Troubleshooting Manual for Petroleum Refining Processes

This short publication summarizes some of my responses related to some questions about practical and theoretical questions related to petroleum refining processes.

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.

Practical questions about hydroprocessing units and my responses

Question 1 - I am Engineer in a residue hydrotreater unit. As this is severe service (360C+), we often see upsets in the feed filter dP and the first reactor dP increases after a few months. Troubleshooting already done on feed composition, unit conditions.

The question concerns the feed tanks upstream the unit. For as there is water and sediments in the feed (from ADU), that should settle into the tank sump and be drained before the tank is lined up to feed the hydrotreater. I did rough estimation via Stokes law and involving the tank height, water, and feed SpGr, viscosity @ T, etc...

It shows that water would need roughly 24h to settle to the tank bottom. 1) Is this approach correct?

2) How to do these concerning sediments?

3) The tanks are old, have been in VGO service before. Last inspections did not reveal any internal corrosion. Now with residue service, SpGr became closer to water, so settling of water would take more time. I scare the tanks are not adequate for this new service.

We also think that even having an automatic BW filter, some of the tank sump comes to plug the reactor catalyst leading at least partly (the rest is coke) to the dP.

Thank you in advance for any feedback on this.

Response - I would like to suggest a two ways approach:

1 - Regarding the Tank and Feedstock filter - The tank was cleaned before the change of service? If not, it's possible to carried out some chemical incompatibility between the residue and the VGO in the ballast, leading to precipitation of asphaltenic compounds which can be plugging the downstream filters, another question is to ensure that the feedstock tanks have backup aiming to allow a frequent (maybe once a year) cleaning stop. Another question is if the feedstock filters have automatic backwash system, considering the hard service of this unity this is fundamental do ensure an adequate operation lifecycle.

2 - Regarding the Hydrotreating Reactors - A special attention needs to be considered to the catalyst grading of the reactors, the feedstock was adequately characterized? If yes, it's possible to determine the best catalyst grading to ensure the desired operational lifecycle including the use of contaminants trap which tends to minimize the pressure drop in the catalyst bed during the operational campaign, mainly in residue hydrotreaters which tends to operate under high severity and feedstock's with high contaminants content (metals, etc.).

Question 2 - What is the importance of analyzing basic nitrogen and non-basic nitrogen in Hydrotreater and Hydrocracking feeds. How does they affect the process and catalysts?

Response - Nitrogen compounds are known as strong inhibitors of activity of hydrotreating catalysts. Basic nitrogen compounds are the main concern related to the catalyst activity, but the non-basic nitrogen compounds can lead to inhibition reactions during the hydrotreating processes. Another relevant impact of the non-basic nitrogen compounds is poor performance of hydrodesulfurization due to the competitive adsorption over the active sites of the catalyst. The most part of nitrogen compounds are aromatics with great chemical stability, being hard to be hydrotreated. To overcome this challenge the nitrogen content in residue hydrotreating/hydrocracking units are controlled using adequate catalyst grading, in some cases with contaminants traps.

In fixed bed hydrocracking units which operates with high nitrogen content feeds, it's common to use a hydrotreating section upstream to the hydrocracking section aiming to protect the hydrocracking catalyst which are normally expensive, furthermore are applied separation vessels between the sections to reduce the nitrogen concentration in the hydrocracking section.

Question 3 - What is the purpose of Multi-catalyst Bed philosophy in hydrotreating of Diesel and Vacuum gas oil cuts?

Besides this, for Multi - Catalyst bed configuration, some x or y % of HDS / HDN / HDA happens in each bed, before getting admitted into next catalyst bed. How and who (factors) controls that x or y%?

Response - The main purpose of the multi-catalyst bed in hydrotreating units is to ensure a volume swell in the reactor leading to the optimization of the processing unit.

Normally, the grading of hydrotreating catalysts involves the use of guard beds in the top of the reactor aiming to control the contaminants concentration using macro porous catalysts which have the function to retain fouling agents like corrosion products, metals, organo-metallic compounds, and diolefins which tends to raise the pressure drop in the reactors and reduce de operational lifecycle of the hydrotreating unit. The following regions of the hydrotreating reactor is filled aiming to maximize the hydrogen uptake, a major part of the HDS and polyaromatic hydrogenation reactions are carried out in the region immediately bellow of the guard bed.

The following zone is normally dedicated to promoting the HDN and monoaromatic saturation reactions, and the following catalyst region is dedicated to promoting hydrogenation reactions once the limitation by nitrogen content tends to be minimized in this section. The percentages of HDS/HDN/HDA in each region relies on the characteristics of the processing unit like the total and partial hydrogen pressure, temperature, quench strategy, the characteristics of the employed catalysts, and from the characteristics of the feedstock.

Question 4 - What exactly causes coking on Diesel and Vacuum gas Oil hydrotreating catalysts (treating Crack and Straight run feeds)?

Response - The coking process in hydrotreating units generally occurs due to the cracking and dehydrogenation reactions in the catalyst beds that are favored by high temperature and due to the feedstock quality, hydrotreating units processing heavier feeds with high concentration of olefinics, polyaromatics, and asphaltenic compounds tends to present higher coke laydown rates. The hydroprocessing units operates under high hydrogen excess aiming to overcome the hydrogen diffusion limitations, once the hydrogenation reactions are exothermic, there is a temperature raising in the catalyst bed which favours the cracking and another side reactions (dehydrogenation) leading to the production of lower added value derivatives and coking deposition over the catalyst surface.

