosity which will settling rates of the water droplets in the oil phase allowing a higher processing capacity of the water-oil separating system. On the other hand, the temperature increase will raise the conductivity of the crude oil, demanding a higher power consumption to promote an adequate desalting. Some studies point out that the adequate temperature for desalting is above 140 oC but it can be optimized according to the characteristics of processed crude oil once heavier crudes tend to present hard separation and desalting characteristics due to the lower density gap between water and crude oil. I believe that you can carry out an economic and energetic integration study aiming to heat the crude upstream the water-oil separating system to a intermediate temperature to improve the water-oil separation and   promote another heating to a higher temperature aiming to maximize the performance of desalting process considering the limitation effect over the conductivity of the crude oil. Based on the information in your question, I believe that the best location for the heat exchanger is between the separators aiming to maximize the performance of the water-oil separating system and avoid an eventual overload of the desalting system.

Question 85 - In our plant, DSN tank exists for cold feeding to CCR unit in case of interruption in u/s hot feed from NHT unit or during total power failure start-up of Refinery. Cold DSN is fed to NHT stripper and thereby to CCR unit via Depentanizer and Re-run columns to produce Hydrogen for NHT start-up. Once Hydrogen is produced, NHT feed in is done. Stripper bottom is recycled back to FSD till DSN is on spec. When DSN is on-spec, feed from DSN tank is isolated. How to avoid poisoning of CCR catalyst while keeping NHT running with off-spec DSN? Further, can DSN from Tank be sent directly to Depentanizer bypassing Stripper in absence HP nitrogen input in Stripper? What are the precautions to be taken for operation of CCR unit?

Response - Well, to respond adequately the question would be fundamental to know the contaminants level of the desulfurized naphtha (DSN) and desulfurized off-spec naphtha.

Some licensors recommend that the naphtha fed to the catalytic reforming units should be hydrotreating aiming to achieve a zero concentration of sulphur in order to ensure a very low concentration of other contaminants like nitrogen, oxygenates, etc. It's not clear in the question if this is carried out in this case, but a side effect of this deep desulfurization is that the absence of sulfur in the catalytic reforming feed can led to metal catalyzed coke (MCC) in the tubes of fired heater which migrate to the reactors and can affect the catalyst circulation and the performance of the unit. To avoid this, the licensors recommend the injection of a controlled amount of organic sulphur (like DMDS) to the feed aiming to ensure minimum sulphur concentration (between 0,25 to 0,50 ppm). It's fundamental to know the concentration of sulphur and nitrogen in the off-spec feed quoted in the question to conduct an adequate decision making process, the contamination by sulphur is reversible and the catalyst performance can be recovered after some regeneration cycles after the contamination stops. Regarding the nitrogen, the high concentration can affect the acid function of the catalyst which can be compensated through a higher chloride dosage in the regeneration section but the main concern here is the salt formation (NH4Cl) in the cold areas of the unit like condensers, separation and compression sections.

Among the actions to protect the catalyst against temporary high contaminants concentration it's possible to reduce the process severity and the flow rate of the unit as well as raise the dosage of chlorination agent (normally Perchloroethylene), as previously mentioned, to maintain the acid function of the catalyst.

Regarding the stripper bypass, I don't have knowledge about this operation mode once this action will add sour gas (high H2S concentration) to the catalytic reforming feed and will overload the depentanizer column with a very low initial boiling point naphtha which can lead to higher gas production as well as raise the benzene concentration in the reforming naphtha. Furthermore, a higher concentration of H2S in the feed can produce deleterious effects over the catalyst as mentioned above.

Question 86 - How can cycle time be optimized in needle coke production while refurbishing a fuel grade coker?

Response - Considering the growth of electric vehicles and consequently the demand for batteries, the search to produce needle coke tends to rise. The needle coke production demands specific operating conditions of the delayed coking units like higher pressure in the reactors (around 4,0 to 8,0 kgf/cm2), higher temperature (500 to 510oC), higher recycle ratios (above than 60 %), and higher cycle time. A typical cycle time for fuel grade cokers is around 16 to 24 hours while e needle grade coker will demand cycle times around 32 to 36 hours aiming to ensure that the gas will be released from the solid coke creating the cavities which will ensure the anisotropic structure, required to the needle coke. It's impossible to produce needle coke applying conventional feeds like vacuum residue, it's necessary to feed high aromatic streams to the delayed coking unit. The decanted oil from FCC is one of the best alternatives to produce high quality needle coke.

