Kurt Detrick (UOP)
The feed to an HF Alkylation unit typically contains about 10 – 20 wppm of Sulfur. Nearly all of the sulfur in the HF Alky feed makes ASO, which mostly stays in the acid phase in the reactor. This ASO is removed from the acid in the Acid Regeneration (or Rerun) column and so most of the sulfur typically is rejected from the unit in the ASO product stream. However, during normal operation of the unit, as much as 20% of the sulfur in the feed to an HF Alkylation unit can wind up in the alkylate, so if the feed to the HF Alkylation unit increases above the typical 10 – 20 wppm, the amount of sulfur in the alkylate willincrease. Also, the following things can cause more than 20% of the sulfur in the HF Alky feed to wind up in the alkylate product:
- Acid carryover from the Settler. Incomplete settling of acid in the settler (due to high velocities or emulsion) can cause acid to be carried over with the hydrocarbon feed to the main fractionator. This will allow the ASO to go to the bottom of the main fractionator with the alkylate and will cause higher sulfur concentration in the alkylate.
- ASO leaving the top of the Acid Regenerator or Rerun column. In some unit configurations, if the Acid Regenerator or Rerun column is over-stripping, flooding, or otherwise operating in a way that causes ASO to go out the top of the column, this ASO (with the associated sulfur) can get into the main fractionator and then it will wind up in the alkylate.
- Internal Regeneration. When internal regeneration (regeneration of the acid in the Iso stripper or main fractionator) is done, all of the sulfur in the feed will wind up in the alkylate. Some units have stopped the use of internal regeneration for this reason.
Kurt Detrick and Daryl Dunham (UOP)
Gasoline corrosivity should not be caused by organic fluorides in the alkylate. Organic fluorides themselves are not corrosive, and organic fluorides will not break down to HF in tankage at ambient temperatures.
However, particles of iron fluoride scale can be present in the alkylate product. The tiny scale particles are created in the acid areas of the unit and some of these particles can eventually work their way down to the bottom of the main fractionator where they leave with the alkylate product. In the bone-dry alkylate stream, these particles are benign, but when they come in contact with water, they can hydrolyze and this can reduce the pH of the water to a point where corrosion of carbon steel can start to occur (about 5.5 – 6.0 pH). There has been corrosion reported in some alkylate tanks due to this phenomenon. Most HF Alkylation units take care of this problem by either managing the pH in the bottom of the alkylate tank or by treating the alkylate before it leaves the unit.
Jim Norton and Chris Steves (Norton Engineering)
Corrosion in alkylate tank bottoms has been seen due to the breakdown of organic fluorides, and many refiners have used KOH heels in these tanks to avoid this corrosion. The KOH heel must be periodically sampled and replaced as needed to ensure that it is not all consumed by trace HF.
Soni O. Oyekan (Prafis Energy Solutions)
The question posed is specific for hydrogen chloride management and it is necessary to broaden it to encompass the numerous challenges associated with chloride induced fouling, corrosion and reliability issues in catalytic reformers, fired heaters and downstream hydroprocessing units. Most of the chloride guard beds and adsorbents are used around the catalytic reformer as it is in that process unit that we need to intentionally add significant amounts of organic chloride for catalyst reactivation and reforming performance.
Foremost, the oil refiner should be operating as ideally as possible limiting nitrogen and sulfur to specific levels in the catalytic reformer feed. That is the naphtha hydrotreater should be operating as required and meeting reformer feed contaminants specifications. Secondly, chloride addition to the reactor section is as per the catalyst/technology licensor specifications and excess organic chloride is not being added. Thirdly, that the catalyst state is adequate, and the catalyst is neither contaminated with metals nor close to the end of life as either of those conditions would necessitate greater catalyst chloride addition rates and lead to increased hydrogen chloride in the net gas and in the debutanizer or stabilizer feed.
Fourthly, water management for the catalytic reformer is practiced satisfactorily so as not to lead to greater use of organic chlorides than necessary in the catalytic reformer in order to maintain catalyst performance. The basis for ensuring that the naphtha hydrotreater is operating satisfactorily and optimal water/chloride management and catalysts state are being maintained in the catalytic reformer is use and maintenance of sound catalyst and process unit monitoring programs. A number of the key performance indicators to be monitored have been discussed separately over the past years at AFPM Q&A Forums. The key performance indicators are, however, still important items as we discuss chloride guard bed locations and practices for effective management of chloride guard beds and adsorbents.