By this reason, for hydrotreating units processing unstable feeds (cracked feeds and residue) its fundamental an adequate design and operation of the quench and temperature control system of the hydrotreating reactors in order to ensure an adequate temperature control through the catalyst bed without hot points.

Question 5 - What is the meaning and importance of analyzing wt % H2 content in Vacuum gas oil feed and product?

Our VGO hydrotreater product specification is said to have min 1.0 % Delta hydrogen content w.r.t feed design value of 11.8wt %. If feed wt % H2 content is more than feed design value, will it make any difference to VGO product %wt H2 content?

Response - The hydrogen content in VGO (Vacuum Gas Oil) is an important characterization factor to determine the crackability of this stream as FCC feed.

Vacuum gas oils with high crackability normally presents high content of paraffin's and tends to present high hydrogen content while refractory VGO presents lower hydrogen contents indicating high concentration of aromatics compounds in the FCC feedstock. The hydrogen content can be used to estimate the yield of FCC products, once higher the hydrogen content in the FCC feedstock more hydrogen can be distributed in the FCC products and improve the final added value of the molecules.

Once the objective of residue upgrading technologies is to improve the relation H/C in the hydrocarbon molecules, there is two ways to achieve this goal. Through carbon rejection which occurs in Delayed Coking and Solvent Deasphalting Process for example or through hydrogen addition like is carried out in residue hydrotreaters and hydrocrackers. The specification of a minimum hydrogen content in the VGO is applied to ensure adequate performance to FCC units once low hydrogen content feeds indicates low crackability feedstocks which can lead to higher yields of low added value products (Decanted Oil) and high coke deposition rates over the FCC catalyst which can lead to extremely high temperatures in the hot sections of FCC units (Higher Delta Coke). Due this fact, it's important to consider the side effects of cracked feeds (Coker Gas Oil, for example) in FCC performance once these streams tend to present lower hydrogen content and, therefore lower crackability.

The raise in the hydrogen content in the VGO fed to the hydrotreater will raise the hydrogen content in the VGO product which will present a higher crackability and better performance as FCC feed. But is always important to consider what is the objective of the process, in some cases can be attractive to maximize the bottom yield in FCC units (Carbon black production for example).

Question 6 - Our Diesel Hydrotreater feed is containing CCR 0.01 wt ppm and Total Aromatics of 31.5%, if CCR of feed is in such low values, indicating coke forming tendencies of feed is very low, then what causes coking on diesel hydrotreating catalysts?

Our Vacuum gas oil hydrotreater feed contains CCR 0.85% and Asphaltenes <300ppm and product CCR is <500ppm. The route for reduction of CCR in VGO hydrotreater is by formation of coke on catalysts or is there any other mechanism of reduction of CCR without coke formation?

Response - Despite the relatively low CCR of the feed, the coking deposition process in hydrotreating catalytic beds depends on a several factors. Considering your information, the aromatics concentration in the feed of diesel hydrotreating unit is considerable and these compounds tends to present high coke lay down rates over the catalyst. It's important to ensure that the hydrogen partial pressure of the reactor is adequate as well as if interbed quench strategy is adequate for your processing unit, if not it's possible to create favorable conditions to hot points in the catalytic bed which favors cracking and dehydrogenation reactions which are responsible for the coke deposition over the catalysts.

Regarding the VGO hydrotreater it's possible to try to reduce the asphaltenes concentration through blending the feed stream with lighter streams, but it's necessary to make a deep study about the chemical compatibility of the streams aiming to avoid the asphaltenes precipitation in the pre-heating system or even in the catalytic bed which can reduce the operating cycle of the unit due to high pressure drop in the hydrotreating reactors. Again, it's fundamental to carry out a study to analyze if the VGO hydrotreating unit is operating under adequate operational conditions like hydrogen partial pressure, temperature, interbed quench strategy, etc. in order to ensure that the processing unit is adequately optimized and prepared to deal with the available feedstock.

Question 7 – What is the importance of maintaining specific range of temperature to Hot HP separator in hydrotreating units?

Relation of H2 gas and salts i.e.: Ammonium Bisulfide and Ammonium chloride solubility with temperature of the gas/liquid mixture upstream of Hot HP separator?

Response - This is a key issue in the hydrotreating units, especially those units processing heavier feedstocks which tends to present higher contaminants content like sulphur and nitrogen that will produce the NH4HS (Ammonium Bisulphide) and NH4Cl (Ammonium Chloride) under specific conditions.

For hydrotreating units processing lighter feeds, like naphtha, the hydrotreated stream tends to present physical properties very different from the water and a single separation vessel can be applied. For heavier feeds like LCO (Light Cycle Oil), and Gasoils two separating vessels are applied once the hydrotreated feed present physical properties closer with the water leading to a harder separation process which demands higher interface area, in these cases are applied a High Pressure (HP) separation vessel (Usually a vertical vessel) where it's separated the recycle gas (H2, H2S, and NH3) from the liquid phase which concentrates water, hydrocarbons and dissolved H2S and NH3. The liquid phase is fed to a Low Pressure (LP) separation vessel (Usually a horizontal vessel), as aforementioned this is the common configuration for high severity hydrotreating units.