Question 87 - Lean Amine (MDEA) from Amine Regenerator bottom pre-heats the feed to regenerator and then is pumped to users while passing through Air Fin cooler (AFC). A slip stream is sent to pre-coat filter and filtered MDEA joins the rundown. We are facing high MDEA rundown temp (>75 degC) due to internal fouling of the rundown Air fin cooler (AFC) which is designed to reduce MDEA temp from 104 degC to 60 degC. MDEA rundown results is as follows: Fe (5-6 ppmw), HSS (1.5-2.0), Na (400-500 ppmw), TSS (200-400 ppmw). There is no trim cooler at downstream of AFC. There is a HSS removal from MDEA skid which was recently commissioned in Mar'23. What could be the reason for fouling of AFCs? Is there any reference of online cleaning of Lean Amine AFC without impacting the quality?

Response - Considering the analysis results for lean amine presented in the question it seems that your processing unit is suffering corrosion and fouling process due to high concentration of Heat Stable Salts (HSS). It's possible to see a high iron concentration (5 to 6 ppmw) in the lean amine which evidentiate corrosion process and the concentration of HSS is around 3 to 4 times of the recommended by the literature (0,5 wt % maximum). The heat stable salts (HSS) tends to hold the iron in solution avoiding the creation of passivating film which exposes the metal to corrosion process, the sodium concentration indicates that is not occurring caustic over injection to neutralize the heat stable salts, a suggestion is analyze if the current dosage is sufficient. Another suggestion is to analyze the efficiency of the lean amine filter and, according to the result, change the filter element mesh to stricter values (a typical case is to reduce the filter mesh from 10 to 5 micron). Regarding the online cleaning process of AFC, I don't have knowledge about these technologies.

Question 88 - With the most profitable refiners focusing on the production of basic chemicals such as aromatics, olefins, and polyolefins, what catalyst and reactor technology is key to this focus?

Response - Considering the growing demand by petrochemicals and the hostile and falling market for fossil fuels is expected an movement by the big players of the downstream industry aiming to promote closer integration between refining and petrochemicals assets in order to raise the yield of petrochemical intermediates like ethylene, propylene, and BTX in their refining assets. For downstream players with sufficient capital power it's possible to expect the announcements of crude to chemicals capital investments like recently announced in Asia and Middle East. Considering this context, is expected that the conversion processes in general will reach a more strategic role in the refining assets and processing units like FCC, Steam Crackers, Catalytic reforming units, and hydrocrackers should deserve more attention from refiners and is expected more capital investments in revamp and green projects in the next years, the same can be said about the catalysts applied in these processing units regarding new technology development and operating costs related to a more intensive use of these catalysts.

Question 89 - What are the positive and negative implications on coker operation associated with the introduction of FCC slurry oil into the feed?

Response - The FCC slurry oil is a highly aromatic stream and can produce high quality coke like the needle coke which presents high added value, but it depends on several process variables adjustments like operating pressure of the reactors and higher cycle time. Due to the characteristics of the FCC operation, the FCC slurry oil tends to present high ash concentration caused by inefficient separation between the catalyst and hydrocarbons in the FCC separating vessel leading to concentration of catalyst powder in the heavier fractions which are recovered in the bottom of the main fractionator of the FCC unit.

Among the negative effects of fed FCC slurry oil to delayed coking units it's possible to quote the lower liquids and higher coke yields in the processing unit, higher tendency of coke deposition in the fired heaters leading to shorter operating life cycle of the delayed coking unit, it's possible to occur higher powder accumulation in the bottom of the main fractionator.

The best way to minimize the deleterious effects of the FCC slurry oil in the delayed coking unit is filtering the stream before fed the stream to the delayed coking unit as well as ensure adequate settle time in the FCC slurry tank aiming to ensure adequate separation between the solids (catalyst powder) and the hydrocarbon phase, this process can be optimized using chemical products like ash reductors which can be dosed in the processing unit (FCC).

Question 90 - What is the typical particle size for HCGO Backwash filters? The frequency of the HCGO backwash filter has increased significantly when the wash oil amount is decreased in the unit. I have checked the data sheet which states that coke particle size>20 microns. Have you encountered any issues with HCGO backwash filters before? If the problem is not related to the filter size, what other factors might be causing it?