The locations for chloride guard beds are based on a variety of chloride management objectives and they are listed below:
•Upstream of booster compressors in continuous catalytic reformers product separation section
•For the gas and liquid products downstream of Re-contacting drums in continuous catalytic reformers
•Feed to the stabilizer
•Recycle gas and net hydrogen gas
•The off gas or refinery fuel gas
•The LPG product stream of the debutanizer
•The reformate
•The regenerator effluent gas of fixed bed cyclic regenerative reformers to meet RMACt 2 HCl regulation
•The reactor section of some cyclic regeneration reformers for excess chloride control
•The effluent gas from the catalyst (platinum/metals)) reduction section of a CCR
•In the reformer Debutanizer overhead section
As can be determined some of the chloride guard adsorbent locations are for managing specific chloride containing streams such as the off gas or refinery fuel gas to protect fired heater burner tips and regenerator effluent gas to meet RMACT 2 HCl regulation. Typically, the primary stream treated extensively in the refinery is the H2 make up or net hydrogen gas. As per past AFPM discussions, this stream must be treated adequately to enhance reliability of downstream Hydrotreating units, PSA units and equipment in refineries.
The majority of the adsorbents used are promoted alumina type despite formation of byproducts such as green oil and organic chloride that can eventually render them and other adsorbents ineffective. Various chloride guard bed operating strategies are in use including single guard beds with aggressive adsorbent replacement frequencies, lead/lag guard beds and effective monitoring to limit green oil, organic chlorides formation and HCl breakthrough or adsorbent saturation in order to effect timely adsorbents replacements.
Other adsorbents such as mixed oxides, zeolitic materials, “mixed” beds of zeolitic materials and mixed beds of promoted alumina and zeolitic materials in varying proportions and loading configurations are also in use.
Current practices for maximizing the use of chloride guard bed adsorbents include the following:
•Operate the naphtha hydrotreaters and catalytic reformers adequately.
oEffectively operate the NHT to minimize nitrogen, sulfur and water entering the catalytic reformers as resultant ammonia, hydrogen sulfide and water in the catalytic reformers would displace catalyst chloride leading to greater organic chloride usage and higher than necessary recycle gas, net gas, Debutanizer feed, off gas, LPG and reformate HCl.
oAmmonia would also lead to ammonium salts deposition in the product separation section and cause fouling and corrosion.
•Use appropriate, active catalysts in the catalytic reformers to ensure minimizing chloride use as discussed previously. Replace catalyst as often based on catalytic performance and the physical and chemical characterizations of the catalyst.
•Use lead/lag chloride guard beds strategy and replace the adsorbents frequently as required as breakthrough HCl in the treated net gas is likely to be too late due to green oil and organic chloride formation.
•Minimize water and Hydrogen sulfide concentrations in the net gas as that would interfere with effective HCl removal by the adsorbent due to competitive adsorption via the first practice item above.
•Use high performance adsorbents that are either capable of retarding the rates of formation of byproducts such as green oil and organic chlorides or capable of adsorbing organic chlorides.
Steve Philoon and Ka Lok (UOP)
In recently designed CCR Platforming units, chloride treaters may be used on a number of streams with the treaters placed in a variety of locations.
In the case of net gas treating the chloride treaters may be placed upstream of the net gas compressor or downstream of any recontact section and chiller-based product recovery system but upstream of a PSA hydrogen purification system (preferred). If chloride is a potential source of problems in the fuel gas system, then a treater can be placed on the Reformate Stabilizer Off gas line to fuel gas. Similarly, overhead liquid stream can be treated to remove chlorides from the LPG stream. For the principle liquid product, chloride treaters can be placed feed to the stabilizer column or on the reformate product line itself. The location selected may influence the sizing of the treater vessels and\or the stabilizer feed\bottoms exchanger and will influence on the adsorbent selection.
Use of molecular sieve has the advantage of reducing the risk of green oil formation over the activated alumina-based absorbent.
For treaters in gas service simply monitoring the outlet of the treater for HCL is not always effective. As the adsorbent (usually alumina) increases in chloride concentration, it can become acidic and catalyze the reaction of HCL with the light olefins present in the net gas. The products of these reactions are light organic chloride compounds that will not be detected by HCL detector tubes or laboratory tests for HCL. For this reason, UOP recommends that the refiner determine the change-out frequency of the alumina based on the calculated chloride loading on the alumina.
Mike Windham and Ka Lok (UOP)
The key to minimizing ammonium chloride salt build-up in the Stabilizer system is to minimize organic nitrogen in the feed. Salt accumulation is generally either in the Overhead Condensers or on the trays in the top portion of the tower. The most commonly used method to dispose of the salts is on-line water washing. We have heard of at least one refiner using a NALCO salt dispersant with success.
Steve Philoon and Ka Lok (UOP)
This topic has been discussed in 2008 NPRA Q&A session, below is some of the key points.