In both cases the range of operating temperature aims to avoid the deposition of ammonium salts which tends to produce precipitations and corrosion which can reduce the lifetime of the process unit and led to accidents, generally the salt precipitation and the corrosion process is controlled through water injection in the post reaction section under temperatures above the precipitation temperature of the salts. In the literature is available some precipitation charts for the Ammonium Bisulphide (NH4HS) and NH4Cl (Ammonium Chloride). The Ammonium Bisulphide can precipitate in temperatures close to 140 oC while the Ammonium Chloride precipitate close to 250 oC, this is especially concerning for processing units with high chloride concentration in the feedstocks once the water injection needs to be made at higher temperature leading the necessity of the use of noblest metallurgy in the processing unit aiming to minimize the corrosion process. It's fundamental a deep characterization of the feedstock of the hydrotreatiing unit in order to measure the expected concentration of Ammonium Bisulphide and Ammonium Chloride based on the concentration of Sulphur, Nitrogen and Chloride in the feed, with this information it's possible to design an adequate water injection system to prevent the salt deposition and corrosion process.

Question 8 - Our Vacuum gas oil hydrotreater feed blend has got three feeds, Light Vacuum Gas Oil/Heavy Vacuum gas oil from VDU & Heavy Coker gas oil from Delayed Coker unit (Blend distillation IBP to FBP in the range of 260 to 600 deg C and specifically Temperature to distill 5% volume is approx. at 350degC). After hydrotreating, in the fractionation column we draw Diesel and Naphtha. As the feed doesn't contain Naphtha and Diesel cuts specifically & being hydrotreater but not hydrocracker, how does generation of Naphtha and Diesel in the fractionation column is happening (Both product streams combined Approx 14%)? Is it purely by thermal cracking at the hydrotreater conditions of Vacuum Gas Oil? How should I understand this?

Response - Despite being a Residue Hydrotreater, it's possible to carry out some cracking reaction in hydrotreating units under some operating conditions. It's important to remember that thermal cracking reactions are favored by high temperature which is normally the case of residue hydrotreaters specially in the end of operational campaign.

Furthermore, some references quote that can be observed conversion rates varying from 10 to 20 % of the feed stream for hydrotreaters while is expected conversion rates between 20 to 50 % for mild hydrocrackers and above 50 % to severe hydrocrackers, considering these parameters and the information from the question the condition can be considered "normal" at a preliminary analysis.

Despite this, please consider checking the severity conditions of the processing unit (especially the temperature and interbed quench strategy) aiming to identify operating conditions which can favors the thermal cracking reactions.

Question 9 - In Diesel Hydrotreater, what factors decides to go with Hot HP Separator / Cold HP Separator / Cold LP Separator configuration or only CHPS/CLPS configuration or only CHPS configuration?

Response - For hydrotreating units processing lighter feeds, like naphtha, the hydrotreated stream tends to present physical properties very different from the water and a single separation vessel can be applied. For heavier feeds like LCO (Light Cycle Oil), and Gas oils two separating vessels are applied once the hydrotreated feed present physical properties closer with the water leading to a harder separation process which demands higher interface area, in these cases are applied a High Pressure (HP) separation vessel (Usually a vertical vessel) where it's separated the recycle gas (H2, H2S, and NH3) from the liquid phase which concentrates water, hydrocarbons and dissolved H2S and NH3. The liquid phase is fed to a Low Pressure (LP) separation vessel (Usually a horizontal vessel), as aforementioned this is the common configuration for high severity hydrotreating units which is normally the case of modern Diesel Hydrotreating units that process unstable feeds (LCO, Coker Gasoil, etc.)

Regarding the configuration of HHPS (Hot High Pressure Separator), CHPS (Cold High Pressure Separator), CLPS (Cold Low Pressure Separator) the choice tends to be driven by energy consumption analysis. The use of HHPS configuration offers the possibility to energy savings once the hydrotreated stream can be fed to the stripping section without reheating need. Furthermore, the HHPS configuration allows a reduction in the dimensions of air-cooler system and the water injection can be made only after the hot separator which allows the use of noblest metallurgy materials in a small section of the processing unit. The advantages of a CHPS are the low contaminants concentration in the recycle gas which can affect the performance of the hydrotreating reactions due to the reduction of hydrogen partial pressure, in some cases this fact tends to be balanced through the higher consumption of make-up hydrogen.

Considering only the performance, the configuration of HHPS and CLPS tends to be chosen for high severity processing units which process heavier feeds, but it's important to consider another factor like dimensions of the processing unit, composition of make-up hydrogen, etc. Furthermore, according to the historic of the refining asset, the use of low pressure separators tends to be chosen to avoid exposing the operators to high pressure systems.

Question 10 - In Hydrotreaters for Stripping of H2S from Hydrocarbon stream, Naphtha Hydrotreater unit uses simple steam reboiler at the bottom of the stripper, whereas Diesel and Vacuum Gas Oil hydrotreaters uses direct steam injection in strippers. why? Why can't we use reboiler system in Diesel and Vacuum gas Oil hydrotreaters strippers instead of direct steam injection?

Response - Normally, there is some limitations in the heat charge in a reboiler due to the thermodynamic and heat transfer restrictions that imposes limits in the maximum temperature reached in a reboiler, by this reason, this strategy is applied for lighter products like naphtha while heavier derivatives like diesel and VGO requires the use of live steam route which is more effective in to promote the stripping performance due to the higher temperature and the reduction in the partial pressure of the hydrocarbons achieved in the strategy of live steam injection. Another factor which that is considered in the design choice is regarding the tolerance of water content in the final hydrotreated product, normally the downstream processes of a naphtha hydrotreating unit do not deal with high water content in the naphtha like the catalytic reforming units for example, leaving to the choice of reboiler strategy in these cases.