Response - Unfortunately, issues related to HCGO (Heavy Coker Gas Oil) filters are not uncommon in the refineries. If the HCGO feed is coming cold from a storage tank, it's possible to face high pressure drop due to high viscosity, due to be a chemically unstable feed the the storage tank should rely with a nitrogen blanketing to avoid the olefins oxidation which will raise the pressure drop in the filters due to gums and polymers precipitation. Another key point is the operating performance of the main fractionator of the delayed coking unit. Disturbances in the delayed coking unit can cause coke fines dragging to the downstream processes, an alternative is to check if the wash oil flow rate of HCGO fractionating section is adequate, lower flow rates can cause coking deposition and worse quality of the stream. Other points to be checked are the backwash flow rate as well as a good characterization of the particles, especially to determine if the main issues are the organic or inorganic particles. If there was a significant reduction in the feed temperature to the HCGO filter system this also can lead to poor performance and the process needs to be checked to get around this issue.

Question 91 - In the event of an emergency shutdown of the FCCU, coke in the reactor and reactor cyclone may be peeled off. On the other hand, a significant period of time will be required to remove the coke. In the event of an emergency shutdown, how should we determine whether or not coke removal is necessary?

Response - Coke deposition in the internals of fluid catalytic cracking units like separator vessels are, unfortunately, a common issue, especially for those units processing atmospheric residue (RFCC) which presents high carbon content in the feed. During emergency shutdowns, the instability and transient operation conditions can lead to coke detachment from the cyclones, for example, leading to the poor separation efficiency of the catalyst which will raise significantly the particulate emissions to the atmosphere, leading inevitably to the shutdown of the processing unit for maintenance considering the current environmental regulations and ESG policies. Some refiners are applying the strategy to install devices in regions where historically are found coke deposition capable to sustain and attach the coke, avoiding the falling over the equipment internals. These devices can be designed according to the historical of coke deposition in the processing unit and are installed during the planned shutdown for maintenance. Regarding the question of remove or not the coke during an emergency shutdown, this should be a technical and economic decision considering the performance of the processing unit previously of the shutdown, availability of maintenance resources and materials, and the operating life cycle of the processing unit.

Question 92 - Why should water boot be vertical? We want to implement water wash of overhead condenser, but 3 phase separator is bottleneck, especially water boot. We cannot extend it vertically, to get more residence time for water (space limitations). Potential option is to extend boot horizontally - to increase its DIA to bi as wide as vessels, or to change boot with another smaller vessel, to provide residence time. Are these options even applicable or must hold on to already proven solution, like vertical boot?

Response - Three phase separation vessels can be vertical or horizontal, vertical vessels are mainly applied when there is a large amount of vapor to be separated from a small amount of the liquid fluid. According to the literature (the classic article from W.Y. SVRCEK & W.D. MONNERY, 1994 is a good reference) describes that the boot is applied to three phase separation vessels when the heavy liquid (in this case, the water phase) amount is lower than 15 to 20 % in weight and a weir is applied when the heavy liquid volume is high. Raising the boot diameter can be a good strategy to improve the residence time and reach a better separation between the heavy and light liquid, but the boot diameter should not be higher than 50 % of the vessel diameter aiming to ensure adequate liquid hold up time. The minimum boot diameter can be found in the literature and the length of the boot should not be less than 900 mm to provide adequate hold up time to the heavy phase and residence time to separation. As previously mentioned, the specialized literature presents guidelines capable of helping to solve this question in a good way, once the modification will be conducted, maybe it is a good moment to verify the whole design of the separation vessel aiming to debottleneck this system. Another suggestion is related to the washing water injection system which will be installed, it's important to take care with this system especially considering the adequate dispersion device (sprinkler nozzle or static mixer) aiming to ensure adequate dispersion of the water and avoid corrosion-erosion phenomena in the pipe or heat exchanger.

Question 93 - A TGT (Tail Gas Treater) that uses a steam heater to raise the temperature is experiencing a drop in quencher pH after shutdown maintenance and startup. In this TGT, passivation was done at shutdown for shutdown maintenance, and pre-sulfurization was done at startup and startup. The TGT reactor has thermometers at the top, middle, and bottom, and since startup, the temperatures at the middle and bottom have been the same, indicating that the catalyst in the lower layer is losing activity. The operation to raise the pH of the quencher with caustic soda had to be repeated two to three times a day, and by raising the H2S/SO2 of the SRU upstream to 5, the SO2 leak from the reactor was reduced, and the quencher pH drop is now only once a week. No unusual operations have been performed. The catalyst has been in use for 10 years, which is not a short period of time, but we do not believe it is a period of use during which a rapid decline in activity occurs. If anyone has experienced similar problems or has knowledge of the possible causes of the problems, it would be helpful.