VCFE Cleaning
Pulling the VCFE tube bundle can be very difficult. Removing and re-installing the bellows is also a difficult task. Care should be taken not to damage the exchanger when pulling or installing the tube bundle. In addition, setting the bundle in the horizontal will, in most instances, cause damage such as tube to tubesheet leaks. UOP does not recommend removing the bundle unless absolutely necessary. It is recommended to wash the exchanger with water in order to remove ammonium chloride salts. Washing with reformate is recommended if the fouling is due to gums or PNA. Removing the bundle should only be done as a last resort.
Cleaning of the tube side is typically successful using high-pressure (up to 10,000 psi) water blasting. This cleaning (hydroblasting) can be done without removing the bundle, as both ends can be made accessible. If the tubeside of the VCFE is plugged, remove the top flange and expose the tubes. The high-pressure water jet lance can be inserted in order to clean the tubes. The shell side is more difficult to clean. Removal of the bundle and hydroblasting has, in most cases, been able to clean away shell-side deposits. But in at least one case, was not effective at cleaning the outside diameter of tubes in the center of the bundle. In-situ attempts at cleaning by washing with reformate or permanganate solutions have produced varying results. Use of hot solvent will help the solubility.
On-line washing procedure for the hot side of CFE by adjusting the last reactor outlet temperature hasbeen done by refineries. The effectiveness of this procedure yields mixed results for improvements.
Troy Small and Ka Lok (UOP)
Typical operating conditions in a reformer do not result in platinum volitization. However, it is possible for Platinum to become volatile and come off the catalyst at very high temperatures. One place this can occur is in the Chlorination zone of the CCR Regeneration Tower, where slipping coked catalyst into an oxygen rich atmosphere can result in very high temperatures. To prevent this, the refiner should make sure that the regeneration tower is operated according to the design.
UOP's experience is that these questions often arise as the result of an assay difference with the reclaimer rather than volatilizing platinum. The assay differences can be due to unrepresentative sampling or poor/biased analyses.
Soni O. Oyekan (Prafis Energy Solutions)
In order to fully answer the questions, it is relevant in my brief response to separate the precious metal catalyst platinum management into five distinct stages to cover a platinum catalyst manufacture to spent catalyst platinum reclamation life cycle. The stages that are pertinent for our review are:
•Reforming catalyst manufacture by the catalyst supplier and platinum settlement
•Reforming catalysts storage and catalyst loading
•Catalyst as used in the reforming units
•Catalyst dumping and transfer to platinum reclamation company
•Platinum settlement with the platinum reclamation company
It must be clearly understood that platinum losses can occur at any of the stages of the catalyst cycle. Some of the losses are due to contractual agreements as agreed upon in the first and fifth stages as a consequence of platinum settlement. The platinum or precious metals manager for an oil refining company should have the necessary expertise to aid in minimization of platinum losses for the oil refiner for the first and fifth stages above. In the fresh catalyst manufacture stage, the agreement with the catalyst manufacturing company for platinum settlement could stipulate 98 % to 99.5 % platinum return for the settlements. The platinum settlement requires that the oil refiner and catalyst manufacturer or platinum reclamation companies for the platinum settlement have appropriate analytical data (platinum assay, LOI for solid content) to permit effective conduction of the platinum settlement. Some oil refiners conduct platinum settlement with the catalyst suppliers, and some do not. I recommend conducting fresh catalyst platinum settlements to establish a reference initial platinum in use for the specific process unit that would be utilizing the fresh catalyst load and that the nominal platinum concentrations not be relied on as indicative of the reference fresh catalyst platinum. In the years that I managed precious metals for two oil refiners as a refinery technologist, several excess platinum troy ounces were returned to my companies’ platinum pool accounts after fresh catalyst platinum settlements with the catalyst manufacturers. In addition, the fresh catalyst platinum settlement data provided a good reference basis for the subsequent platinum inventory in the reactors after the catalyst loading.
In the spent catalyst platinum reclamation, a similar legal agreement could stipulate another 98 to 99 % platinum settlement with some additional platinum percent penalties for coke, catalyst alumina state (alpha or delta) and metals impurities. Thus, based on the two platinum settlements for fresh and spent catalyst for a catalyst life cycle, platinum losses due to contractual agreements and lack of the appropriate level of platinum management expertise by the oil refiner could lead to platinum losses in the range of 3 to 5 wt. % for the oil refiner.