Another consideration that normally drives the choice is the generation of sour water in the strategy of live steam injection for stripping tower, as mentioned above this is normally the choice only for heavier streams due to the costs, environmental, and operational issues associated with sour water stripping units.

Question 11 - In any Hydrotreater unit, it is observed from our Material Balance sheets, that H2 dissolution is more in Hot HP Separator liquid than in Cold HP Separator liquid, even though the temperature is high in HHPS with marginal high pressure on the vessel compared to CHPS? Why is it so?

It is observed that Diesel Hydrotreater HHPS H2 dissolution is found to be on higher side compared to VGO Hydrotreater HHPS H2 dissolution, even though both pressure and temperature at HHPS of Diesel hydrotreater is less than VGO hydrotreater? Why is it so?

Response - The hydrogen solubility is a key parameter in hydroprocessing processes, and the facts pointed in the questions are related to the solubility behavior of hydrogen. The hydrogen solubility in hydrocarbons is increased by pressure and temperature, this explains the phenomenon described in the first question.

In the second question, it's expected that the hydrogen solubility is higher in lighter hydrocarbons like diesel than VGO (Vacuum Gas Oil) once the hydrogen solubility is reduced in the presence of heavy aromatics and heteroatoms, which is characteristic of VGO.

Question 12 - What is the main role of support material of the hydrotreater catalyst CoMo/NiMo? Does Support material participate in the hydrogenation reactions for HDS/HDN/HDA?

Hydrotreater catalyst at times said to be acidic and at times to be neutral as per the documents, which is true? Acidic nature for the catalyst is due to catalyst material or due to support material?

Response - The catalyst carrier or support offers mechanical resistance, high superficial area aiming to ensure an adequate distribution of the active phase (metals), and it's responsible to control the acid function of the catalyst which is desired to be low in the hydrotreating units. The support normally doesn’t have catalytic activity for hydrogenation reactions which is essentially carried out in the metal sites. Another function of the support in hydrotreating catalysts is to ensure an adequate pore distribution aiming to minimize the catalysts plugging due to coke or metals deposition which can lead to short operating lifecycle of the hydroprocessing units, this is an especial concern in residue hydrotreating units.

The catalysts applied in most severe services normally present acid and hydrogenation characteristics especially those applied in residue hydrotreating or hydrocracking processes. Catalysts applied in hydrocracking processes can be amorphous (alumina and silica-alumina) and crystalline (zeolites) and have bifunctional characteristics more pronunciated once it's desired that the cracking reactions (in the acid sites) and hydrogenation (in the metals sites) occurs simultaneously. The active metals used to this process are normally Ni, Co, Mo and W in combination with noble metals like Pt and Pd.

It’s necessary a synergic effect between the catalyst and the hydrogen because the cracking reactions are endothermic and the hydrogenation reactions are exothermic, so the reaction is conducted under high partial hydrogen pressures and the temperature is controlled in the minimum necessary to achieve the desired conversion of the feed stream. Despite these characteristics, the hydrocracking global process is highly exothermic, and the reaction temperature control is normally made through cold hydrogen injection between the catalytic beds.

As described above, the acid function in hydrocracking catalysts is take place in the acidic support which can be amorphous silica-alumina (ASA) and/or a zeolitic material while the hydrogenation reactions are carried out in the metal sites.

Question 13 - For Diesel hydrotreaters. How to finalize the catalyst from only CoMo & only NiMo and combination of NiMo/CoMo? Why is it said to be that NiMo catalyst consumes more H2 than CoMo catalyst?

Response - The catalyst grading of the diesel hydrotreater reactors relies on the feed stream quality, especially related to the contaminants content like sulfur and nitrogen as well as the participation of cracked streams like LCO, Coker Gas oil, etc. which are harder feeds to hydrotreating process. For feed streams with high content of these compounds it's applied a catalyst grading in the hydrotreating reactors with increased presence of high active catalysts like NiMo over alumina.

Once the CoMo is less active than NiMo catalysts, the first is applied to improve sulphur removal and olefins saturation while the NiMo catalyst is responsible for promoting nitrogen removal and aromatics saturation. The filling of the reactor (downflow reactors) normally starts with guard beds to protect the active catalysts against contaminants like metals (Ni and V) followed by the heteroatoms and unstable compounds saturation in the following beds in order to ensure an adequate temperature control in the catalyst beds. A relatively common configuration is to use a wide pore NiMo catalyst in the guard bed followed by a blending of CoMo and NiMo in the first catalytic bed aiming to promote sulfur removal and aromatics saturation followed by a NiMo bed aiming to promote the hydrodenitrogenation reactions followed by a last catalytic bed with a catalyst with high dehydrogenation performance (CoMo). Again, the catalyst grading configuration relies on the feed stream quality, design characteristics of the processing unit, and hydrotreating goals (specifications of the hydrotreated stream).

Regarding the higher hydrogen consumption of NiMo catalysts, as described above these catalysts are more chemically active than CoMo and are responsible for nitrogen removal and aromatics saturation which are more refractory contaminants, leading to a higher hydrogen consumption to achieve hydrotreating goals.

Question 14 - What is the difference between Type1/Type2/Brim/Hybrim catalysts?

What is Direct and Indirect desulphurization route in HDS reaction in hydrotreaters, what factors affects the routes or pathways?