Response - The low pH informed in the quench water indicates higher concentration of SO2 in the hot gas from the hydrogenation reactor that feeds the quench tower due to the sulphuric acid formation in the water phase. This fact can be understand as a symptom of poor performance of the hydrogenation section that is reinforced by the absence of temperature variation between the middle and bottom catalytic beds of the hydrogenation reactor informed in the question.

Once was mentioned a recent maintenance shutdown, I understand that the catalyst preservation and activation procedure during the shutdown and start up steps should be verified aiming to identify eventual failures. It seems that the catalyst was damaged or poisoned during the shutdown or start up procedure, leading to activity loss which could explain this behaviour.

Questão 94 – How much sodium is the limit expressed in mg/kg that can be loaded into a Visbreaker system without causing excessive scaling or coking in the oven?

Response - Unfortunately, I don't have much experience with the visbreaking process but considering the similarities with delayed coking process I believe that we can use the same considerations. Normally, it's considered a safe range of sodium concentration in the feed of delayed cokers between 10 to 15 ppm (mg/kg) in order to control the deleterious effects regarding coke laydown in the fired heaters tubes. The sodium concentration should be controlled at lower level as possible, but this is a trade-off nowadays considering the necessity to inject caustic soda in the crude upstream and downstream of the desalting system aiming to control the salt deposition in the crude oil distillation units as well as downstream units like hydrotreaters and delayed cokers.  There is a growing concern related to chloride concentration in the feed of processing units which can lead to severe corrosive process due to salt deposition, this is a special issue for refiners processing crudes with high concentration of nitrogen which can lead to formation of ammonium chloride (NH4Cl) in higher temperatures than it's normally observed. It's necessary to carry out a deep study to determine the adequate level of caustic soda injection in the crude to control the chlorides without prejudice to run length of thermal cracking processes like visbreaking and delayed coking units once excessive dosages will accelerate the fouling in the fired heaters.

Question 95 - Is it feasible to process slack wax in a delayed coker? If so, what are the potential benefits and operational challenges?

Response - The literature presents some reports about slack wax processing in thermal cracking refining units like delayed coking units. It's important to consider that the composition of slack wax is normally high in paraffins (between 80 to 90 % in mass) with low sulfur content (lower 1,0 % in mass), these characteristics will lead to good yields of light compounds like naphtha with high octane index due to the high olefins content which is expected from thermal cracking processes. The gasoil fractions tends to concentrate paraffinic molecules which will lead to difficulties to meet cold start requirements of diesel pool considering the current quality requirements, but this can be easily solved through adequate hydrotreating/hydrodewaxing processing units. In my point of view, the main challenge is related to the design of the delayed coking units which are normally designed considering a balance between the bottom and top yield of the main fractionator. Process highly paraffinic feeds like slack waxes will produce high amount of distillates like naphtha and gasoil fractions which will overload the top systems of the main fractionating tower and the pumping systems, on the other hand the bottom systems will operate probably below the minimum flow rate of pumps, heat exchangers and separation sections of the main column, creating significant operating and reliability issues. My advice is to carry out a blending of slack wax with conventional feeds to the delayed coking units like vacuum residue, asphaltic residue, and decanted oil from FCC in order the balance the benefits and side effects to processing highly paraffinic feeds like slack waxes in delayed coking units.

Question 96 - What could be done for de-aromatization of unconverted oil (UCO or off test) from hydrocracker unit? Could hydrotreating be effective or not? In the other words, is it possible to saturate PNA's (which they are in the UCO) by hydrotreating?

Response - Polynuclear aromatics compounds (PNA) are refractory to hydrotreating reactions due to their molecular arrangement and chemical stability, demanding severe and specific design and operating conditions to reach good saturation of these molecules. It's totally possible to reach a good performance in the PNA saturation in hydroprocessing units, but as previously mentioned it's necessary to supply adequate conditions for this. Normally, the hydrotreating technology licensors of processing units dedicated to process high PNA content recommends the operation under high hydrogen partial pressure (higher than 55 kgf/cm2) and high activity catalyst like Ni-Mo which presents a high hydrogenation potencial that will lead to a good performance to aromatics saturation as well as high WABT (Weighted Average Bed Temperature). It's important to taking into account that the aromatics saturation is limited thermodynamically and there is a temperature limit where the aromatics saturation performance will achieve the equilibrium and the performance will be not improved simply raising the WABT, at this point raise the temperature will only cause a severe deactivation of the catalyst. Under this scenario, if possible, the refiner can try to raise the hydrogen partial pressure to supply better thermodynamic conditions and displace the aromatics saturation equilibrium. In summary, it's totally possible to reach Polynuclear Aromatics Saturation in hydrotreaters, as long as adequate operating conditions in the processing unit.

