Tips for Problem Solving of Hydrotreating and ...

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        This short publication summarizes some of my responses related to some questions about practical and theoretical questions related to hydrotreating and hydrocracking processing units.

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.

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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.

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).

using synthetic gas mixtures are considered to be bench scale.”

Spokesmen who affirm that industrial-scale reliability is now available, as well as spokesmen who deny it, often quote the NAE panel’s requirement of 1 year of operation at the 100-Mw scale. They argue whether or not this has been acheived in a particular unit, and whether or not the experience in this unit is generally applicable.

Not much attention has been paid to the other important ingredient by the panel as necessary to insure industrial process availability: the requirement that technical and economic data must be available to permit design of full-scale units to meet specific local and regional conditions.

The NAE panel did its work at a time when the chemistry of sulfur oxides scrubbing appeared far simpler than it does today. The panel considered 16 stack gas control procedures. A reflection of the subsequent advance in knowledge is the fact that 10 of the 16 were not represented by presentations at a meeting that EPA held in Atlanta in early November of 1974 to review the status of control technology. The Atlanta meeting considered 13 processes, of which 7 were not on the list of the 1970 NAE panel.

It should also be remembered that there have been expensive large-scale development failures in sulfur oxide emission control (see Table 11–1). One process, limestone injection into a boiler followed by a scrubber, that EPA urged upon utilities as late as early 1972 (Walsh 1972), is no longer being offered for sale.

The record would stand as an indictment of the engineering profession were it not for the fact, now evidient, that the engineer was compelled to press forward into design and construction of scrubbing equipment of unprecedented size in absence of adequate chemical knowledge. Never before had the chemical engineer been asked to treat such a large flow of gas even for a chemistry that was well understood. It is not surprising, therefore, that many of the early disappointments involved failure of large

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