Is there any relation for catalyst selection and route preference for HDS? How does route or pathway makes any difference in final product s specification?

Response - This classification is related to the Mo-S2 in hydrotreating catalysts. In Type I structures there is a strong interaction between the active phase and the carrier (Al2O3) mainly the interaction between the Mo and Oxygen from the support.

In Type II structure there are only weak interactions between the active phase of the catalyst with the carrier, the literature describes that Type II structure tends to present higher catalytic activity than Type I structure once the strong interaction with oxygen raises the required energy to promote the desulfurization reactions in the Type I catalysts.

The BRIM catalyst family was introduced by Haldor Topsoe company in 2003 and, among other improvements, presents higher dispersion of the active phase over the catalyst carrier leading to higher catalytic activity. The HyBRIM catalysts is an improvement of the BRIM catalyst where the interaction between the active phase and the carrier is optimized leading to a higher catalytic performance according to the licensor.

Regarding the desulfurization route:

1 - Direct Desulfurization - The whole atmospheric residue (or the hydrotreating feed) is fed to a hydrodesulfurization unit, and the sulphur compounds are treated according to hydrodesulfurization reactions.

2 - Indirect Desulfurization - The heavier fraction is separated from the atmospheric residue (or another stream which is the goal of the desulfurization process) from a separation process like vacuum distillation unit or through carbon rejection routes like Solvent Deasphalting (SDA). Once the sulfur and other heteroatoms tend to concentrate in the heavier fractions of the crude oil, this process indirectly reduces the sulfur content of the light fractions.

Question 15 - In hydrotreater, what is the difference between direct desulphurization reaction route and indirect desulphurization reaction route for HDS reactions?

Is it type of catalyst (NiMo/CoMo) or the nature of the sulfur molecule that decides the reaction route of HDS reactions?

Response - I believe that the question is regarding the desulfurization of heavy oils. The desulfurization of heavy fractions can be divided in two routes:

1 - Direct Desulfurization - The whole atmospheric residue (or the hydrotreating feed) is fed to a hydrodesulphurisation unit and the sulphur compounds are treated according to hydrodesulphurization reactions.

2 - Indirect Desulfurization - The heavier fraction is separated from the atmospheric residue (or another stream which is the goal of the desulfurization process) from a separation process like vacuum distillation unit or through carbon rejection routes like Solvent Deasphalting (SDA). Once the sulphur and other heteroatoms tend to concentrate in the heavier fractions of the crude oil, this process indirectly reduces the sulphur content of the light fractions.

The chemical characteristics of the sulfur compounds have a direct effect on its removal performance. Desulfurization of compounds that contain aliphatic sulphur, i.e. thiols and sulfides, is easier than desulfurization of compounds that contain aromatic sulphur, i.e. thiophenics. Due to this fact, the hydrodesulphurisation of heavier fractions requires higher operating severity than the process units operating with lighter fractions.

The characteristics of the catalysts affects the performance of the desulfurization process, Ni-Mo catalysts are more chemically active then the Co-Mo catalysts. For this reason, the Ni-Mo catalysts are employed for hydrotreating feeds with high nitrogen and refractory sulphur compounds (Thiophenics) content while the Co-Mo catalysts are employed to treating less refractory feeds.

Question 16 - Loading of catalyst in Hydrotreaters units follows Dense and Sock loading, despite having many advantages of Dense loading except high pressure drop.

It is observed in plants, that 1st bed grading and bulk catalyst to be in Sock loading and following beds top catalyst layers of little height in Sock loading followed by Dense loading for remaining bed height?

Q1. During what instances we choose to go for Sock loading?

Q2. How to choose inert balls size and quantity on catalyst bed support grid and on outlet collector?

Response - Normally, the dense loading is preferred once minimize the void spaces in the catalytic bed leading allowing a better flow distribution as well as higher catalyst mass in the reactor leading to a better performance during the operating run.

The advantage of sock loading process is the lower pressure drop through the catalytic bed, this can be a decision factor in processing units which limitations in dynamic equipment, but even under this scenario this issue tends to be relevant in the end of run, not in the start of run. Under normal conditions, the dense loading is preferred than sock loading process.

Regarding the choice of inert balls size, bed support grid, and outlet collector these devices have great impact over the total pressure drop and performance of the reactor, the design needs to follow the recommendations of technology licensors considering the specificities of each processing unit allied with the best engineering practices once high pressure drop can lead to the collapse of the support grid, causing an unplanned shutdown of the processing unit.

Question 17 - What is the purpose of Hot H2 Stripping in Diesel and Vacuum gas oil hydrotreater units at 360degC, during shutdown steps for catalyst replacement work?

Significance of maintaining H2S and not maintaining H2S in the HP loop during Hot H2 stripping? What are the consequences?

Response - The use of hot hydrogen stripping in hydrotreating units is related to the cleaning the reactor internals previous the maintenance services (hydrocarbon removal) as well as the coking removal over the catalysts which also help the spent catalyst draining step. Another purpose of the hot hydrogen stripping is to reduce the H2S content in the reaction section, previously the access of the maintenance workforce, leading to safer work conditions.

The hot hydrogen stripping can also be used to minimize the pressure drop of the catalyst bed through the removal of "soft" coke which lay down over the catalyst, over the time this coke can be converted into "hard" coke, which is more difficult to be removed, normally requiring a catalyst regeneration step.