Question 97 - Our aviation kerosene silver corrosion test is good, but copper corrosion is bad. What is the possible reason for this situation?

Response - This is a very interesting question considering that silver is more susceptible to corrosion than copper to sulphur compounds. Nevertheless, it's important to analyse the steam flow rate to the stripping tower in the kerosene hydrotreating unit, a poor performance of the stripping section will drag hydrogen sulfide to the product which will lead to failure to the corrosion tests (both copper and silver). It's important to analyse each step of the production process regarding the contamination with elemental sulphur or mercaptans, in processing units which relies on a caustic treating process to produce jet fuel kerosene, the caustic dragging can be a cause of corrosion in the derivative.

Another key point to be checked is the presence of sulphate reducing bacteria (SRB) in the storage tanks which can cause failures in the corrosive tests once these bacteria excretes H2S. In this case, the issue can be solved using biocide agents in the tank as well as ensuring a cleaning and draining routine of the kerosene storage tanks which will minimise the risks of SRB presence in the tanks bottom.

Question 98 - There is any safe limit for chlorine for DCU feed?

Response - According to my experience, some refiners are considering a chloride content below than 5 ppm (in mass) in the delayed coking feed as a safe limit to avoid corrosion issues, especially in the cold areas of the processing unit like the main fractionator top section.

The corrosion control is more effective considering the chloride content in the sour water accumulated in the vessel in the top of  the main fractionator once this is the main area which suffers with  chloride salt deposition, mainly the NH4Cl. The chloride content is related with the salt deposition temperature which needs to be controlled above the operating temperature of the top section of the main fractionator, the chloride content in the top vessel of the main fractionator should be controlled below than 4,0 ppm.

A good way to ensure these chloride limits in the delayed coking units is to control the performance of the desalting section of the crude oil distillation unit defining adequate chloride levels in the desalted crude and monitoring this parameter frequently as a critical parameter.

Special attention should be done to refiners processing high nitrogen content crude oils once the high nitrogen concentration will lead to higher concentrations of NH4Cl.

Question 99 - Regarding the operation of an UOP Fixed bed reformer (3 reactors).There is any limit of higher Delta T's of fixed bed reformer?

Response - The Delta Temperatures in semiregenerative naphtha catalytic reforming units should be controlled aiming to ensure an adequate lifecycle of the catalyst. It's is well knowledge that a rise of 30 F will the reactions rates, but the limitation here is the higher coke deposition over the catalyst which will significantly reduces the reformate and hydrogen yields. With a naphthenic feed it's really expected that the Delta T's will increase once the naphthene dehydrogenation are quick and highly endothermic reactions. Tecnically, there is no a limit to the reactors Delta T, but as previously mentioned it's necessary to control the Delta T in order to minimize the deactivation of the catalyst and reach a short lifecycle which can make the operation of the processing unit uneconomical. The best way to control the performance and the health of the catalyst of a semiregenerative naphtha reforming is through the WAIBT (Weighted Avarege Inlet Temperatures) parameter which can be calculated as presented below:

WAIBT = (Weight Fraction of Catalyst in the Reactor 1) x (Inlet Temperature of the Reactor 1) + (Weight Fraction of Catalyst in the Reactor 2) x (Inlet Temperature of the Reactor 2) + (Weight Fraction of Catalyst in the Reactor 3) x (Inlet Temperature of the Reactor 3)

The literature defines the Delta WAIBT as a good way to determine the end of life of the catalyst of semiregenerative naphtha catalytic reforming units, this parameter is calculated as presented below:

Delta WAIBT = WAIT - SOR WAIBT

Where WAIBT is the current Weighted Average Inlet Temperature while the SOR WAIBT is the Weighted Average Inlet Temperature in the Start of Run of the processing unit, when the Delta WAIBT is higher than 30 to 40 degrees, the catalyst is in the end of their lifetime.