Major additional significant losses could occur in stages 2 to 4 listed above. These combined areas of catalyst loading, in unit catalyst usage, catalyst dumping, and precious metals management are so intertwined and extensive that I strongly recommend securing the services of experienced technical experts who understand clearly the three major catalytic reforming technologies – semi regen, cyclic and continuous catalytic regeneration reformers and how their operations could greatly contribute to significant platinum losses. If you also own paraffin isomerization units and other process units that use platinum catalysts seek the assistance of a technical expert who fully understands platinum or precious metals management as well as the operations of the relevant oil refining process units that utilize platinum catalysts. An excellent oil refining expert could also work with your engineers and other relevant oil refinery personal on proactive steps for cost efficient catalyst management, process monitoring, and optimization and equipment management to minimize platinum losses.
Sandie Brandenberger and Ka Lok (UOP)
Directionally the higher O2 level is better for metal dispersion during oxidization. UOP recommends a minimum of 5% with typical maximum O2 content of 8-10 mol% based on the seal oil combustion limits. Dry seals or nitrogen purge seals allow higher O2 content without explosive conditions. If there is a history of coke ball formation or damage to reactor internals, maintaining a maximum of 3 mol% oxygen during the oxidation step should be considered and should be determined on case-by-case basis. When O2level is lower than desirable level, extending the hold time during the oxidization step will directionally improve metal dispersion.
Peter Eckels and Ka Lok (UOP)
If the coke content is very low in comparison with the coke burn capacity of the CCR unit, the operation can be limited in one of a few ways. A minimum gas flow is required to ensure the catalyst is properly dried before leaving the regenerator and returning for reduction. Sufficient flow must be maintained to protect the electric heaters and heat the catalyst for chlorination at low coke regeneration conditions. In some cases, the regeneration vent gas valve or makeup air valve to the regenerator is/are not in a stable control range. These are the consideration factors operating the CCR in standby mode.
There are several basic means of operation if normal White Burn cannot be maintained continuously. These may not all be available for all units.
Grey Burn mode is an operation with a mixture of nitrogen and air to the bottom of the Regeneration Tower to overcome the low flow limitation of the electrical heater. But the oxygen concentration for oxychlorination will be lower than normal operating conditions, reducing effectiveness of metals redistribution. Oxygen control could be harder in this operation.
Black burn catalyst circulation with regeneration allows an even laydown of coke on the catalyst inventory. Catalyst regeneration is operated in intermittently when coke on catalyst reaching the target level. The CCR will have to start with black burn mode first before switching to white burn mode and therefore a small portion of the whole catalyst inventory may not be regenerated in white burn mode. This operation mode uses 100% of catalyst inventory in the system to build up coke and therefore it maximizes the time interval between intermittent regeneration. This mode allows operation to monitor chloride and coke levels on catalyst and enables operation adjustment accordingly.
No catalyst circulation allows coke laydown of coke on catalyst inventory only in the reactor stacks. The catalyst regeneration is done in intermittent manner. This operation mode allows the regeneration in white burn mode continuously. The advantage of this operation mode is operation simplicity because operation does not require switching between black burn and white burn modes. However, the time interval between regeneration could be shorter due to not all the catalyst inventory in the system is available to build up coke.
Soni O. Oyekan (Prafis Energy Solutions)
The current challenges of low coke naphtha operation for the CCRs have been forced due to low severity operations. The low severity operations have been caused by factors such as ethanol blending and diesel to gasoline price incentives which have led to lower endpoint cut naphtha and low octane severity operations for catalytic reformers. Operating the catalyst regenerator section in the reactivation of catalyst with low coke can lead to significant catalyst and equipment damage. In addition, over several cycles of catalyst circulation, significant catalyst activity declines would occur due to inadequate catalyst reactivations. The greatest challenges, however, in low coke naphtha reforming in CCRs are operations with attendant risks of significant catalyst and equipment damage in the catalyst regenerator.
Catalyst circulation is typically recommended to ensure minimizing the chances for stagnant catalyst layers in transfer pipes that can become plugged leading to possible additional significant challenges in the reactor section. The methods utilized to manage the low coke operation is usually the preference of a specific oil refinery staff and what they are comfortable with. I also recommend catalyst circulation to minimize catalyst transfer plugs and to permit catalyst fines removal. Catalyst circulation would also permit getting a good assessment of the catalyst coke and when to re-start the regenerator section. One of the CCR technology licensors has recommended that their catalyst regenerator equipment could be modified to permit operating at lower coke levels in the low range of 2.0 to 2.5 wt. % coke and you can avail yourself of their services.
To use a proactive process based novel inventions to optimize low coke naphtha operations in CCR units, please review Oyekan, S. O., Rhodes, K. D., Newlon, N. K., US Patent 8,778,823, July 2014 assigned to Marathon Petroleum Company and Oyekan, S. O., Robicheaux, M. G., US Patent Application 2014/0138282 A-1, May 2014.