Despite the benefits, the hot hydrogen stripping needs to be carried out carefully once the process can lead to the reduction of the metals in active phases of the catalyst if a very low H2S concentration is applied in the reaction system. By this reason, the catalyst licensors recommend a minimum concentration of H2S in the recycle gas (normally above 1.000 ppm) during the hot hydrogen stripping process to minimize the risk of catalyst deactivation by metal reduction.

Question 18 - Is it advisable to active a regenerated and/or reactivated hydrotreating catalyst with only feed?

Response - This is not recommended once the sulphiding process requires an adequate concentration of sulphur capable to promote the conversion of metal oxides into metal sulphides which is the active phase of the hydrotreating catalysts. It's difficult to ensure and control the concentration of sulphur in the catalytic bed with only the available sulphur in the feed, this can lead to the permanent deactivation of the catalyst due to the metal reduction.

By this reason the sulphiding process applying a sulphiding agent (DMDS or TBPS, for example) with carefully controlled procedure, especially related to the temperature control. The sulphur is heated in the presence of hydrogen generating H2S which is able to carry out the sulphiding reactions of the catalyst metals and generating the active phases (MoS2, CoS, NiS, and WS2).

Question 19 - Emergency depressurization Valve in Diesel and Vacuum gas oil Hydrotreater unit brings down the design operating pressure to half of it on its actuation in 15 minutes as per design Philosophy of the valve. What is the importance of this 15 minutes, why not less than or more than 15 minutes?

Response - The Emergency Depressurization System (EDS) is one of the most important process safety systems of a hydroprocessing unit and requires adequate criteria for design to avoid a worsening of the scenario of operational emergencies.

According to the literature, there is two criteria to design an EDS system: Low Rate Depressurizing System - Capable of reducing the pressure from the operating level to close 7,0 kgf/cm2 in 60 minutes. In this case the system is designed to reduce the operating pressure into 50 % of the designed pressure in 15 minutes with the objective of cut the sequence of hydrotreating reactions and block the raising of temperature and pressure;

High Rate Depressurizing System - Capable of reducing the pressure from the operating level to 20 % of the designed pressure in 15 minutes, this system is normally applied in high severity processing units considering the emergency scenarios like temperature runaway, fire, and leak in the reaction system.

The reason of 15 minutes as design parameter is to avoid the excessive mechanical stress of the processing unit through controlling the depressurizing rate which can lead to an excessive stress over critical systems like reactor internals if is much high or can limit the capacity of the processing unit to deal with emergency scenarios if the depressuring rate is too low.

Another side effect related to high depressuring rate system is the risk of damages in the flare operating systems, the capacity of the flare system to deal with the processing unit during emergency shutdowns in hydrotreating units needs to be considered during the design step in order to minimize the risks of damages in the flare tubulation system which can lead a worsening in the emergency scenario.

The depressuring rate is controlled through a restriction orifice (one or more) downstream of the emergency depressuring valve, this system needs to be adequately tested during the start-up of the processing unit to ensure that the EDS will act adequately under real emergency scenarios.

Question 20 - How does hydrotreated Vacuum gas oil product specification with respect to Sulphur content/Nitrogen content/ Aromatics/ product distillation (D5 %) impacts the FCC unit operations and its product yields?

Response - These parameters can be translated as the quality of feedstock to the FCC unit.

The sulfur content normally can’t be reduced in the FCC unit, so the sulfur content in the FCC product streams will be proportional to the feed content, this is especially critical with current regulations of gasoline (Tier 3) which can limit the use of cracked naphtha in gasoline pool or demands higher hydrotreating capacity/severity from the refiner.

The Nitrogen content is a key parameter to FCC feedstock once can be a severe poison to the FCC catalyst leading to the deactivation of acid function of the catalyst (basic nitrogen), furthermore will produce deleterious impact over the quality of FCC products.

Regarding the aromatics content in the FCC feed, aromatics compounds tend to be refractories to the catalytic cracking process, by this reason the crackability of the FCC feedstock tends to be lower for feeds with high aromatics concentration which will led to high bottom yield (normally with lower added value) in the FCC units. For the same reason, higher distillation temperatures of the feed tend to reduce the crackability of the FCC feedstock, leading to higher yields of bottom streams and high catalyst consumption since the contaminants like sulfur, nitrogen and metals tends to concentrate in the heavier fractions of the crude oil.

Question 21 - 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 shortage of market supply.

Question 22 - It is well understood that the Minimum Pressurization Temperature is the temperature below which the steel is assumed to be brittle due to H2 Embrittlement in hydrotreating units.

How to understand conceptually, why beyond MPT value, H2 embrittlement is not an issue or reactor is not susceptible to loss of ductility, even though pressure after MPT value will be higher than before?

Response - This is strictly related with the material behavior. The Minimum Pressurization Temperature (MPT) or Minimum Safe Pressurization Temperature (MSPT) establishes the minimum temperature that the reactor material still present toughness to avoid internal tension which lead to a transition from a ductile to fragile behavior, this is especially important in unplanned shutdown of the hydroprocessing units.

Above the MPT, despite the higher pressure, the reactor material present adequate toughness to avoid internal tensions capable to produce failures due to the hydrogen embitterment. It's important to know the MPT envelop of the reactor material to define the limits of the reactor and decision making during operational emergencies. Normally, it's recommended that the maximum pressure applied to the reactor do not exceed 25 % of the design pressure below the MPT.