Question 100 - How to decide the turndown of plant during design phase?

Response - This decision is normally supported by a review of the economic analysis which supported the initial investment decision. Changes in the consumer market or capital investments from the competitors can led the management team to reduce the capacity of a process plant or, in some cases, stop the enterprise istill in the design step.

Another factors can be changes in the regional or global economy which can represent severe threats to the investments which will raise significantly the investment risks.

Question 101 – How to prevent fouling in feed/effluent exchanger of sour water stripper unit?

Response - Unfortunately, fouling is a common operational issue in sour water stripping units.

The first step to understand and solve the fouling issues is to understand the origin of the fouling, in the literature is widely known that are the following types of fouling:

1 – Particulates

2 – Hydrocarbons

3 – Salts

4 - Elemental sulfur

5 – Polymers

The particulates are normally related to corrosion products from upstream sections or coke fines (if the sour water from delayed coking units is being treated), this can be solved using an adequate filtering system to the feed of the sour water treating unit as well as identifying adequately the most critical regions of corrosion and take actions to control this phenomenon.

The hydrocarbons are the most common source of fouling in sour water stripping units and the best way to control this is to manage the separation between the hydrocarbon and water phases in the feed, it is necessary to control adequate residence time to allow this separation. The best design practice is to install a feed tank upstream to the sour water stripping unit to allow adequate residence time with interface level control to minimize the hydrocarbon dragging to the sour water stripping unit. Among the operational variables, the pH can be controlled aiming to avoid the emulsion stabilization, the pH tends to raise along the stripping column considering that the H2S is easily stripped than NH3 causing fouling issues in the bottom section of the column (and then in the feed/effluent exchanger). The fouling caused by salts is normally related to excess of calcium or magnesium in the sour water, this issue can be avoided by controlling the hardness of the washing water of the upstream processing units like distillation and hydrotreaters through an adequate water treating process. The fouling caused by elemental sulfur can be faced by refiners dosing agents to control cyanide in the sour water (like ammonium polysulphide), a overdosage can reduce the pH and cause precipitation, in this case, it's necessary to define and control the adequate amount of polysulphide. The polymer fouling is caused normally by reasons like hydrogen cyanide polymerisation and amine degradation. The hydrogen cyanide can be neutralized using ammonium polysulphide, but it's necessary to control the correct amount as previously quoted to avoid elemental sulfur precipitation and the amine degradation fouling requires a deep analysis of the upstream process of amine treating units aiming to identify foaming and another operating issues. Another design strategies like the use of fouling resistant trays in the stripping columns and the instalation of by-passes in the feed/effluent exchangers can help to improve the lifecycle of the sour water stripping units. In summary, the first step to define a strategy of solution of fouling in sour water stripping units is identify the main source of fouling among the quoted fouling origin and then define the best strategy for this case. A very good source which can help to understand and solve this issue is the article from Mr. Phillip Le Grange which was published by PTQ Magazine in the Q2 2019 issue.

Question 102 – How do you manage the pressure in the Fractionator within a Delayed Coker Unit? From what I understand, both the Fractionator and Coke Drum pressures are regulated by the inlet pressure of the Wet Gas Compressor. However, during drum switch operations, significant fluctuations in the Fractionator top pressure make it challenging to maintain stable control. Are there any control strategies or philosophies that could be applied to effectively manage the pressure in the Fractionator and Coke Drum in the Delayed Coker Unit in addition to inlet pressure control of WGC?

Response - The control of a fractionator of a delayed coking unit is really a challenge for the process control experts. The batch characteristics of the delayed coking units lead to sudden variations to the mass and energy flow rates to the fractionator during the coke drum switch, where the hydrocarbon vapor flow to the column is practically cutted off. Normally, these disturbances are minimized through the anticipatory manipulation of circulating refluxes, side flow rates, and other heat sources capable to restore the energy and mass balance in the fractionator and minimize the disturbances, this control strategy demands a lot of agility and training from the operating personal and failures can lead to the production of off-spec streams which will cause financial losses and operating issues in the downstream units like hydrotreating units. Due to these characteristics the conventional feedback process control strategy is not adequate to manage the operational challenges of delayed coking fractionators. In the literature it's possible to find publications describing advanced process control strategies for delayed coking fractionators using multivariable predictive control and inferential quality estimators in order to minimize the disturbances caused by the process variations in delayed coking units. An excellent reference about this topic is the article published in the Q3 2016 issue of PTQ Magazine by Mr. Dinesh Jaguste which describes in detail the implementation of this control strategy.