Question 23 - In Diesel and VGO hydrotreaters, after feed pump failure, feed introduction is advised to be done within 15minutes, in case feed introduction is not done in 15minutes, it is advised to take normal shutdown by cooling of the reactor. What's wrong in feed introduction after 15minutes at high temperature?

Response - This recommendation normally is related to concern of catalyst deactivation due to the combination of high temperature and absence of H2S environment due to the lack of feed by long period (above 15 minutes).

It's possible to keep the processing unit under circulation and monitoring the H2S concentration in the recycle gas in order to guarantee the minimum concentration to avoid the desulfurization process of the catalytic bed, but this is necessary to be analyzed case by case once this procedure relies on the severity of the processing unit as well as the feed composition (straight run diesel + VGO, for example).

Question 24 - How to confirm the runaway reactions in Diesel and Vacuum Gas Oil hydrotreater units reactors? Is it localized phenomena at particular location or gets spread on wide area of the catalyst?

As the 1st bed hydrotreater bed exotherm for our unit already stays at above 40degC during normal operation, I would like to understand how to confirm the runaway phenomenon and take necessary actions like actuation of EDPS. Kindly guide.

Response - This is one of the key points during the operation of hydrotreating units, especially those processing chemically unstable feeds like VGO or delayed coking gas oils and can reach a great part of the catalytic bed. The temperature runaway in hydrotreating reactors is a phenomenon where the catalytic bed presents a sudden and uncontrolled temperature overshooting produced by exothermic reactions which can be provoked by the reasons below:

1 - Sudden flow rated reduction of the feed: This can lead to a hot points in the catalytic bed and, in extreme cases, damage the catalyst due to sintering of the active phase;

2 - Inefficient control of the fired heater: The fired heater control needs to respond adequately in case of overheating of the reacting section;

3 - Change in the feed composition: A significant raise in the composition of chemical unstable compounds of the hydrotreating unit can lead to a temperature raising of the catalytic bed once the rate of highly exothermic reactions raises significantly. It's necessary to keep the feed composition as stable as possible;

4 - Failure or deficient quench gas flow: The quench injection is responsible for keep under control the temperature of the catalytic bed as well as supply additional hydrogen to the hydrotreating reactions, for this reason it's necessary to ensure that this system is well designed and operated as the design requirements aiming to minimize hot points in the catalytic bed which can led to temperature runaway of the reactors;

5 - Sudden Change in the Capacity of Recycle Compressor: This can lead to a drastic reduction in the quench flow rate to the reactors and produce hot points in the catalytic bed, the change in the capacity needs to carry out in a smooth way to avoid sudden variations in gas flow rate through the catalytic bed;

6 - Methanation Reactions: This phenomenon is a concern especially in processing units operating under high severity (hydrocracking units, for example) and is related with the dragging of CO2 and CO to the reactors which combined with the operating conditions (temperature and pressure) can favor methanation reactions which are highly exothermic and will produce temperature runaway in the catalytic bed;

The main characteristic of the temperature runaway of the catalytic bed is a sudden and abnormal raise in the reactor temperature. For this reason, an adequate temperature monitoring of the catalytic beds is fundamental to identify the temperature runaway and allow mitigation actions in an adequate moment. One of the main side effects of the temperature runaway is the significant raise of coke laydown rate and the sintering of active phase of the catalyst which can produce a raise in the pressure drop in the reaction section, this can indicate that you have a problem with temperature runaway in the hydroprocessing unit. In summary, check if your processing unit is presenting uniform temperature distribution through the catalytic bed and if the pressure drop is raising under abnormal rate and if you are facing with some of the 6 reasons of temperature runaway above.

Question 25 - What are the ways or methods to ensure that sulphiding of the Diesel and VGO hydrotreater units catalyst is completed? What are the check points?

1. Is it purely by injecting the stoichiometric required DMDS into the reactor that decides sulphiding is completed?

2. Is there any thumb rule for quantity of DMDS that needs to be dosed during first stage of sulphiding at 220degC or quantity of DMDS that needs to be injected for second stage of sulphiding at 330 C?

Response - Nowadays, the modern and high performance Type II hydroprocessing catalysts demand the wet sulphiding method where normally is applied the DMDS as sulphiding agent due to his high H2S content in comparison with other compounds like DMS.

In this case, the checkpoints of the sulphiding process are the monitoring of the following parameters:

1: Injection rate: This should be monitoring comparing the real flow rate in relation with the theoretical flow rate through the level variation from the DMDS drum;

2: Hydrogen concentration in the Recycle Gas: This parameter is normally monitored through an online analyzer and the hydrogen concentration should be higher than 80 % during the sulphiding process. The monitoring of H2S content in the recycle gas is another fundamental parameter which needs to be monitored frequently during the sulphiding process.

3: Reactor Temperature: This parameter should be controlled to avoid temperature runway which can lead to coking deposition and catalyst damages.

4: Water Content in the Separator Vessel: Is expected water formation as by product during the sulphiding process, if the sour water level in the separation drum does not raise, this can be a signal of the sulphiding process is facing some trouble according to the step of the sulphiding process.

During the first stage of the sulphiding process, the temperature should be controlled between 180 to 230 oC (approximately) in order to ensure the decomposition of the sulphiding agent. Under this step, the temperature rise should be controlled under a rate of 15 to 20 oC/hour and the total sulfur content needs to be controlled between 1,5 to 2,0 % in mass (considering the sulphiding agent and the sulfur content in the hydrocarbon). The second phase of the sulphiding process is carried out under temperatures between 310 to 340 oC still under a total sulfur content of 1,5 to 2,0%, the reaction will reduce the H2S content in the recycle gas around to 2.000 ppm, after some hours the catalyst will stop to consume H2S in the sulphiding process and the H2S concentration will rise in the recycle gas, after this step the sulphiding process can be considered ended when is not observed sour water formation in the separation drum.

Question 26 - Why does NiMo catalyst has more hydrogenation potential than CoMo? What nature of it is enabling it to do so?

Why does Nitrogen molecules in hydrotreating units takes the path of hydrogenation followed by hydrogenolysis instead of direct hydrogenolysis unlike sulphur molecules which gets hydrogenolysis directly?

Between CoMo and NiMo, Which catalyst has more deactivation rate and why?

Response - The behaviour of NiMo and CoMo catalysts is strictly related to the chemical interaction between the metals and carrier (Type I and Type II catalysts) in the catalyst. The hydroprocessing reactions takes place in the active sites of the catalyst which is generally accepted to be located in the sulfur vacancies of the on the edges of MoS2 crystallites, these vacancies is significantly increased when the catalyst is promoted with Co or/and Ni. The Co-Mo-S phase is similiar to MoS2 structures with promoter atoms located in the edges of a tretragonal pyramidal geometry at the edge planes of the MoS2 while to Ni promoted catalysts, Ni can be present in three forms after the sulfidation: Ni3S2 crystallites over the support, nickel atoms on the edges of MoS2 structures, and nickel cations at octahedral or tetrahedral sites in the alumina. These different arrangement and interaction between the promoters (Ni and Co) with the MoS2 structures and the support leads to the different behaviour for CoMo and NiMo for hydrotreating reactions, being the CoMo more selective for sulfur removal under relatively low hydrogen consumption while the NiMo catalyst is more selective for hydrogenation and hydrodenitrogenation under higher hydrogen consumption rates.

The reactivity of sulfur compounds to the hydrotreating reactions tend to be higher than the nitrogen compounds once nitrogen in generally concentrated in the cracked and heavier fractions of the crude oil and great part of these nitrogen compounds have six or five pyridinic ring which are unsaturated, for remove nitrogen from these heterocyclic compounds it's necessary to hydrogenate the ring containing the nitrogen before to broke the carbon-nitrogen bond (hydrogenolysis), this is necessary due to the high energy of the carbon-nitrogen bonds in these rings. In the sulfur compounds case, the most part of the sulfur atoms are concentrated in thiophenic molecules that present relatively low energy bonds carbon-sulfur and can directly result in sulfur removal without necessity to saturate the heteroatom ring. By this reason, in hydroprocessing units treating heavier feeds which can concentrated refractory sulfur compounds like dimethyldibenzothiophene, the catalyst blending requires to rely on NiMo bed aiming to promote the hydrogenation function of the catalyst in order to minimize the steric hindrance of the sulfur molecules and improve the reactivity and consequently the efficiency of the hydrotreatment.

Related to the deactivation rate, this depends on the feed quality and severity of the processing unit but is expected than NiMo catalysts tends to have a higher deactivation rate than the CoMo catalysts once this catalyst (NiMo) is applied to treat heavier and cracked feeds which is notable refractories to hydroprocessing reactions.

Question 27 - About Processing Used Lubricating Oil in Hydroprocessing units. How long the Catalysts lasts? It's necessary to put metal trap separately?

Response - This topic is receiving increasing attention nowadays dragged by the necessity to improve the circularity potential of the crude oil derivatives. The hydroprocessing route of used lubricating oil re-refining tends to grow in the next few years due to the toxic byproducts generated by the other routes like the Meiken process.

Regarding the available hydroprocessing technologies there is the Hylube process developed and commercialized by UOP Company. In this process are applied two hydroprocessing reactors under pressure around 60 bar, the first reactor is act like a catalytic guard reactor where the porosity of the catalyst is designed to retain impurities, especially metals, and protect the second reactor which is responsible for the most part of the hydroconversion of the feed. Due to the characteristics of the used lubricating oils, it is expected that any hydroprocessing unit dedicated to process this feed rely on metal traps to protect the active catalyst and ensure an adequate lifecycle to the processing unit.

Interesting and complete references about this topic are the book Refining Used Lubricating Oils by James Speight and Douglas Exall and the book Design Aspects of Used Lubricating Oil Re-refining by Firas Awaja and Dumitru Pavel.

Question 28 - 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 29 – 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 30 - 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 31 - 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 32 - Is CCR value directly linked with asphaltene content? In hyrocracker feed specification why limit is fixed for both CCR and Asphaltene content? Why both parameters are 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 33 - 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 34 - 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 35 - 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 36 - 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 37 - 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 38 - 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 39 - 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 40 - 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 41 - 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 42 - 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 43 - 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 44 - 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 45 - 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 46 - 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 47 - 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 48 - 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 49 - 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 50 - 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 51 - 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 52 - 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 53 - 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 54 - 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 55 - 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 56 - 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 57 - 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 58 - 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 59 - 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 60 - 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 61 - 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 62 - 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 63 - I am currently working on fixed bed platforming unit. The unit is Semi-regenrative 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 ranges between 120-140F. The gases production decreases and hydrogen purity also decreases. While sulphur and nitrogen are in limit and moisture and HCL is also in limit

What could the reason's 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 the 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 63 - 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 odours.

Question 64 - We have a problem in one of the main towers (capacity = 150.000 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.

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