Question 103 – What potential limitations might arise when attempting to reduce the coke drum pressure to the design values in Delayed Coker Units? Additionally, what types of investments might be necessary if we aim to further decrease the pressure?

Response - The pressure in the coking drum is defined by the pressure in the top vessel of the main fractionator where the pressure is controlled and added to the pressure drop between the fractionator and the transfer line between the coke drum and the fractionator. The main effects related to the reduction of the pressure in the coke drum is a higher volumetric flow rate of the gases inside the coke drum once it will make easier the vaporization of the heavier fractions of the feed. This will lead to these consequences: - Raise the risks of foam formation inside the coke drum, leading to the higher risks of coke deposition in the transfer line and the bottom of main fractionator; - Reduction in the naphtha and light coker gas oil (LCGO) yields raising the yield of heavy coker gas oil (HCGO) once the residence time of the feed in the coke drum will be shorter; - Reduction in the coke production and quality raising the risk of shot coke production, mainly for extremely heavy feeds. As previously described, the decision to severely reduce the operating pressure of the coke drum should consider the side effects related to quality degradation of the streams and coke as well as their impacts over the downstream processes like the hydroprocessing units. The main systems which need to be verified in a reduction in the operating pressure of the coke drums are the capacity of the main fractionator aiming to deal with the additional volumetric flow rate of gases as well as the adjacent systems like the transfer line.

Question 104 – I am facing a problem in vibration of control valve controlling atmospheric residue at 10 bar pressure,175 'C temperature and 620-630 m3/hr flow. Instrument team dropped the valve and found internals parts damaged, and they explained that the reason of damaging the valve internals due to low opening of 15-16%.

The attended and fixed the valve but same vibrations still coming.

Can anyone explain the exact reason and solution for this?

Response - Considering the problem description and process data, this vibration in the valve seems to be related to cavitation phenomena. Cavitation phenomena in control valves occur when the pressure reaches values below or near to the vapor pressure of the process fluid, this condition allows the vaporization of the lighter fractions of the fluid creating bubbles which collapse as the pressure increases in the pipeline, leading to noise and vibration. Another effect of the cavitation is severe damages to the valve internals due to the erosion caused by the bubbles that collapse which remove material from the valve walls. A design strategy to improve the valve resistance to cavitation is to apply anti-cavitation internals with higher hardness aiming to avoid material losses in case of cavitation of the valve. Regarding the process actions which can be taken to avoid the cavitation phenomena it's possible to quote: - Increase the downstream pressure by throttling a downstream valve or installing an orifice; - Reduce the differential pressure using two valves in series; - Installing a bypass line in the control valve; - Verify the valve design, if the valve is not adequately designed the operation conditions will facilitate the cavitation phenomena. The information from the instrumentation personnel regarding the low operating opening (15 - 16 %) can make sense; - Place the control valve in a lower elevation or where the fluid temperature is lower can minimize the cavitation occurrence once raise the pressure in the valve and reduce the vapor pressure, respectively.

Question 105 - Regarding visual inspection of Vacuum Distillation Bottom Product of Used Oil Re-refining, the sample appears to be more viscous after heating, and less viscous on cooling! Is there an explanation for that?

Response - If it is real, this is a very uncommon situation. According to my knowledge about used lubricating oil re-refining, the bottom of the vacuum distillation column is composed of an asphaltic residue which represents close to 13 % of the total used lube oil feed to the plant. This Residue contains high quantity of polymers and metals and can be used for asphalt blending, production of paving asphalt, bitumen protective covering or as fuel in the cement factories, and the expected behavior of this hydrocarbon is a lower viscosity as the temperature increases. The inverse behavior of viscosity of a liquid with temperature increase is not common and is not expected for hydrocarbons. Among the liquids with this uncommon viscosity behavior (higher viscosity with higher temperature) we can quote the sulfur between 150 oC to 200 oC and some polymers once the polymerization reactions occur by condensation and addition mechanisms. The condensation mechanism generally, involves the formation of water or other small molecule, unlike the process of addition polymerization that simply combines the monomers such as ethylene to produce polyethylene, heating the liquid mixture, accelerates the mechanism of polymerization which results in rapidly increasing the viscosity until the liquid components interact to produce solid long-chain polymers. My suggestion is to carry out an adequate characterization and viscosity analysis in the laboratory to identify the viscosity behavior with the temperature variation of the sample.

Question 106 - How to reduce sox emissions from FCC regenerator by changing which operating conditions..? Like Catalyst circulation reduction, combustor temperature reduction, ROT reduction etc.

Response - This is an increasing concern for any refiner operating fluid catalytic cracking units (FCC) taking into account the necessity to reduce the environmental footprint of the refining processes in compliance with the current regulations. Unfortunately, act only in the process variables will get poor results and severely impact the performance and profitability of the FCC unit, the most widely and effectives ways to reduce the SOx and NOx emissions in FCC units are:

1: Managing the sulfur content in the feed stream - This can be achieved selecting low sulfur crudes or hydrotreating the FCC feed streams;

2: Applying catalyst additives capable of reducing the SOx and NOx emissions in the regenerator - There are several FCC catalyst licensors which developed commercial solutions of these additives like BASF and Grace Company. It's important to taking into account that the most part of the SOx reduction additives will convert the SOx into H2S which will be removed in the amine treating step, but it's possible to the high H2S concentration can raise the occurrence of Carbonyl Sulfide (COS) in the LPG and Propylene streams once this is produced by the hydrolysis reaction between H2S and CO2. Considering this fact, the refiner needs to manage this threat through and adequate management of the lifecycle of carbonyl sulfide adsorbing bed to avoid troubles in the chemical grade propylene specification;

3: In partial burning processing units, the use of low NOx burners can ensure significant reduction in the SOx and NOx emissions;

4: Applying a flue gas scrubbing system - There are some process technologies developed to this purpose in commercial stage;

5: Reduce the processing capacity from the FCC aiming to control the SOx and NOx emissions rate - This is a drastic alternative, but can be considered according to the severity of the local regulations faced by each refiner.

As described above, there is not an easy way to achieve an effective reduction in the SOx and NOx emissions from FCC units and the most part of the available solutions requires significant capital investment. Again, just acting on the operating variables is not an effective way to reduce the SOx emissions and will produce significant impact over the performance and profitability from the processing unit.

Question 107 - Integrated refinery and cracker plant, processing Naphtha, Ethane, Propane, LPG as feed, experiencing high Vinyl Acetylene in their Mixed C4 feed. Severity controlled, however, VA concentration is very high. We are interested in understanding the likely reasons for high Vinyl Acetylene formation? What are the parameter need to be checked to control Vinyl Acetylene concentration? What parameters should be monitored to maximize the throughput of the butadiene unit while limiting Vinyl Acetylene contamination?

Response - Unfortunately, this is a relatively common operating issue in naphtha steam cracking units. Normally, the C3/C4 streams from the steam cracker contain low concentrations of acetylene derivatives known as MAPD (including vinylacetylene) which needs to be removed in order to avoid catalyst poisoning of downstream processes or undesidered side reactions. As informed in the question, once the severity of the naphtha steam cracking process is under control, it's important to evaluate the quantity of vinylacetylene is produced in the process plant, according to the literature is considered normal a production of 0,5 to 2,0 wt % in the C4 stream which can impact the efficiency of the butadiene solvent extraction step. To minimize this issue, normally the operators install a selective hydrogenation step downstream of the steam cracking in order to convert the MAPD compounds into C2 fraction preserving the olefins. If the process do not has a selective hydrogenation step, it's necessary to carry out a technical and economic study to install this process once is very hard to control the MAPD formation manipulating process variables of the steam cracking, the best way is control the process severity which is under control considering the question information and presents the side effect of economic impact related to the olefins yield. A very good reference about this issue is the article published in the Q2 2024 issue of PTQ Magazine by Edgar Jordan, Charlotte Fritsch, and Joachim Haertlé.

Dr. Marcio Wagner da Silva is Process Engineering Manager at a Crude Oil Refinery based in São José dos Campos, Brazil. He earned a bachelor’s in chemical engineering from the University of Maringa (UEM), Brazil and a PhD. in Chemical Engineering from the University of Campinas (UNICAMP), Brazil. He has extensive experience in research, design and construction in the oil and gas industry, including developing and coordinating projects for operational improvements and debottlenecking to bottom barrel units, moreover Dr. Marcio Wagner earned an MBA in Project Management from the Federal University of Rio de Janeiro (UFRJ), and in Digital Transformation at Pontifical Catholic University of Rio Grande do Sul (PUC/RS), and is certified in Business from Getulio Vargas Foundation (FGV).